re engineering the chemical processing plant

RE-ENGINEERING THE CHEMICflL

PROCESSING PMNT Process Intensification

edited by

flndrzej ftankiewicz DSM Research

Geleen, and Delft University of Technology

Delft, The Netherlands

Jacob fl. Moulijn Delft University of Technology

Delft, The Netherlands

Copyright © 2004 by Marcel Dekker, Inc. All Rights Reserved.

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Preface

The book you are about to read will introduce you to modern ways of re-engineering the chemical processing plant by means of Process Intensification (PI).

The story behind this book had begun with the paper Process Intensification:Transforming Chemical Engineering, which we published in the millennium issueof Chemical Engineering Progress (January 2000). After a pretty enthusiasticresponse to our paper by the chemical engineering community, Marcel Dekkerproposed to us writing a book on that subject. After some discussions we came tothe conclusion that it was not a good idea to write the entire book ourselvesbecause, as you will see next, Process Intensification is a very broad discipline andincludes many diverse expertise fields. So, instead of writing all chapters on ourown, we have invited a number of prominent experts in various areas of ProcessIntensification, both from industry and from academia, to contribute to what nowhas become the world’s first book on that subject.

The principal aim of this highly practice-oriented book is to illustrate thecurrent developments and the frontline research in the area of ProcessIntensification. The book is primarily intended for engineers, technologists andresearchers in chemical, biochemical and engineering companies, who areinvolved in process design and development and are interested in learning moreabout equipment and techniques that may bring quantum-leap improvements totheir technologies. Also for others working in the forefront of process design and


development it is intended to be inspiring, in particular for the chemical engi-neering community in the universities and the National Laboratories. We hopethat it will contribute to a better image of the chemical industry and even play arole in attracting more high-quality, motivated students to the discipline. Thebook may also be beneficial to R&D managerial personnel who wish to have abroader understanding of the principles and methodology of ProcessIntensification and gain the up-to-date knowledge of the emerging novel equip-ment and processing methods that could help to achieve technological break-throughs in the processes at their companies.

The book has a certain logical structure that can be inferred from scanningthe individual chapter headings. Chapter 1 introduces the reader into the genesis,philosophy and principles of Process Intensification and discusses its dimensionand structure. It provides general information on process-intensifying equipmentand methods and gives some examples of their application on the commercialscale. The three subsequent chapters describe selected types of the PI-equipment.Most of that equipment have already been successfully implemented on the com-mercial scale or is ready for implementation. Chapters 2 and 3 are devoted to therotating equipment, rotating packed beds and spinning disk reactors, in which theuse of high gravity fields leads to spectacular miniaturization of the processingunits. Chapter 4 in turn describes the technology, design and application of com-pact and multifunctional heat exchangers. The next three chapters show howbringing certain structures in various scales of chemical processing environmentcan boost process efficiency, by dramatically improving mixing, heat and masstransfer. Various types and scales of such structuring are presented: microreactorsin Chapter 5, large-scale structured catalysts and reactors in Chapter 6 and inline mixing equipment in Chapter 7. Following that “hardware” part of thebook, its next four chapters focus on some important methods that can be used forintensification of chemical processes. Chapter 8 presents the application aspectsof functional integration of reaction and separation into reactive separation sys-tems, or integration of different separative techniques into hybrid separations. InChapter 9 the modeling issues of the reactive separation systems are discussed.Chapter 10 discusses some aspects of the integration of reaction and heat transferin multifunctional reactors, while Chapter 11 focuses on the application of pro-cess synthesis principles to the optimal design of integrated chemical processingplants. The final three chapters of the book address more general issues of Pro-cess Intensification. Chapter 12, based on the experiences within DSM, showshow the PI-principles can be applied in the industrial environment for re-designing and development of process-intensive chemical plants, while Chapters13 and 14 focus respectively on safety and sustainability aspects of PI.

The chemical industry skyline in the 21st Century is changing. New highlyefficient devices have already begun replacing the tens-of-meters high reactors andseparation columns. In the still denser populated world inhabited by the still more


educated and environment-conscious society, there will be no room (literally andfiguratively) for the huge, inefficient chemical factories of today, generating tensof tons of waste per each ton of the useful product. As a part of the society-drivenchanges miniaturization and, in general, intensification of chemical and biochem-ical plants, will become inevitable.

We are well aware that the present book does not cover all developments inthe field of Process Intensification. It has not had such ambitions. With this col-lection of contributions by the leading experts in the field, we have tried to focuson the main developments and main issues only, hoping that they will give thereader sufficient flavor of PI and will encourage him/her to further studies on howto re-engineer a chemical processing plant basing on the “smaller-cheaper-safer-slicker” principles of Process Intensification. Both contributors and editors willbe very glad to hear from the reader if we indeed have succeeded. Also sugges-tions for a possible next edition are welcome!

Andrzej StankiewiczJacob A. Moulijn


Contents

PrefaceContributors

1. Process Intensification: History, Philosophy, PrinciplesAndrzej Stankiewicz and A. A. H. Drinkenburg

2. Chemical Processing in High-Gravity FieldsDavid L. Trent

3. The Spinning Disc ReactorC. Ramshaw

4. Compact Multifunctional Heat Exchangers: A Pathway to Process IntensificationB. Thonon and P. Tochon

5. Process Intensification Through Microreaction TechnologyWolfgang Ehrfeld

6. Structured Catalysts and Reactors: A Contribution to Process IntensificationJacob A. Moulijn, Freek Kapteijn, and Andrzej Stankiewicz


7. Inline and High-Intensity MixersAndrew Green

8. Reactive and Hybrid Separations: Incentives, Applications,BarriersAndrzej Stankiewicz

9. Reactive Separations in Fluid SystemsE. Y. Kenig, A. Górak, and H.-J. Bart

10. Multifunctional Reactors: Integration of Reaction and Heat TransferDavid W. Agar

11. Process Synthesis/IntegrationPatrick Linke, Antonis Kokossis, and Henk van den Berg

12. Process Intensification in Industrial Practice: Methodology and ApplicationRemko A. Bakker

13. Process Intensification for Safety Dennis C. Hendershot

14. Process Intensification Contributions to Sustainable DevelopmentG. Jan Harmsen, Gijsbert Korevaar, and Saul M. Lemkowitz


Contributors

David W. Agar Lehrstuhl für Technische Chemie B, University of Dortmund,Dortmund, Germany

Remko A. Bakker DSM Fine Chemicals Austria, Linz, Austria

H.-J. Bart Institute of Thermal Process Engineering, University of Kaiserslautern,Kaiserslautern, Germany

A. A. H. Drinkenberg DSM Research, Geleen, The Netherlands

Wolfgang Erhfeld Ehrfeld Mikrotechnik AG, Wendelsheim, Germany

A. Górak Lehrstuhl für Thermische Verfahrenstechnik, University of Dortmund,Germany

Andrew Green BHR Group Limited, Cranfield, England

G. Jan Harmsen Delft University of Technology, Delft, The Netherlands

Dennis C. Hendershot Rohm and Haas Company, Bristol, Pennsylvania, U.S.A.


Freek Kapteijn Delft University of Technology, Delft, The Netherlands

E. Y. Kenig Lehrstuhl für Thermische Verfahrenstechnik, University of Dortmund,Dortmund, Germany

Antonis Kokossis University of Surrey, Surrey, England

Gijsbert Korevaar Delft University of Technology, Delft, The Netherlands

Saul M. Lemkowitz Delft University of Technology, Delft, The Netherlands

Patrick Linke University of Surrey, Surrey, England

Jacob A. Moulijn Delft University of Technology, Delft, The Netherlands

C. Ramshaw Department of Chemical and Process Engineering, University ofNewcastle upon Tyne, Newcastle upon Tyne, England

Andrzej Stankiewicz DSM Research, Geleen, and Delft University of Technology,Delft, The Netherlands

B. Thonon Greth, CEA–Grenoble, Grenoble, France

P. Tochon Greth, CEA–Grenoble, Grenoble, France

David L. Trent The Dow Chemical Company, Freeport, Texas, U.S.A.

Henk van den Berg Faculty of Chemical Technology, University of Twente,Enschede, The Netherlands, and Ghent University, Ghent, Belgium


1

Process Intensification: History, Philosophy, Principles

Andrzej Stankiewicz and

A. A. H. Drinkenburg

DSM Research, Geleen, The Netherlands

1. INTRODUCTION

Process intensification (PI) is currently one of the most significant trends in chem-ical engineering and process technology. It is attracting more and more of theattention of the research world. Four international conferences, several smallersymposia/workshops every year, and a number of dedicated issues of professionaljournals are clear proof of it. A number of commercial-scale applications of thePI principles have already taken place. But how did it all begin?

2. A BIT OF HISTORY

According to Miriam-Webster’s Collegiate Dictionary, the word intensive has prob-ably its origins somewhere in 15th century. And it was not many years later, rightat the peak of the Renaissance, when Georgius Agricola published his famous bookDe Re Metallica (1), the book that is commonly regarded as the first comprehen-sive textbook on the engineering of mining and metallurgy. De Re Metallica isrichly illustrated with woodcuts showing equipment and processing methods usedin the times of Agricola. In many of those woodcuts clear elements of processintensification–oriented thinking can be found. One example is shown in Figure 1,


FIGURE 1 Sixteenth century technology of gold retrieval from gold ore. (FromRef. 1.)


which illustrates the process of retrieving gold from gold ore. The technology ispretty simple. The ore is crushed by the stamp, “C,” ground in the mill, “F,” andmixed with mercury in vessels “O.” Gold is extracted from the ore by mercury andis later separated from it by pressing the mixture through a leather or cloth filterbag (not shown in the drawing). Taking a closer look at the woodcut, one noticesthat the stamp, the mill and the stirrers for mixing the ore with mercury are alldriven by the same water wheel, “A,” via the common axle, “B,” and a number ofvarious gears. Speaking the language of the 21st century, one could say, “A mar-velous example of a green, energy-based, highly integrated processing plant!” (Onedare not, however, call it a sustainable technology. Not only are the gold reservesunsustainable, but the operations involving mercury are not environmentallyfriendly, as we all know today.)

Yet there is another aspect to Agricola’s woodcut. As one may have noticed,some of the equipment shown (“O”–“S”) exhibits a striking resemblance to theequipment used in the chemical process industry almost 450 years later (seeFigure 2). Were the contemporaries of Agricola so ingenious, or are we merelysatisfied with the inventions of past centuries? At the dawn of the third millennium,in-series stirred tanks still remain the most common chemical processing system.An attempt to break this domination of the stirred-tank technology by the inventionand introduction of the static mixer (2,3), is one of the finest and earliest modernexamples of process intensification. Here, the technological leap was achievednot by the improvement of the stirring itself but, quite the opposite, by abandoningthe mechanical stirring as a method of mixing fluids! This reveals one of the mostimportant features of PI—the changes it brings are drastic in nature, revolutionaryrather than evolutionary.

In the scientific literature, the term process intensification started to appearin the mid-1960s and early ’70s, mostly in East European publications concerning

FIGURE 2 Four and half centuries have passed, yet almost no fundamentaldifferences can be seen between the technology of 1556 and that of 2002.


metallurgical processing (4–7) (an interesting coincidence, considering thatAgricola’s book was also on metallurgy, not on chemical processing). Of course,all of those papers understood process intensification as simply equivalent toprocess improvement. Also, in the first chemical industry–oriented articles (all ofEast European origin, by the way), the term process intensification had that samemeaning (8–10).

The birth of process intensification as a chemical engineering discipline cameseveral years later in the United Kingdom and was marked by the paper publishedin 1983 by Colin Ramshaw from the ICI New Science Group, who described theirstudies on the application of centrifugal fields (so-called “HiGee”) in distillationprocesses (11). A few months later the Annual Research Meeting, entitledProcess Intensification, was held at UMIST, Manchester (12). Interestingly, thefirst paper presented at that meeting concerned processing of gold ore usingintensive methods—a strange coincidence, indeed.

Both in the paper by Ramshaw and in the report from the UMIST conference,first definitions (or rather descriptions) of process intensification can be found.Ramshaw (11) describes PI as “devising an exceedingly compact plant whichreduces both the ‘main plant item’ and the installation’s costs,” while accordingto Heggs (12) PI is concerned with order-of-magnitude reductions in processplant and equipment. In one of his subsequent papers, Ramshaw writes about typ-ical equipment volume reduction by two or three orders of magnitude (13).

Until the early 1990s, process intensification was mainly a British disciplineand was focused primarily on four areas: the use of centrifugal forces, compact heattransfer, intensive mixing, and combined technologies (14). It was also the Britswho organized the first international conference on PI (15). By that time, however,process intensification had already become an international business, for manyresearch centers in different countries had entered the field. In Holland, for instance,Delft University of Technology, together with DSM, carried out research on structured reactors (16). Another group in Delft investigated centrifugal adsorp-tion technology (17). In France, Greth CEN institute in Grenoble carried outextensive studies on compact heat exchange equipment (18). In Germany, researchon microtechnology flourished in the Institut für Mikrotechnik Mainz (19), whilein China a special center at Beijing University was established to carry out R&Dactivities in the area of high-gravity processing (20). In the United States a numberof research institutes started PI-related studies too, e.g., Pacific Northwest NationalLaboratory in the field of microchannel heat exchangers (21) and MIT in the fieldof microreactors (22). Also early on, a number of chemical companies got involvedin process intensification. This involvement resulted in the first successful commercial-scale applications, such as the methyl acetate process of EastmanChemical (23), the hydrogen peroxide distillation system of Sulzer (24), and thehypochlorous acid process of Dow Chemical (25).


The end of the 20th century and the beginning of the 21st have seen a fastgrowth in PI-related activities in both industry and academia. In the UK the ProcessIntensification Network was launched, gathering a large number of industrial andacademic participants; a similar network has been established in the Netherlands.Four international conferences and several smaller symposia on PI have beenorganized so far. Process Intensification, traditionally tied up with the commoditychemicals sector, has begun entering new areas, such as bioprocessing and fer-mentation (26–28) and, quite recently, fine chemistry (29). The definition of processintensification has changed accordingly. It is no longer exclusively regarded asdrastically smaller equipment/plants (although equipment compactness remainsits most obvious feature). Process intensification, as it is widely understood now-adays, comprises novel equipment, processing techniques, and process developmentmethods that, compared to conventional ones, offer substantial improvements in(bio)chemical manufacturing and processing (30).

The question may arise why it took so long for PI to come into the picture.One possible answer is the enormous expansion of the process industry in thethird quarter of the 20th century, expansion in market size but certainly also inplant scale. There were very few incentives at that time for the very risky intro-duction of new technologies. In the fourth quarter of the century, much effort wasspent on modeling, optimization, and control, resulting, among other things, inthe well-known onionskin methodologies of process development (e.g., Ref. 31).Although very worthwhile at the time, these sequential, onionskin methodologies(first the reactor, then separation/purification, then heat integration, then processcontrol, safety, etc.) hindered the thinking in terms of integrated equipment. Moreincentives were found in environmental engineering, also a topic of the secondhalf of the 20th century, that developed from the end-of-pipe solutions to prob-lems to the integrated process solutions backsourcing the problems.

3. THE PHILOSOPHY AND OPPORTUNITIES

OF PROCESS INTENSIFICATION

The philosophy of process intensification has been traditionally characterized byfour words: smaller, cheaper, safer, slicker. And indeed, equipment size, land usecosts, and process safety are among the most important PI incentives. But processintensification can (and should) also be placed in a broader context—the contextof sustainable technological development. Several years ago DSM published apicture symbolizing its own vision of process intensification (32), in which sky-scraping distillation towers of the naphtha-cracking unit are replaced by a compact,clean, and tidy indoor plant (see Figure 3). The importance of PI for sustainabledevelopment and its role in the company’s responsible business has been furtherstressed in a recent publication by the company’s CEO, Peter Elverding (33). Here,


process intensification was the highest-rated activity of DSM within the known“Triple-P” (profit–planet–people) triangle, as shown in Table 1.

From this general philosophy of process intensification follow concreteopportunities that PI offers to chemical enterprises, as shown in Figure 4. Theseopportunities exist primarily in four areas: costs, safety, time to market, and com-pany image.

FIGURE 3 DSM’s vision of process intensification.

TABLE 1 Process Intensification in the Profit–Planet–PeopleTriangle of DSM

Triple P

Profit Planet People

Process intensification •• ••• ••Green routes • ••• ••Recycling • ••• ••Energy efficiency • ••• •

Source: Ref. 32.


FIGURE 4 Main benefits from process intensification.

3.1. Costs

Process intensification leads to substantially cheaper processes, particularly interms of:

Land costs, resulting from much higher production capacity and/or numberof products per unit of manufacturing area

Other investment costs, resulting from cheaper, compact equipment, reducedpiping, reduced civic works, integrated processing units, etc.

Costs of raw materials, due to higher yields/selectivitiesCosts of utilities, in particular costs of energy, due to higher energy efficiencyCosts of waste processing (less waste generated in process-intensive plants)

Figure 5 shows the estimated savings in some DSM technologies, after applyingthe PI principles (grass-roots situation).

3.2. Safety

Process intensification drastically increases the safety of chemical processes. It is obvious that smaller is safer. In Table 2 some of the more severe chemical disasters of the past century are listed. The table shows clearly how disastrousconsequences may arise from the large inventories when something goes wrong.And of course, one may not claim that process intensification would have preventedall those tragedies. Yet a study done at AIChE showed that methyl isocyanate (MIC),


the poisonous intermediate that was released at Bhopal, could be generated andimmediately converted to final products in continuous reactors that contained atotal inventory of less than 10 kg of MIC (34)! But process intensification offersnot only smaller equipment but also much better possibilities for keeping processesunder control, for instance, via extremely efficient heat removal from exothermicreactions (one speaks about heat transfer coefficients exceeding 20,000 W/m2K)or via fully controlled gas–liquid flow in structured catalysts that prevents liquidmaldistribution and hot-spot formation. Furthermore, intensification of the pro-cessing plant often leads to elimination of one or more of its components, whichalso has a direct advantageous effect on process safety (“What you do not havecannot leak”).

3.3. Time to Market

Process intensification also offers substantial improvements to those sectors ofthe chemical industry in which time to market plays a crucial role, e.g., the finechemical and pharmaceutical sectors. Ramshaw (35) discussed how processintensification could shorten the time to market in case of a low-tonnage pharma-ceutical process. The idea consists in developing a continuous lab-scale processand using it directly as the commercial-scale process. One must not forget thatliquid flow of only 1 milliliter per second means, in continuous operation, circa30 tons per year, which is quite a reasonable capacity for many pharmaceuticals.

FIGURE 5 Estimated savings in some DSM technologies, achieved by apply-ing PI principles to process and plant design (grass-roots situation).


TABLE 2 Some 20th Century Disasters in the Chemical Industry

EstimatedPlace Date Chemicals amount Casualties

Oppau/ September 21, Ammonium sulfate, 4,500 tons exploded ca. 600 dead, 1500 injuredLudwigshafen 1921 ammonium nitrate

Flixborough June 1, 1974 Cyclohexane 400-ton inventory, 28 dead, 89 injured40-ton escaped

Beek November 7, Propylene (mainly) �10,000-m3 inventory, 14 dead, 107 injured1975 5.5 tons escaped

Seveso July 10, 1976 2,4,5 Trichlorophenol, 7-ton inventory, No direct casualties, dioxin 3 tons escaped ca. 37,000 people

exposedSan Juan, November 19, LPG �10,000-m3 inventory ca. 500 dead, 7000

Mexico City 1984 injured (mainly outside the plant)

Bhopal December 3, Methyl isocyanate 41 tons released ca. 3,800 dead, 2720 1984 permanently disabled

Pasadena October 23, Ethylene, isobutane, 33 tons escaped 23 dead, 130–300 injured1989 hexene, hydrogen


The advantages of such an approach over the conventional one, based on thescale-up philosophy of “stirred tank → bigger stirred tank → even bigger stirredtank, always batch” are twofold:

Process development takes place only once, with no scale-up via a pilotplant to the industrial scale. The scale-up of a batch process in stirredtanks is not straightforward, especially in the case of reactions with alarge heat effect or a strong viscosity effect, and therefore can be trouble-some and time-consuming.

All the administrative (FDA) procedures involved in the legal approval ofthe production technology of the drug take place only once: The lab-scale technology is the commercial-scale technology.

In consequence, the start of commercial production can be greatly speeded up,in some cases even by several years. Time to market will be shortened and the patentlifetime of the drug will be much more effectively utilized (read: utilized longer).

3.4. Company Image

More and more chemical companies do recognize the fact that their image, theirreputation, plays a very important role in successful business. A proper image ofthe company is necessary to ensure public support for its activities. A study donein the United States showed that only the tobacco industry and the nuclear energysector had a worse reputation than the chemical industry. The situation in Europeis probably not very much different. On the other hand, process intensification,deeply anchored in the philosophy of sustainable development, in safe and envir-onmentally friendly processing, presents perhaps the simplest, the most obviouskey to the improved image of the chemical industry.

4. TECHNOLOGICAL BREAKTHROUGHS

AND CREATION OF SHAREHOLDER VALUE

In the reality of the global markets of the 21st century, not only do chemical com-panies compete with each other, they also have to compete with other sectors ofthe economy by proving to their shareholders that the revenues they receive fromthe chemical business are as good as or better than from other fast-growing sec-tors, such as software and servicing. But the chemical process industry strugglesto create value for its shareholders. As John Goldhill of Arthur D. Little Inc.wrote, the “chemical industry has lagged other industries in creating value forshareholders for at least the past 10 years” (36). And indeed, when looking atstock index developments in various types of enterprises in the period 1997–2002(Figure 6), one notices that the value of chemical shares grew substantially slowerthan in other sectors. The strategy of growth in the chemical process industry at


the present time is based mainly on mergers, splits, takeovers, and modifying thestructure of the product portfolio. It is basically a strategy of growth via trade, notvia technological innovation. In most chemical companies nowadays, opportuni-ties are sought in cost reductions via optimization of the primary business workprocesses (e.g., “operational excellence”) and via opening up bottle necks in theexisting production facilities.

Unfortunately, neither of these activities can make a company very attrac-tive to shareholders. In the optimization of work processes, a critical limit in costreductions will soon be reached, and competitors will also follow more or less thesame path, so the company’s competitive advantage will only be temporary unlessa shakeout takes place. Opening up bottlenecks, squeezing out yet another few percent from existing plants, is also not the way to convince investors that thecompany is capable of delivering an adequate growth in earnings. One of the mostobvious solutions to that problem lies in innovations and technological break-throughs, because only innovations and technological breakthroughs can ensurea sustainable technological advantage, cost leadership, and growth potential.Innovations and technological breakthroughs are what process intensification isall about.

FIGURE 6 History of the Standard & Poor stock index in the chemical sectorand other selected sectors of the economy, 1997–2002. (From: www.bloomberg.com.)


www. bloomberg.com

www. bloomberg.com

5. PROCESS INTENSIFICATION TOOLBOX

The toolbox for process intensification is schematically shown in Figure 7. Itincludes process-intensifying equipment (PI hardware) and process-intensifyingmethods (PI software). Obviously, in many cases overlap between these twodomains can be observed as new methods may require novel types of equipmentto be developed and, vice versa, novel apparatuses already developed sometimesmake use of new, unconventional processing methods. In Figure 7, examples ofboth PI hardware and PI software are shown. Many of them will be discussed indetail in other chapters of this book. Here, we give only a brief overview of themore important PI items.

5.1. Process-Intensifying Equipment

As already mentioned, one classic example of technological breakthroughs inprocess engineering was the invention and commercialization of static (motion-less) mixers, examples of which are shown in Figure 8. Nowadays, static mixersnot only offer a more size- and energy-efficient method for mixing or contactingfluids. In the SMR static mixer reactor by Sulzer, mixing elements are made of

FIGURE 7 Process intensification toolbox.


heat transfer tubes (Figure 9). Thanks to that, the SMR units can successfully beapplied in processes in which simultaneous mixing and intensive heat removal/supply are necessary, e.g., in nitration, neutralization, and polymerization re-actions.

For the cases when efficient mixing has to be coupled with a solid-catalyzedreaction a whole family of open-crossflow-structure catalysts has been devel-oped. The best known of them are the so-called KATAPAK®s, commercialized bySulzer. One of them, KATAPAK-M® is shown in Figure 10. It has good mixingproperties and can simultaneously be used as the support for catalytic material.

FIGURE 9 Static mixer reactor developed by Sulzer. (Courtesy: SulzerChemtech.)

FIGURE 8 Various types of static mixers. (Courtesy: Koch-Glitsch.)


KATAPAK®s are applied in catalytic distillation and in some gas-phase exothermicoxidation processes traditionally carried out in fixed beds. In these processes KATAPAK®s exhibit very good radial heat transfer characteristics (37).

In nonreactive distillation processes structured packings are also widelyused. One of the most recent and most promising types is the Super X-Pack devel-oped by Nagaoka International Corporation, shown in Figure 11. This wire-basedpacking is claimed to be able to reduce the height of a distillation column by afactor of 5 compared to a conventional tray design, as shown in Figure 12 (38).

FIGURE 10 Sulzer’s KATAPAK-M®. (Courtesy: Sulzer Chemtech.)


Also, the Super X-Pack can save up to 80% energy due to a substantially lowerpressure drop.

Heterogeneous catalytic processes can often be intensified by the use ofmonolithic catalysts (39). These are metallic or nonmetallic bodies forming amultitude of straight, narrow channels of defined uniform cross-sectional shapes(Figure 13). In order to ensure sufficient porosity and to enhance the catalyticallyactive surface, the inner walls of the monolith channels are usually covered with athin layer of washcoat, which acts as the support for the catalytically active species.

The most important features of the monoliths are:

Very low pressure drop in single- and two-phase flow, one to two orders ofmagnitude lower than in conventional packed-bed systems

FIGURE 11 Super X-Pack, developed by Nagaoka International Corp.(Courtesy: Nagaoka International Corp.)


FIGURE 12 Height reduction of a distillation column claimed by Super X-Pack. (Courtesy: Nagaoka InternationalCorp.)


High geometrical areas per reactor volume, typically 1.5–4 times higherthan in the reactors with particulate catalysts

Very high catalytic efficiency, practically 100%, due to very short diffusionpaths in thin washcoat layer

Stankiewicz (40) gives a spectacular example of reactor size reduction by a fac-tor of ca. 100, as a result of replacement of the conventional system with a mono-lithic reactor operated horizontally in a pipeline.

For highly exothermic reactions the so-called HEX reactors present a verypromising option. The basic common feature of all HEX reactors is much morefavorable heat transfer conditions in comparison with conventional reactors (heattransfer coefficients typically 3500–7500 W/m2K, heat transfer areas up to 2200 m2/m3). A HEX reactor developed by BHR Group Ltd. (Figure 14) was ableto decrease the by-product formation in one of ICI Acrylics’ processes by 75%(41) and to decrease the processing time in a Hickson & Welch fine chemicalprocess from 18 hours to 15 minutes, saving 98.6% of batch time (42).

Even higher values of heat transfer coefficients than those in the HEX reac-tors can be achieved in microreactors. Here, values up to 20,000 W/m2K are rep-orted (43). Microreactors (Figure 15) are chemical reactors of extremely smalldimensions that usually have a sandwich-like structure, consisting of a number ofslices (layers) with micromachined channels (10–100 µm in diameter). The layersperform various functions, from mixing to catalytic reaction, heat exchange, or sep-aration. Integration of these various functions within a single unit is one of the mostimportant advantages of microreactors. The very high heat transfer rates achievablein microreactors allow for isothermal operation of highly exothermic processes (alsoimportant in carrying out kinetic studies). The very low reaction-volume-to-surface-area ratios make microreactors potentially attractive for carrying out reactionsinvolving poisonous or explosive reactants (think about partial oxidation reactions).

FIGURE 13 Monolithic catalysts.


Also, microchannel heat exchangers have channel sizes around or lower than1 mm and are fabricated via silicon micromachining, deep X-ray lithography, ornonlithographic micromachining. The reported values of heat transfer coefficientsin the microchannel heat exchangers range from ca. 10,000 to ca. 35,000 W/m2K(21,44).

High heat transfer coefficients, though not as high as in the previous case,are also achievable in spinning disk reactors (Figure 16). This type of reactor hasbeen developed by Ramshaw’s group at Newcastle University and is primarilyapplied to fast and very fast liquid–liquid reactions with large heat effect, such as nitrations, sulphonations, and polymerizations. At very short residence times(typically 0.1 s), heat is efficiently removed from the reacting liquid at heat trans-fer rates reaching 10,000 W/m2K. The spinning disk reactor investigated in oneof SmithKline Beecham’s processes offered a 99.9% reduction in reaction time,99% reduction in inventory, and 93% reduction in impurity level (45).

Rotational movement and centrifugal forces are used not only in spinningdisk reactors. The earlier-mentioned high-gravity (HiGee) technology, started atICI’s New Science Group in the late 1970s as a spinoff of a NASA research proj-ect in deep space (microgravity environment), has developed into one of the mostpromising branches of process intensification. It consists of intensifying the mass

FIGURE 14 HEX reactor developed by the BHR Group. (Courtesy: BHR GroupLtd.)


FIGURE 15 An example of a microreactor. (Courtesy: Institut für MikrotechnikMainz.)

transfer operations by carrying them in rotating packed beds, in which high cen-trifugal forces (typically 1000 g) occur. This way, not only mass transfer but alsoheat and momentum transfer can be intensified. The rotating bed equipment, orig-inally dedicated to separation processes (such as absorption, extraction, distilla-tion), can also be applied to reacting systems (especially those mass transferlimited). It can potentially be applied not only to gas–liquid combinations but alsoto other phase combinations, including three-phase gas–liquid–solid systems. TheHiGee technology has already been successfully applied on a commercial scale,for deaeration of flood water in Chinese oil fields (20), where conventional vacuumtowers of ca. 30-m height have been replaced by the rotating machines of ca. 1-mdiameter. The earlier-mentioned hypochlorous acid technology of Dow Chemicalpresents another example of a successful application of rotating packed beds (25).Also, successes have been achieved in the crystallization of nanoparticles. In thegroup of Chong Zheng, very uniform 15- to 30-nm crystals of CaCO3 are madein a rotating crystallizer at processing times 4–10 times shorter than in the con-ventional stirred-tank process (46).

Another interesting example of process-intensifying equipment, also under-going commercialization, is the centrifugal adsorber. This is a new continuousdevice for carrying out ion exchange and adsorption processes. By using a cen-trifugal field for establishing countercurrent flow between the liquid phase andthe adsorbent, very small adsorbent particles (10–50 µm) can be used. This allowsfor extremely compact separation equipment (see Figure 17) with very short con-tact times and high capacities (typically 10–50 m3/h, (17)).


Other examples of interesting PI hardware include the supersonic gas–liquid reactor developed at Praxair Inc. (47) and the jet impingement reactor ofNORAM Engineering and Constructors (Hauptmann et al. (48)). The former isbased on using a supersonic shock wave to disperse gas into very tiny bubbles ina supersonic inline mixing device, while the latter uses a system of specially con-figured jets and baffles in order to divide and remix liquid streams with highintensity. Also, rotor/stator mixers (49) are dedicated for processes requiring veryfast mixing on the micro scale. They contain a high-speed rotor spinning close toa motionless stator. Fluid passes through the region where rotor and stator inter-act and experiences highly pulsating flow and shear. Inline rotor/stator mixersresemble centrifugal pumps and therefore may simultaneously contribute topumping the liquids.

5.2. Process-Intensifying Methods

As seen in Figure 7, three well-defined categories of PI software can be distin-guished:

Novel processing methods, such as integration of reaction and one or moreunit operations in so-called multifunctional reactors and integration oftwo or more separation techniques in hybrid separations

Use of alternative forms and sources of energy for chemical processingNovel methods of process/plant development and operation

Multifunctional reactors can be described as reactors that, alongside chem-ical conversion (and for the sake of it), integrate at least one more function (usu-ally unit operation) that conventionally would have to be performed in a separate

FIGURE 16 Spinning disk reactor.


piece of equipment. A pretty widely known example of the integration of reac-tion and heat transfer in a multifunctional unit are reverse-flow reactors (50).Also, a number of interesting reactor concepts for combining endo- and exother-mic reactions have been developed (51,52).

Reactive separations present probably the most significant class of multi-functional reactors, of which reactive distillation is one of the better-known andcommercially applied examples. The multifunctional reactor is in this case apacked distillation column, in which the packing material acts simultaneously asthe catalyst carrier. Chemicals are converted on the catalyst and reaction productsare continuously separated by fractionation (thus overcoming equilibrium limita-tions). Besides the continuous removal of the reaction products and higher yieldsdue to the equilibrium shift, the advantages of catalytic distillation units consistmainly in lower energy requirements and lower capital investment (53). Currently,numerous studies are being carried out in the field of reactive distillation modeling,as reviewed recently by Taylor and Krishna (54). Also, research on new internalsfor catalytic distillation columns attracts a lot of attention. Reactive distillationoriginates from and finds most applications in the hydrocarbon processing. Recently,interesting papers on the application of reactive distillation in fine chemistry beganto emerge (55). The reverse process to the reactive distillation, i.e., reactive con-densation, has also been studied (56).

Reactive extraction processes involve simultaneous reaction and liquid–liquidphase separation and can be effectively utilized to obtain significant improvementsin yields of desired products and selectivities to desired products in multireactionsystems, thereby reducing recycle flows and waste formation. The combination of

FIGURE 17 Progress in size reduction of adsorption equipment, up to themost recent centrifugal adsorption technology. (From Ref. 118.)


reaction with liquid–liquid extraction can also be used for separation of waste by-products that are hard to separate using conventional techniques (57,58). Reactiveextraction can also be used for selective separation of amino acids (59).

Reactive crystallization, or precipitation, has been investigated by numer-ous research groups. Processes of industrial relevance include liquid-phase oxi-dation of para-xylene to terephthalic acid, the acidic hydrolysis of sodiumsalicylate to salicylic acid, and the absorption of ammonia in aqueous sulfuricacid to form ammonium sulfate (60). A very special type of reactive crystalliza-tion is diastereomeric crystallization, widely applied in the pharmaceutical indus-try for the resolution of enantiomers (61). Another fine example of reactiveprecipitation is the earlier-described production of nano-size particles of CaCO3in high-gravity fields (46).

Reactive absorption is probably the most widely applied type of a reactiveseparation process. It is used for production purposes in a number of classicalbulk-chemical technologies, such as nitric or sulfuric acid. It is also often employedin gas purification processes, e.g., to remove carbon dioxide or hydrogen sulfide.Other interesting areas of application include olefin/paraffin separations, wherereactive absorption with reversible chemical complexation appears to be a prom-ising alternative to the cryogenic distillation (62).

Numerous investigations are being carried out in reactive adsorption pro-cesses, for instance, in chromatographic reactors, which integrate continuouscountercurrent chromatographic separation with chemical reaction (63,64), and inperiodic separating reactors, which are a combination of a pressure swing adsor-ber with a periodic flow-forced packed-bed reactor (65). This allows achievinghigher conversions and better yield by separating educts and products of an equi-librium reaction from each other. In the simulated moving bed reactor (SMBR), themovement of the bed with regard to the reactant inlets/outlets is usually realizedin a rotating system. One of the more interesting developments here is the rotat-ing cylindrical annulus chromatographic reactor, shown in Figure 18 (66). In thisdesign the inlets of the mobile phase are uniformly distributed along the annularbed entrance, while the feed stream is stationary and confined to one sector. As aresult of the rotation of the reactor, the selectively adsorbed species take differenthelical paths through the bed and can be continuously collected at fixed locations.

Another interesting example of reactive adsorption is the so-called gas–solid–solid trickle flow reactor, in which adsorbent trickles through the fixed bedof catalyst, removing selectively in situ one or more of the products from the reac-tion zone. In the case of methanol synthesis this led to conversions significantlyexceeding the equilibrium conversions under the given conditions (67).

A huge research effort is devoted nowadays to membrane reactors. Themembrane can play various functions in the reactor systems (68); it can, forinstance, be used for selective in situ separation of the reaction products, as aresult of which an advantageous equilibrium shift can be achieved. It can also be


applied for a controlled distributed feed of some of the reacting species, either toincrease the overall yield/selectivity of the process or to facilitate the mass trans-fer (e.g., direct bubble-free oxygen supply/dissolution in the liquid phase via hollow-fiber membranes (69)). The membrane can also be used for the in situseparation of catalyst particles and even homogeneous catalysts from the reactionproducts (70). Finally, the membrane can incorporate catalytic material, thusbecoming itself a highly selective reaction-separation system. Membranes aremore and more frequently employed in the life sciences sector, in manufacturingof pharmaceuticals, in combination with a bioreactor in which enzymatic reactiontakes place (71).

Multifunctional reactors may also combine reaction and phase transition. Awell-known example of such a combination is reactive extrusion. Reactive extrud-ers have been used increasingly in polymer industries. They enable reactive pro-cessing of highly viscous materials without a need for using large amounts ofsolvents, as is the case in stirred-tank reactors. Most of the reactions carried outin extruders are single- or two-phase reactions. Recently, however, new types ofextruders have been investigated with catalyst immobilized on the surface of thescrews, which enables carrying out three-phase catalytic reactions (72).

FIGURE 18 Rotating cylindrical annulus chromatographic reactor.


Fuel cells present another widely investigated type of multifunctional reac-tors, in which chemical reaction is integrated with the generation of electricpower (73). Simultaneous gas–solid reaction and comminution in a multifunc-tional reactor has also been investigated (74).

Among hybrid separations, the integration of membranes with another sep-aration technique presents the most important category. In membrane absorptionand stripping the membrane is used as a permeable barrier between the gas andliquid phases. By using hollow-fiber membrane modules, large mass transferareas can be created, which results in compact equipment. Besides, absorptionmembranes offer operation independent of gas- and liquid-flow rates, with noentrainment, flooding, channeling, or foaming (75,76).

In membrane distillation, two liquids (usually two aqueous solutions) heldat different temperatures are mechanically separated by a hydrophobic mem-brane. Vapors are transported via the membrane from the hot solution to the coldone. The most important (potential) applications of membrane distillation are inwater desalination and water decontamination (77–79). Other possible fields ofapplication include recovery of alcohols (e.g., ethanol, 2,3-butanediol) from fer-mentation broths (80), concentration of oil–water emulsions (81), and removal ofwater from azeotropic mixtures (82). Membrane (pervaporation) units can also becoupled with conventional distillation columns, for instance, in esterifications orin production of olefins, to split the azeotrope (83,84).

Membrane chromatography systems include microporous or macroporousmembranes that contain functional ligands attached to their inner pore structure,which act as adsorbents. In this sense, membrane chromatography is a hybridcombination of liquid chromatography and membrane filtration. Its most impor-tant potential applications include separations of biomolecules, such as proteins,polypeptides, and nucleic acids (85,86).

Among hybrid separations not involving membranes, adsorptive distilla-tion (87) offers interesting advantages over conventional methods. In this tech-nique a selective adsorbent is added to a distillation mixture. This increasesseparation ability and may present an attractive option in the separation ofazeotropes or close-boiling components. Adsorptive distillation can be used, forinstance, for the removal of trace impurities in the manufacturing of fine chemi-cals (it may allow for switching some fine chemical processes from batchwise tocontinuous operation).

Several unconventional processing techniques using alternative forms andsources of energy have been investigated thus far and are of importance forprocess intensification. The use of the centrifugal fields instead of gravitationalones was discussed earlier in this chapter. On the other hand, the formation ofmicrobubbles (cavities) in the liquid reaction medium as a result of ultrasoundwaves has opened new possibilities for chemical syntheses. These cavities can bethought of as high-energy microreactors. By their collapse, “microimplosions”


take place with very high local energy release (temperature rises up to 5000 K,and negative pressures up to 10,000 bar are reported (88)). This may have variouseffects on the reacting species, from homolytic bond breakage with free radicalsformation, to the fragmentation of the polymer chains by the shock wave in theliquid surrounding the collapsing bubble. In case of solid-catalyzed (slurry) sys-tems, the collapsing cavities can additionally affect the catalyst surface—this can,for example, be used for in situ catalyst cleaning/rejuvenation (89). Sonochemistryhas also been investigated in combination with other techniques, e.g., electrolysis,in the case of the oxidation of phenol in wastewater (90).

The use of solar energy in chemical processing has also been investigated.Studies describe, for example, the cycloaddition reaction of a carbonyl compound toan olefin carried out in a solar furnace reactor (91) or oxidation of 4-chlorophenolin a solar-powered fiber-optic cable reactor (92). The concept of using solar lightfor the synthesis of �-caprolactam was evaluated, and it was shown that the returnon investment was better than for the conventional technology (93). Solar reac-tors can also be used advantageously in water treatment plants (94).

The use of microwave dielectric heating offers significant advantages forchemical synthesis (95–97). Microwave heating was shown to enable some organicsyntheses to proceed up to 1240 times faster than by conventional techniques (98).

The employment of electric fields to augment process rates and to controldroplet size is known for a range of processes, including paint spraying, cropspraying, and coating processes. In these processes the electrically chargeddroplets exhibit much better adhesion properties. Electric fields can also enhanceprocesses involving liquid/liquid mixtures, in particular liquid–liquid extraction,where rate enhancements of 200–300% were reported (99). Bioseparations (e.g.,DNA separation) present another area in which electric fields can be advanta-geously applied (100).

Interesting results have been reported concerning the so-called gliding arctechnology, i.e., the use of plasma generated by the formation of gliding electricdischarges (101–103). In this technology, gliding electrical discharges are pro-duced between electrodes placed in the fast gas flow. They offer a low-energyalternative for conventional high-energy-consuming high-temperature processes.The applications tested so far in the laboratory and on industrial scale include:methane transformation to acetylene and hydrogen, destruction of N2O, reform-ing of heavy petroleum residues, CO2 dissociation, activation of organic fibers, airpurification from volatile compounds, natural gas conversion to syngas, and SO2reduction to elementary sulfur.

A number of other methods, not falling within any of the earlier-mentionedcategories, may prove useful for process intensification. Some of them, such as supercritical fluids, are already known and have been applied in other indus-tries (104,105). Because of their unique properties, especially the high diffusioncoefficient, supercritical fluids are attractive media for mass transfer operations,


FIGURE 19 Process intensification via process synthesis: methyl acetate plant of Eastman Chemical.


e.g., extraction (106), and for chemical reactions (107,108). On the other hand,cryogenic techniques (distillation or distillation combined with adsorption (109)),nowadays almost exclusively used for production of industrial gases, may in thefuture prove attractive for some specific separations in manufacturing of bulk orfine chemicals.

Among novel methods of process/plant operation, the use of unsteady-state(periodic) operation of chemical reactors has been studied for more than threedecades (110). In many processes studied on the lab scale, the intentional puls-ing of flows or concentrations led to a clear improvement of product yields/selectivities (111). Purposeful pulsing of the feed in trickle-bed reactors has beenshown to bring significant improvement in mass transfer rates, in catalyst wetting,and in the radial uniformity of liquid flow (112). The commercial-scale applica-tions of the periodic operation are scarce and are practically limited to the reverse-flow reactors discussed earlier. One of the main reasons is that a stationary processhas the advantage of providing constant production and product purity, without theneed for additional investments to synchronize nonstationary with stationary partsof the process. Further developments in the field of advanced process control maydefinitely change this picture, especially where the time constant of the pulsingmode is small—synchronizing will not be problematic.

Finally, in order to get a more or less complete picture of the process inten-sification toolbox, the PI-oriented methods for process/plant development mustbe mentioned. Among them, process synthesis (PS) definitely plays the mostimportant role (113–115). Process synthesis is in some sense a sister discipline ofprocess intensification that aims at the development of a cost-optimal processconcept based on the required functionalities. It includes diverse levels of activi-ties, starting from basic conceptual plant design (often based on the “out-of-the-box” approach), through the selection of optimal pieces of equipment and optimalinterconnections between them (plant integration), up to cost estimates. Processsynthesis permits early assessment and evaluation of the manufacturability ofproducts resulting from potential new chemistries. A textbook example of a com-mercial application of process synthesis is the methyl acetate plant of EastmanChemical (23), in which a task-oriented integration of reaction and separation ina multifunctional reactor reduced the number of pieces of equipment from 28 to 3(see Figure 19). In recent years, process synthesis has gone beyond its traditionalfield of applications and has entered new sectors, such as bioprocessing and phar-maceuticals manufacturing (116,117).

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2

Chemical Processing in High-Gravity Fields

David L. Trent

The Dow Chemical Company, Freeport, Texas, U.S.A.

1. INTRODUCTION

The use of high-gravity, or centrifugal, fields for chemical processing has generatedmuch interest in recent years. Fluid acceleration creates an environment in whichmass transfer rates are two to three orders of magnitude higher than ratesachieved in more conventional equipment, such as packed towers and stirredtanks. Heat transfer is also enhanced. Short contact time and fast transfer ratesallow a reduction in equipment size and in-process inventory. Many chemicalprocesses could benefit from these unique properties by reducing the cost of con-struction, reducing working capital, improving safety, producing less waste, etc.In addition, the use of high-gravity fields may provide solutions to processingproblems more effectively and more economically than conventional equipment.

Before exploring the wide range of applications, a brief review of the his-tory of development, a discussion of the process fundamentals, and an introduc-tion to mechanical design issues will help to set the boundaries for use ofhigh-gravity fields in chemical processing.

2. HISTORICAL DEVELOPMENT

The use of centrifugal fields in chemical processing dates to the beginning of theindustry with such physical transport operations as pumping, compression, andsolid/liquid separations. Extending this use into mass and heat transport operations


such as liquid/liquid extraction, gas/liquid interactions, reactions, crystallization,and heat transfer, however, is a much more recent development. Although thecommercial application of high-gravity fields has been limited in these recentareas of interest, the potential to improve existing chemical processes and developnew processes remains high.

The extension of the use of centrifugal fields into the commonplace ofchemical processing had its start in 1945 with the commercial application of acentrifugal liquid extractor to the recovery of penicillin (1). The liquid extractor,based on earlier patents by Podbielniak (2), employed perforated concentricplates for contacting the two countercurrent liquid streams.

Pilo and Dahlbeck (3) introduced an apparatus employing a variety of rotorinternals and liquid distributors for “countercurrent contact of two fluids havingdifferent specific gravities” in their 1960 U.S. patent. Although the authors suggestuse of the apparatus for gas scrubbing, distillation, heat exchange, and reactions,commercial application was not exploited. In 1966 Podbielniak (4) described acentrifugal device with concentric perforated plates for gas/liquid contact. Todd(5) extended the concepts of Pilo and Podbielniak to a multistage device in 1969.

In 1981 Ramshaw and Mallinson (6) provided for filling the rotor with ahigh-surface-area material such as glass beads or wire gauze in order to effectimproved mass transfer. They also described a rotor for distillation, complete withvapor/liquid contacting, condenser, and boiler. Following this patent, consider-able academic interest developed in an attempt to exploit the high mass and heattransfer rates, high-throughput capability, and short contact times afforded bycountercurrent fluid contact using centrifugal fields.

Sustained commercial application, however, did not occur until 1997, whenZheng et al. (7) described the successful stripping of oxygen from water used insecondary oil field recovery. In 1999 Trent et al. (8,9) introduced the first commer-cial application involving simultaneous absorption, reaction, and stripping. Both ofthese involve gas/liquid contact using a woven wire screen for the rotor internals.

3. PROCESS FUNDAMENTALS

In order to fully appreciate the application opportunities available from use ofhigh-gravity fields, an understanding of what happens within the rotor would behelpful. Since many of the models used to describe chemical processes in moreconventional equipment do not apply to high gravity, much of the informationavailable is empirical. However, where possible the traditional models are beingmodified to match the observed behavior. The process fundamentals of interesthere are mass transfer, heat transfer, fluid distribution and holdup, flooding, pres-sure drop, power requirements, and rotor internals. Although this last item couldbe considered with the mechanical design discussion, it is included here becausethe rotor internals can have a significant impact on the other process variables.


Before addressing these process fundamentals, however, a description ofthe basic process equipment would be helpful. Figure 1 shows a schematic of asimplified rotating packed-bed (RPB) contactor. This RPB illustrates countercur-rent gas/liquid operation. Liquid enters at the eye of the rotor, being distributedon the rotor packing at the inside diameter. The centrifugal force of the spinningrotor accelerates the liquid radially outward. Gas enters the stationary housingand passes through the rotor from outside to inside. The gas exits at the eye of therotor, while liquid drains from the housing. Seals on the drive shaft and on therotor ensure that the gas moves through the rotor.

Figure 1 shows a gas continuous configuration. A liquid continuous arrange-ment is also possible (10). For liquid/liquid extraction the preceding descriptionapplies if we consider the liquid to be the heavy phase and the gas to be the lightphase (11). In both of these scenarios the light phase enters through the drive shaftand channels radially in the rotor end plate to the outer periphery of the rotor fordistribution into the heavy phase (see Figure 2).

3.1. Hydrodynamics

Understanding the flow of liquid and gas through a rotating packed bed is import-ant to understand the performance results achieved. Liquid flow involves two

FIGURE 1 Schematic of gas continuous operation of a rotating packed-bed gas/liquid contactor.


components, liquid introduction to the packing and flow through the packing. Flowthrough the packing occurs in the radial direction with very little tangential or axialspreading (see Figure 3). A slight curvature in the radial flow results from thedirection of rotation. The degree of curvature and spreading is primarily a functionof rotor speed and liquid viscosity and less a function of packing type and liquidflow rate (12,13). Gas flow does not impact liquid flow through the rotor (14).

The flow pattern described results in incomplete wetting of the packing atthe outer diameter of the rotor (13). Thus, not all of the packing surface area isutilized for mass transfer operations. For maximum use of the packing surfacearea, the ratio of the outside diameter to the inside diameter should be minimized.Scale-up from small-diameter to large-diameter rotors typically provides moreefficient use of the packing. In spite of this incomplete wetting of the packing,very high mass transfer rates are achieved.

Enhanced mass transfer performance results from the initial contact of the liq-uid feed with the rotor. Studies at the Higravitec Center of Beijing University ofChemical Technology (HCBUCT), using a video camera attached to the rotor,revealed a breakup of the liquid feed into smaller droplets that filled the void spaces

FIGURE 2 Schematic of liquid continuous operation of a rotating packed bedfor gas/liquid and liquid/liquid contacting.


of the packing. This effectively increases the interfacial surface area of the liquidbeyond that of the surface area of the packing. At a packing depth of about 7–10 mm,most of the liquid has been accelerated to the rotor speed and wets the packing.Liquid flow continues mostly along the packing surface, but some additionaldroplets fly across the void spaces in the packing. The most intense mixing and masstransfer occur in the inlet zone of the packing. The degree of mass transfer enhance-ment at the inlet is a function of the type of packing (porosity, shape of packing struc-ture, etc.), rotor speed, method of liquid distribution, and liquid properties (15).

Liquid distribution on the rotor affects the initial contact zone performance.Some of the variables to consider include angle of impingement, velocity of the liquid spray, and acceleration of the liquid via rotation. Optimum performancerequires full axial wetting of the packing, whereas full tangential wetting is not nec-essary. The use of nozzles rotating in the same direction and at the same speed as therotor gave poor results for a hypochlorous acid process due to little surface area orliquid-side mass transfer enhancement (8). Although nozzles rotating in the oppos-ing direction of the rotor would be expected to provide the best mass transfer per-formance, the increased cost of manufacture of the equipment may not be justified.

Gas distribution in the rotor has not been studied as thoroughly as that of theliquid. Gas entering the rotor at the outside diameter accelerates radially inwarddue to the reducing diameter. Gas tangential velocity relative to that of the rotorvaries depending on the rotor packing. With parallel flat plates the low frictionaldrag (high slippage) makes the gas spiral inward with a path length much longerthan the radial thickness of the packing (Figure 4a). The gas path length approachesthe radial thickness of the packing as the packing surface area increases and the

FIGURE 3 General flow pattern of liquid in a gas continuous rotor.


porosity decreases, due to the increased drag of the rotor on the gas (Figure 4b).This behavior significantly impacts gas-side mass transfer performance (16).

The thickness of the liquid film on the rotor packing helps determine masstransfer rates. Film thickness can be shown to be inversely proportional to rotorspeed to the 0.8 power (17). Visual measurements using a video camera attachedto the rotor show a water film thickness of 20–80 microns on foam metal packingand 10 microns on wire gauze packing (15). Theoretical models estimate similarfilm thickness values (13,18,19). Film flow is expected to be laminar. In additionto rotor speed, liquid flow rate and fluid properties affect the film thickness (14).

FIGURE 4 Relative path of gas (a) in a rotor of low resistance (e.g. parallel flatplates), and (b) in a rotor of high resistance (low porosity, high surface area).


The foregoing discussion focuses on gas continuous operations, in whichthe bulk of the rotor volume is filled with gas. Liquid continuous operation canbe accomplished by collecting the liquid at the outside diameter of the rotor andchanneling it back for discharge at a diameter slightly larger than the feed radialposition, as illustrated in Figure 2 (20). Gas introduced into the liquid pool at theouter diameter moves countercurrent to the liquid inside the packing of the rotor.Because of the hydraulic pressure of the liquid from the centrifugal force, the gasbubbles start out small and steadily expand. Contact of the bubbles with the rotorpacking results in breakup of the bubbles, to maintain a high interfacial surfacearea. References 21 and 22 illustrate this approach for water deaeration and a cen-trifugal field bioreactor, respectively.

Liquid–liquid contact in an RPB involves introduction of the heavy liquidat the inside diameter of the rotor and the light phase at the outside diameter. Thetwo liquid phases move countercurrent to one another within the rotor packing.Centrifugal force causes the heavy liquid to move radially outward. This dis-places the light liquid, which moves radially inward. The design of the rotor pack-ing influences the contact between the two liquid phases (23,24).

3.2. Flooding

Rotating packed-bed devices handle high volumes of fluids in a small equipmentvolume, compared to packed towers, due to the acceleration of gravity. TheSherwood flooding correlation for packed towers (25) is expressed as a plot of

Early RPB researchers discovered that this flooding correlation for packed tow-ers applied equally well to RPBs when the gravity term (g) was replaced by cen-trifugal acceleration (r�2). As acceleration increases, the gas flooding velocity(UG) increases in order to maintain the same value of the first term. Since the ratioof liquid (L) to gas (G) flow remains constant, liquid flow increases commensur-ately with gas flow. Most researchers observed higher gas velocities before theonset of flooding than predicted by the Sherwood correlation (17,26,27).

Measurement of flooding by traditional means of observing a sharp pres-sure increase as gas rates increase is not effective with the RPB (28). Floodingcan be determined experimentally by adjusting rotor speed and holding gas andliquid rates constant. Flooding will occur at the point of maximum pressure dropover a range of rotor speeds (28) or by observing the increased pressure dropchange (inflection point) as a function of decreased rotor speed (27,29).

In liquid–liquid contact, two types of flooding can occur. The design of thecentrifugal contactor defines the throughput capability before flooding occurs.The second type of flooding relates to the principal interface moving into either

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the light-phase or heavy-phase takeoff. The movement of the principal interfaceis controlled by the back pressure of the light-phase takeoff, which is a functionof the rotor speed and phase density difference (24).

3.3. Residence Time

Liquid residence time in the packed rotor varies as a function of packing depth,packing type, rotor speed, and liquid properties (26). Two basic approaches havebeen applied to the measurement of liquid in the rotor. The first measure is theaverage residence time of the liquid within the rotor, and the second is the liquidholdup on the packing. Due to the flow patterns described previously, not all ofthe rotor packing is wetted and not all of the liquid resides on the packing sur-face. As a result, average residence time and liquid holdup are distinct measuresof liquid flow, contrary to the experience with packed towers.

Tracer methods, both visual and electrical conductivity sensors, have beenapplied to measure the residence time of the liquid in the rotor (15,26). Measuredliquid residence time ranges from about 0.2 seconds to about 1.8 seconds. Timedecreases as the rotor speed increases, as liquid flow rate increases, and as theradial position increases. Gas flow rate and liquid viscosity (narrow test range)have little impact on residence time (15,19,26).

Since liquid does not completely wet the packing and since film thicknessvaries with radial position, classical film-flow theory does not explain liquid flowbehavior, nor does it predict liquid holdup (30). Electrical resistance measure-ments have been used for liquid holdup, assuming liquid flows as rivulets in theradial direction with little or no axial and transverse movement. These data canthen be empirically fit to film-flow, pore-flow, or droplet-flow models (14,19).The real flow behavior is likely a complex combination of these different flowmodels, that is, a function of the packing used, the operating parameters, and fluidproperties. Incorporating calculations for wetted surface area with the film-flowmodel allows prediction of liquid holdup within 20% of experimental values (18).

Liquid holdup in liquid–liquid extractors must be defined for both theheavy and light phases. The light-phase outlet pressure is used to control the rel-ative liquid holdup of the two phases. Higher light-phase outlet pressure increasesthe light-phase holdup. This pressure has been correlated with the phase densitydifference, rotor speed, and rotor dimensions (24). In addition, packing charac-teristics of volumetric surface area and porosity influence liquid holdup andthroughout capability (23).

3.4. Mass Transfer

Developing correlations to describe mass transfer in rotating packed beds hasproven to be a challenge. Penetration theory (31), film-flow theory (32), and modi-fied surface-renewal theory (12) are some examples of leveraging previous work


to the rotating packed bed. At issue is the need to include a multiplicity of variablesas well as to understand the hydrodynamics of gas/liquid contact within the rotor.Each modeling approach compared experimental data to the models, with reason-able fit. However, these expressions have not yet been validated on multiple-RPBdesigns and operating conditions. A key to developing a generalized model formass transfer performance is the proper understanding and treatment of the fluidflows within the rotor, as discussed previously in Section 3.1.

The combination of high surface area, high velocities, thin films, andintense mixing in the packing provides an environment for intensive mass trans-fer, resulting in values for height of transfer unit (HTU) of 1.5–4 cm (26). Masstransfer has been described using HTU, number of transfer units (NTU), masstransfer coefficient (kL, kG), and volumetric mass transfer coefficient (kLa, kGa,kSa). To accommodate the variation in packing surface wetting with radial distance,an area transfer unit (ATU) has been proposed (33). Another proposed method ofevaluation uses a volume transfer unit (VTU) to account for the entire volume ofthe packed rotor (34). Although the ATU and VTU methods may have merit inevaluating RPB performance, these methods make comparison with other trans-fer devices based on HTU more difficult.

The possible physical design parameters affecting mass transfer includepacking and packing supports. Atomization of the liquid as it impacts the spin-ning rotor packing creates high-surface-area liquid drops, in addition to the filmwetting of the high-surface-area packing. This atomization results in significantmass transfer apart from the packing surface. As a result, low-surface-area pack-ings produce equivalent, if not better, volumetric mass transfer coefficients thando high-surface-area packings (35,36). This implies that low-surface-area pack-ing with high porosity can effectively replace high-surface-area packing, contraryto the experience with packed towers. The result is lower-cost packing, reducedpressure drop, and higher throughput. Packing supports at the inside diameter ofthe rotor generally provide a positive effect on mass transfer (36). Due to theatomization of liquid exiting the rotor, additional mass transfer occurs in thespace between the rotor and the housing (37).

Operational parameters of importance to mass transfer include rotor speed,liquid rates, and gas/liquid ratios. Mass transfer increases proportionately to rotorspeed, decreases with increasing liquid flow, and increases with gas/liquid ratio(17,26,36). Although most references present rotor speed as revolutions perminute (rpm), expression as either tangential velocity (�r) or multiples of gravity(r�2) provides a better basis for comparison among the different rotor designs andfor scale-up.

Gas-side mass transfer in rotating packed beds does not show the samelevel of enhanced performance as liquid-side mass transfer. Average volumetricgas mass transfer values for a wire screen packing increased with gas flow ratebut decreased with increased rotor speed. Compared to a packed tower, the RPB


mass transfer coefficient (1–8 s�1) was similar when operated at similar super-ficial gas velocities (�1 m/s). However, when gas velocities were increased (4–12 m/s) in studies with parallel flat plates as the rotor, the mass transfer coef-ficient also increased, to a high of 45 s�1 (16). In a commercial-scale RPB, averagevolumetric gas-side mass transfer coefficients of 40–50 s�1 were achieved usinga wire screen packing and gas velocities of 4–5 m/s (9). The factors affecting gas-side mass transfer are less understood than those of liquid-side mass transfer.

Liquid–liquid mass transfer in centrifugal extraction contactors shows simi-lar trends on performance as the gas–liquid contactors. Mass transfer improves athigher rotor speeds, higher solvent ratio, and higher phase density difference.Since the light-phase outlet pressure controls the liquid holdup of both phases,decreasing the light-phase outlet pressure decreases the light-phase holdup andincreases the number of transfer units (24). Packing characteristics of pore size,porosity, and volumetric surface area also play a role in performance (23). Singlecentrifugal extractors have achieved up to 10 theoretical stages of extraction (38),but they could achieve up to 20 stages with suitable rotor design (11).

Liquid–solid mass transfer has also been studied, on a limited basis.Application to systems with catalytic surfaces or electrodes would benefit fromsuch studies. The theoretical equations have been proposed based on film-flowtheory (32) and surface-renewal theory (39). Using an electrochemical cell withrotating screen disks, liquid–solid mass transfer was shown to increase with rotorspeed and increased spacing between disks but to decrease with the addition ofmore disks (39). Water flow over naphthalene pellets provided 4–6 times highervolumetric mass transfer coefficients compared to gravity flow and similar super-ficial liquid velocities (17).

3.5. Pressure Drop

Gas pressure drop through the RPB rotor is an important consideration whencomparing the performance of the RPB with other mass transfer devices, such asa packed tower. Numerous studies on pressure drop in RPB rotors employing avariety of packings have yielded some surprising differences from conventionalpacked towers. For example, lower pressure drop for wetted packing compared todry packing has been reported (26,40,41). Not all researchers observed this phe-nomena, because pressure drop was found to be a function of packing type, rotordesign, gas rates, liquid rates, and rotor speed (41).

Pressure drop has been reported for a number of rotor internals, includingcorrugated structured packing (28), foam metal (26,40,42), rectangular and ellip-tical cylinder plastic grains randomly packed (41), wire screen (43), and glassbeads (17). In spite of the variation in porosity from 0.38 to 0.95 and in volumet-ric surface area from 500 to 4000 m2/m3, all of these studies showed similaritiesof increased pressure drop as rotor speed increased and gas rates increased.


In addition to the differences in packing type and characteristics, these stud-ies also had a variety of rotor and housing designs for the gas and liquid inlet andoutlet. All of them measured a gas pressure drop from the inlet gas piping orhousing to the outlet gas pipe. In an effort to develop models of the pressure drop,most have attributed the pressure drop to centrifugal and frictional forces in therotor, using a film-flow model (43). Pressure drop is proportional to the square ofrotor speed (26,42). Compared to conventional packed towers, the pressure dropis lower per NTU (26) and about 15 times higher at flooding conditions (28).

To account for the differences in machine configuration and to betterexplain the pressure-drop observations (e.g., lower pressure drop with onset ofliquid flow), a model based on conservation of mass and momentum, in particu-lar gas angular momentum, was developed (40). This model divided the pressuredrop into four increments that included the gas inlet to the machine housing, therotor, the eye of the rotor, and the gas exit nozzle from the machine.

Although the models for pressure drop have a basis in theory, all are fitempirically to data generated from specific equipment. Application of these models may not be relevant to machines of different configurations and packingtypes (9).

In liquid–liquid contactors, pressure drop is defined by the light phase. Theheavy phase enters at near atmospheric pressure and is accelerated by the rotor toits discharge pressure. The pressure drop of the light phase is a function of phasedensity difference, rotor speed, rotor diameter, and location of the principal phaseinterface (24).

3.6. Heat Transfer

Most of the experimental work on heat transfer in centrifugal fields has been doneon spinning discs, which is the subject Chapter 3 in this book. A brief review ofthe enhanced heat transfer rates is relevant here. Studies on a smooth, flat spin-ning disc show heat transfer coefficients as high as about 20 kW/m2K. The co-efficient is highest at the inlet to the disc, due to disturbances as the liquid isaccelerated to the angular velocity of the disc. The heat transfer coefficientincreases with increased rotor speed but decreases with increased radial position.Higher-viscosity fluids decrease the heat transfer rate. Heat transfer is generallyhigher than predicted from film or penetration theory (44).

Modifying the surface of the disc allows enhancement of the heat transferby introducing liquid film instabilities (waves). Best results come from the com-bination of thin films and large instabilities as revealed from the study of four sur-face geometries: smooth, sprayed metal, and two types of concentric grooves. Ingeneral, increased rotor speed gives higher heat transfer. However, with thegrooved disc, high speeds cause liquid separation from the disc, resulting in lowerheat transfer (45).


Heat transfer involving non-Newtonian fluids has not been studied in rotat-ing devices. Models have been developed for gravity-driven heat transfer forpower-law fluids (46). These models may be useful as a starting point to evaluateperformance in higher-gravity fields.

Another method of introducing heat to fluids in rotating devices involvesthe generation of eddy currents by rotation through a stationary magnetic field. This approach was successfully used in a polymer devolatilization process (47).

3.7. Power

Estimating power consumption in rotating systems depends on several factors,including acceleration of the liquid, windage effect of gas drag on the rotor, fric-tion in the bearings and seals, and gas pressure drop (33,48). As would be expec-ted from power consumption in such devices as centrifugal pumps, the largestpower component involves acceleration of the liquid to the angular velocity of the rotor at the outer diameter. Gas pressure drop actually decreases power consumption in the rotor. Frictional losses are defined by the design of the rotor,bearings, and seals.

3.8. Rotor Internals

The rotor packing has an impact on all of the previously mentioned process fun-damentals. Hydrodynamics, especially at the liquid inlet to the packing, is a func-tion of packing porosity and volumetric surface area (12,13). These same packingproperties influence pressure drop, residence time, and flooding velocity (26).Liquid-side mass transfer performance is best with wire gauze as compared toglass beads or parallel flat plates (8). Gas-side mass transfer is better with paral-lel flat plates than in wire gauze (16). Flat plates provide the best medium for heattransfer (34).

4. MECHANICAL DESIGN

Since the development of high-gravity fields requires rotating equipment, themechanical design is very important when considering operating performance,cost of design and fabrication, ease of maintenance, and overall reliability.Although most public reports on RPB studies describe the particular RPB designused in the reported studies, very little information has been published on themechanical design principles. Original equipment manufacturers of rotating equip-ment provide an effective resource for proper design and fabrication of RPBs. Thefollowing discussion outlines some of the basic issues to be considered in themachine design. Overriding all of the following discussion is the need to design astable rotor with minimal vibration under the desired operating conditions.


4.1. Cantilevered Versus Centerhung Rotor

Once the rotor dimensions of inside diameter, outside diameter, and axial heighthave been determined from the operating performance requirements, the rotororientation on the shaft must be determined. Two options are available, can-tilevered, also called overhung, and centerhung. These relate to the position of therotor relative to the shaft and the support bearings. The cantilever design placesthe rotor at the end of the shaft, while the centerhung design positions the rotorin the middle of the shaft, with bearings on either side of the rotor. Often thedetermining factor for selection is the ratio of axial height (AH) to outside rotordiameter (OD). The conservative approach limits cantilever selection to AH/OD �0.5, though designs with ratios up to about 0.85 are possible. Numerous examplesof rotating equipment, such as pumps, compressors, and centrifuges, can be foundfor each design configuration. Figure 5 illustrates the vertical-shaft cantileverdesign; Figure 6 illustrates the horizontal-shaft cantilever design; and Figure 7illustrates the horizontal-shaft centerhung design.

In addition to the rotor dimensions, other considerations for selection ofrotor shaft position include impact on operating performance, cost of manufacture,maintenance, and number and type of seals. The operating performance is notexpected to deviate significantly based on rotor position. The possible considera-tions include rotor imbalance due to flooding of the housing and liquid distributionon the rotor. In general the centerhung design is considered more stable, but it hasa slightly higher cost of manufacture due to the split case housing, is more diffi-cult to maintain, and requires two shaft seals instead of one. Standard equations forfatigue and rigidity are used to determine shaft diameter for both orientations.

4.2. Horizontal Versus Vertical Shaft Orientation

The centerhung design is restricted to a horizontal shaft orientation. A vertical-shaft cantilever design is expected to have slightly lower maintenance costs thanthe horizontal-cantilever design. Both cantilever options should have similardesign and fabrication costs. Flooding of the housing due to insufficient liquiddrainage would be less of a problem with the vertical-shaft arrangement withrespect to rotor imbalance. Liquid distribution on the rotor can be influenced bygravity more on the vertical shaft, but the effect should be minimal. Reference 7illustrates both the centerhung and cantilever horizontal-shaft arrangements anddiscusses an application for use of the vertical-shaft cantilever design.

4.3. Seals

Two types of seals are needed to prevent fluid leakage from the housing and toensure that gas passes through the rotor countercurrent to the liquid. Seals on theshaft as it passes through the housing can be of a design appropriate for the fluids


FIGURE 5 Pilot-scale RPB illustrating the vertical-shaft cantilever design withdirect motor drive. (Photo courtesy of The Dow Chemical Company.)


being handled. Mechanical seals, lip seals, and packing glands are some suitableexamples. As mentioned previously, a centerhung rotor requires two shaft seals,whereas the cantilever rotor requires only one. To seal the rotor to prevent gasbypassing, labyrinth seals and liquid ring seals are options. Figure 1 shows theposition of seals for a vertical-shaft cantilever design.

FIGURE 6 Pilot-scale RPB illustrating the horizontal-shaft cantilever designwith direct motor drive. (Photo courtesy of Higravitec Center of Beijing Uni-versity of Chemical Technology.)


4.4. Power Train

Options for connecting the motor drive to the shaft depend on the shaft orienta-tion. A vertical-shaft cantilever design would prefer a belt drive to reduce the costof manufacture of the support structure and to facilitate maintenance. A horizon-tal shaft has the additional option of direct coupling. Variable speed can beaccomplished through a gearbox or preferably through variable frequency controlon the motor. In addition to the power requirements discussed previously, thestartup power to overcome the torque of the rotor must be considered. This start-up power is related to the time required to reach the desired rotor speed.

4.5. Liquid Distribution

As discussed previously, proper liquid distribution on the rotor is critical to per-formance, but it is also important to prevent rotor imbalance. Rotor imbalance

FIGURE 7 Commercial water deaeration RPB using the horizontal-shaft cen-terhung design and direct motor drive. (Photo courtesy of Higravitec Centerof Beijing University of Chemical Technology.)


from liquid maldistribution is especially a problem with high-viscosity fluids.Liquid distributor pipes extending the axial length of the rotor, depending on the pipe diameter and number, can be a source of additional gas pressure drop.Gas flow around the distributor pipes can also be a source of vibration. In the caseof a vertical-shaft arrangement, the liquid head in the vertical distributor pipemust be considered to ensure equal liquid distribution on the axial length of the rotor.

The rotor can be used to assist liquid distribution either by attaching the dis-tributor pipes to the rotor or by introducing the liquid onto the rotor and allowingthe centrifugal force to move the liquid to the packing. In the former option the liquid must enter the RPB through the shaft, requiring machining of a channel in theshaft and an additional seal. In the case of liquid–liquid extraction, the shaft musthave at least one channel for introduction of one of the liquid phases (see Figure 2).

4.6. Rotor Packing

The selection of the type of rotor packing depends largely on the performancerequirements. However, there are some mechanical design considerations. Examplesof packing include woven wire screen, pellets randomly packed, foam metal, andstructured packing. The materials of construction must have physical properties suf-ficient to withstand the hydraulic forces created by the accelerating liquid. The pack-ing must be dimensionally stable during operation to avoid rotor imbalance issues.

Some packing materials may require supports to keep them in place. Properdesign of the supports will consider porosity to prevent flooding, strength, impacton fluid distribution, and pressure drop.

4.7. Multiple Rotors

Several designs involving multiple rotors have been proposed. To accommodatethe need of additional transfer units in countercurrent gas/liquid contact, a verti-cal shaft with at least two rotors and appropriate internals to conduct the gas fromthe bottom to the top and liquid from top to bottom can be built (5). Another vari-ation allows for heat transfer in addition to the mass transfer. By providing rotorsfor condensation and for boiling and multiple packed rotors for gas/liquid con-tacting, a self-contained distillation column on a single shaft is envisioned (6).Obviously, these multiple-rotor devices are more complex from a mechanicaldesign and construction perspective. However, they offer some interesting possi-bilities for reducing plant size by combining multiple unit operations and addi-tional stages of separation in one piece of equipment.

5. APPLICATIONS

The operating and design principles given previously provide a basis for under-standing the performance enhancement available to a wide variety of applications.


These include standard mass transfer operations, such as absorption and strip-ping, but also include reaction systems. The following section highlights some ofthe applications for which data are available and suggests other opportunities forexploitation of the intensified mass and heat transfer capabilities.

5.1. Absorption

Absorption of a component of a gas stream into a liquid is a common practice inthe chemical industry to affect cleanup of vent gases, conduct chemical reactions,purify products, or to recover products from process streams. The enhanced masstransfer capability of RPBs provides the opportunity to perform absorption pro-cesses in smaller equipment, to lower inventories, to shorten startup and shutdowntimes, and to lower pressure drop (48). Figure 8 provides a visual comparison ofthe size of a conventional absorber tower next to three RPBs that handle theequivalent gas and liquid flows (9).

An example of industrial relevance is the removal of sulfur dioxide (SO2)from vent gases by absorption into water or a lime slurry (48). In the waterabsorption process, both gas-film and liquid-film resistance to mass transferoccurs. As a result the overall mass transfer rate is proportional to gas-flow rateand acceleration but inversely proportional to liquid-flow rate. Due to the fastreaction of SO2 with lime, this system is only gas-film diffusion limited. Theoverall mass transfer rate is largely unaffected by gas- or liquid-flow rate and isproportional to acceleration, but to a lesser extent than the water absorptionprocess. In both cases the overall mass transfer rate is reportedly much higherthan the corresponding conventional packed towers.

In another study of gas-side mass transfer–limited absorption involvingSO2 absorption into a sodium hydroxide solution using a wire screen packing, theoverall mass transfer coefficient was found to be lower than reported data forpacked towers (16). Replacing the wire screen packing with two parallel rotatingplates significantly enhanced the mass transfer performance.

Absorption of hypochlorous acid into water, a liquid-side mass transfer–limited process, showed HTU values as low as 4 cm, with a strong dependenceon liquid-flow rate. Heat of absorption removal was identified as a potential issuewith absorption in rotating beds (9).

5.2. Stripping

Removal of volatile components from the liquid phase to a gas phase has been theobject of much study in RPB devices. One of the early successful applicationswas oxygen removal from water for use in secondary oil field recovery and boilerwater feed (7). The oil field application demonstrated oxygen removal from 6–14ppm to less than 50 ppb in both 50-T/h and 300-T/h RPBs using natural gas forstripping. The packing had 92% porosity and 500-m2/m3 volumetric surface area


FIGURE 8 Commercial use of RPB technology in HOCl process. (Photo cour-tesy of The Dow Chemical Company.)


and was constructed of wire mesh. Comparison with conventional vacuum des-orption in a packed tower combined with chemical reduction agents showedlower cost and equipment size for the RPB approach. The proposed applicationon oil platforms promises further advantages of weight reduction over conven-tional process equipment.

The boiler feed water deaerator reduced oxygen content to less than 7 ppbby the use of exhaust steam (low pressure). This system operated at lower tem-perature (110�C), used lower-pressure steam (0.05-MPa gauge), and achieved theoxygen specification without the use of chemical reducing agents as compared toconventional thermal desorption in a packed tower (7).

The preceding approaches to water deaeration used a gas continuousprocess in the RPB. A liquid continuous RPB has been designed and tested forthis application as well (21). The liquid continuous process allows design of theRPB for reduced power requirements, but it does require higher-pressure gas toovercome the hydraulic head of the liquid. The schematic in Figure 2 shows theliquid takeoff near the eye (inside diameter) of the rotor, thus recovering thepower needed to accelerate the liquid. In the case of oil platform water deaerationusing produced methane gas, boosting the pressure of the available gas would notbe necessary. As with the gas continuous process, mass transfer is enhanced byincreasing rotor speed and increasing gas-flow rate. Sampling at various radialpositions in the polyurethane foam packing revealed the possibility of liquidback-mixing within the rotor that reduced the mass transfer efficiency, i.e., fewertransfer units than expected (21). Further work on the hydrodynamics of the gasand liquid interaction may be warranted in order to realize the full potential of thisenergy-saving approach for stripping (49).

A novel example of stripping in rotating packed beds is the stripping ofresidual monomer and solvent from polymers (47). In polystyrene production,conventional vacuum desorption achieves residuals reduction to about 500 ppm.Steam-stripping technology is available to reduce residuals to about 200 ppm.Compared to steam stripping, the RPB technology is expected to reduce capitalcost, energy costs, and equipment size and to eliminate the potential for side reac-tions of steam with the polymer. A pilot-scale devolatilizer, called an Accelerator,and a larger demonstration unit showed the viability of this approach. Data col-lected at 5–10 mm Hg pressure followed the equilibrium curve for residualstyrene and solvent. This indicates that the mass transfer capability is more thanadequate to achieve equilibrium conditions in the short residence time in therotor. As a further effort to minimize process costs, the devolatilization rotatingpacked bed was combined with a centrifugal pelletizer. Heating of the rotor andpolymer was accomplished by eddy currents generated by placing magnets on eitherside of the rotor. Since the high viscosity of the polymer melt requires higher g-force to achieve thin-film flow over the packing compared to earlier gas/liquidapplications, the packing must have sufficient compressive strength to withstand


the generated forces. A reticulated foam metal packing of high porosity (�90%)and high surface area (�500 m2/m3) was used (47).

Air stripping of volatile organic compounds from groundwater shows thepossibility of using RPB technology for either continuous operation or intermittentremedial operation. Since a wide variety of processes (membranes, air stripping,biological activity, chemical oxidation, and carbon adsorption) are available toremove volatile organics from water, the selection of RPBs will depend on the performance requirements and the relative cost compared to the alternatives.Tests on air stripping of jet fuel components from groundwater show the viabilityof RPB use (33). Both a wire gauze packing and a reticulated foam metal packing proved effective in removing compounds such as benzene, o-xylene,toluene, 1,2,4-trimethylbenzene, and naphthalene. A demonstrated number of trans-fer units as high as 12 gave corresponding height-of-transfer-unit values of 2–3 cm.

Another stripping application actually involves absorption and reaction aswell. Chlorine gas absorbs into sodium hydroxide aqueous solution, reacts to pro-duce hypochlorous acid (HOCl), and is then stripped using excess chlorine gas.The primary measure of performance of this operation is the recovery of strippedHOCl. This study showed the importance of liquid distribution (type of spraynozzle), gas/liquid ratio, and type of packing (wire gauze preferred over glassbeads or flat plates). Above a minimal g-force, little performance improvementwas seen. Low-surface-area wire gauze packing (660 m2/m3) was just as effectiveas high-surface-area (2800 m2/m3) packing (8). Scale-up to commercial operationof this process showed a doubling of the HTU for this gas-side mass transfer–limited stripping. The actual pressure drop in the commercial scale RPB was halfthe expected value. This same correlation, empirically based on centrifugal andfrictional factors of film flow, effectively modeled the pilot RPB (9).

5.3. Distillation

Distillation combines absorption and stripping in one device. Rotating packedbeds perform distillation by use of external condensers and reboilers, as in con-ventional towers (29), or by use of internal heat exchangers as part of the rotor (6).Up to 20 theoretical plates were demonstrated in a rotor of 800-mm diameter (50).Distillation was demonstrated on a 3-tons/h pilot plant separating an ethanol/propanol mixture at total reflux. The pilot plant consisted of two RPBs, one forstripping and one for rectification, along with external reboiler and condenser,respectively (51). Retrofit of existing distillation towers with an RPB has beenproposed as a means to adding separation stages (52).

Another pilot distillation study employed only one RPB along with exter-nal condenser and reboiler. The cyclohexane/n-heptane mixture was separated atrates up to 9 tons/h at total reflux. The system provided up to 6 transfer units(NTU) of separation in a 21-cm packing depth. The primary variable affecting


separation was rotor speed. Two operating pressures and a range of gas loading(reboiler duty) provided additional data for analysis (29,42).

5.4. Heat Transfer

Heat input or removal in rotating systems is best accomplished using plates toseparate the heat transfer fluids from the process fluids (45). Since spinning disctechnology is discussed in Chapter 3 of this book, this section will cover only theapplication of heat transfer in conjunction with rotating packed beds and some ofthe issues related to further development needs.

In the 1950s Hickman developed a centrifugal vapor compression evapora-tor for seawater desalination (53). This device consisted of multiple spinningdiscs. Seawater sprayed on one side of the disc evaporated, while the centrifugalforce removed the residue from the plate surface. The vapor was compressed andreturned to the opposite side of the plate, where condensation provided the heatfor evaporation and the desired freshwater for recovery. Overall heat transfercoefficients of 18 kW/m2-K are about three times higher than those achieved insteam turbine condensers.

A high-intensity heat pump, called Rotex, has been developed takingadvantage of the enhanced heat and mass transfer performance of rotating discs(44). This single device carries out the processes of evaporation, condensation,absorption, and heat transfer to a working fluid.

The higher heat transfer coefficients experienced by Hickman led to theconcept of placing a peripheral reboiler and core condenser on either side of arotating packed bed (50). This concept would be useful for distillation applica-tions that need reflux and boilup. The internal exchangers as part of the rotorwould decrease the required heat transfer surface area but would involve add-itional design and fabrication complexity.

Although heat exchangers on either end of a packed rotor are an option forreplacing external heat exchangers for distillation, the problem of heat transferwithin a porous packed bed remains. Heat input can be achieved by use of eddycurrents (47), microwaves, or sonic energy. Thus operations such as evaporation,stripping, and endothermic reactions can be envisioned. Heat removal, however,is more problematic. Exothermic reactions must be conducted adiabatically with-in the rotor, unless a suitable means of extracting the heat of reaction can bedeveloped. One approach could be alternating packing and heat transfer plates.This raises the complexity of design and fabrication but could provide the neededcooling to approach isothermal operation. A simpler method of evaporative cool-ing is possible if the evaporation is compatible with the chemical process.

5.5. Adsorption

Centrifugal adsorption technology (CAT) allows the use of very small adsorbentparticles (microns) to increase the mass transfer efficiency. Application to ion


exchange, volatile organic removal from water, recovery of pharmaceutical pro-teins, and production of fine chemicals are examples of potential commercialinterest. Process advantages include low inventory, low contact time, steady-stateoperation, and relatively small equipment (54).

The CAT mode of operation involves introduction of the adsorbent near theaxis of the rotor, allowing the centrifugal force to move the particles radially out-ward. Liquid introduced at the outer periphery of the rotor moves countercurrentto the adsorbent and is removed at the axis of the rotor. Adsorbent slurry collectsat the periphery and is conducted to the rotor axis for discharge. Experimentsusing activated carbon to adsorb n-butanol from water revealed that the degree ofback-mixing is the dominant factor in performance. Back-mixing is a function ofrotor speed, density difference between the phases, and the particle diameter (54).

The hydrodynamics of two-phase flow in CAT were compared to two-phaseflow under gravity using a large-diameter (1.3-mm) particle in water with a smalldensity difference and a small-diameter (81.8-micron) particle in water with alarge density difference. The throughput capacity of the CAT was higher than pre-dicted from the homogeneous-flow model, though the model works well for thegravity-flow column. Pressure drop estimates were used to predict void fractionsin the range of 0.7–0.8. Higher rotor speeds resulted in higher void fraction (55).

5.6. Liquid–Liquid Extraction

The use of centrifugal fields for liquid–liquid extraction was perhaps the firstcommercially successful application of rotating packed beds. Podbielniak modi-fied a patented vapor–liquid contactor (2), using a perforated spiral passagewayas the rotor packing, to solve problems with penicillin recovery in 1945 (1).Penicillin broth forms stable emulsions that require centrifugal force to break.Solvent extraction was effective only at low pH, which caused penicillin deg-radation. Multiple stages were needed to affect the necessary concentration. Inaddition, the fermentation liquor varied significantly from batch to batch and plantto plant. Conventional countercurrent solvent extraction, mixer-settlers, and mixer-centrifuge combinations could not effectively solve these problems without prod-uct loss. The centrifugal solvent extractor achieved 98% product recovery by takingadvantage of its low liquid holdup, short residence time, high centrifugal force,and multistage countercurrent contacting.

Continuous glycerin washing of soap produced by saponification has beendemonstrated in a countercurrent centrifugal extractor (38). The device achievesphase separation with as little as 0.02 specific gravity difference and accomplishesup to 10 theoretical stages of extraction. Some of the advantages over prior opera-tions reportedly include flexibility in feed, low holdup, less waste due to more efficient separation, simple operation, rapid startup, and small space requirements.

Rotors filled with ceramic foam instead of perforated cylinders have beentested for liquid extraction of trace contaminants from water (23). The test solution


used a C12 alkene to extract 1,2-dichloroethane from water. Flow patterns are sim-ilar to the previous applications, with the heavy phase introduced at the innerdiameter of the rotor and the light phase at the outer diameter. Countercurrentflow is achieved as the heavy phase moves outward, displacing the light phaseinward. The results indicate an optimum pore size for the ceramic foam, decreas-ing height of transfer unit with increasing rotor speed, and increasing holdup withincreasing dispersed phase flow.

The ability of the centrifugal extractor to solve difficult liquid–liquid separ-ation problems, as illustrated in the previous examples, has allowed its use in awide range of extraction applications. The long history of use has given it a gen-eral acceptance in chemical manufacturing—an acceptance not shared by thebroader application of gas–liquid interactions.

5.7. Crystallization

In the reactive precipitation process of reacting CO2 with Ca(OH)2 slurry to pro-duce nanoparticles of CaCO3, the controlling steps of the process are absorptionof CO2 and dissolution of solid Ca(OH)2. The degree of supersaturation dependson the reaction rate and controls the nucleation rate and, therefore, the particlesize. The intense mass transfer and micromixing capability of the rotating packedbed provides the environment to produce CaCO3 particles of size 15–30 nm witha narrow size distribution. Reaction time reduces 4- to 10-fold, compared tostirred-tank reactors. Rotor speed, gas–liquid ratio, and initial calcium hydroxideconcentration influence reaction rate. An increase in rotor speed reduces the aver-age particle size. Addition of growth inhibitors also helps to control particle sizeand size distribution (56).

High-gravity reactive precipitation (HGRP) has been extended to the pro-duction of aluminum hydroxide and strontium carbonate (57). Aluminum hydroxidefibrils precipitate from the reaction of sodium meta-aluminate (NaAlO2), water,and carbon dioxide and are formed in diameters of 1–10 nm and lengths of50–300 nm. Rotor speed, gas- and liquid-flow rates, and initial reactant concen-trations control particle size. Strontium carbonate particles of 40-nm mean dia-meter and narrow size distribution have been produced from the liquid–liquidreaction of strontium nitrate and sodium carbonate.

Crystallization that occurs during evaporation can potentially be intensifiedby use of vapor recompression and spinning discs. In this scenario, the evapor-ated vapor is compressed and then condensed on the bottom of the discs to heatthe crystallizing fluid (58). This approach may permit operation at higher tem-peratures, lower surface area, and less time.

Recrystallization of an active pharmaceutical ingredient on a spinning disc,employing a solvent/antisolvent approach to induce rapid precipitation, results inthe desired small particles (1–15 microns) and narrow particle size distribution (59).


5.8. Reactions

A significant number of chemical reactions of commercial interest have reactionrates limited by heat or mass transfer rather than the intrinsic chemical kinetics.Many of these processes could be intensified by application of the RPB or spin-ning disc technology. Indeed, application of centrifugal fields to reacting systemsmay represent its greatest potential. Improved yield performance may also bepossible (50). Spinning disc technology provides the advantage of heat transfercapability. Thus, temperature control of exothermic reactions such as styrenepolymerization and intense mixing allow reduction in reaction time compared toconventional batch reactors, especially in the latter stages of reaction, where vis-cosity is higher (60,61).

Polycondensation reactions that are equilibrium controlled, such as poly-esters, could benefit from the thin films generated in the RPB or spinning disc. Inthese reactions, removal of the coproduct of polymerization, e.g., ethylene glycol,is necessary in order to advance the polymerization, a task that becomes increas-ingly difficult as the reaction proceeds. The thin films and short residence time ofhigh-gravity devices aid the evaporation and may permit operation at higher tem-peratures than conventional reactors (50).

As mentioned previously, RPB and spinning disc technology may providebenefits for reactions that are mass or heat transfer limited, i.e., for fast kineticreactions. Unfortunately, the true chemical kinetics are often unknown. A small-scale RPB or spinning disc may prove to be a useful screening tool to determinethe intrinsic chemical kinetics. In one such study, six different reactions of inter-est in the manufacture of pharmaceuticals were screened using a spinning discreactor (59). Three of the six reactions were found to be limited by liquid–liquidmixing. These include a phase-transfer-catalyzed Darzen’s reaction, a crystalliza-tion, and a highly exothermic condensation reaction. In the Darzen’s reaction thereactant inventory was reduced 99% and the impurity level decreased 93% ascompared to the conventional reactor. Crystallization achieved mean crystal sizeof 3 microns, with a narrow size distribution. The highly exothermic reaction hadexcellent temperature control.

A lab-scale RPB has also been used to investigate reactions that are masstransfer constrained in conventional reactors. Testing a reaction that involvesrelease of a volatile organic as part of molecular weight buildup revealed overallprocess reaction rates equivalent to conventional reactors in time frames morethan two orders of magnitude lower. Here the high surface area and surfacerenewal capability of the RPB helped to overcome the transfer limitations acrossthe liquid boundary. Similar results were seen on other reaction processes thatwere constrained by a liquid–surface interaction (62).

Reactions can be combined with other unit operations, as in the example ofreactive stripping in the production of hypochlorous acid (HOCl). An RPB was


used to absorb chlorine gas into an aqueous sodium hydroxide solution (caustic)where reaction to HOCl occurs. Since HOCl is unstable in the presence of thecoproduct sodium chloride, stripping of the HOCl into the gas phase is necessaryto recover a stable product. The reaction of chlorine with caustic is essentiallyinstantaneous: therefore the process reaction rate is liquid-side mass transfer con-trolled. Stripping of HOCl, on the other hand, is believed to be gas-side masstransfer controlled. The intense mass transfer capability of the RPB allowed 10%higher yields while using less than half the stripping gas as compared to conven-tional spray tower operation. This study showed low-surface-area, high-porositywire screen packing to perform better than glass beads or parallel flat plates.Packing support design and liquid distributor type influenced performance.Operating parameters of importance included rotor speed and the gas–liquid ratio(8). Scale-up issues and performance of the commercial HOCl operation (9) arediscussed in Section 6.

Fermentation reactions are often limited by oxygen transfer rates. Theenhanced mass transfer achieved in centrifugal fields applied to bioreactions shouldbe expected to increase productivity. A centrifugal field bioreactor (CFBR) demon-strated higher productivity in the overproduction of lipase with Staphylococcuscarnosus as compared to conventional fermenters (22). Both batch and semibatchfermentation in the CFBR showed no influence on the biological activity ofgrowth or exoprotein synthesis. Lipase productivity rates were proportional tooxygen transfer rates, which were 10 times higher than in shaken cultures. TheCFBR process involved feed of air and liquid to the outer diameter of the rotor,with takeoff at the center. Air was dispersed in the liquid by either a sieve drumor a multilayer-sintered screen. The inward radial movement of the gas helped tosuspend the bacteria in the culture against the centrifugal force. An external cir-culation loop for the liquid allowed heat exchange and product analysis. Sincemany fermentation reactions are characterized by foaming, the CFBR wasequipped with a foam breaker—a stator with needles positioned at the insidediameter of the rotor.

5.9. Other Applications

Centrifugal fields in an electrochemical cell facilitate the removal of gas bubblesfrom the electrodes, thus reducing the voltage requirement. A rotating chlorine cellshowed a drop in voltage from 3.17 V to about 2.8 V at 3-kA/m2 current densitywhen accelerated to 200 g (50). Demonstration of a rotating air cathode providedgreater voltage drop at higher current density as compared to a stationary cell (50).

Dedusting or demisting in rotating devices provides opportunity to removesmall particles at very high throughput. A “mop-fan” built with flexible fibers ina conventional fan housing effectively removes 50% of two-micron particles of slaked lime dust in conjunction with water spray on the rotor. Inline rotary


demisters employing reticulated metal foam are used for oil separation from airon turbines (50).

6. SCALE-UP AND COMMERCIAL USE

Although considerable studies on the use of centrifugal fields in chemical pro-cessing have been reported for lab- and pilot-scale operations, little informationis public on scale-up criteria for either performance parameters or equipmentdesign. Three examples of commercial use of centrifugal fields are available forreview. These include liquid–liquid extraction, water deaeration, and reactivestripping for hypochlorous acid production.

Commercial application of centrifugal fields encounters considerableresistance from both the technical and business communities due to both real andperceived risks. The real risks involve reliable mechanical design of rotatingequipment, which includes seals, bearings, and rotor stability. Perceived risks on process performance may derive from a lack of understanding of the processfundamentals and how performance may change with the scale of operation.Overcoming the tendency to “use what we know and understand” represents achallenge that goes beyond the effort at technology development.

Convincing the technical community to accept the risk of rotating equip-ment for chemical processing may be easier on applications with clear perform-ance advantages over conventional process equipment. A good example is theliquid–liquid solvent extraction of penicillin (1), in which the low residence timeand ability to handle emulsions and solids allowed 98% product recovery. The10% higher yields and 50% reduction in stripping gas for the HOCl reactive strip-ping process provides another example of performance advantage (8).

In addition to the lower operating costs associated with enhanced perform-ance, the business community is interested in lower capital investment and assur-ance that the process will reliably perform as designed in terms of productcapacity, on-stream time, and product quality. The smaller size of the centrifugalequipment may satisfy the capital investment question. This was the primarydriver for implementation of water deaeration in China (7). Lower capital is alsoa driver for the polymer devolatilization application (47). The question of reliableperformance is best addressed through convincing the technical community andleveraging the considerable industry experience with design and manufacture ofrotating equipment, such as pumps, compressors, and centrifuges.

6.1. Scale-Up Criteria

A number of parameters can be considered for scale-up of rotating packed beds,including rotor packing, liquid distribution, flooding, pressure drop, rotor speed,HTU, NTU, temperature, and pressure. Since the same packing material (same


porosity and volumetric surface area) is used on scale-up as in the pilot equip-ment, HTU is expected to scale directly. Results from water deaeration (7) andHOCl production (9) indicate this may be true for liquid-side mass transfer–controlled operations but not for gas-side mass transfer–controlled operations.

Flooding of the rotor in gas continuous operations would be expected at thesame gas velocities as in the pilot RPB. Flooding normally occurs at the insidediameter of the rotor due to the lower cross-sectional area and higher velocities.However, a check of the porosity of the outer packing support may be necessaryto ensure that no flooding occurs. If scale-up of rotor speed is based on constantrotor tangential velocity instead of constant g-force, then the gas velocity at flood-ing will be lower with larger-diameter rotors. Throughput capability and back-pressure control of the light-phase takeoff (24) control flooding in liquid–liquidor gas-dispersed systems.

Using a pressure drop model based only on centrifugal acceleration andfrictional drag, the HOCl scale-up overpredicted the pressure drop of the com-mercial RPB by a factor of 2 (9). A more rigorous approach to pressure drop cal-culation that takes into account the conservation of angular momentum and theinlet and outlet zones of the rotor and housing (40) should provide more pre-dictable scale-up performance.

Rotor speed has an impact on mass transfer performance, flooding, andpressure drop. Rotor speed on scale-up can be determined based on maintainingconstant tangential velocity (r�) or constant acceleration (r�2). Rotor speed willbe higher for constant-acceleration scale-up. Impact on both process performanceand equipment design must be understood in making this determination. Scale-upbased on constant acceleration is conservative for mass transfer and flooding per-formance, while constant tangential velocity is conservative for pressure drop.

Liquid distribution may be an important parameter, as demonstrated in theHOCl process, where different liquid distributors provided significantly differentresults (8). The initial contact of the liquid with the rotor influences the masstransfer performance of the RPB in gas continuous operations (15). Although theuse of a packing support at the inside diameter of the rotor would be expected toimpact this initial liquid contact with the rotor, experiments did not show anyreduced mass transfer performance (36).

As mentioned earlier, the same rotor internals used in pilot tests should beused upon scale-up. The rotor dimensions of inner diameter and axial height are determined by maintaining a constant superficial gas velocity at the rotor eye.The radial packing depth, and thus the outer diameter, is based on the number oftransfer units required. Adjustments in packing depth and packing type may benecessary to achieve the desired liquid holdup or residence time, e.g., for chemi-cal reaction (26).

As with any chemical operation, the physical properties of the fluids, suchas density, viscosity, and heat capacity, must be known. If chemical reaction is


involved, then the intrinsic kinetics must be understood. Of particular relevanceis the influence of mass transfer rates on overall reaction rates.

6.2. Scale-Up Design

Once the rotor dimensions have been determined along with the operating condi-tions, attention shifts to the mechanical design of the RPB. The major concernwith rotating equipment is maintaining stable operation, i.e., limiting vibration.Excessive vibration results in premature seal and bearing failure, poor processperformance, metal fatigue, and increased maintenance costs and downtime.

Proper mechanical design principles determine the option of cantilevered orcenterhung rotor, the shaft diameter, the type and position of bearings, and sealdesign. The drive train, whether belt driven or direct coupled, is determined bythe power requirements and the shaft orientation. The housing must be sufficientto contain the temperature and pressure of the operation and to provide adequateinlet and outlet nozzles for the process fluids.

6.3. Commercial Examples

Two commercial examples of rotating packed-bed operation are water deaerationfor the Chinese oil fields (7) and HOCl reactive stripping in the United States (9).These two cases illustrate nicely the range of process conditions and design features available for successful scale-up. Water deaeration (Figure 7) uses adirect-coupled drive on a horizontal-shaft centerhung rotor to process a low gas-to-liquid operation (�3 : 1 vol/vol). The HOCl process (Figure 9) employs a belt-driven, vertical-shaft cantilever rotor to contact a high gas-to-liquid ratio.

The water deaeration process employed a staged scale-up program. From the lab operations, a 50 tons/hour (T/h) pilot RPB was built and tested in the oilfield. Using natural gas for stripping at a gas/liquid ratio of over 2, the desired oxygen content in the exit liquid of less than 50 ppb was demonstrated. This suc-cessful demonstration led to the installation of a full-scale commercial RPB toprocess 300 T/h. This unit has rotor dimensions of 600-mm ID, 1000-mm OD, and700-mm AH. The wire screen packing has high porosity (92%) and low surface area(500 m2/m3). The rotor spins at a modest 750 rpm. Performance matched that of thelab and pilot units, achieving a typical 30-ppb oxygen content. Two 250-T/h unitshave been designed for installation on an oil platform to process seawater, provid-ing advantages in size and weight as compared to conventional technology (7).

Figure 9 shows the commercial-scale RPB for the reactive stripping processfor HOCl production. Figure 8 provides a visual impression of the process intensi-fication that occurs using RPBs. The three RPBs shown in the lower left of thepicture process the same volume of gas and liquid as the tall absorber tower to theright. The scale-up factor from the pilot unit of over 400 : 1 yielded a rotor ofslightly less than 2 m in diameter. Performance of product yield from raw materials


met or exceeded that of the pilot unit. Pressure drop half of that expected impliesthe need for better predictive correlations. Due to the higher gas-handling capa-bility and the conservative scale-up design, much higher capacities are anticipatedas compared to the design. The liquid-side mass transfer performance as meas-ured by chlorate formation showed performance equivalent to or better than thatof the pilot RPB. However, the gas-side mass transfer, as represented by HOClstripping, showed a doubling of the HTU to about 8 cm. The mechanical relia-bility after two years of operation indicates no issues due to RPB operation. TheRPB is very easy to start up and shut down (9).

These two successful commercial applications of rotating packed bedsprove that scale-up from pilot-scale equipment can achieve the desired processperformance in commercial-scale operations. In addition, the mechanical reli-ability of the rotating equipment is in line with the experience with other rotating

FIGURE 9 Commercial RPB for HOCl production using the vertical-shaft cantilever design with belt drive. (Photo courtesy of The Dow ChemicalCompany.)


devices. Thus, the risk concerns of performance and reliability can be managedacceptably. Both the technical and business communities can have confidencethat future applications will meet expectations of process performance and on-stream time.

7. FUTURE

Extension of the use of centrifugal fields into chemical processing, beyond thephysical movement of fluids, has shown limited niche application in the past, inspite of considerable research activity. Specialty application of centrifugal fieldsto liquid–liquid extraction has enjoyed success for more than 50 years. Advant-ages stem from operation at low density differences, breaking of emulsions,short contact times, and higher efficiencies as compared to other liquid–liquidextractors.

The commercial use of rotating equipment for the broader field of gas–liquid operations has only a five-year history. Numerous examples of possibilitiesin the areas of absorption, adsorption, stripping, distillation, reactions, crystalliza-tion, and other operations have been referenced. The chief objections to the use ofcentrifugal fields have been associated with the risks of scale-up and the operationof rotating equipment. The two commercial applications of water deaeration andHOCl reactive stripping demonstrate the ability to reliably scale up processesinvolving a wide range of gas–liquid loadings. Process performance in both casesmet or exceeded design criteria, with good operating reliability.

Further application will likely require significant cost or performanceadvantages over more conventional process technology. Considerable commer-cial experience will be needed before centrifugal fields will enjoy commonacceptance among both technical and business interests in the chemical industry.To gain that status, projects must be selected carefully to ensure that advantagesare realized over alternative technologies.

The most likely opportunities for exploitation will come from mass transfer–limited reactions and the combination of unit operations in one device. Examplesof reactions mentioned earlier include polymerization, condensation reactions,crystallization, and heterogeneous catalysis. Combined unit operations are illus-trated by reactive distillation, polymer devolatilization with pelletization, and theuse of heat exchangers (reboilers and condensers) with distillation.

In addition to research on process applications, research to define the fun-damental performance characterizations is needed. A number of empirical corre-lations have been developed for pressure drop, residence time, power, flooding,etc. More generalized theoretical expressions for these parameters that accurate-ly predict performance on a wide range of rotor designs and sizes would be verybeneficial to confidently scale-up the technology.


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3

The Spinning Disc Reactor

C. Ramshaw

University of Newcastle upon Tyne, Newcastle upon Tyne, England

1. INTRODUCTION

When considering the options for intensifying reactions that involve multiple fluids, it is helpful to identify the shortcomings of the conventional equipmentthat is currently in use. In this context, perhaps the most frequently used item isthe stirred vessel fitted with a cooling jacket, shown in Figure 1. A turbine impellergenerates a circulation comprising two toroidal vortices, and the turbine torque isnormally prevented from driving a free vortex by the use of wall baffles, asshown. If a gas–liquid reaction is involved, then the gas is usually injected directlybelow the impeller via a suitable sparging arrangement. The popularity of thestirred vessel is due to its perceived simplicity and adaptability, coupled with thefact that it is superficially straightforward to scale-up from the laboratory beakerthat was used when the process was being developed. Unfortunately, it suffersfrom several serious problems, as indicated later.

In the normal case of a geometrically similar scale-up, it can be readilyshown that the surface area per unit volume varies inversely with the vessel dia-meter. Thus larger vessels are more difficult to cool, since the heat generated bya reaction in a potential runaway situation is proportional to the vessel volume,whereas the surface area available to dissipate a given heat output is decreased.Vigorous reactions may require the reactor to be “detuned” by operating withmore dilute feedstock in order to reduce the full-scale reaction intensity. This


could influence the reaction temperature trajectory and compromise the yield andselectivity.

A further unfortunate characteristic of the stirred vessel is that its mixingcapability is also a strong function of its size. Scale-up usually proceeds on thebasis of a constant impeller tip speed, and since the mean circulation speed in thevortices is broadly proportional to the tip speed chosen, the circulation time isproportional to the vessel diameter. Thus the turnover time of the vessel contentsincreases at the larger scale and the macro mixing performance deteriorates.

These fundamental shortcomings of the stirred vessel have generated a con-siderable degree of uncertainty when fine chemical or pharmaceutical processesare being developed for full-scale operation. This has led the relevant regulatingauthorities, e.g., the U.S. Food and Drug Administration, to insist on a processvalidation at laboratory, pilot, and full scale. Since each validation entails signifi-cant administration and delay, the procedure can hold up the implementation ofcommercial production by several years. Because a new metabolically activemolecule will be patented as soon as possible and certainly before clinical trialsand process development, this delay significantly erodes the time available underpatent cover to recoup a company’s R&D expenditure and make a profit from apotential “blockbuster” drug.

1.1. The “Desktop” Continuous Process

The predominant culture that prevails for the production of fine chemicals/drugs,with an output of up to (say) 500 tons per year, is to operate batchwise. As alreadynoted, this stems from the fact that the process is almost always developed froma batch-operated beaker or flask. However, it is worth observing that an output of500 tons/year of active substance corresponds to a continuous process flow rateof around only 70 mL/second. This allows various items of intensified equipment

FIGURE 1 Stirred vessel showing circulation pattern.


to be assembled and operated continuously literally on a “desktop” to meet theproduction demand. The decision to switch from batch to continuous processingimmediately confers an intensification benefit because the peak batch processloads (e.g., heat output, liquid removal, etc.) are distributed in time, so equipmentsize can be reduced. Thus with a new process, the laboratory scale becomes thefull scale when allowed to run continuously and the scale-up delays describedearlier are largely avoided. This strategy is generating considerable industrialinterest as the commercial pressure to bring new molecules to market rapidly con-tinues to increase.

A further factor that favors continuous “desktop” manufacture is its poten-tial impact on the overall business process of making and marketing fine chem-icals. Thus it goes without saying that with very short process residence times, theoperation can be much more responsive so that grade changes can be effected inseconds rather than hours. This facilitates just-in-time manufacture, which canlead to dramatic reductions in the capital costs associated with the multiple gradesof stock that may be needed rapidly to satisfy demanding customers.

1.2. Exploitation of Centrifugal Fields

Approximately two-thirds of the unit operations performed in process engineer-ing involve multiphase contact (e.g., distillation, gas/liquid reaction, boiling). Inthe absence of an imposed acceleration field, the system fluid dynamics are dom-inated by surface forces so that the interfacial area developed is relatively small,and, with no buoyancy force, there can be no countercurrent interfacial motion.When these conditions prevail, the intensity of the operation is very low, with lit-tle if any process performance (e.g., reaction, separation, heat transfer) beingexhibited. This scenario leads naturally to the suggestion that a high-accelerationfield would stimulate the generation of smaller bubbles, higher flooding veloci-ties, and more intense shear stresses.

This “Higee” strategy has been championed over many years because of itsprofound and beneficial impact upon many important multiphase operations:

AbsorptionDistillationBoilingCondensationLiquid extractionParticle disengagementHeat pumpsEtc.

One particular embodiment of this approach is the rotating packed bed,which was originally conceived as the “Higee” equivalent of a packed column (1),


shown schematically in Figure 2. The equivalent stage height within the toroidalpacking is about 1.5 cm for a gas film—limited system, compared with about 60 cm in a conventional packed column. Equivalent flooding velocities may be estimated from the Sherwood plot and can be very high, even for packing witha specific surface area exceeding 1000 m�1. Since the Higee duty was origi-nally envisioned as being purely orientated to mass transfer, no specific heattransfer capability was provided. However, the spinning disc reactor (SDR) maybe regarded as an alternative to the “Higee” rotor. It can act as a mass transfer/contacting device (possibly with multiple discs) or as a particularly intensegas–liquid reactor (when fitted with heating/cooling provision). Its attraction liesin the high heat and mass transfer rates that can be stimulated between the discand the thin liquid film generated on its surface, and between the film and theadjacent gas. The performance and applications of the SDR are considered indetail later.

As might be expected, the enhanced acceleration field is established on apermanent basis within a rotor that receives and discharges the working fluid. Thealternative approach, which relies upon a permanent vortex field, i.e., a cyclone,

FIGURE 2 The “Higee” contactor (continuous gas phase).


will not be considered here, primarily because the fluid residence time within avortex cannot be easily controlled independent of the desired acceleration. At thisjuncture it is helpful to consider the basic physics of motion in a rotating system.

1.3. Free Motion of a Particle Around an Axis

Consider the free (frictionless) motion of a particle P of mass m rotating arounda fixed axis O on a smooth surface, as shown in Figure 3. The particle is con-strained to move in a circular trajectory by a light string that exerts an inward ten-sion T and generates a corresponding acceleration. This can be estimated asfollows: In time �t, P moves along an arc that subtends an angle �� to the axis.The angular velocity of P is given by � � d��dt and its speed is v � r�, wherer is the length of the string. During the time �t the change of velocity of P is�v � v sin (��) and as �� → 0, �v is directed along the radius toward O.

The acceleration is

Hence,

acceleration � r�2

and the string tension needed to maintain this is

T � m�2r

dv

dt

v d

dtv� �

sin ( )

� �as � → 0

FIGURE 3 The free motion of a particle around an axis.


If the string is broken, then the acceleration ceases and the particle leaves its cir-cular trajectory and continues along a tangent at a velocity v.

If the string is slowly shortened then:

1. The particle moves toward O in a spiral trajectory involving many turns.2. Work must be done in order to overcome the string tension T.

For the proposed frictionless system, this work input results in an increasein the particle’s kinetic energy. Conversely, if the string were lengthened, the par-ticle’s velocity would decrease. Figure 4 shows the spiral trajectory of the parti-cle and the corresponding velocity diagram. The radial and tangential velocitycomponents are dr�dt, v, respectively, giving a resultant VR that is a tangent to thespiral trajectory.

Noting that dr�dt << v, the angle between v and VR is �, wheretan � � (dr�dt)�v. As the particle moves inward towards O, the component of Talong the spiral trajectory is responsible for increasing its speed.

Hence, with an inward tension deemed to be negative we have:

or

(1)

This confirms that v increases for an inward spiral trajectory (i.e., whendr�dt is negative). From Eq. (1) we have

and integration from v1r1 to v2r2 gives

dv

v

dr

r��

dv

dt

dr

dt

v

r�� ⋅

mdv

dtT m

v

r

dr

dt v� ��sin ( )� �

2 10⋅ ⋅ →as

FIGURE 4 Spiral trajectory as a particle moves toward the center of rotation.


i.e.,

(2)

This is an important result because it shows that in a frictionless or free situation,angular momentum is conserved.

The same conclusion can be reached from energy considerations, notingthat the work done by the string tension as r is reduced is equal to the gain in par-ticle kinetic energy, i.e.,

Hence

leading again to Eq. (2).The practical consequence of Eq. (2) for rotational fluid flow can be quite

dramatic, as demonstrated by the high wind speeds that may be generated nearthe center of free vortex flows—e.g., tornadoes and typhoons.

1.4. Flow Over a Rotating Surface

If we now consider the behavior of a liquid film on a rotating disc, the motion isno longer “free,” because the film is influenced by the disc via the shear forcegenerated at the solid/liquid interface. Liquid supplied to the inner region of thedisc is first brought up to the disc’s rotational velocity by the tangential shearforce and then moves radially outward, to be discharged from the disc periphery.In a stationary frame of reference, the liquid trajectory is therefore a spiral witharms separated by a radial distance given by

In a rotating frame of reference (i.e., that of an observer anchored to the disc),the trajectory is nearly radial. Since this reference frame is most relevant when we consider the disc/fluid interaction, it is helpful to evaluate the flow on thisbasis.

We shall assume that frictionless flow occurs through a closed radial chan-nel that is fixed to the rotating disc. Thus as the fluid moves outward, the only

2

⋅ dr

dt

� �v

rdr dv

� �� T dr mv

rdr d m

vmv dv

2 2

2

v r v r2 2 1 1�

ln lnv

v

r

r2

1

1

2

�


interaction with the disc/channel is via a tangential force. In a time dt, the changeper unit mass in the fluid angular momentum as it moves a distance dr is

(3)

Note that since the fluid is forced to rotate at the disc speed, � is constant.The change in angular momentum given by Eq. (3) is brought about by the torquegenerated by a tangential force F exerted on the fluid at the radius r by the radialchannel and acting for a time dt. Since dM � Fr dt, the tangential accelerationimposed by the channel on the fluid as it moves radially is

(4)

This is the Coriolis acceleration, which is imposed on particles moving in a rotat-ing reference frame—e.g., liquids on a rotating surface or winds in the earth’satmosphere. The resultant acceleration experienced by a particle is a combina-tion of the radial and tangential components, making an angle � to the radius,where

In general

and the unconstrained flow will be largely radial. This can be readily confirmedduring the operation of a spinning disc reactor or a rotating packed bed, becauseany deposits from the fluid flow lie close to the radius vector.

2. THE SPINNING DISC REACTOR

As pointed out earlier, a spinning disc, or more generally a rotating surface of revo-lution, is an alternative to the “Higee” rotating packed bed. It is particularly effec-tive when high heat fluxes or viscous liquids are involved. The object is togenerate a highly sheared liquid film when a liquid is supplied to the unit at ornear its center. The film is initially accelerated tangentially by the shear stressesestablished at the disc/liquid surface. This causes the liquid to approach the disc’sangular velocity and then move outward as a thinning/diverging film under theprevailing centrifugal acceleration. The phenomenon was studied in detail byWoods (2), who photographed the behavior of a fully wetting dilute film of ink as

dr

dtv<<

tan � �� 222� �

dr

dtr

v

dr

dt

F

m r

dM

dt

dr

dt� �

12

dM d r r r dr� �( )� �2


it traveled over a spinning glass disc. Care was taken to supply the liquid from acentral axisymmetric distributor in a particularly uniform manner. After calibration,the local film thickness was inferred from the density of the photographic imageat that point.

Even when great care was taken to ensure that the liquid feed was intro-duced to the disc in an axisymmetric manner with the minimum disturbance, thesmooth inner film always broke down into an array of spiral ripples, as shown inFigures 5 and 6. These spiral structures then broke down further until the wavepattern became utterly chaotic, provided that the disc was big enough. It is knownthat liquid film flow over a surface is intrinsically unstable, and the phenomenonhas been studied by several workers (3–7). It appears to be qualitatively equiva-lent to the breakdown of a smoke plume rising from a lighted cigarette, where achaotic zone is generated about 20 cm above the source. The behavior can also beobserved when a liquid film flows over a stationary surface such as a windowpaneor a dam spillway.

Woods concluded that two types of wave existed: nearly two-dimensional(2D) and three-dimensional (3D). The amplitude of the two-dimensional spiralwaves grew rapidly, and therefore a theory based on the assumption of small ampli-tude is not valid across the whole disc. A transition from 2D to 3D waves occurredonce their amplitude reached about three to four times the local film mean thick-ness. Higher liquid flow rates stimulated a more rapid breakup of the wavelets.Only about 1% increase in liquid surface area was ascribed to the presence ofwaves. Thus any improvement in mass/heat transfer performance generated bythe waves is due to the additional shear they induce. It will be appreciated thateven in the absence of ripples, highly sheared thin liquid films, such as those thatcan be readily generated on a spinning surface, provide an ideal fluid dynamicenvironment for the rapid transmission of heat, matter, and momentum. This isdue to the short diffusion path length involved for transfer between the adjacentgas phase to the liquid film and thence to the disc surface. These characteristicsof a spinning disc (or more generally a rotating surface of revolution) make itideal for performing any intrinsically rapid physical or chemical transformationin a liquid, even if it is viscous. Typical examples include polymerization, pre-cipitations, and rapid exothermic organic reactions. Some of these are describedin more detail later.

2.1. The Nusselt-Flow Model

While the fluid dynamics of the actual film-flow process across the disc is daunt-ingly complex, a very approximate interim flow model may be based uponNusselt’s treatment of the flow of a condensate film. This assumes that the flow isstable (i.e., ripple free), that there is no circumferential slip at the disc/liquid sur-face, and that there is no shear at the gas/liquid interface. The treatment is based


FIGURE 5 Liquid film behavior on a rotating disc, with Q � 19 cm3/s and � (a) 100, (b) 200, (c) 300, (d) 400, (e) 500, (f) 600 rpm.


FIGURE 5 (cont.)


FIGURE 6 Local liquid film behavior on a rotating disc, with (a) Q � 19 cm3/s, � 100 rpm; (b) Q � 19 cm3/s, � 200 rpm; (c) Q�13 cm3/s, � 400 rpm;(d) Q�19 cm3/s, � 500 rpm; (e) Q�19 cm3/s, 600 rpm; (f) Q�19 cm3/s, � 600 rpm.


FIGURE 6 (cont.)


on the schematic representation given in Figure 7. Part b represents the local filmat a radius r. The shear stress on the annular plane at a distance y from the disc provides the radial acceleration for the fluid lying between y � y and y � s. Thus aforce balance on the film lying between r � r and r � r � dr, with zero shearstress at the gas–liquid interface, gives

(5)

The boundary conditions are:

1. u � 0 at y � 0 since there is no fluid slip at the disc/liquid interface.2. du�dy � 0 at y � s since there is no shear stress at the gas/liquid

interface.

� �2r s ydu

dy( )� �

FIGURE 7 (a) Sketch of a liquid film on a rotating disc. (b) Detail of a liquidfilm on a rotating disc.


Hence

(6)

The average film velocity is given by

The maximum film velocity (at y � s) is

(7)

Referring to Figure 7a, the liquid is supplied to the disc at a radius ri and amass flow rate M. It is deemed to instantaneously acquire and maintain the discangular velocity as it moves over the disc to be discharged at its periphery. At aradius r the mass flow rate is given by

(8)

Eliminating Uav from Eqs. (7) and (8) gives

(9)

Inserting Eq. (9) into Eq. (8) gives

(10)

Hence the average time required for the liquid to travel from ri to ro is

(11)

If we consider a typical example of water flowing over a disc under the fol-lowing conditions:

M � � � �� 3 10 10 102 3 3kg/s N-s/m kg/m2 3� �

tdr

U Mr r

oo iR

R

i

� � �0 3

4

12 2

2 2

1 3

4 3 4 3∫

( )� ��

�

/

/ /

�M w

r2 2

2

1 3

1 3

12 ��

−/

/

Uw r M

r wav ��

�

�

�

2

2 2 2

2 3

3

3

2

/

sM

wr�

3

2 2 2

1 32 3�

�

−/

/

M U s r� � �av 2

Uw rs

Umax .� ��

�

2 2

21 5 av

Us

u dyw rss

av � �1

30

2 2

∫�

�

ur

syy

� � �

�

2 2

2


then from Eq. (11) the average liquid transit time on the disc is 0.25 s, and fromEq. (9) the film thickness at the disc edge (provided that the film does not breakup into rivulets) is 28 microns. A more viscous liquid, such as a polymer (say,� � 10 N-s/m2), would have a thickness at the disc periphery of 600 microns anda transit time of about 5 seconds.

As already noted, the foregoing calculations must be regarded as a guideonly, since the films are intrinsically unstable, with waves being amplified as theliquid proceeds to the edge of the disc. It will be appreciated that this process pro-ceeds more rapidly with relatively inviscid liquids.

2.1.1. Mass Transfer

A conservative estimate of the disc’s mass transfer performance may be obtainedfrom the Nusselt model, assuming that there is no film mixing as it proceeds tothe edge of the disc. For unsteady diffusion into a finite stagnant slab, the plotshown in Figure 8 from (8) gives the relative concentration distribution within theslab at various times, with a zero initial concentration and a surface concentrationC0 imposed at time t � 0. The parameter on the curves is the Fourier number, Fo,where

(12)

and

D � solute diffusivity within the film

te � exposure time of the film surface

s � film thickness

As can be seen from Figure 8, if Fo � 0.02, the concentration changeswithin the film are confined largely to the surface layer and the local mass trans-fer coefficient is given by the Higbie penetration theory (9) as

(13)

For the previous example of a polymer flowing over the disc, a typicalFourier number may be calculated from:

D

t

s

e

��

�

10

5

6 10

9

4

�

��

m s

s

m

2/

kD

tLe

�

1 2/

Fo �Dt

se

2

r ri o� � � �� 5 10 0 25 1002 m m s (955 rpm)1.


Thus Fo � 0.014.Equation (7) shows that the film surface velocity is given by

Hence from Eq. (11),

(14)tM

r re � �32

9

3

4

2

2 2

1 34 3

14 3 ��

( )/

/ /

U UM

rmax ./

/� �1 59

32

2 2

2

1 3

1 3av ⋅

−�

� ��

FIGURE 8 Concentration distribution at various times in a slab �� x � � forzero initial concentration and surface concentration C0.


Inserting this into Eq. (13) gives

(15)

However, it must be noted that as the film flows over the disc, the film thick-ness progressively decreases, provided the liquid fully wets the disc. As thisoccurs, the concentration profiles normal to the disc plane are compressed, there-by causing a proportionate enhancement of the solute diffusion rate beyond thatpredicted by penetration theory. Thus the local value of kL can be corrected toaccount approximately for the steepened concentration gradients by multiplyingby a factor s1�s, where s1 is the film thickness at a radius r1 as given by Eq. (9).The corrected local value of kL is then

(16)

At the point of film formation, where r � r1, Eq. (16) shows that kL � �.However, the average value of kL over the disc surface is given by

(17)

This requires numerical integration. As pointed out at the outset, these estimatesof the mass transfer performance are likely to be conservative as the disturbanceof the film by ripples has been neglected. This will reduce the exposure time sig-nificantly, particularly with inviscid liquids.

2.1.2. Heat Transfer

The Nusselt model was originally developed to correlate the performance ofvapor condensers. In this case, the latent heat of condensation is discharged at the gas–liquid interface and subsequently conducted through the draining con-densate film, the conduction path length being the local film thickness. When aliquid film is heated or cooled on a spinning disc, the conduction path length isless (about 50% of the thickness) because all of the sensible heat does not haveto be conducted through the entire film. Since the thermal diffusivity of most liquids is typically of the order of 10�7 m2/s, compared with a mass diffusivity ofaround 10�9 m2/s, the Fourier numbers involved in the heat transfer version ofFigure 8 are approximately 100 times their mass transfer equivalent. This impliesthat the heat transfer process involves the whole liquid film rather than merely

kr r

k r drL Lr

r

av�

�

12

22

12

1

2

( ) ∫

kD M r

r r rL �

�

��

( )

1 2 2 2

2

1 6

1

2 3

4 314 3 1 2

2

3

1/ / /

/ / /

kD M

r rL �

�

��

( )

1 2 2 2

2

1 6

4 314 3 1 2

2

3

1/ /

/ / /


a thin layer near the disc surface. The Higbie model for heat transfer is thereforeinappropriate.

For the larger Fourier numbers involved in the heat transfer it is reasonableto represent the film temperature profile approximately by a quadratic expression:

(18)

A, B, and C are constants determined by these boundary conditions:

It can be shown that

(19)

and

(20)

Since the film temperature gradient perpendicular to the disc will be muchgreater than that in the radial direction, the local heat flux (Q) into the film willbe controlled by the value of dT�dy at the disc surface. Hence

Thus the effective film coefficient is

(21)

For our earlier example with water on a 0.5-m-diameter disc, Eq. (21)implies that the heat transfer film coefficient at the periphery is 43 kW/m2k, withthe predicted film thickness of 28 microns. For this estimate to be realistic it isessential that the film wet the disc and not break up into rivulets. This dependsupon a force balance at an incipient “dry-out point,” as indicated in Figure 9. Atthe film stagnation point the film momentum is potentially destroyed by the actionof the component of the surface forces parallel to the disc. Thus for an average filmvelocity Uav we must satisfy the following condition for rivulet maintenance:

T U s( cos )1 2� �� av

hQ

T T

ksw s

��

�2

Q kdT

dy

ks

T Ty

w s� � ��

0

2( )

dT

dy

T T

s

y

sw s�

�21

( ) −

T T T Ty

sT T

y

sw w s w s� � � � �22

2( ) ( )

1 0

2

3 0

.

.

.

T T y

T T y s

dT dy y s

w

s

� �

� �

� �

at

at

at�

T A By Cy� � � 2


where T is the surface tension per unit length, � is the contact angle, � is the liq-uid density, and s is the local film thickness.

Coherent films are less likely as they become thinner and their velocitydecreases. An inspection of Eqs. (9) and (10) reveals that U2

av �s is proportional to�2/3M5/3. Hence the tendency to form rivulets is less at higher disc speeds and liq-uid flow rates and increases with large T and small �.

2.1.3. Film-Flow Instability

The existence of the wave structure within the film is of major practical interest, aswas highlighted by some elegant experimental work conducted by Brauner and Maron (7). They monitored the instantaneous local film thickness of a liquid flowing down a stationary inclined plane using a capacitance technique.

FIGURE 9 Schematic showing liquid film “dry-out” on a rotating disc.


Simultaneously the local mass transfer coefficient was measured between the discand the liquid, using the limiting electrolytic current method. Their plots are repro-duced in Figure 10. It was clear that the passage of a ripple was associated with asignificant enhancement of the mass transfer coefficient, as a consequence of theflow field associated with ripple propagation. An analogous phenomenon may beobserved when sand particles are disturbed by wavelets in shallow seaside pools.While the phenomenon has considerable theoretical interest, its immediate practi-cal implication is very important because it suggests that the disc heat and masstransfer performance could be enhanced still further by appropriately engineeringthe disc surface profile. Some experimental results are discussed next.

FIGURE 10 Simultaneous time traces of local instantaneous film thicknessand transfer rate.


2.2. Heat/Mass Transfer Performance

An early application of the SDR to heat transfer duty was developed by Hickman(10), who was interested in the desalination of brackish water using a vapourcompression evaporator. A sketch of his initial arrangement is shown in Figure 11,where brackish water flowing on one side of a rotating disc assembly was evapo-rated by condensing steam on the other. The very high overall heat transfer coef-ficients that could be achieved (up to 45 kW/m2K) ensured that the pressure ratiodemanded from the vapor compressor was minimal, thereby establishing veryefficient operation. In a further development of the idea (11), a series of discassemblies was mounted on one vertical shaft that was enclosed within a tower,as shown in Figure 12. It is significant that Hickman did not report any problemsassociated with the deposition of crystal scale, even though his experimental runslasted for several hundreds of hours. Radial rather than spiral stains were, how-ever, exhibited, which suggested that the influence of the Coriolis accelerationwas minimal. Since the temperature difference between the condensate and theevaporating brackish water was only 1–2�C, it is presumed that the equivalentsupersaturation was insufficient to cause significant crystal nucleation.

The ability of the spinning disc to operate with very small driving tempera-ture and concentration differences can improve the thermodynamic efficiency ofthe overall process system. This is clearly the case with the Hickman vapor com-pression evaporator, and it is also exemplified in the applications described next.In general, the power needed to rotate the spinning disc assembly is a small frac-tion of that saved by virtue of the establishment of an intensified fluid dynamicenvironment.

2.2.1. The Rotating Electrolytic Cell

A laboratory-scale rotating chlor-alkali membrane cell was constructed and testedsome years ago in ICI. The electrodes comprised closely spaced catalyzed discs thatwere separated by a Nafion membrane. The anolyte and catholyte concentrationscorresponded to those in the brine and sodium hydroxide solutions used in thestandard (FM21) industrial version of the membrane cell. As can be seen fromFigure 13 (taken from Ref. 12), while the industrial cell voltage at a current densityof 3 kA/m2 was 3.17 V, that of the rotating unit was a function of the applied accel-eration, falling to about 2.75 V at 100 “g.” At a higher current density, the benefitsof enhanced acceleration were even more marked. It will be recognized that the enhanced buoyancy forces generated by the high acceleration can eliminate the polarization effects associated with bubbles that adhere to the electrodes/membranes or remain in the electrolytes. In principle a compact rotor comprising a bipolar cell assembly of closely spaced discs is capable of an exceptionally highchlorine production rate while operating at exceedingly competitive voltage.


FIGURE 11 Schematic of the single-element Hickman still.


FIGURE 12 (a) Schematic of the multiple-disc Hickman still. (b) Full-scaleHickman still.

2.2.2. The “Rotex” Absorption Heat Pump

The main factor that has been responsible for the slow adoption of absorption heatpumps for heating and air conditioning duties has been their high capital costcompared with that of vapor compression equivalents. This is due largely to thecycle complexity, as shown in Figure 14, which displays the four principal cycleelements, all of which involve vapor–liquid systems:

1. Condenser2. Evaporator3. Generator/boiler4. Absorber


FIGURE 12 (cont.)


The only single-phase item involved is the solution heat exchanger, whichis intensified by the use of laminar flow in a matrix of fine channels. A sketch ofthe single-effect Rotex (13) design is shown in Figure 15, where it can be seenthat a hermetically sealed rotating disc assembly fulfills the four functions listed.The working fluid consists of a water solution of either mixed alkali metalhydroxides or lithium bromide.

The evaporator receives low-grade heat from the circulating ambient air andvaporizes the refrigerant at low pressure, the vapor being promptly absorbed atthe absorber disc immediately opposite. Working fluid from the absorber sump,now rich in refrigerant, is returned to the generator via a solution pump and asolution heat exchanger. The latter consists of a matrix of closely spaced metalfoil, which, as discussed earlier, gives very efficient heat transfer in a small vol-ume. The heat of condensation and absorption is removed from the condenser/absorber disc assembly by circulating water that enters and leaves via a mechan-ical seal. The working fluid is pumped around the cycle by a pitot tube assembly,with the tubes dipping into a peripheral liquid trough. Since the Rotex machineoperates with a horizontal axis, the pitot tube arm is counterweighted to resist thefrictional torque exerted by the trough.

Information recently released (14) shows that the double effect air condi-tioning version of Rotex has achieved a coefficient of performance of 1.0 at a tem-perature lift of 35�C using lithium bromide solution. This unit is about to enterfield trials. Its high performance is entirely due to the intensity of the heat andmass transfer environment generated on the liquid film flowing over the discs.

It is also worth noting that Alfa Laval has developed a process evaporatorfor fruit juice and milk concentration using a nested stack of cones. Figure 16

FIGURE 13 Voltage characteristics of a rotating chlorine cell.


shows the arrangement, in which steam is caused to condense outside the coneswhile the process fluid is concentrated on the film flowing on the inside surface.In this context it may be observed that the disc and the cone are specific exam-ples of the general case of a rotating surface of revolution. The acceleration caus-ing the outward movement of the liquid film is the resolved component of �2ralong the surface in question.

Koerfer (15) performed an interesting study with a series of perforated andsmooth rotating discs 600 mm in diameter at speeds up to 600 rpm. The masstransfer performance was measured using the oxygen/water system, with theresults shown in Figure 17. Very good performance was recorded with the perfor-ated discs, and this was attributed partly to the short exposure time of the film asit negotiated each perforation and partly to the extra film area created.

Interestingly, the film behavior was much more predictable when it flowedover the disc surface containing the raised lips arising from the punching oper-ation. Film flow on the alternate side tended to “leak” through the disc, particu-larly at lower liquid flows, presumably due to the Coanda effect as liquidnegotiated the rounded edge of the holes.

As part of a general development to use spinning discs in an intensifiedabsorption heat pump, Aoune and Ramshaw (16) measured both the local andaverage heat transfer performance on smooth rotating surfaces. The disc surface

FIGURE 14 Single-effect absorption heat pump cycle.


temperature was estimated by extrapolating the values given by the digitized out-put from thermocouples embedded at depths of 1 mm and 9 mm in a brass disc.The local film temperature was measured by a thermistor contained in a stylusthat could be traversed over the disc surface.

Using water, the heat transfer coefficient on the 50-cm-diameter disc regu-larly exhibited a minimum value at a radius of about 17 cm. On the other hand,with the use of a water/60% monopropylene glycol mixture, no minimum wasobserved and the absolute performance was much poorer than that obtained withwater. This behavior is attributed to the tangential fluid slip generated as the feedliquid is brought up to the rotor’s angular velocity. This slip appears to be moremarked with low-viscosity liquids, which seems intuitively reasonable.

FIGURE 15 (a) Rotex design concept. (b) Rotex prototype.


Mass transfer studies were then performed, based upon the aeration of liquidthat had been previously stripped. The local liquid oxygen concentration was estab-lished by carefully abstracting a film sample via a quill that could be traversed overthe disc. An Orbisphere oxygen analyzer was used. Rather poor agreement wasobtained between the experimental results for water (kL � 4 → 10 m/s � 10�4) andthe Higbie predictions (approx 1 m/s � 10�4) based upon total exposure time ofliquid on the disc. Clearly, liquid mixing within the film generates exposure timesthat are much shorter than the liquid residence time on the disc.

Another study, by Jachuck and Ramshaw (17), explored the influence ofsurface profile upon the heat transfer performance of a spinning disc. Using asmooth disc as a benchmark it was shown that disc surfaces disrupted with metalpowder or grooves gave a significantly improved performance—presumably dueto the better film mixing. The best performance at modest disc speeds wasobtained with “undercut” grooves (Figure 18), which were originally conceivedas a technique for improving the circumferential distribution of any radialrivulets. At higher disc speeds, the film radial velocity was such that liquid wasprojected off the disc, thereby compromising the heat transfer process.

FIGURE 15 (cont.)


FIGURE 16 Alfa Laval concentrator.

2.3. Reactor Applications

2.3.1. Strategic Considerations

At the most basic level, the SDR is an extremely effective gas–liquid contactingdevice. This makes it ideal for performing many intensified heat or mass transferoperations and, as will be discussed later, it may be deployed as an evaporator oran aerator/desorber. However, its principal application in the process industry islikely to be as a very high-performance reactor. Since the reactor is the heart of


FIGURE 17 Mass transfer performance of a rotating disc (O2/H2O system).


any process, the SDR can radically improve the economics and efficiency ofmany key processes, both in the commodity and in the fine chemical area.

In order to exert full control over the progress of a chemical reaction or phys-ical transformation, the fluid dynamic environment must be sufficiently intense soas to ensure that the mixing and heat transfer rates are faster than the intrinsicchemical kinetics. This concept is shown diagrammatically in Figure 19, whichillustrates the progress of a reaction represented simply as A � B → C, with thereactants A, B traveling in plug flow along a tubular reactor. When the interdiffu-sion of A, B is slow compared with the reaction rate, then C is produced near theoriginal plane of A�B separation. This represents a total loss of control on twocounts:

1. The A�B stoichiometric ratio varies wildly across the reactor diameter.Therefore the selectivity for the desired product C is likely to be com-promised because a more realistic reaction scheme will usually includemany side reactions.

FIGURE 18 Types of disc grooves tested.


2. Most of the reaction to C occurs in the immediate neighborhood of theplane of A–B separation. Thus only a small fraction of the availablereactor volume is utilized and an opportunity for intensification is lost.

On the other hand, when mixing is fast, the A�B ratio is uniform and con-trol over the product spectrum can be maintained. All the reactor space is used tomaximum effect. Since the intrinsic kinetics are allowed free rein, the reactor is able to operate at the maximum intensity permitted by the specific chemicalsystem.

While it should be self-evident that a rational reactor design demands aknowledge of both the fluid dynamic environment and the detailed process kinet-ics, the latter are rarely available. In many instances this leads to the severe limi-tation of many important reactions by an inadequate fluid dynamic intensity.Some of these are known to be fast, e.g., liquid-phase nitrations, while others are(incorrectly) assumed to be slow, e.g., most polymerizations. In these circum-stances the pragmatic approach is to use a high-intensity reactor for each systemand then to assess the impact upon the space–time productivity. Obviously, anintrinsically slow system is resistant to further acceleration and this will rapidlybecome evident. One significant qualification of this contention involves the very

FIGURE 19 The influence of mixing and reaction rates on reactor behavior.


short residence time in the SDR compared with its conventional counterparts. Incertain reactions the process temperature is restricted to one that avoids productbreakdown in the time available. Since the residence time in the SDR when per-forming a polymerization reaction is up to 10 seconds rather than the severalhours involved in a conventional stirred vessel, we must re-examine the process’stemperature trajectory. A higher operating temperature may well be acceptablebecause the undesired breakdown components may not have time to be generated.However, the higher temperature will reduce the liquid viscosity and acceleratethe reaction. The lower viscosity will reduce the residence time still further.Therefore the SDR can exploit a process operating envelope that is much largerthan what is accessible to conventional technology.

With regard to the processing of viscous liquids, by far the most importantapplication relates to the manufacture of polymers. The key processes are:

1. Condensation reactions2. Radical reactions3. Devolatilization

The progress of a condensation reaction is controlled by an equilibriumwith a volatile product, which, if continuously removed, drives the reaction for-ward. Unfortunately, as polymerization proceeds, the liquid viscosity increases,rendering the removal of the volatile component much more difficult. The batchstirred vessel, which is conventionally used for polymer manufacture, has a limi-ted ability to remove a volatile component from the increasingly viscous polymermelt. On the other hand, the SDR can maintain effective mass transfer and, as willbe shown in Section 2.3.2, can achieve in one pass (taking several seconds) thesame increment in polymerization as would conventionally require tens of min-utes. The SDR with one or more discs on the same shaft is therefore capable ofperforming polycondensation extremely rapidly. The short residence time alsofacilitates rapid changes of product grade with minimum wastage.

The rate of a polymerization that proceeds via a series of radical reactions iscontrolled by the micromixing environment within the polymer melt. Once againthe stirred vessel is a poor means to achieve the high desired intensity, whereas theSDR has an impressive capability in this respect. It is well known that UV radia-tion is a very effective means of radical generation, and this technique has beenproposed in the past for stimulating certain radical polymerizations. However, theradiation extinction distance in a polymer melt is only a few millimeters, so a poly-merization reactor comprising a stirred vessel having a diameter of several metersis not a rational option. On the other hand, the thin polymer films that can be cre-ated and maintained on the SDR allow all of the polymer to receive a continuous,uniformly high radiation dose and thereby maintain a very high reaction intensity,as described later for the manufacture of butyl acrylate.


The industrial manufacture of polymers is rarely taken to completion, andthis requires unreacted monomer to be removed from the product—when its vis-cosity is highest. This devolatilization procedure is notoriously difficult becauseit usually involves the vacuum stripping of a stirred vessel’s contents for manyhours. Just as the SDR promotes the removal of the volatile component of a con-densation reaction, it is also effective in dramatically accelerating thedevolatilization process.

2.3.2. Polymerization

Polystyrene. The manufacture of polystyrene from various grades of prepoly-mer has been performed (18,19) on a 36-cm brass SDR using the arrangementshown in Figure 20. A series of concentric grooves was machined in the disc sur-face in order to improve liquid mixing within the film. The reaction operates viafree radicals, which were initiated in this case using benzoyl peroxide. In the firstinstance a series of batch runs was performed in a conventional laboratory-scalestirred vessel in order to produce a calibration curve (Figure 21) of conversion

FIGURE 20 Schematic of an SDR styrene polymerizer.


versus time. This vessel was then used to produce about 200 mL of prepolymerat a range of conversions that was supplied to the inner spinning disc surface overa period of about 30 seconds. The SDR was heated from below by a stationaryradiant ring, and the polymer produced was collected in a cooled annular troughsurrounding the disc. The styrene was diluted with about 16% w/w toluene in orderto reduce the viscosity.

Figure 21 also shows the increment in polymer conversion in one pass overthe disc as a function of the initial conversion in the preliminary batch. It can beseen that the equivalent batch time that can be ascribed to one pass on the discincreases (up to 58 minutes) as the initial conversion increases to 63%. Thisimplies that the benefits of the SDR become more marked as the polymer visco-sity increases. It is envisaged that the process can be scaled up either by using alarger disc or by mounting several discs on one shaft. The latter approach (i.e.,several discs in series) does, however, involve the problem of transferring poly-mer from the peripheral collection trough to the center of the next disc. An alter-native may be to operate discs on one shaft in parallel. For the experiments justdescribed, the feed rate was roughly 5–10 mL/s, which is equivalent to an outputof up to 250 tons/year on a continuous basis, though at this early stage this shouldnot be considered the ultimate limitation.

FIGURE 21 Free-radical polymerization of styrene.


The fundamental reasons for the high performance of the SDR are still amatter for debate. The significance of micromixing and the consequent improvedprobability of radical interaction has already been mentioned. However, anotherfactor is expected to be the divergent character of the flow on the disc. This maybe expected to align the polymer molecules and thereby encourage the juxtapos-ition of the reactive groups.

Polycondensation. The reaction between maleic anhydride and ethyleneglycol has been studied as an example of polycondensation (19). Since the reac-tion proceeds on an equilibrium basis, in order to drive it to completion the waterproduced must be eliminated from the increasingly viscous polymer melt. Thegrooved brass 36-cm disc described earlier for the polystyrene experiments wasused at a temperature of 200�C and a disc speed of 1000 rpm. As before, theexperimental procedure involved the establishment of a benchmark batch cali-bration against which the subsequent disc runs could be compared. A typical acidnumber plot versus batch time is presented in Figure 22. As the acid numberdecreases, the conversion to polymer increases. The water of reaction was removedfrom the polymer film by maintaining a large nitrogen purge to the vapor space.This technique, rather than the application of a vacuum, was the preferred methodfor reducing the water vapor partial pressure.

It can be seen that the increment in polymerization following one pass inthe SDR corresponds to many minutes of reaction in the small batch reactor usedas a reference. This is particularly encouraging because the mass transfer intensityin the laboratory stirred reactor is likely to be much greater than its industrial-scale

FIGURE 22 Time savings in the SDR for polyestirification.


equivalent, and it therefore provides a demanding benchmark for the spinningdisc performance.

2.3.3. Fine Chemical Manufacture

The intense heat and mass transfer environment that can be established within theliquid film flowing over the disc allows high selectivities and conversions to beachieved when fast liquid-phase reactions are performed. Very encouraging resultswere achieved in an industrial study of a phase-transfer-catalyzed (p-t-c) Darzen’sreaction to produce a drug intermediate (20). In comparison with the currently usedbatch processes, the ptc reaction on the SDR had a 99.9% reduced reaction time, a99% reduced inventory, and a 93% reduced impurity level. A more recent study hasinvolved a 20-cm-diameter SDR with a catalytically activated surface to performthe rearrangement of �-pinene oxide to campholenic aldehyde (21), which is animportant intermediate used in the fragrance industry. The comparative perform-ance of the batch reactor and the SDR is shown in Table 1. For equivalent conver-sion and selectivity, the unoptimized SDR gave a much higher throughput than theequivalent batch reactor and avoided the need to separate a catalyst slurry from theproduct. Figure 23 shows the variation of selectivity as a function of disc speed andfeed flow rate; Figure 24 gives the conversion levels achieved. While conversionfalls from 100% at the higher flows and speeds, presumably due to the reduced liq-uid residence time on the disc, the selectivity increases. Thus it might be expected that a larger disc (or a sequence of small discs) could combine high con-version and high selectivity. The batch reactor performance is summarized inFigure 25, where it can be seen that 100% conversion requires 5 minutes (� 1 sec-ond on the disc) and a maximum selectivity of 65% is reached.

2.3.4. Precipitation/Crystallization

The operation of crystallizers and precipitators is critically dependent upon thesupersaturation environment prevailing within the crystal magma because thisinfluences both the nucleation of new particles and the growth of those that

TABLE 1 Comparison of the Best SDR Runs with Batch Results forConversion of �-Pinene Oxide to Campholenic Aldehyde

Batch process SDR (continuous)

Process time (s) 300 1Processed feed 1.2 kg/h 209 kg/hConversion (%) 100 100Selectivity (%) 64 62Note Catalyst separated from No loss of

the product mixture catalyst


FIGURE 23 Selectivity towards campholenic aldehyde at 85�C at various feed rates.


FIGURE 24 Conversion of �-pinene oxide at 85�C and various disc speeds.


FIGURE 25 Batch reaction: conversion and selectivity towards campholenic aldehyde.


already exist. In general, the growth rate has a first-order dependence on super-saturation for diffusion-limited systems and 1.5- to second-order where the surfaceintegration resistance is significant (22). On the other hand, the nucleation rate,whether primary or secondary, has a higher-order dependence on supersaturation,typically in the range 2–9 (23), with the lower values being obtained mainly withlow-molecular-weight solutes. Hence it will be recognized that high supersatur-ation can be readily achieved on SDRs. This feature makes them attractive forproducing very small or nanosized particles, which are currently of intense indus-trial interest.

Supersaturation can be generated in several ways. Perhaps the simplesttechnique is merely to cool a solution saturated at a higher temperature.Alternatively, supersaturation can be created by removing the solvent or addingan antisolvent for systems where the solubility is only a weak function of tem-perature. Finally, supersaturation can be created by reaction—between either twoliquids or a liquid and a gas. In all these cases the intense environment createdwithin an initially crystal-free liquid film moving in plug flow over the disc cangenerate very high supersaturation and consequently small and fairly uniformcrystals. This characteristic of the SDR may be attractive in several industries(e.g., pharmaceuticals and coatings), where the product quality is intimately relat-ed to the fineness of the crystals and the tightness of the size distribution. Theconcept has recently been tested (24) in a spinning cone precipitator (Figure 26),which shares most of the characteristics of a spinning disc except that the cen-trifugal acceleration vector is not aligned to the cone surface. Barium sulfate wasgenerated by mixing equimolar solutions of BaCl2 and Na2SO4 in a central reser-voir. A thin liquid/slurry film flowed to the cone rim, from where it was collectedand subsequently analyzed in a Malvern Mastersizer. Equivalent batch experi-ments were performed, for the purpose of comparison, in a 50-mL agitatedbeaker. At a supersaturation of 500, defined as

the cone produced crystals at 6000 rpm that had a Sauter mean diameter of 3.2 microns, compared to 6.85 microns from the batch runs. However, at a supersaturation of 5000, the batch yielded a Sauter mean diameter of 0.75 microns,compared with 0.18–0.32 microns from the disc. The particle size distributionsreproduced in Figure 27 highlight the cone behavior at 8000 rpm more starkly,with a decided shift to 0.1–1 microns, compared with 1–10 microns in the batch.

Thus it can be seen that for many systems, spinning precipitators hold outthe prospect of generating the crystal size distributions that have considerableindustrial interest. This view is further reinforced by some earlier work bySmithKline Beecham (20), which crystallized an unnamed product (API) on a

Molar concentration of Na SO or BaCl

Molar solubility of BaSO2 4 2

4


15-cm spinning disc. Figure 28, reproduced from their work, compares the prod-uct size distribution from the SDR with that of the standard industrial material.Crystallization was induced by adding an antisolvent, and it can be seen that avery significant impact could be made on the normal product size distribution,which was smaller and narrower on the disc. The stainless steel disc was subjectto crystal scaling after a few runs. However, a thin layer of PTFE suppressed thiswithout significantly impairing the disc’s heat transfer performance.

2.4. Comparative Spinning Disc Reactor Costs

It should be borne in mind that the SDR is most effectively exploited when it is runon a continuous basis. The industrial units constructed to date have had disc diam-eters up to 30 cm and have been capable of processing around 30 g/s of feedstock.This corresponds to a continuous annual output of (e.g., polymer) 1000 tons/year.For a typical pharmaceutical product, a 15-cm disc could process about 7 g/s,equivalent to an annual output of 200 tons. With conventional stirred-vessel tech-nology, a roughly equivalent unit to the 30-cm SDR is a 2000-L batch reactor

FIGURE 26 Spinning cone reactor layout.


fabricated in 316 stainless steel, complete with computer control and a serviceunit delivering a disc temperature in the range �20�C–150�C. The cost of such areactor is estimated to be £250,000 with a range of £200–£400K.

Comparative SDR costs have been supplied by the most experienced manu-facturers (Triton Chemical Systems Ltd), which has been collaborating closelywith Newcastle University during the last 5 years of SDR development.

2.4.1. Spinning Disc Reactor Layout/Specification

The SDR system is available as a “desktop” or a “floor-standing” unit. The latteris easier to use at pressure and is more versatile for use with ancillaries such asfeed units. The desktop system is restricted to the core system shown within thecircle in Figure 29, which is a schematic representation of the floor standingarrangement. Experience has shown that SDR design and manufacturing detailsare very important in ensuring that the unit operates satisfactorily over its full per-formance range. Thus care must be taken with respect to the feed arrangementsto ensure that instantaneous stoichiometric ratios are held constant and that feedliquids are mixed exceedingly rapidly on the disc. A sophisticated feed system isprovided to ensure that the delivery of two liquid feeds is accurate. If this level ofcomplexity is not required, then there may be a significant cost saving. As shownin Figure 29, feed liquids are supplied by means of two ram injectors accuratelydriven by precision step motors. Product is drained from the disc casing into a

FIGURE 27 BaSO4 size distribution for a supersaturation of 5000.


lower receiving vessel, which can be isolated during SDR operation. This allowsproduct removal without interfering with the progress of the reaction.

Specifications for all systems include:

a. Constructed with 316L stainless steel contact parts for the reactor(other materials are available).

b. Constructed with a vessel to take 3-atm pressure or vacuum (other pres-sures available). Vessel has twin-wall construction to allow heating orcooling of the jacket.

c. Supplied complete with:i. Temperature measurements at the disc.

ii. Electronics and interfaces for:1. Up to 8 measurements of temperature.2. Up to 8 measurements of pressure etc.

iii. An advanced computer control and data-logging system.

The main variables on a specific spinning disc reactor are:

d. The material of construction.e. The range of ambient conditions under which the reaction take place

(temperature and pressure).f. The range of disc speeds available.g. The degree of instrumentation required.

FIGURE 28 Recrystallization of an API: comparison of size distributions.


2.4.2. Prices

These prices are for the spinning disc reactor as just defined in points a to c, includ-ing advanced control and data-logging software with computer and flat screen. Thefollowing ancillaries are usually necessary for the system, and these may exist, orTriton can recommend units or can supply them integrated into the system.

a. Feed systemsb. Heat transfer fluid system(s) for the disc/vessel/feedersc. Reaction atmosphere control systems, e.g., pressure/vacuum/gas feed

Example prices are given in Table 2. As can be seen from the table, the costof the SDR system is significantly less than that for a stirred vessel with a simi-lar productive capacity. However, cost considerations are likely to be much lesssignificant than the competitive edge that SDR technology is likely to bring interms of improved selectivity and product quality.

FIGURE 29 Schematic of the SDR manufactured by Triton.


3. CONCLUSION

It will be apparent from the examples cited that centrifugal fields in general, androtating surfaces in particular, can exert a profound influence upon process engineering. Our experience at Newcastle University has repeatedly demonstratedthat the greatest gains are achieved when the SDR is presented with the mostsevere process conditions. While our original target was simply to reduce equip-ment size and installed cost, it rapidly became apparent that reduced size alonewould be insufficiently persuasive for the acceptance and adoption of SDR tech-nology. However, it has now been demonstrated that in addition to yielding

TABLE 2 Spinning Disc Reactor Costs

M/C type Description Price

Type K-01 Small benchtop system, speeds from 100 to2000 rpm, suitable for disc temperature to 100�C

15-cm disc £40,00030-cm disc £60,000

Type P-01 Free-standing system, speeds from 200 to 4000 rpm, suitable for disc temperatures to 100�C

15-cm disc £75,00030-cm disc £90,00070-cm disc £120,000

Options Price

For either system, typical options are:a. Heat/cool system, to provide cooling at the walls to £17,500

20�C and heating/cooling at the disc for 5�C to 120�Cb. Injector system, 2000-mL capacity, with twin walls to allow £5,000

heating cooling; feed rates from 0.07 to 7 mL/s; fast down, with index and control system

c. Stirrers for the injectors £1,000d. Cooled deflector ring system £1,400e. Receivers, for operation up to 10 bar

1 L £9802 L £1,4004 L £1,960

f. Additional instrumentationMeasurement of heat transfer fluid flow rate £1,400Measurement of temperature difference across disc £400Added temperature measurements (per channel) £100Added pressure measurement (per channel) £400


a better product (i.e., tighter size/molecular weight distributions, improved purity,etc.), the SDR can transform the whole business process.

As with any radical step, the main obstacle to its implementation is the nor-mal human resistance to change. In order to overcome this it is vital that chemistswho are developing new processes be as fully conversant with SDRs as with theirbeakers and flasks. There is therefore a latent market for miniature versions of thisequipment throughout the world’s industrial and teaching laboratories. Once thecultural block is overcome and company decisionmakers fully appreciate thebreadth of impact in prospect, then the process industry will be fully prepared tomeet the challenges of the new century.

REFERENCES

1. Mallinson R, Ramshaw C. European Patent No. 2568B, 1969.2. Woods WP. The hydrodynamics of thin liquid films flowing over a rotating disc.

Ph.D. dissertation, Newcastle University, Newcastle, U.K., 1995.3. Thomas S, Faghri A, Hankey W. Experimental analysis and flow visualization of a

thin liquid on a stationary and rotating disc. Trans Asme—J Fluids Eng 1991; 113(March):73–80.

4. Wasden FK, Dukler AE. A numerical study of mass transfer in free-falling wavyfilms. AIChE 1990; 36(9):1379–1390.

5. Oron A, Davies SH, Bankoff SG. Long-scale evolution of thin liquid films. RevModern Physics 1997; 69(3):931–980.

6. Brauner N, Maron DM, Dukler AE. Modeling of wavy flow in inclined thin films inthe presence of interfacial shear. Chem Eng Sci 1985; 40(6):923–937.

7. Brauner N, Maron DM. Characteristics of inclined thin films, waviness and the asso-ciated mass transfer. Int J Heat Mass Trans 1982; 25(1):99–110.

8. Carslaw HS, Jaeger JC. Conduction of Heat in Solids. 2d ed. Oxford UniversityPress, Oxford, 1959:101.

9. Higbie R. The rate of absorption of pure gas into a still liquid during short periods ofexposure. Trans Am Inst Chem Eng 1935; 31:365.

10. Saline Water Conversion Report for 1959. U.S. Dept. of Interior Office of SalineWater, 1959:40.

11. Saline Water Conversion Report for 1957. U.S. Dept. of Interior Office of SalineWater, 1957:7.

12. Ramshaw C. The opportunities for exploiting centrifugal fields. Heat RecoverySystems CHP 1993; 13(6):493–513.

13. Ramshaw C, Winnington TL. An intensified absorption heat pump. Proc Inst Refrig1988; 85:26–39.

14. Gilchrist K, Lorton R, Green RJ. Process intensification applied to aqueous LiB� rotating absorption chiller with dry heat rejection. 7th UK National Conferenceon Heat Transfer, Nottingham, U.K., Sept. 10–12, 2001.

15. Koerfer M. Hydrodynamics and mass transfer of thin liquid films flowing on rotat-ing perforated discs. Departmental Report, Chemical Engineering Department,Newcastle University, Newcastle, UK, 1986.


16. Aoune A, Ramshaw C. Process intensification: heat and mass transfer characteristicsof liquid films on rotating discs. Int J Heat Mass Trans 1999; 42:2543–2556.

17. Jachuck RJJ, Ramshaw C. Process intensification: heat transfer characteristics oftailored rotating surfaces. Heat Recovery Systems CHP 1994; 14(5):475–491.

18. Boodhoo KVK, Jachuck RJJ, Ramshaw C. Process intensification: spinning disc poly-merizer for the manufacture of polystyrene. In: Ramshaw C, ed. 1st InternationalConference on Process Intensification in the Chemical Industry, Antwerp, Dec. 1995.

19. Boodhoo KVK, Jachuck RJJ, Ramshaw C. Spinning disc reactor for the intensifica-tion of styrene polymerisation. In: Semel J, ed. 2nd International Conference onProcess Intensification in Practice, Antwerp, Oct. 1997.

20. Oxley P et al. Evaluation of spinning disc reactor technology for the manufacture ofpharmaceuticals. IEC Res 2000; 39(7):2175–2182.

21. Vicevic M, Jachuck RJ, Scott K. Process intensification for green chemistry:rearrangement of �-pinene oxide using a catalyzed spinning disc reactor. 4th Inter-national Conference on Process Intensification for the Chemical Industry, Brugge,Belgium, Sept. 10, 2001.

22. Mullins JW, ed. Crystallization. 2d ed. Butterworths, London, 1972:162.23. Mullins JW, ed. Crystallization. 2d ed. Butterworths, London, 1972:162.24. Hetherington P, Scalley MJ, Jachuck RJ. Process intensification: continuous produc-

tion of barium sulphate using a spinning cone precipitator. 4th International Confer-ence on Process Intensification for the Chemical Industry, Brugge, Belgium, Sept. 10,2001.


4

Compact Multifunctional Heat Exchangers:A Pathway to Process Intensification

B. Thonon and P. Tochon

Greth, CEA–Grenoble, Grenoble, France

1. INTRODUCTION

In the first part of this paper, the different technologies of compact heat exchangersare presented and their range of application is given. The second part presents thestate of the art for heat transfer and fluid flow characteristics for single-phase,evaporation, condensation, and heat and mass transfer. The last part presentsapplications of compact multifunctional heat exchangers.

As proposed by Shah and Mueller (1), heat exchangers may be character-ized by the compactness factor, in m2/m3, and it is generally admitted that valuesgreater than 700 m2/m3 characterize compact heat exchangers. Often, compactheat exchangers also refer to nontubular heat exchangers, even if shell-and-tubeheat exchangers can have high compactness factors. For the heat exchanger con-sidered, the hydraulic diameter ranges from less than 1 mm to 10 mm. There aremainly two types of compact heat exchangers: the plate type (primary surfaceheat exchanger) and the plate–fin type (secondary surface heat exchanger). Thesetwo types of heat exchangers are described and new technologies are presented.

In the process industry, there are only four basic operations: reaction, sep-aration, mixing, and heat transfer. The traditional unit operation is to perform eachtask in one or more pieces of equipment sequentially, for example, heat transfer inheat exchangers and reaction in reactors. Combining two or more tasks in one piece


of equipment is implemented only for control or enhancement purposes; one exam-ple is cooling in a jacketed stirred-tank reactor. But this is not an intensified process,because it requires a batch operation, which has poor efficiency, and the reactioncannot be controlled effectively. Examples of multifunctional heat exchangers are:

Multistream heat exchanger (heat transfer between more than two fluids)Reactor heat exchanger (reaction and heat transfer)Reflux condenser (heat transfer and separation)

This chapter presents the state of the art in compact heat exchanger technol-ogy and provides heat transfer and mass transfer characteristics of these devicesand their use in the process industry.

Heat transfer is commonly required in the process industry for heating, cool-ing, vaporizing, or condensing. In most cases, only two streams (one hot and onecold) are in thermal contact within the heat exchanger. The most commonly usedheat exchanger is the shell-and-tube heat exchanger, which has poor heat transferperformance and requires a significant volume and ground area. Compact heatexchangers and enhancement technologies allow reducing the heat exchangervolume, to increase its effectiveness and to reduce capital and operating costs.

Compact heat exchanger technologies are sufficiently advanced, but theiruse and acceptance in the process industry are not yet widespread. Reasons arethe lack of awareness of their benefits and the absence of reliable design methodsand investigations under actual operating conditions. Compact heat exchangersinclude plate heat exchangers as well as plate–fin heat exchangers, which arecharacterized by hydraulic diameters between 1 and 10 mm. But recent develop-ments in manufacturing techniques, such as printed circuit heat exchangers anddiffusion-bonded and superplastic-formed heat exchangers, allow reaching hydraulicdiameters below 1 mm. These heat exchangers offer compactness greater than1000 m2/m3 and are suitable for industrial processes.

Rapid advances in range of design and operational reliability have madecompact heat exchangers attractive for many applications in various industries.Their high performance has already made their use widespread in the automotive,aerospace, air conditioning, refrigeration, and electrical equipment industries forsingle-phase and phase-change duties. In the automotive industry, plate-type heatexchangers are used as heaters, evaporators, and condensers (2), and since the1970s, the volume/heat capacity ratio has been divided by a factor of 2 (Figures 1and 2). This improvement has been achieved because of a radical change in heatexchanger technology and by the adoption of mass production systems integrat-ing innovative technologies.

Compact heat exchangers produced individually are generally more expensivethan a conventional shell-and-tube unit, and their payback time will be longer.But taking space, weight, and convenience into account compact heat exchangerscan be used cost effectively in a wider range of applications than the niches cur-rently being used in the process industry (3).


2. COMPACT HEAT EXCHANGER TECHNOLOGY

2.1. Classification of Compact Heat Exchangers

Heat exchangers can be classified in many different ways, such as according totransfer processes, number of fluids, surface compactness, flow arrangements,heat transfer mechanisms, type of fluids (gas–gas, gas–liquid, liquid–liquid, gas–two-phase, liquid–two-phase, etc.), and industry. Heat exchangers can also beclassified according to the construction type and process function (Figure 3). Refer

FIGURE 1 Progress in evaporator technology. (From Ref. 2.)

FIGURE 2 Progress in condenser technology. (From Ref. 2.)


to Shah and Mueller (1) for further details. In the follo1wing sections, only non-tubular heat exchangers will be described.

2.2. Plate Heat Exchangers

Plate heat exchangers (PHEs) were formerly used for milk pasteurization andgradually became the standard choice for heat treatment in the liquid-food indus-try. Actually, pasteurization must be considered a biological reaction, because thenative composition of the liquid is denatured during the heat transfer process. Inpractice this denaturation leads to fouling. The facility of dismantling plate heatexchangers is one of the main reasons for their extensive use in the food industry.Furthermore, because the heat transfer coefficients are high, the fluid path lengthwill be shorter and relatively well defined. Due to the lack of large dead areas in thechannels, the corresponding residence time distribution is short and homogeneous.

Eventually, with the development of larger plates, their use began to growquickly in the chemical, petrochemical, district heating, and power industries, butessentially for single-phase duties. The concept of phase change in PHEs originatedin the 1970s for ocean thermal-energy conversion (OTEC) applications; the work-ing fluids were Freon R22 or ammonia (4). These first studies on evaporation andcondensation have been used for the development of PHEs in the refrigeration

FIGURE 3 Classification of heat exchangers.


industry (5–10). Now PHEs have come to be used more often in the processindustry (11–12), but their use is still not widespread.

In terms of technology, PHEs are made from corrugated plates (Figure 4)that are pressed together. The plate size ranges from 0.02 m2 to over 3 m2 with con-ventional pressing technology (Figure 5), but can reach up to 15 m2 for explosion-formed plates (Figure 6). The hydraulic diameter lies between 2 and 10 mm formost common plates, but free passages and wide gap plates exist for viscous fluidapplications. Typically, the number of plates is between 10 and 100, which gives5–50 channels per fluid. Furthermore, the use of high-quality metal and manu-facturing techniques makes lead plate heat exchangers less prone to corrosionfailure than shell-and-tube units (13).

To ensure tightness, three technologies are available: gaskets, semiweldedor totally welded, and brazing. The gasketed PHE is the most common type, withthe gasket material selected as a function of the application (temperature, fluidnature, etc.). Temperatures up to 200�C and pressure up to 25 bars can beachieved by such heat exchangers. For applications where gaskets are undesirable(high pressure and temperature or very corrosive fluids), semiwelded or totallywelded heat exchangers are available (Figure 7). The last variant is the brazed-plate heat exchanger. The plate pattern is similar to conventional gasketed units,but tightness is obtained by brazing the pack of plates. For common applicationscopper brazing is used, but for ammonia units nickel brazing is possible. Thistechnology leads to inexpensive units, but the plate size is generally limited toless than 0.1 m2. The drawback is that the heat exchanger cannot be opened, andfouling will limit the range of application.

FIGURE 4 Plate heat exchanger. (Courtesy of Alfa-Laval Vicarb.)


2.3. Spiral Heat Exchangers

The spiral heat exchanger consists of two metal sheets that are welded together andthen rolled to obtain spiral passages. The passages can be either smooth or corru-gated; in some cases, studs or spacers are introduced between the metal sheets.These devices have two functions: (1) to adjust the spacing and (2) to induce tur-bulence and increase heat transfer. The general flow configuration can be crossflow(single or multipass) or counterflow, depending on the configuration of the inlet andoutlet distribution boxes. The heat transfer surface ranges from 0.05 m2 for refrig-eration applications (Figure 8) to 500 m2 for industrial processes (Figure 9). Spiralheat exchangers are often used for phase-change applications, because the geome-try of the hot and cold stream channels can be adapted to the process specifications.

Recent developments in manufacturing technologies (laser welding) haveallowed the manufacture of cost-effective recuperators based on the spiral con-cept (Figure 10) or the folded-plate recuperator (14–15).

FIGURE 5 Corrugated plates. (Courtesy of Alfa-Laval Vicarb.)


2.4. Plate-and-Shell Heat Exchangers

The basic principle of these heat exchangers is to insert a bundle of plates in ashell (Figure 11). On the plate side, the fluid flows inside corrugated or embossedchannels (more often in two passes). On the shell side, the flow is similar to shell-and-tube heat exchangers, and baffles can be inserted. This technology can beused for revamping an application, because the shell can be kept identical to thatfor a bundle of tubes. These heat exchangers are often used in the process indus-try as boilers (boiling on the shell side) because high pressures can be reached veryeasily on the shell side. Furthermore, a large gap on the shell side allows the useof dirty services, because cleaning is possible via removal of the bundle of plates.

2.5. Plate–Fin Heat Exchangers

Aluminium plate–fin heat exchangers (PFHEs) were initially developed in the 1940s to provide the aerospace industry compact, light, and highly efficientheat exchangers for gas/gas applications. Because the mechanical characteristicsof aluminum are increased at low temperatures, this technology has been usedsince 1950 for the liquefaction of natural gases. Nowadays, aluminum plate–fin heat exchangers are extensively used in applications such as air separation,hydrocarbon separation, and industrial and natural gas liquefaction (16). Theplate–fin heat exchanger offers process integration possibilities (12 simultaneous

FIGURE 6 Explosion-formed plate. (Courtesy of Packinox.)


FIGURE 7 Various welded plate heat exchangers.


different streams and more in one single heat exchanger) and high efficiency underclose temperature approach (1–2�C) in a large variety of geometric configurations.

The brazed plate–fin exchanger consists of stacked corrugated sheets (fins)separated by flat plates, forming passages that are closed by bars, with openingsfor the fluid inlet and outlet (Figures 12 and 13).

FIGURE 8 Spiral heat exchanger for the refrigeration industry. (Courtesy ofSpirec.)

FIGURE 9 Spiral heat exchanger for the process industry. (Courtesy of Kapp.)


In its simplest form, a heat exchanger may consist of two passages, with thecooling fluid in one passage and the warming fluid in the other. The flow direc-tion of each of the fluids relative to one another may be countercurrent, cocurrent,or crossflow.

The fins and the parting sheets are assembled by fusion of a brazing alloyto the surface of the parting sheets. The brazing operation happens in a vacuumfurnace in which the brazing alloy is heated to its point of fusion. All parts in

FIGURE 10 Laser welding of a spiral recuperator. (Courtesy of ACTE.)

FIGURE 11 Shell plate heat exchangers. (Courtesy of ACM and Barriquand.)


contact are bonded by capillary action (Figure 14). Once the brazing alloy hassolidified, the assembly becomes one single block. All passages for flow distri-bution and heat transfer of the streams are contained in the internal geometry ofthe block. Inlet and outlet headers with nozzles for the streams are fitted, by weld-ing, around the openings of the brazed passages. These nozzles are used for con-necting the heat exchanger to existing plant pipework.

Numerous fin corrugations have been developed, each with its own specialcharacteristics (Figure 15). Straight fins and straight perforated fins act like par-allel tubes with a rectangular cross section. Convective heat exchange occurs dueto the friction of the fluid in contact with the surface of the fin. The channels ofserrated fins are discontinuous, and the walls of the fins are offset. For air flows,

FIGURE 12 Components of a brazed aluminum plate–fin heat exchanger.(From Ref. 12.)


louver fins are extensively used; for process applications (single- and two-phase),continuous or offset strip fins are used.

For higher-temperature applications or when aluminum is not acceptable,stainless steel (temperatures up to 700�C) or copper materials can be used. Forvery high temperatures (gas turbine heat recovery; T � 1200�C), a ceramicplate–fin heat exchanger has also been developed (17) (Figure 16).

For high-pressure applications in the hydrocarbon and chemical processingindustries, a titanium compact heat exchanger has been developed by Rolls-Laval. This heat exchanger consists of diffusion-bonded channels that are createdby superplastic forming of titanium plates (18). This heat exchanger can handlehigh pressure and corrosive fluids and is suitable for marine applications.

2.6. Flat Tube-and-Fin Heat Exchangers

The concept of flat tube and fins in heat exchangers has been developed in theautomobile industry for engine cooling and air conditioning (19–21). In suchapplications one of the two fluids is air and the other is either water or a refriger-ant. The nonequilibrium of the heat capacities of the two fluids leads to the adop-tion of different enhancement technologies for both fluids. Generally on the airside the surface is finned (plain or louver fins), and on the other side the fluidflows in small-diameter channels (Figures 17 and 18). The technology is based on

FIGURE 13 Plate–fin heat exchangers. (Courtesy of Nordon Cryogénie.)


assembling aluminum elements, by either mechanical expansion or brazing. Forconventional applications, the pressure can be up to 20 bars. Recently, heatexchangers with operating pressures up 140 bar have been manufactured (22) forcar air conditioning systems, using carbon dioxide as refrigerant.

2.7. Microchannel Heat Exchangers

Microchannel heat exchangers are compact heat exchangers where the channelsize is around or lower than 1 mm. Such heat exchangers have been developed forsevere environments, such as offshore platforms (23). New applications are alsoarising for high-temperature nuclear reactors (24). To manufacture such smallchannels, several technologies are available (25): chemical etching, micro-machining, electrodischarge machining, etc.

The most common one is the printed circuit heat exchanger developed bythe Heatric Company. The channels are manufactured by chemically etching into

FIGURE 14 Stainless steel brazed plate–fin heat exchanger. (Courtesy ofNordon.)

FIGURE 15 Different fin geometries.


a flat plate. The plates are stacked together and diffusion bonded. These heatexchangers can support pressures up to 500–1000 bar and temperatures up to900�C (not simultaneously with high pressure). The typical size of the channelsis 1.0 � 2.0 mm, and the plate size can be up to 1.2 � 0.6 m (Figure 19). The processing technique is as flexible as for plate–fin heat exchangers, and crossflow

FIGURE 17 Condenser. (Courtesy of Livernois.)

FIGURE 16 Ceramic recuperator. (Courtesy of Céramiques et Composites.)


and counterflow configurations are employed. The main limitation of themicrochannel heat exchanger is the pressure drop, which is roughly inversely pro-portional to the channel diameter. For high-pressure applications, the pressuredrop is not a constraint; but for other fields of application it will be the main bar-rier to the use of such heat exchangers.

More recently, Chart-Marston has developed the Marbon heat exchanger (26).This heat exchanger is made of stainless steel plates stacked and bonded together(Figure 20). Several configurations are possible: (1) shell and tube, and (2) plate–fin.The use of such a heat exchanger as a chemical reactor is under consideration,and the thermal and hydraulic characterization has been undertaken as European-funded project (27).

Very compact heat exchangers are also used for cooling electronic devicesor microreactors (Figure 21). In these heat exchangers the channel size ranges from50 �m to 1 mm. Single phase and boiling are encountered in such applications(28). Applications in the chemical processing industries are also foreseen (29).These units can be very small in size and the heat duty per unit volume is veryhigh, up to15 kW/cm3.

2.8. Matrix Heat Exchangers

Perforated, or matrix, heat exchangers are highly compact and consist of a stackof perforated plates made of high-thermal-conductivity material, such as copperor aluminum, alternating with spacers of low thermal conductivity, such as plas-tic or stainless steel. The pack of alternate low- and high-thermal-conductivity

FIGURE 18 Extruded aluminum minitube.


plates are bonded together to form leakproof passageways between the streams(Figure 22). The main bonding technique adopted is diffusion bonding; moreinformation can be found in Ref. 30.

Such heat exchangers have been developed for cryogenic and low-temperature applications (31) and for fuels cells (32). They are suitable for a largerange of operating conditions, but there is very little information on their thermaland hydraulic behavior. Furthermore, as the heat is transferred by conduction inthe plate, the temperature distribution is not homogeneous.

2.9. Selection of Heat Exchanger Technology

The selection of the technology of compact heat exchangers depends on the oper-ating conditions, such as pressure, flow rates, and temperature, as well as on otherparameters, such as fouling, corrosion, compactness, weight, maintenance, andreliability. Table 1 summarizes the major limits for the different types of compactheat exchangers.

FIGURE 19 Printed circuit heat exchanger. (Courtesy of Heatric.)


3. SINGLE-PHASE FLOW

3.1. Flow Pattern

For corrugated heat exchangers, the flow is almost three-dimensional, and thevelocity field is difficult to measure. Flow visualization (33–34), realized in ahigh-scale channel, clearly shows a recirculation area downstream of the corru-gation edges (Figure 23). These areas are large at low Reynolds number (left ofpicture). But the transition to turbulent flow (Figure 24), which occurs atRe � 200, reduces the size of these areas. Local information (35–38) on the heat

FIGURE 20 Marbond heat exchanger. (Courtesy of Marston.)

FIGURE 21 Microchannel heat exchangers.


transfer and velocity fields indicates that the heat transfer coefficients are linkedto the flow pattern and to the mixing intensity in the channel; variations of �50%were measured. Furthermore, the size of the recirculation areas could be limitedby a proper design or by choosing appropriate operating conditions. The work ofAmblard (39) and Hugonnot (34) has shown that choosing the appropriate corru-gation angle, and with the Reynolds number above a critical value, a quasi–plugflow can be obtained. This indicates that corrugated heat exchangers could be suit-able for chemical reactions. The angle of corrugation has some influence on theglobal flow pattern. If we consider the entire channel as a two-dimensional medium,the flow behavior can be studied. At this scale the flow can be considered homogen-eous if there is the same flow rate through the channel width. Thonon et al. (37)have shown that for low aspect ratio and low corrugation angle (� � 30�), there isflow inhomogeneity, up to �15% of the flow rate distribution. But at high corruga-tion angle or for higher aspect ratio (Ar � 2), the flow is almost homogeneous.

In plate–fin heat exchangers, the flow structure depends on the fin geometry.Continuous fins can be assimilated to rectangular channels and the flow is almostidentical to pipe flows. For offset strip fins or louvered fins, there is a high degreeof mixing (40), and the flow becomes turbulent even at low Reynolds number(Figure 25).

3.2. Heat Transfer and Pressure Drop

For corrugated heat exchangers, extensive information is available in the literature(37–42); these studies have shown that the major geometric parameter is the

FIGURE 22 Matrix heat exchanger.


corrugation angle. The enhancement of heat transfer, compared to a smooth chan-nel, is up to six times greater. But at the same time, the pressure drop can beincreased by a factor of 100. Other geometric parameters are also influential, andmanufacturers are continuously improving the plate’s design.

For plate–fin heat exchangers in single-phase flow, the heat transfer coeffi-cients are related to the developed heat transfer surface, and the area ratio mustbe taken into account. As related to the projected surface, the overall heat transfercoefficient is very high. Heat transfer and pressure drop can be estimated fromcorrelations (43–44), but these correlations give only an estimate of the perform-ance, because local modification of the fin geometry will affect heat transfer andpressure drop.

For microchannel heat exchangers there is a large discrepancy between various experimental reports (45). Recent studies (46) have shown that down to

TABLE 1 Operating Conditions of Compact Heat Exchangers

Maximal Maximal pressure temperature Number of

Technology (bars) (�C) streams Fouling

Aluminum plate–fin 80–120 70–200 �10 Noheat exchanger

Stainless steel plate–fin 80 650 �2 Noheat exchanger

Ceramic plate–fin heat 4 1300 2 Noexchanger

Diffusion-bonded 500–1000 800–1000 �2 Noheat exchanger

Spiral heat exchangers 30 400 2 YesMatrix heat exchangers 1000 800 �2 NoFlat tube-and-fin 200 200 2 No

heat exchangerBrazed-plate heat 30 200 2 No

exchangerWelded-plate heat 30–40 300–400 �2 Yes/no

exchangerPlate-and-shell heat 30–40 300–400 2 Yes/no

exchangerGasketed-plate heat 20–25 160–200 �2 Yes

exchangerGraphite-plate heat 7 180 2 Yes

exchangerPlastic-plate heat 5 200–250 �2 Yes/no

exchanger


0.5 mm, there is no significant deviation, compared to large tube correlations, butthe size of the channels has to be accurately measured. More work is still neededto fully understand heat transfer and fluid flow in submicronic channels.

3.3. Fouling

The design of heat exchangers under foulant conditions results in oversizing, thussubstantially raising the cost of plants. Fouling is also responsible for processinefficiencies, due to increased thermal resistance. In water-cooling applications,particulate and precipitation fouling are frequently responsible for the decrease inheat transfer performance. Hence, the thermal and hydraulic performances needto be well understood if the heat exchange capability of practical equipmentneeds to be accurately predicted.

FIGURE 23 Flow structure in a corrugated channel: laminar flow (Re � 150).

FIGURE 24 Flow structure in a corrugated channel: turbulent flow (Re � 5000).


Plate heat exchangers are frequently used in industrial processes becausethey are more compact and have higher thermal performance than conventionalshell-and-tube heat exchangers, and it is generally admitted that plate heat exchan-gers are less prone to fouling than conventional shell-and-tube heat exchangersdue to the higher level of the shear stress. Measured fouling resistances (47–50)clearly indicate that fouling resistance values are about 10 times lower in corru-gated channels than on a plain surface and that the geometry (Figures 26 and 27)and the fluid velocity (Figure 28) are the major influential parameters. For instance,it has been shown that the fouling resistance is almost inversely proportional tosquare of the fluid velocity. This implies that the fluid velocity has to be con-trolled properly while operating plate heat exchangers.

FIGURE 25 Flow structure in plate–fin heat exchangers.

FIGURE 26 View of the deposit (30� corrugation angle at a velocity of 0.5 m/s).


While sizing plate heat exchangers, the great sensitivity of fouling to thefluid velocity and channel geometry precludes the use of a single value for thefouling resistance. If the TEMA fouling resistance values are applied to plate heatexchangers, excess heat transfer surface will be required, which can lead to poorefficiency. It is often recommended that the fouling margin not exceed 25% of the

FIGURE 27 View of the deposit (60� corrugation angle at a velocity of 0.5 m/s).

FIGURE 28 Void fraction in a corrugated heat exchanger.


extra heat transfer surface. Furthermore, the inverse velocity dependence of thefouling resistance needs to be taken into account at the design stage. If the extrasurface required for fouling is provided by adding plates, a maximum heat dutyis reached; but adding more plates will finally reduce the heat duty. In design pro-cedures, this needs to be taken into account. In the case of severe particulate foul-ing conditions, 60� corrugation angle plates should be selected rather than lowercorrugation angles, and a minimum fluid velocity of 0.3 m/s is suggested.

4. PHASE-CHANGE HEAT TRANSFER

4.1. Two-Phase Flow Characteristics

During vaporization or condensation, thermal and hydraulic performances dependessentially on the two-phase flow structure. Furthermore, as very often in indus-trial processes, the heat exchanger operates in thermosyphon or under natural circu-lation; knowledge of the pressure drop and liquid holdup is of major importance.But there is very little information on two-phase flow characteristics in compactgeometries.

Carey (51) has studied pressure drop and void fraction in different types ofcompact heat exchangers and has outlined the differences with plain tube geom-etries. Kreissig and Muller-Steinhagen (52) and Margat et al. (53) have shownthat the principles of the methods developed for plain tubes can be used but need tobe adapted. The main results of these studies are that the liquid holdup is signifi-cantly affected by the mass flow rate (Figure 28); the liquid holdup is underesti-mated by conventional correlation; the two-phase flow multiplier can be estimatedfrom a Chisholm-type correlation.

Winkelmann et al. (54) have studied air–water flows in a corrugated heatexchanger. Flow visualization and two-phase pressure drop measurements havebeen performed. The flow visualizations have shown that the flow pattern is complexand that a wavy or a film flow occurs in most cases (Figure 29). The two-phasepressure drop depends on the total flow rate and vapor quality, and Chisholm-typecorrelation is proposed. More work is required to characterize the flow structurein compact heat exchangers and to develop predictive methods for the frictionalpressure drop and the mean void fraction.

4.2. Vaporization

Vaporization is the most common unit operation to be found in the process indus-try; the use of compact heat exchangers as evaporators began 40 years ago for theconcentration of sugar or salt solutions. Nowadays, compact heat exchangers areused in several industrial processes, and this is particularly true for plate–fin heatexchangers, which are closely integrated in distillation and separation processesof natural and industrial gases. In most cases, evaporation takes place in an upward


flow and the fluid enters subcooled. The heat exchanger operates either in naturalor forced circulation. In the refrigeration industry, the heat exchangers are oftenlocated directly after the expansion valve, so a two-phase mixture is fed at thebottom of the heat exchanger. The main problem with such a configuration is toobtain a homogeneous phase distribution among the channels. Some manufactur-ers insert a distribution device in the inlet port, and it has been shown that a sig-nificant improvement in thermal performance can be achieved. For large compactheat exchangers operating under two-phase flow at the inlet, each phase of themixtures is fed independently in the channel in order to ensure an effective phasedistribution.

The design of compact heat exchangers for vaporization duties requiresknowledge of the heat transfer coefficients, and it is generally admitted that thebasic mechanisms occurring during flow boiling are similar to those encounteredin plain tubes (55–57), but no general predictive method is available. For plateheat exchangers, most of the data published have been obtained with pure refriger-ants, and the operating conditions are rather different than those encountered inthe process industry. The general trend of these studies is that the heat transfercoefficients are significantly higher than those obtained in conventional plaintubes, but the pressure drop is also increased. Concerning boiling of mixtures, com-pact heat exchangers provide high single-phase heat transfer coefficients; hencethe vapor phase will be well mixed, and no major degradation of the heat trans-fer coefficient should be observed.

Concerning plate–fin heat exchangers, the design of such units is much morecomplicated because up to 12 different fluids can flow in the heat exchangers.

FIGURE 29 Two-phase flow visualization in plate heat exchangers.


For distillation and separation processes, the fluids encountered are mixtures, andspecific pinch methods have been developed (58–59); a temperature approach ofless than one degree can be obtained. To optimize the number of layers per fluidand the type of fins, the heat transfer coefficients need to be accurately predictedin the different boiling zones (subcooled, nucleate, convective, dryout, etc.).Therefore, tests were performed with cryogenic fluids, hydrocarbons (60–61), andrefrigerants (62) in order to measure the heat transfer coefficients under actual flowand geometric conditions. Tests with cryogenic fluids are required because thebehavior of such fluids is significantly different than that of organic fluids.

Microchannel heat exchangers are used for boiling applications, but there isa lack of data for process fluids. Studies using water or refrigerant at low pressureshave outlined differences in the flow pattern and in the heat transfer coefficients(63–64). This comes from the fact that the bubble diameter is limited by the chan-nel size. For low and intermediate vapor qualities, the heat transfer coefficients areincreased as compared to plain tubes of larger diameters. But for higher vapor qual-ities, partial dryout may occur and will reduce heat transfer. Because microchannelheat exchangers might operate under large temperature differences, estimation ofthe critical heat flux is important; but most of the correlations were obtained for flatplates and single tubes, so their extension to microchannels is doubtful (65).

To select and apply a boiling method to compact heat exchangers, severalfacts must be taken into account. The method must be based on the fundamentalmechanisms occurring, because purely empirical and curve-fit methods cannot begeneralized. The basic assumptions for developing such a method is that bothnucleate boiling and convective boiling occur and that the dominant mechanismdepends on geometric and operating conditions (Figures 30 and 31). The effect of

FIGURE 30 Vaporization in a plate–fin heat exchanger.


geometry must be taken into account in the convective area, and that of the oper-ating conditions in the nucleate regime. It is assumed that nucleate boiling corre-lations, established for plain tubes, are sufficiently accurate and general to beapplied for compact heat exchangers. In the convective regime, the heat transfercoefficient is reported to the single-phase liquid heat transfer coefficient, and anenhancement factor is introduced to take into account the liquid–gas interactions.The main questions are to evaluate the single-phase heat transfer coefficient of thegiven geometry, especially at low Reynolds numbers, and to characterize theenhancement factor. The fundamental problem is to learn whether enhancementfactors developed for plain tubes can be applied successfully to compact geom-etries. Work is being carried out in several R&D organizations, and knowledge onboiling in compact geometries should be improved in the near future.

4.3. Condensation

Condensation occurs in many industrial processes, but rarely with pure fluids.The fluids encountered are mixtures, and noncondensable gases are often present;this makes the condensation process very complex.

In compact geometries the heat transfer coefficient depends on the two-phaseflow pattern (51–67). For low condensation rates, the heat transfer is gravity controlled, and the heat transfer coefficient depends on the liquid film thickness.For higher condensation rates, the heat transfer coefficient depends on the vaporshear effect, and for small passages the liquid–vapor interaction leads to high heattransfer coefficients.

FIGURE 31 Vaporization in plate heat exchanger.


In the case of mixtures or in the presence of a noncondensable gas, the con-densing vapor must diffuse through the gas to the phase interface. For this to hap-pen a partial pressure gradient toward the phase interface is necessary. The partialpressure of the vapor falls from a constant value at a rather large distance from thephase interface to a lower value at the interface. Correspondingly, the accompany-ing saturation temperature also falls toward the interface. Therefore, during con-densation, the condensing vapor arrives by diffusion at the condensate surface, andit is the thermal resistance in the vapor that limits the process. Hence in order toimprove the heat transfer, one must reduce the thermal resistance on the vaporside. Several factors can enhance the condensation process by reducing the vapor-side resistance. During condensation of mixtures or of vapors that contain non-condensables, the heat transfer on the vapor side can be improved by raising thevapor velocity. It has been shown that the heat transfer coefficient can be improvedby approximately 30% by increasing the vapor velocity. The use of a finely undu-lated surface can also achieve significant augmentations in heat transfer duringcondensation. It has been shown that corrugation can promote turbulent equili-brium between the phases and thus contribute to the increase in heat transfer.

Compact heat exchangers are characterized by small hydraulic diameters(1–10 mm), and there is no reliable design method to estimate heat transfer coef-ficients during condensation in such small passages. In the available literature,condensation of mixtures and of vapor in the presence of noncondensables hasbeen studied, but essentially for conventional geometries (plain tubes), and onlyfew results have been published with fluids representative of actual process condi-tions (hydrocarbons). A 2002 study (68) has shown that in laminar regime therewas a significant mixture effect (decrease in the heat transfer coefficient), but inturbulent regime the mixtures behave as pure fluids (Figure 32). This outlines thehigh degree of mixing encountered in such heat exchangers, which is not surpris-ing, since the geometric pattern of the plates is similar to modern column packing.

5. HEAT TRANSFER AND MASS TRANSFER

5.1. Macromixing

Macro- and micromixing are two major issues for compact multifunctional heatexchangers. Macromixing is closely linked to the flow pattern and has often beenstudied using flow visualization techniques. An effective macromixing will givehigh transfer rates and homogeneous flow distribution in the channel. The typicalscale of macrostructures ranges from half of the channel height to one-tenth of thechannel height. Micromixing is associated with much smaller scales and willaffect the reaction rate.

In this section, we will focus on two technologies able to produce both heattransfer and mixing: the corrugated heat exchanger and the plate–fin heat exchanger.


These heat exchangers have a flexible design and high heat transfer performanceand might be suitable for combined heat transfer and mass transfer services. Thethermal-hydraulic performances of the two technologies were presented in previ-ous sections. Flow visualization has often been used for analyzing flow structures,but it only gives access to qualitative information or time-average measurements.Using advanced numerical methods allows simulating single-phase flow in com-plex geometries, such as those encountered in compact heat exchangers. Thedetermination of the mixing ability for water at a Reynolds number equal to 2000is managed under the aspects of the computational fluid dynamics (CFD) method.An extensive review of turbulence models useful for compact heat exchangersimulation is available in the literature (70). The most appropriate model for eachselected geometry will be discussed next.

For corrugated heat exchangers, the flow is almost three-dimensional.Analyzing the stream function inside the 60� heat exchanger (Figure 33), we cansee that flow is mainly in the direction of the flow inlet, which is a characteristicof subchannel flow. However, a small part of the fluid is deviated by the channels:furrow flow. So for this angle, quasi–plug flow can be obtained, which is in

FIGURE 32 Mixture effect during condensation.


accord with the work of Amblard (39) and Hugonnot (34). The flow can be con-sidered homogeneous since there is the same flow rate through the channel width.This indicates that high teta corrugated plate heat exchangers could be suitablefor chemical reactions where macromixing is needed. Analyzing the stream func-tion inside the 30�-angle heat exchanger (Figure 34), we can see that the flow fea-ture is equally in the direction of the flow inlet (subchannel flow) and inside thefurrow (furrow flow). This result is in good agreement with Gaiser (35).

For a two-dimensional wavy channel, a numerical analysis performed atGRETh (69) has shown that the mixing efficiency of corrugated channels is excel-lent because one corrugation is nearly sufficient to have perfect mixing (Figure 35).

For finned passages, all the described phenomena (Figure 36) are in goodaccord with the regime predicted by the literature (70). Indeed, the flow hits thefront edge of the rectangular obstacle and separates immediately. The shear layerreattaches to the wall and splits in two parts: one part flows upstream, creating arecirculating and high-shear area; the other is convected downward by the meanflow. With time, the shear layer becomes unstable near the reattachment point,and it oscillates; it generates a vortex production inside the bubble and a growth of the recirculating area. Due to impinging shear layer instabilities, the

FIGURE 33 Stream function inside the 60�-angle plate heat exchanger.

FIGURE 34 Stream function inside the 30�-angle plate heat exchanger.


FIGURE 35 Numerical analysis of the mixing efficiency in a corrugated channel.


long bubble is broken, and an eddy is liberated and convected toward the trailingedge. At the same time, the bubble length decreases. This vortex shedding occurson both faces of the obstacle, and a Von Karman street is formed in the wake ofthe fin.

In plate–fin heat exchangers, the flow structure depends on the fin geometry.For offset strip fins, there is a high degree of mixing (25), and the flow becomesturbulent even at low Reynolds numbers (Figure 37). Indeed, after six rows of fins(e.g., about 20 mm), the flow is homogeneous, so the macromixing in the plate–finheat exchanger is very efficient.

5.2. Micromixing

Micromixing technologies have only recently been applied to the design ofminiaturized devices for chemical applications, so called microreactors. The maincomponents of such microreactors are mixers and heat exchangers.

FIGURE 36 Closeup of the flow pattern inside the offset strip fin geometry.

FIGURE 37 Evolution of a passive scalar inside the offset strip fin geometry.


Two ranges of applications can be distinguished:

The laminar regime for highly viscous fluids (such as glue), slow reactions,and long residence time

The turbulent regime for classical gas- or liquid-phase reactions, fast reac-tions, and low residence time

The present work will only focus on the turbulent regime, which is charac-terized by eddies with a wide range of length and time scales. The largest eddiesare typically comparable in size to the characteristic length of the mean flow. Thesmallest scales are responsible for the dissipation of turbulent kinetic energy,which is related to the turbulent fluctuation of the velocity. The higher the fluctu-ation, the smaller the scales in the flow and the better the micromixing efficiency.So for highly turbulent flows (high velocity or large hydraulic diameter), themicromixing will be very efficient, without any inventive turbulent generators.However, for compact heat exchangers where the hydraulic diameter is not large(a few millimeters) and the velocity not very high (maximum 1 m/s), enhance-ment techniques are needed in order to lower the turbulent-transition limit.

The mixing efficiency is determined by both the value of the pressure dropsand the turbulent energy dissipation, � (m2/s3):

where � is given by

and L is the total length of the passage, � is the voidage of the passage, �p is thepressure drop through the passage, and � is the density of the fluid. The higher themixing efficiency, the better is the micromixing.

For corrugated heat exchangers, using the numerical simulation for both60� and 30�, we can evaluate � directly from the computation and � from classi-cal correlation (5–7) (Table 2). The lower the angle, the lower the mixing effi-ciency. Indeed, while lowering the angle, the flow pattern moves from subchannelflow (helicoidal motion) to furrow flow (duct flow), which reduces the mixing.The mixing efficiency is very low (below 10%), so corrugated heat exchangers,which are able to produce high heat flux, e.g., high macromixing, are not able toproduce micromixing.

For plate–fin heat exchangers, using the numerical simulation for the OSFgeometry, we can evaluate �directly from the computation and � from classicalcorrelation (8) (Table 2). The OSF geometry is able to produce both macro- andmicroturbulent scales. For the Reynolds number considered, the fully turbulent

�

��

�pu

L

��

��


regime is not achieved (one must increase the velocity or the plate width by a fac-tor of 2), but this geometry is efficient for heat transfer and mixing. It could be agood concept in the area of temperature-controlled reactions.

5.3. Two-Phase Flow Heat Transfer and Mass Transfer

Numerous industrial operations involve a heat transfer between a liquid phase anda gaseous phase, with an important mass transfer effect, either as desorption-evaporation or as absorption-condensation. Here are some examples: reconcen-tration, by evaporation, of solvents, toxic industrial effluents; production, byabsorption, of industrial aqueous acid solutions; reversible or irreversible chemicalreactions (oxidation, hydrogenation, sulfonation); purification of permanent gases(air, smoke) by scrubbing of soluble vapors; desorbers and absorbers for heat pumps,where these two operations occur simultaneously.

In these multifunctional processes, heat transfer and mass transfer are twocombined and simultaneous functions, and the objective is to substantially improvethese functions in order to save energy, to increase the process efficiency, and toreduce the size and cost of industrial plants. Corrugated pads are often used in thedehumidification process or in chemical heat pumps, but a higher efficiency couldbe reached by using diabatic units, where the wall could exchange heat with theliquid film.

6. APPLICATIONS

6.1. Feed/Effluent Heat Exchangers

Feed/effluent heat exchangers are used in many industrial processes to warm upthe fluid before the reactor and to cool it down after treatment at high tempera-ture. The conventional design of such heat exchangers is based on shell-and-tubeunits. But to increase the thermal effectiveness of the heat exchangers, therequired heat length becomes very important, and high pressure drop will occur.

TABLE 2 Micromixing Performance for Plate HeatExchangers (PHE) and Offset Strip Fin (OSF) Heat Exchangers

PHE

30� angle 60� angle OSF

Turbulent intensity (%) 6.2 11.0 10� (m2/s3) 0.43 2.01 17.1� (W/kg) 13.1 51.0 24.1� (%) 3.2 5.7 71


The alternative to conventional shell-and-tube units is to use plate-type heatexchangers (71). The advantages of PHEs are a high thermal effectiveness for alow pressure drop, compactness, and a low propensity to fouling (72). In the refin-ing industry, large PHEs are used in hydroteater (HDT and HDS); the aim ofthese treatments is to reduce the naphtha contaminants (sulfur, hydrogen, arsenic,lead, etc.).

With the use of such units, two main problems arise: two-phase flow distri-bution and fouling. At the inlet of the effluent, the flow is a two-phase mixture(gas oil � hydrogen) with vapor qualities up to 20%. To ensure a homogeneousflow distribution, the two phases are fed independently at the inlet of each channel(up to 200 channels in parallel). These devices allow reaching high thermal effec-tiveness and a low-temperature approach. Concerning fouling, because these heatexchangers are totally welded and cannot easily be cleaned, tests have been per-formed under laboratory conditions and on site to measure the fouling resistance.Full-scale tests, realized at the Belgium Refining Corp in Antwerp (Figure 38), haveshown that after two years the thermal performance remained constant. A moreprecise study was also performed using the Alcor fouling apparatus. A PHE andshell-and-tube unit were connected in parallel; under similar operating condi-tions, the PHE exhibited no fouling while the shell-and-tube heat exchangerreached fouling resistance values up to 3 � 10�4 m2 K/W.

Work is still required to optimize the design of such PHEs because the heattransfer mechanisms are complex. On the feed side, a complex mixture is vapor-ized; on the effluent side, the mixture is condensed with the presence of noncon-densable gases.

6.2. Process Evaporators

Concentration of liquid by evaporation is widely used in industry, and a largevariety of techniques have been adopted. Two cases are considered: (1) the efflu-ent is desolved in water (salt solution, for example); (2) the effluent and the wateract as a mixture (water–acid solutions, for example). In the first case, the con-centration process is very efficient and high concentration can be obtained. Inspecific cases, crystallization of the effluent can be achieved. Most of the techno-logical developments have been obtained on the projects dealing with desalina-tion of seawater or the sugar industry. For mixtures, the liquid effluent to beconcentrated is partially evaporated in a heat exchanger; at the outlet, the vaporphase is richer with more volatile compounds and the liquid phase is richer withless volatile compounds. The evaporator acts as a first stage of distillation. Theeffectiveness of the evaporation process depends essentially on the mixture-phaseequilibrium. In most of the cases, the heating fluid is steam. As a consequence,the channel geometry must me adapted on each side of the heat exchanger toachieve optimal performance. Furthermore, as in most cases, the evaporating fluid


has high fouling capacities, so the evaporating side must be cleanable and thedead zone should be avoided.

Different types of evaporators are used in concentration processes: (1) flashevaporation through a discharge valve, (2) horizontal tubular or plate reboilers(submerged or falling film), (3) vertical tubular or plate evaporators (climbing orfalling film), (4) specific evaporators (direct contact, scraped surface, etc.). Compactheat exchangers are mostly used as vertical evaporators with either falling film(Morgenroth et al. (73)) or climbing film (Brotherton (12), Patel and Thomson (11)).A special high-capacity compact falling-film evaporator has been developed forsugar plants. The use of falling-film evaporators allows one to reach higher heattransfer coefficients than with climbing-film evaporators, especially for low tem-perature differences. Compared to conventional systems, the overall heat transferperformance can be up to two to five times higher, as a function of the fluid

FIGURE 38 Process flow sheet. (Courtesy of Packinox.)


viscosity. Plate heat exchangers are also used as evaporator in sugar plants, witha similar level of enhancement. All the tests realized (lab scale, pilot and indus-trial units) have shown that plate heat exchangers, even with small hydraulicdiameters, are not more sensitive to fouling than plain tubes. Other applicationsin concentration processes of such compact heat exchangers should be developed,especially in the area of mixture concentration.

6.3. Integrated Heat Exchangers in Separation Processes

Distillation and separation processes for purifying products requires up to 40% ofthe overall energy demand in chemical processes (Trambouze (74)), and greatefforts have been made by engineers to reduce this energy consumption. The inte-gration of the reboiler and the condenser in distillation columns has already beenstudied, but the entire potential of integrated heat exchangers has not been achieved(Lyon et al. (75)). As an example, plate heat exchangers have been installed in anisopropanol dehydration plant as reboilers and condensers, and they have beenfound to be very effective from the viewpoint of heat transfer.

Condensers and evaporators used in distillation processes are generallybased on the horizontal shell-and-tube design, and this conventional design leadsto large units, which have generally low energy efficiency. The integration is notoptimal, and the heat exchanger is still a part of the unit. Introducing compact heatexchangers in place of shell-and-tube units will permit greater compactness andlower energy consumption. Because the heat exchanger operates purely in coun-tercurrent, heat is available at a higher temperature than with horizontal shells. Forlong-term developments, diabatic distillation and separation units must be studied.

Reflux condensers are often used in the top of a distillation column, as afirst stage to separating the lighter and heavier components. Because the liquidflow and gas flow are countercurrent, a critical gas flow rate exists above whichflooding occurs. Although considerable work in the literature has been devoted tofree-falling films (mostly in pipes), the issue of flooding in countercurrent gas/liquid flow has not been settled yet. Furthermore, the literature concerning flood-ing in narrow passages is extremely poor. Thus, for the case of compact con-densers, there is no reliable tool for engineering-type predictions of flooding.From the viewpoint of heat transfer, reflux flows are more complex to study thandownward flows. Heat and mass transfer occur simultaneously between thephases, and heat transfer is present between the liquid phase and the walls.Literature provides neither experimental data nor a reliable prediction method forthese special system conditions. Studies about condensation inside channels havebeen carried out only for concurrent downward flow of pure vapors in tubes withlarge diameters. At present the design parameters for reflux condensation canonly be estimated by the classical annular-flow model presented in heat transferhandbooks. However, this model is not applicable under the given conditions, and


it is known from practical applications that the real condensation rate in compactheat exchangers deviates drastically from predictions according to the classicalmodel. Modeling reflux condensation (dephlegmation) in a compact plate–fin heatexchanger has been carried out (Urban et al. (76)), and the critical aspect raised is the distribution of the liquid film on the fins and the importance of the flow-distribution device. An integrated heat exchanger distillation column has beendeveloped by Nakaiwa et al. (77), and a plate technology has been adopted. Theanalysis outlines the energy savings that can be obtained as well as the highercompactness.

Aluminum plate–fin heat exchangers are often used as condensers in distil-lation and separation processes, but they require nonfouling and noncorrosive fluids. In the chemical industry, stainless steel or welded-plate heat exchangershave been used as top condensers of distillation columns, because they can beeither directly installed inside the column or closely integrated outside (Figure 39),

FIGURE 39 Overhead reflux condensers.


thereby avoiding extra piping and a pump for the condensate. In these compactheat exchangers, condensation can take place in downward or reflux flows. Forreflux condensation, flooding may occur at high gas-flow rates, and the design ofnontubular heat exchangers is essentially based on correlations and methodsdeveloped for plain tubes or larger hydraulic diameters. Studies on reflux flows innoncircular and small-hydraulic-diameter channels are required for a better designof such apparatus.

Looking more closely at the basic structure of compact heat exchangers, itcan be noticed that the geometry is quite similar to the packing for distillationcolumns. Because the wetting area is significantly higher than for plain tubes andhigher mixing occurs in liquid films, an additional rectification effect may occurduring reflux condensation. This phenomenon needs to be evaluated because itmay reduce the height of the distillation column.

6.4. Reactor Heat Exchangers

The development of integrated chemical reactor heat exchangers requires sizingtools for aiding the design and operation of the process. The thermal performanceof these heat exchangers is of prime importance for a global analysis of energyefficiency. Furthermore, a local analysis of flow and heat transfer conditions is alsorequired for a better characterization in terms of chemical reactors (mixing inten-sity and residence time distribution).

At bHr Group (Phillips et al. (78)), tests performed on fast exothermic reac-tions have shown that energy savings up to 40% could be achieved and that theamount of by-product was significantly reduced. Extension of this work to com-mercial compact heat exchangers is currently being considered; the first resultsindicate that chemical heat exchangers (CHEs) could be suitable as continuouschemical reactors.

Catalytic plate reactors already exist, but their range of application isextremely limited (Jachuck and Ramshaw (79)). The basic idea is to bring intocontact the heat source (catalytic reaction) and the heat sink (cooling medium).The catalyst is coated as a thin layer on one side of the plate, and on the other sideflows the coolant fluid. For dehydrogenation applications, a plate-type catalyticreactor has been developed (Arakawa et al. (80)); the benefits are its higher flex-ibility, via the control of the process temperature, and its higher productivity, viathe reduction in by-product formation. An alternative solution is to pack smallballs of catalyst between two plates. Plate catalytic reactors can operate underhigh heat fluxes and probably allow innovative reaction schemes (reaction withpure oxygen or under pressure).

At ECN, an example of coupling reactions has been studied for reformingof methanol. Combustion and reforming of methanol are done in two catalytic


compartments of the reactors. The catalyst has been coated on the internal surfaces of a stainless steel plate–fin heat exchanger. A monolith structure is usedfor each reacting channel, and a model has been developed to calculate the heattransferred between the channels. This heat exchanger reactor has been developedfor integration in a fuel cell plant.

The design of a heat exchanger reactor has to match several objectives:

The residence time must be sufficient to ensure a complete reaction withinthe heat exchanger.

The fluid temperature must be controlled, which implies high heat transfercoefficients.

If the feed and the reactant are not premixed and well dispersed, a sufficientturbulent intensity must be generated by the channel geometry.

The pressure drop must be acceptable.

The chemical process gives the enthalpy of reaction, the flow rate, the reactiontime, and the required reaction temperature. The first step in the sizing procedureis to calculate the required number of channels for the heat exchanger. Then thepass arrangement is selected in order to achieve the highest possible Reynoldsnumber within an acceptable pressure drop. For example, if the total number ofchannels is fixed by the residence time: channels in series will induce high veloc-ities and high pressure drop; channels in parallel will induce low velocities andlow pressure drop. The second step is to estimate the heat transfer coefficient andto check that the heat flux can effectively be controlled by the secondary fluid (thelower heat transfer coefficient should be on the reaction side).

The use of compact heat exchangers, where the channel characteristicdimension is between 1 and 10 mm, allow high heat transfer and mass transfercoefficients, even for low Reynolds numbers. The limitation will come from themixing intensity, which may not be sufficient to ensure droplet breakdown and toavoid droplet coalescence, which will directly affect the reaction. Preliminarystudies have shown that corrugated heat exchangers and OSF plate–fin heatexchangers are in turbulent flow even at low Reynolds numbers (Re � 300) andthat they should be suitable for chemical reactions. Investigation of the fluid flowis of prime interest for such applications, and advanced numerical methods pro-vide local and instantaneous values that can be used to characterize the chemicalreaction. These advanced CFD methods may also be used to develop specificcompact heat exchangers for chemical reactions as well as mixing devices.

Microchannels are also envisioned as a structure for heat exchanger reac-tors. If the channel dimensions range between 100 and 500 �m, the area per unitvolume is very high and allows catalytic reactions within the heat exchanger. Anexample of such a heat exchanger reactor is given by Rebrov et al. (81). Severalstudies are ongoing for applications in fine chemicals, reaction screening, andmicro hydrogen reformers (29).


7. CONCLUSIONS

In this chapter, the different types of compact heat exchangers have been reviewedand their applications described. Single-phase flow applications are now commonin the process industry; due to their high thermal effectiveness, compact heatexchangers are real alternatives to conventional shell-and-tube units.

For multiphase applications, the use of compact heat exchangers is still notwidespread, but in some industries these heat exchangers are widely used. In par-ticularly, in the refrigeration industry, plate heat exchangers are often used as boil-ers and condensers. A great opportunity to transfer knowledge and technology canbe applied to compact heat exchangers. On the one hand, within the new environ-mental requirements, mixtures will replace the conventional refrigerants (pure flu-ids), and the refrigeration industry is not used to dealing with mixtures. Transfer ofknowledge from the process industry, which is used to dealing with mixtures, couldhelp the refrigeration industry. On the other hand, because the refrigeration indus-try already uses compact heat exchangers, the transfer of technology to the processindustry will be fruitful. New applications for compact heat exchanger should alsoarise in environmental systems (Shah et al. [82]). Heat exchangers can also be con-sidered an active component in the process and not only a utility. For instance, heatexchanger reactors (Phillips et al. [27]) or diabatic distillation columns (Nakaiwa et al. [77]) may be designed using compact heat exchanger technology.

Process intensification (PI) is described as a key for future development inprocess plants (Green [83]), and because the cost of energy is now decreasing inEurope, the search for compactness in equipment is the goal to be achieved. There-fore, adopting compact heat exchangers is probably the most effective way tointensify a process. To support and develop intensive technologies, there is a needfor basic studies on heat transfer and mass transfer in compact geometries (R&Dprojects) and also targeted actions on specific applications (demonstration projects).As outlined at the Compact Heat Exchangers for the Process Industry conference(Shah [84–86]), the development of new products must be realized in collabora-tion with the process industry, and the reliability of compact heat equipment is thefirst goal to achieve. Furthermore, manufacturers must propose manufacturingstandards and design methodologies. Finally, the high potential of compact heatexchangers clearly matches the objectives of process intensification, and a muchwider use should emerge within 5–10 years.

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5

Process Intensification ThroughMicroreaction Technology

Wolfgang Ehrfeld

Ehrfeld Mikrotechnik AG, Wendelsheim, Germany

1. MICROTECHNOLOGY AS A KEY FOR THE ADVANCED

DESIGN OF CHEMICAL PLANTS

Since the middle of the 20th century, general technological progress has beendominated essentially by a unique strategy of success, which constantly aims atcomprehensive miniaturization and integration of functional elements in technicalsystems. The most outstanding development took place in microelectronics, whereintegrated circuits with hundreds of millions or even billions of transistors havebecome products of our daily lives. More recently, micromechanical, microoptical,microfluidic, and many other microdevices have become the basis for a multibilliondollar business, the market for microtechnology (1–3). The products of microtech-nology have achieved a key position in information, communication, entertain-ment, automotive, and medical technologies.

In the chemical and pharmaceutical industries, biochemists were the first or,at least, the fastest to make an interdisciplinary move into the promising field ofmicrotechnology. Terms like lab on a chip, microarrays, microfluidics, and micrototal analysis systems have become familiar to all working in the life sciences andat the front line of genomics, proteomics, glycomics, metabolomics, and all the other“omics” in this area. They regard biomolecules to some extent like a source of dataand, consequently, have no problems applying information-based technologies to


innovations in their own field. In the same way, large information technology (IT)companies have started to enter the area of life sciences, utilizing their traditionalstrengths in combination with the concepts of microreaction technology.

Researchers in chemical engineering are also intensively analyzing the pos-sibilities potentially offered by the general strategy of miniaturization and inte-gration to realize a radical change in design philosophy for modern chemicalplants. This research and development work started at ICI in the late 1970s, andthe term process intensification was used to characterize the novel concept. Themain intention was to achieve much lower investment, operating, and maintenancecosts for chemical plants, without decreasing their production capacity, by meansof a dramatic reduction in plant size; they aimed at a reduction factor of 100 oreven 1000 (4,5).

This intention may look more like a dream than a serious concept. However,the technological progress even in standard plant items has proven, beyond doubt,that this concept has a realistic basis. One may just consider, on the one hand, astandard stirred-tank reactor with a cooling jacket having a volume of about 10 m3

and, on the other hand, a potentially equivalent reactor for the same productioncapacity consisting of a static mixer and a compact heat exchanger having a volume of about 0.1 m3 (6). This simple comparison demonstrates the superior-ity of continuous operation over batch processing with regard to specific plantvolume and its importance in process intensification. Many other potential exam-ples exist, such as spinning disk reactors, vortex scrubbers, reactor-mixing sys-tems, and, of course, multifunctional reactors, which integrate reactions and unitoperations.

There is no doubt that the ultimate development of process intensificationleads to the novel field of microreaction technology (Figure 1) (7–9). Because ofthe small characteristic dimensions of microreaction devices, mass and heat trans-fer processes can be strongly enhanced, and, consequently, initial and boundary con-ditions as well as residence times can be precisely adjusted for optimizing yield andselectivity. Microreaction devices are evidently superior, due to their short responsetime, which simplifies the control of operation. In connection with the extremelysmall material holdup, nearly inherently safe plant concepts can be realized. More-over, microreaction technology offers access to advanced approaches in plant design,like the concept of numbering-up instead of scale-up and, in particular, the pos-sibility to utilize novel process routes not accessible with macroscopic devices.

As a matter of fact, microfabrication methods have to be introduced intochemical engineering in order to profit from the potential advantages of micro-reaction technology. Although this is a difficult hurdle, a few chemical companieshave successfully started to utilize microreaction technology for commercial syn-theses of fine and special chemicals. Nevertheless, much effort must still be spentto transfer further promising research results into commercial application and to


get away from traditional strategies in chemical engineering. In the following,a comprehensive analysis covering these aspects will be given.

2. EFFECT OF MINIATURIZATION ON UNIT OPERATIONS

AND REACTIONS

2.1. Enhancement of Heat Transfer

and Mass Transfer Processes

Diffusion, thermal conductivity, and viscosity are physically similar phenomenathat involve the transport of a physical quantity through a gas or liquid. The driv-ing forces for the corresponding transport fluxes of mass, energy, and momentumare the gradients in concentration, temperature, and velocity, respectively, wherein all three cases the fluxes are in the same direction as the gradients. For givendifferences in these properties, a decrease in the characteristic dimensions resultsin an increase in these gradients and, correspondingly, in higher mass and heattransfer rates as well as in higher viscous losses. Accordingly, mixing and heatexchange systems with extremely high transfer rates per unit volume can be real-ized by miniaturization; on the other hand, however, the effect of viscous losseshas to be taken into account.

Besides the effect of decreasing linear dimensions on the correspondinggradients, the effective surface area for exchange processes has to be considered.With decreasing characteristic dimensions, the surface-area-to-volume ratio of

FIGURE 1 Evolutionary development through process intensification.


the system increases. This results in a corresponding enlargement of the specificinterface area, i.e., of the area per unit mass or unit volume, for transfer processes,so, in connection with the enhancement of the gradients, i.e., the driving forcesfor heat and mass transfer, extremely efficient mixers and heat exchangers can berealized by miniaturization. Furthermore, the amount of material in a system isreduced with the reciprocal third power of the characteristic dimensions, and,consequently, the response time of the microdevice is extremely reduced so thatin most cases large differences concerning temperatures and concentrations arediminished immediately.

It was demonstrated in many cases that highly exothermal reactions can beperformed under isothermal conditions using the channels of micro heat exchan-gers as reaction volumes (Figure 2) (10). Pioneering work on this subject startedin the late 1980s, when micro heat exchangers with extremely high transfer ratesper unit volume were produced by means of advanced mechanical micromachin-ing methods (Figure 3) (11). Meanwhile, specific heat transfer rates of more than 20 kW per cm3 have been achieved, and a broad spectrum of materials has beensuccessfully applied.

A wide variety of micromixers are also available that allow mixing times inthe submillisecond range (8,12). They utilize mainly the concept of multilamination,where two streams of fluids are split into a large number of small substreams and fedalternately into an interdigital flow system, where they merge into a joint stream(Figure 4). Other concepts are based on the principle of splitting, side-to-sidearrangement, and further splitting to generate an increasing number of substreamswith different compositions, as known from large-scale static mixers. Vortex-typemicromixers have also abeen applied.

A wide range of applications for micromixers exists in the fields of gas–liquidsuspensions and liquid–liquid emulsions, with extremely small bubble and dropletsizes, respectively. A high uniformity concerning size distribution is achievable;in particular, the specific power consumption for generating suspensions andemulsions is much lower than in the stirring devices or high-pressure jets usuallyapplied in the macroscopic range (8,13). Accordingly, micromixers are prom-ising tools to improve the performance of phase transfer and other exchangeprocesses.

2.2. Inherent Process Restrictions in Miniaturized Devices

and Their Potential Solutions

As a matter of fact, miniaturization inevitably results in a number of processrestrictions, and completely new problems arise, too. There are, above all, theproblems of blockage of microstructures by solid particles and fouling effects.Moreover, corrosion might be much more dangerous for microscopic than formacroscopic devices.


FIGURE 2 Microreactor for parallel screening of catalysts for partial oxidation of methane.(Source: D. Hönicke, TU Chemnitz.)


Nevertheless, if the solid particles are small enough, they will have no nega-tive effect on the operation of a microreactor. On the contrary, microreactors caneven produce pigments of higher quality, i.e., smaller size and better uniformity,than macroscopic devices. This positive result was obtained experimentally atClariant Company; consequently, a microreactor pilot plant for pigment productionis under construction (14). By means of highly efficient micromixers, Siemens AxivaCompany succeeded in improving the synthesis of acrylate resins. They could avoida detrimental portion of high-molecular-weight resin and, consequently, foulingof the main continuously operating reactor. Evidently, there are at least concretechances to get around some of the problems resulting from small characteristicdimensions.

There is, of course, no possibility of avoiding all problems inherently con-nected with small dimensions. For instance, gravitational forces cannot be efficientlyutilized to transport fluids at small characteristic dimensions, since the effects of

FIGURE 3 Micro heat exchanger produced by means of mechanical micromachining. (a) Platelet with grooves of 30-�m depth and 70-�m width.(b) Assembly of crossflow heat exchanger. (c) Final devices. (Source:Forschungszentrum Karlsruhe.)


surface forces might far exceed those of mass or bulk forces. This problem isimmediately evident when regarding the reflux in a miniaturized distillation col-umn or the settler in a micromixer/microsettler system. To some extent, rotatingdevices can be applied and centrifugal forces can be utilized for material transport.This approach has been demonstrated successfully in microfluidic systems, but itis not a general solution. Consequently, other methods for phase separation arerequired for miniaturized process devices, such as microfiltration to break emulsionsand the utilization of hydrophobic and hydrophilic surfaces or capillary effects.Finally, surface effects will become more and more dominant in chemical reac-tions when the characteristic dimensions are reduced, which may produce advan-tages or disadvantages, depending on the respective type of reaction.

2.3. Consequences for the Selection of Reaction Routes

and Plant Design

The extreme enhancement in mass and heat transfer rates through miniaturizationof process devices results in fundamentally novel design possibilities with respectto selecting alternative reaction routes and plant design. In contrast to macrodevices like large stirring tanks, the starting conditions for a chemical reaction canbe set precisely with respect to time and concentration because of the much fastermixing of educts in a micromixer. The reaction starts at precisely defined time andposition with a spatially uniform composition. Thus, unfavorable reaction conditions

FIGURE 4 Micromixers. (a) Interdigital structure of a multilamination micro-mixer. (b) Principle of split-and-recombine static micromixers. (Source: IMM.)


due to incomplete mixing are minimized that eventually result in undesired sideand secondary reactions and, consequently, losses in yield and selectivity.

The high heat transfer rates achievable in micro heat exchangers and reactorsavoid unfavorable reaction conditions resulting from hot spots or thermal runawayeffects. An optimum temperature or temperature profile for the reaction can bechosen with respect to spatial distribution and time. Thus, a fast-flowing fluid ele-ment can be cooled down or heated up very rapidly, in fractions of a millisecond.Because of the small thermal mass of microdevices, a periodic change of tempera-ture of the reactor can be realized, with a typical time constant of some seconds.All these examples offer possibilities to improve yield and selectivity.

Since micro reactors—except for high-throughput screening in combinatorialmaterials research—are usually operated under continuous conditions, it seemssimple to adjust the optimum residence time by means of a suitable delay loop orchannel that is also favorable with respect to yield and selectivity (Figure 5). How-ever, the flow conditions in microdevices are generally characterized by a lowReynolds number; consequently, a parabolic Hagen–Poiseuille profile will existin long channels and ducts. This flow profile results in an unfavorable broadeningof the distribution of residence time. Special channel configurations allow one toreduce this effect.

2.4. Process Control and Safety

The inherent advantage of precise adjustment of the starting and boundary con-ditions for chemical reactions and unit operations in microdevices provides anovel basis for process control. Taking into account, in addition, the small holdup,

FIGURE 5 Process intensification by setting the optimum residence time.


it is evident that an extremely short response time is a further inherent advantageof microreaction devices with respect to process control.

As a result, there is a unique chance to utilize alternative reaction routes for chemical synthesis that so far have not been applied commercially, for reasonsof safety or difficulties in process control or because it is fundamentally impos-sible to realize such reaction routes using macroscopic devices. This is the case, inparticular, for controlled reactions in the explosive regime (Figure 6) (15). This isaccessible by means of microreaction devices, since, due to their small characteris-tic dimensions, they act like flame retention baffles. Moreover, the small dimen-sions allow reactions to be performed at extremely high pressure, which is ofimportance for chemical processes using supercritical solvents.

2.5. Sustainable Development by Numbering-Up

and Distributed Production

The safety problems connected with the storage of large quantities of educts andproducts remain, of course, unchanged when a conventional plant is replaced by amicroreaction plant with the same production capacity. Nevertheless, this problemmay be reduced by replacing a large plant by several small plants for distributedproduction. In contrast to conventional plants with macroscopic process devices,where scale-up usually results in a considerable reduction of specific investmentcosts, microreaction plants may instead profit from the mass production of micro-devices in reducing specific investment costs. Scale-up for achieving the desiredproduction capacity can be done only at one site, while a plant comprising a large

FIGURE 6 Explosion-proof continuous synthesis in the explosive range. Thereaction system consists completely of flame-retarding microchannels.


number of identical chemical microdevices according to the numbering-up conceptcan be split up for production at several sites. As a result, microreaction technologymay contribute to the strategy of sustainable development by saving resources dueto a higher yield and, in particular, by flexible production on site and on demand.

There are a number of further advantages to the numbering-up concept.Research results can be transferred into production faster, plants can be constructedin a shorter time, and the production capacity can be adjusted more flexibly tovariations in demand. Since mass production of microdevices may result in rela-tively a low cost per piece, novel cost-saving maintenance and repair conceptsbased on disposable elements might be introduced.

3. FROM BASIC PROPERTIES TO TECHNICAL

DESIGN RULES

In contrast to microelectronics, where extremely powerful software tools anddetailed design rules exist for the development of ultralarge-scale integrated cir-cuits, there are no corresponding comprehensive means in microreaction technol-ogy available to date. Such design tools should comprise mathematical modelingof flow and chemical reactions in miniaturized systems as well as specificationsfor suitable materials and simulation of manufacturing processes applicable to therespective microreaction devices.

Since it will take several years to realize such an integral software toolbox,individual approaches with separate steps have to be applied to meet gradually therequirements of microreactor design. Standard software for computational fluiddynamics is directly applicable in this context, and there are also powerful softwaretools for the simulation of special steps in microfabrication processes. However,there has been rather little experience with materials for microreactors, optimiza-tion of microreactor design, and, in particular, the treatment of interdependenteffects. Consequently, a profound knowledge of the basic properties and phenom-ena of microreaction technology just described is absolutely essential for the suc-cessful design of microreaction devices.

For instance, proper design rules must take into account that mixing andheat exchange systems with extremely high transfer rates per unit volume can berealized via miniaturization but that an increase in viscous losses may counter-balance the positive effects. Accordingly, suitable figures of merit must be definedfor micromixers and micro heat exchangers that consider the ratio of mass or heatfluxes to pressure losses. However, the value of such a figure of merit should bealways considered in context with further boundary conditions of the process andthe interdependence of several process properties. Decreasing the characteristicdimensions of a system results, as already explained, in a reduction in the materialholdup and a simultaneous enlargement of the surface-area-to-volume ratio of thesystem. These aspects also determine the speed of mixing and heat transfer and,


consequently, the degree of miniaturization required in a specific case. Such aspectshave to be considered for a favorable design of a microreaction system; in somecases, extremely small dimensions are not necessary to avoid unfavorable re-action conditions resulting from hot spots or thermal runaway effects.

If large amounts of materials have to be transported, a favorable design shouldinstead consider large pumps rather than arrangements with many micropumps,which, in most cases, are commercially unattractive for cost reasons and technic-ally less suitable because of their comparatively low efficiency.

4. MICROFABRICATION OF REACTION

AND UNIT OPERATION DEVICES

4.1. General Requirements

Since the production of chemicals in a continuous process is inevitably connectedto a transport of material, three-dimensional microfabrication processes are requiredin order to realize sufficiently large cross sections for channels and ducts as wellas reaction volumes. Meanwhile, a wide variety of such processes as well asdesign and test methods exist that all essentially originated from either semicon-ductor technology or precision engineering. Thin-film methods, applied to a largeextent in semiconductor technology, are less suitable for the generation of three-dimensional microreaction devices but are widely used for surface processing andprotection as well as for manufacturing sensor elements.

Because of the extremely wide variety of reactions, educts, products, andprocess conditions, a sufficiently broad spectrum of materials is required to realizesuitable microdevices for chemical processes. Metals and metal alloys, plastics,glass, ceramic materials, semiconductor materials like silicon, and various auxil-iary materials for sealing, surface treatment, etc. have been successfully applied forrealizing microreaction devices.

Besides such basic aspects concerning the shape of and materials for micro-reaction devices, costs play a major role in the selection of a microfabricationprocess. In this respect, the number of pieces and the precision that is reallyrequired, as well as aspects like availability and manufacturing experience, must betaken into account. In contrast to the situation some years ago, the prerequisites forcost-effective mass fabrication as well as small-scale production or rapid prototyp-ing have essentially changed. Modern commercial equipment for the production ofmicrodevices is available that allows unreliable and uneconomic laboratory-scalemanufacturing devices to be replaced.

Mathematical modeling of the device function may also help to cut costs,since it allows more realistic specifications to be worked out with regard to func-tional requirements. In addition, mathematical modeling of the process sequencefor microfabrication and assembly will be useful for cost saving. Such hard and


soft aspects will be considered in more detail in the following analysis of micro-fabrication methods for reaction devices.

4.2. LIGA Technology

The LIGA technology allows the production of ultraprecise micro structures withan extreme aspect ratio from a wide variety of materials (16,17). It is based on acombination of deep lithography, electroforming, and molding processes. In thefirst step of the manufacturing sequence, a pattern from a mask or by means of aserial beam-writing process is transferred into a thick resist layer on an electri-cally conductive substrate. Ultraprecise microstructures with an extreme aspectratio can be generated by deep X-ray lithography. Using special epoxy resists like SU 8, which utilizes intrinsic optical waveguide properties of irradiatedcrosslinked regions, favorable results are also achievable by means of UV litho-graphy.

In the second step, the three-dimensional relief-like structure of the resistpolymer generated by deep lithography is transferred into a complementary metal-lic structure by means of electroforming, starting from the electrically conductivesubstrate. Usually a nickel sulfamate electrolyte is applied, but there are also provenelectrolytes available for deposition of other metals and metal alloys.

The metal structure generated by means of electrodeposition may be the finalproduct in some special cases. In general, however, it is used in a third step as amaster tool for a replication process, such as injection molding, casting, or emboss-ing, for mass fabrication of microstructures. A wide variety of mold materials can beapplied for micromolding, e.g., organic polymers, preceramic polymers, and ceramicand metallic powders with organic binders for subsequent sintering, so that mostmaterial requirements for chemical microdevices can be favorably met.

It should be emphasized that the development of the LIGA technology orig-inated from a special requirement in nuclear process engineering. Curved micro-nozzles with characteristic dimensions in the micrometer range were needed asmass products for aerodynamic separation of the uranium isotopes in the frame-work of a large technological development work at the Karlsruhe NuclearResearch Center (Figure 7). Today there are a number of LIGA products that evi-dently have promising markets in the fields of micro-optics and integrated optics,molecular biotechnology, and microactuators. More recently, LIGA componentsand systems have been successfully applied to chemical engineering andmicroreaction technology, respectively. A number of chemical companies and, ofcourse, research institutes utilize devices such as micromixers, micro heatexchangers, and micro bubble columns as well as modular systems with integratedfunctional elements for reaction, heat transfer, mixing, separation, and fluid distri-bution for process development. LIGA devices are also seriously being consid-ered by the chemical industry for the production of fine chemicals.


4.3. Wet and Dry Etching Processes

Wet etching processes are widely used to produce microstructures by means oftransferring resist patterns into various materials. However, for most materials onlyisotropic etching processes exist, so, because of lateral underetching of the resistpattern, only shallow microchannels or other shallow structures can be generated atthe surface of a bulk material. Three-dimensional structures can be manufacturedwhen the pattern is etched completely through thin foils, which then have to bestacked in order to realize deep microchannels with a high aspect ratio (Figure 8).

Isotropic etching has been applied several times for manufacturing micro-reaction devices. The technological expenditure is relatively low, but there aresome restrictions concerning accuracy, surface roughness, and geometricaldesign. The product spectrum comprises various types of heat exchangers, micro-mixers, separators, reaction units, and even integrated devices with several func-tional elements.

Wet chemical anisotropic etching of monocrystalline silicon has been widelyapplied in microtechnology (18,20). This method is based on the dependence ofetching velocity on crystal orientation, so only a few basic geometries can be

FIGURE 7 Double-deflecting micronozzle for aerodynamic separation of ur-anium isotopes manufactured by LIGA technology from nickel. The smallestcharacteristic dimensions achieved in such devices are below 10 �m. (Source:Institute of Nuclear Process Engineering at the former Karlsruhe NuclearResearch Center, now Forschungszentrum Karlsruhe, Siemens.)


realized. Besides silicon, there has been very little manufacturing experience withother monocrystalline, inevitably very expensive, materials. Consequently, wetchemical anisotropic etching is in general not very attractive for manufacturingchemical microdevices because of strong restrictions with respect to shape andmaterial. Nevertheless, the technological expenditure is low, and material prob-lems can also be solved via the deposition of protection layers. A number ofmicrofluidic devices have been manufactured by means of this method, such asmicropumps, microvalves, and flow-distribution systems.

Besides anisotropic etching of monocrystalline materials, another wetchemical etching process exists that uses a special type of photosensitive glass (19).A wafer consisting of such glass is irradiated through a mask with UV light andsubsequently heated to a temperature between 800 and 900 K. This results in acrystallization of the irradiated regions that can be dissolved much faster in hydro-fluoric acid than the nonirradiated parts. This method has been successfully appliedto produce microreaction devices such as mixers, heat exchangers, and micro titerplates from glass.

Precise microstructures with nearly any cross-sectional shape can be gen-erated by means of anisotropic plasma-etching methods, where again silicon is themost important and proven material (18,20). Usually, a mask pattern is transferredinto a thin layer consisting of a material resistant to plasma etching on a silicon

FIGURE 8 Microetched foil of stainless steel for manufacturing micro heatexchangers by stacking and diffusion welding. (Source: Ehrfeld Mikrotechnik,Ätztechnik Herz.)


wafer. Subsequently, silicon is etched by means of a fluorine-containing low-pressure plasma that generates gaseous silicon compounds.

In order to generate microstructures with an extremely high aspect ratio, thedirected etching process is connected with a subsequent deposition process fromthe plasma where the walls oriented in parallel to the etching direction are coveredwith a plasma polymer resistant to the reactive plasma (21). By means of multiplerepetition of directed etching and side wall passivation, channels and other structureswith nearly vertical walls can be realized; accordingly, extremely high aspect ratiosare achievable for nearly any cross-sectional shape (Figure 9).

FIGURE 9 Channel structure of a phase separator generated by ASE deepetching of silicon. (Source: IMM.)


This so-called advanced silicon etching (ASE) process achieves etchingvelocities on the order of 0.2 mm per hour. The ASE dry etching method has, ofcourse, limitations concerning, e.g., material selection and surface smoothness orthe brittleness of silicon, which makes it nearly impossible to use it directly as amold insert in micromolding processes. Nevertheless, it is possible to transferASE silicon structures into complementary metal structures by electroforming.For a number of applications, ASE is evidently a favorable alternative to LIGA inmanufacturing devices for microreaction technology.

4.4. Mechanical Micromachining

In the past few years, impressive progress has been made in so-called mechanicalmicromachining, utilizing technologies based on so-called ultraprecision machining.Complex three-dimensional microstructures have been generated with shape accur-acies in the submicrometer range by means of milling, turning, and grinding(11,22). Three- and five-axis ultraprecision micromilling machines are availableas commercial products. Using diamond tools, an extremely low surface rough-ness of a few nanometers is achievable for nonferrous materials. Progress has alsobeen made in machining stainless steel by using ultrafine-grain hard metal toolsand novel technologies like vibration cutting. In addition, mechanical micro-machining has been successfully applied with brittle materials. Micromixers,micro heat exchangers, and reaction systems have been successfully produced bymeans of this technology (Figure 3).

It is evident that there are hardly any limitations concerning the generationof microstructures for chemical microdevices with complex geometries,extremely high aspect ratio, and high precision from a wide variety of materialsby means of mechanical micromachining. Rather, restrictions may exist whenmanufacturing closely packed channels or other structures, because of the finitesize of the tools. Also, manufacturing costs may become a problem in mass fabrication; but in such a case, mechanical micromachining may be helpful formanufacturing mold inserts for mass fabrication by means of micromolding.Moreover, there are other mechanical methods for high-volume production, likepunching and embossing, that have been successfully applied in fabricating, e.g.,micro heat exchangers.

4.5. Microelectrodischarge Machining

An interesting alternative to standard mechanical micromilling, turning, drilling,and grinding methods is microelectrodischarge machining (EDM), which is virtu-ally unlimited with respect to the geometrical shape of the work piece (23). Materialis removed in a discharge between the electrically conductive work piece and anelectrode by small sparks in a dielectric fluid such as oil or deionized water. Animportant advantage in micromachining is that the forces acting on the work piece


in EDM are extremely low. The disadvantages of micro EDM are a relatively largesurface roughness, limitations in miniaturization because of the finite size of theelectrodes and the spark gap in the electrical discharge, and very long machiningtimes, so this method is essentially used to manufacture mold inserts or prototypes.

The methods of mechanical micromachining and micro EDM have beenextensively applied to the fabrication of components such as micro heat exchangers,mixers, and reaction channels as well as chemical microsystems with integratedheat exchange, reaction, mixing, and distribution elements (Figure 10).

4.6. Micromachining by Means of Laser Radiation

Microfabrication by means of laser radiation covers a wide range of differentmethods (24,25). On the one hand, these are processes where material is removedin an intense electromagnetic field by melting, evaporation, decomposition, photo-ablation, or a combination of these phenomena. On the other hand, generatingprocesses exist where structures are built up from liquid resins, laminated layers,or powders using, e.g., photochemically induced crosslinking of organic compounds

FIGURE 10 Micromixing element generated by microelectrodischarge machin-ing. (Source: Ehrfeld Mikrotechnik, Zumtobel.)


in stereolithography or powder solidification by laser sintering. In addition, weldingby means of laser radiation is of major importance for the connection and assem-bly of microdevices.

There are no restrictions worth mentioning concerning materials in micro-machining by laser radiation, which is a real advantage for chemical microdevices(Figure 11). However, limitations exist to achieving critical dimensions below 10 �m and low surface roughness. Removal of material is also often connected withthe generation of debris, which reduces accuracy. Since laser-based microfabricationprocesses, except lithography, are essentially serial rather than parallel machiningmethods, their productivity is comparatively low. Nevertheless, they offer a hugepotential in rapid prototyping.

Laser-based micromachining processes have been applied to date only on arelatively small scale for manufacturing chemical microdevices (27). This willprobably change relatively soon, since rapid prototyping will become more andmore important for developing novel microreaction devices.

5. IMPLEMENTATION OF MICROREACTION TECHNOLOGY

Microreaction technology has created a novel basis for:

Accelerating screening in combinatorial material developmentRealizing extremely powerful tools for the evaluation of new reaction path-

waysImplementing comprehensively the concept of process intensification for

the production of fine and special chemicals (Figure 12)

FIGURE 11 Part of a static micromixer manufactured by laser ablation fromaluminum oxide. (Source: Ehrfeld Mikrotechnik, Heidelberg InstrumentsMicrotechnologies.)


Following the development route of miniaturization in the life sciences, theimplementation of microreaction technology in combinatorial material develop-ment has been very successful. Companies like Symyx, located in Silicon Valley,whose business is based on the highly effective synthesizing and screening of ahuge number of chemical compounds, have demonstrated that faster developmentand cost savings are achievable by means of microreaction devices. Not only canthe amount of reactants, auxiliary substances, waste, energy, and space be mini-mized, but all the other advantages of microreaction devices mentioned earlier canalso be favorably utilized (see, e.g., Ref. 27). The research work of such companiesis focused on more efficient catalysts, new polymers, high-performance phosphorsfor illumination, and, of course, drug development and many other substances.Promising work in this direction is also being done at universities and govern-ment research centers (7,8,26,28).

Researchers at BASF have shown that microreactors can be utilized thatgive access to operating conditions that cannot be realized by means of macroscopicequipment. They succeeded in improving yield and selectivity in a highly exother-mal two-phase reaction in connection with the synthesis of a vitamin precursor.At Degussa company, a microreactor test facility for proprietary reactions isunder construction. The major focus in this context is the implementation ofmicroreaction devices as powerful tools for process development and, in particu-lar, for the evaluation of new reaction pathways.

Companies like Clariant and Merck use microreactors for production, andthey are obviously convinced that the ultimate development of process intensifi-cation leads to microreaction technology. In contrast to other companies, Clariant

FIGURE 12 Microreaction technology aims at production of (a) information,(b) tools for process development, and (c) chemicals.


has reported about its work and its promising progress (14). Researchers at Clariantassume that about 15% of future production facilities will be based on micro-reaction technology.

However, microfabrication methods that are usually unfamiliar to chemicalengineers have to be introduced to profit comprehensively from microreactiontechnology. This transition from standard manufacturing methods of plant com-ponents to the development and production of microdevices is also inevitablyconnected with the application of special materials that are not yet proven in chem-ical engineering. In addition, novel design rules that have not existed until nowshould be implemented for the long term to speed up the development of noveldevices.

Essential progress is to be expected from the introduction of so-calledmodular microreaction systems. The system developed by Ehrfeld Mikrotechnikcomprises single functional elements for reactions, unit operations, transport,

FIGURE 13 Modular microreaction system consisting of functional elementsfor reactions and unit operations arranged on a base plate. The cube-shapedmodules of stainless steel with built-in microstructures have a side length of25 mm and can be operated at pressures up to 100 bar. (Source: EhrfeldMikrotechnik.)


measurement, and control. The modules can be arranged and connected in a widevariety of configurations and serve as a toolbox for realizing development platformssimilar to microplants (Figure 13). By means of such platforms, the optimum oper-ation conditions of chemical processes as well as favorable plant configurationscan be determined and novel reaction routes tested. There is also a wide range ofapplications in combinatorial chemistry. Since the microplants are usually set upfor continuous operation, they have a comparatively high productivity and can beutilized directly for small-scale production of special chemicals.

6. CONCLUSIONS

Future progress in chemical engineering will be strongly determined by processintensification through microreaction technology. It offers fundamentally novelopportunities to save direct costs in the areas of development, investment, opera-tion, and maintenance as well as to reduce indirect follow-up expenditures in connection with storage, transport, and changes in demand or market trends. A roadmap of microreaction technology for novel process routes and efficientproduction is shown in Figure 14. Nearly all major chemical, chemical engineer-ing, and pharmaceutical companies are interested in or even active in analyzingthe potential of microreaction technology. Moreover, there are a number of power-ful three-dimensional microfabrication technologies that should meet nearly allrequirements concerning geometries as well as materials of microreaction devicesin prototyping and mass fabrication.

However, the implementation of a novel technology needs time. It is neces-sary to prove carefully the potential advantages, to develop a sufficiently broadscientific basis, to implement reliable and cost-effective fabrication of chemicalmicrodevices on an industrial basis, to gain experience in the design, construction,

FIGURE 14 Roadmap of microreaction technology for novel process routesand efficient production.


and operation of microreaction plants, and finally to demonstrate real commercialsuccess. Meanwhile, a lot of effort has gone in this direction, but there is still amultitude of tasks to solve until decisionmakers will be convinced enough of thecommercial prospects of microreaction technology to accept the inevitable finan-cial risks of technological progress.

REFERENCES

A comprehensive overview of the international development work on microreaction tech-nology can be found in the Proceedings of the International Conferences on MicroreactionTechnology, which are listed in the following.

Ehrfeld W, ed. Proceedings of the 1st International Conference on Microreaction Technol-ogy. Berlin: Springer, 1998.

Ehrfeld W, Rinard I, Wegeng R, eds. Process Miniaturization: 2nd InternationalConference on Microreaction Technology, IMRET 2; Topical Conference Preprints.AIChe, New Orleans, 1998.

Ehrfeld W, ed. Proceedings of the 3rd International Conference on Microreaction Technol-ogy. Berlin: Springer, 2000.

Rinard I, ed. 4th International Conference on Microreaction Technology. Topical Confer-ence Proceedings. AIChE Spring National Meeting, Atlanta, GA, March 5–9, 2000.

Matlosz M, Ehrfeld W, Baselt JP, eds. Proceedings of the 5th International Conference onMicroreaction Technology. Berlin: Springer, 2001.

Rinard I, ed. 6th International Conference on Microreaction Technology, ConferenceProceedings. AIChe Spring Meeting, New Orleans, March 10–14, 2002.

The literature cited in this contribution is listed here.

1. Ehrfeld W, Ehrfeld U, Kiesewalter S. Progress and profit through microtechnologies.Proceedings VDE World Microtechnologies Congress, MICRO.tec, Vol. 1, 2000: 9–17.

2. Market Analysis for Micro Systems II, 2000–2005. A NEXUS Task Force Report,2002.

3. Bundesministerium für Bildung und Forschung. Förderkonzept Mikrosystemtechnik2000�, Bonn, Germany, Jan 2000.

4. Stankiewicz AI, Moulijn JA. Process intensification: transforming chemical engin-eering. Chem Eng Prog 2000; (Jan):22–33.

5. Green A, Johnson B, John A. Process intensification magnifies profits. Chem Eng1999; (Dec):66–73.

6. Wood M, Green A. A methodological approach to process intensification. IchemESymposium Series No. 144, 1998:405–416.

7. Jensen KF, Hsing I-M, Srinivasan R, Schmidt MA, Harold MP, Lerou JJ, Ryley JF.Reaction engineering for microreactor systems. Proceedings of the 1st InternationalConference on Microreaction Technology. Berlin: Springer, 1998:2–9.

8. Ehrfeld W, Hessel V, Haverkamp V. Microreactors. In: Ullmann’s Encyclopedia ofIndustrial Chemistry. 6th ed. Weinheim: Wiley-VCH, 1999.

9. Jäckel K-P. Microreaction Technology—Vision and Reality. Plenary Lecture,ACHEMA 2000, Frankfurt.


10. Mayer J, Fichtner M, Wolf D, Schubert K. A microstructured reactor for the catalyticpartial oxidation of methane to syngas. Proceedings of the 3rd International Confer-ence on Microreaction Technology. Berlin: Springer, 2000:187–196.

11. Schubert K, Bier W, Linder G, Seidel D. Herstellung und Test von kompaktenMikrowärmeübertragern. Chem Ing Tech 1969; 61:172–173.

12. Löwe H, Ehrfeld W, Hessel V, Richter T, Schiewe J. Micromixing technology. Pro-ceedings of the 4th International Conference on Microreaction Technology. AIChESpring National Meeting, Atlanta, GA, March 2000.

13. Bayer T, Heinichen H, Natelberg T. Emulsification of silicon oil in water—comparisonbetween a micromixer and a conventional stirred tank. Proceedings of the 4th Inter-national Conference on Microreaction Technology. Atlanta, GA, AIChE SpringNational Meeting, March 2000:167–173.

14. Wochner M. Mikroreaktoren—kleine Ergänzung für grosse Kessel, Clartext No. 3/2002.

15. Hagendorf U, Jänicke M, Schüth F, Schubert K, Fichtner M. A Pt/Al2O3 coatedmicrostructured reactor/heat exchanger for the controlled H2/O2 reaction in theexplosion regime. Proceedings of the 2nd International Conference on MicroreactionTechnology, AIChE Spring Meeting, New Orleans, LA, March 1998, 81–87.

16. Ehrfeld W, Münchmeyer D. Three-dimensional microfabrication using synchrotronradiation. Nucl Inst Meth Phys Res 1991; A303:523–531.

17. Ehrfeld W, Ehrfeld U. Microfabrication for process intensification. In: Matlosz M,Ehrfeld W, Baselt JP, eds. Proceedings of the 5th International Conference onMicroreaction Technology. Berlin: Springer-Verlag, 2001:3–12.

18. Koehler M, Ätztechniken. In: Ehrfeld W, ed. Handbuch Mikrotechnik. München:Carl Hanser Verlag, 2001:279–322.

19. Freitag A, Dietrich TR, Scholz R. Glass as a material for microreaction technology.Proceedings of the 4th International Conference on Microreaction Technology. AIChESpring National Meeting, Atlanta, GA, March 2000:48–54.

20. Rangelow IW, Kassing R. Silicon microreactors made by reactive ion etching. Pro-ceedings of the 1st International Conference on Microreaction Technology. Berlin:Springer, 1998:169–174.

21. Laermer F, Schilp A (Robert Bosch GmbH). Method of Anisotropically Etching Silicon.U.S. Patent No. 5501893, 1996.

22. Weck M. Ultraprecision machining of microcomponents. In: Weck M, ed. Proceed-ings of the International Seminar on Precision Engineering and Microtechnology,Aachen: European Society for Precision Engineering and Nanotechnology, July 2000.

23. Michel F, Ehrfeld W, Koch O, Gruber H-P. EDM for microfabrication—technologyand applications. In: Weck M, ed. Proceedings of the International Seminar onPrecision Engineering and Microtechnology, Aachen, July 2000.

24. Bremus E, Gillner A, Hellrung D, Höcker H, Legewie F, Poprawe R, Wehner M,Wild M. Laser processing for manufacturing microfluidic devices. In: Proceedings ofthe 3rd International Conference on Microreaction Technology. Berlin: Springer,2000:187–196.

25. Gillner A, Klotzbücher T. Lasermikrobearbeitung. In: Ehrfeld W, ed. HandbuchMikrotechnik. München: Carl Hanser Verlag, 2001:105–143.


26. Wiessmeier G, Schubert K, Hönicke D. Monolithic microreactors possessing regularmesopore systems for the successful performance of heterogeneously catalysed re-actions. In: Ehrfeld W, ed. Proceedings of the 1st International Conference on Micro-reaction Technology., Berlin: Springer, 1998:20–26.

27. Jandeleit B, Schaefer DJ, Powers TS, Turner HW, Weinberg WH. Combinatorialmaterials science and catalysis. Angew Chem Int Ed 1999; 38:2494–2532.

28. Claus P, Hönicke D, Zech T. Miniaturization of screening devices for the combina-torial development of heterogeneous catalysts. Catal Today 2001; 67:319–339.


6

Structured Catalysts and Reactors: A Contribution to Process Intensification

Jacob A. Moulijn, Freek Kapteijn,

and Andrzej Stankiewicz

Delft University of Technology, Delft, The Netherlands

1. INTRODUCTION

In this book the chemical plant is focused upon. Therefore, the present chapteremphasizes chemical reactors for the chemical process industry. But it should bemade clear that structured packings and catalysts also have a large potential inconsumer products. Chemical reactors form the heart of a (petro-)chemicals pro-duction plant. Given the large variety of plants it is no surprise that a wide varietyof chemical reactors are used. Catalytic reactors can be roughly divided into random and structured reactors. It is useful to start with a summary of themajor basic concerns (apart from high activity, selectivity, etc.) for catalytic reactors:

Catalyst quality on a microscopic length scale (quality, number of activesites)

Catalyst quality on a mesoscopic length scale (diffusion length, loading,profiles)

Ease of catalyst separation and handlingHeat supply and removalHydrodynamics (regimes, controllability, predictability)


Transport resistance (rate and selectivity)Safety and environmental aspects (runaways, hazardous materials, selectivity)Costs

On each of these, random and structured reactors behave quite differently. Interms of costs and catalyst loading, random packed-bed reactors usually are mostfavorable. So why would one use structured reactors? As will become clear, inmany of the concerns listed, structured reactors are to be preferred. Precision incatalytic processes is the basis for process improvement. It does not make senseto develop the best possible catalyst and to use it in an unsatisfactory reactor. Boththe catalyst and the reactor should be close to perfect. Random packed beds donot fulfill this requirement. They are not homogeneous, because maldistributionsalways occur; at the reactor wall these are unavoidable, originating form thelooser packing there. These maldistributions lead to nonuniform flow and con-centration profiles, and even hot spots can arise (1). A similar analysis holds forslurry reactors. For instance, in a mechanically stirred tank reactor the mixingintensity is highly non-uniform and conditions exist where only a relatively smallannulus around the tip of the stirrer is an effective reaction space.

Catalytic conversion and separation are conventionally carried out in separ-ate pieces of equipment. A combination of functions in single units is an elegantform of process intensification. When one of the functions is a chemical reaction,it is referred to as a multifunctional reactor. A good example is catalytic distilla-tion technology from the CDTech Company. They have introduced elegant tech-nology for desulfurization of oil (2). Structured reactors will play a key role in thedesign of novel processes based on multifunctional reactors (3). A monolith is agood example. Monolithic catalysts are shown in Figure 1.

FIGURE 1 Monolithic structures of various shapes. Square-channel cordieritestructures (1, 3, 5, 6), internally finned channels (2), washcoated steel mono-lith (4).


Depending on one’s point of view, a monolith can be considered a reactoror a catalyst; the borders between catalyst and reactor vanish (4,5). Other struc-tured reactors also deserve attention, but for the “message” it suffices to limit thediscussion to monolithic reactors.

2. OVERVIEW OF STRUCTURED REACTORS

Structured reactors and catalysts are encountered in a large variety (3,6).Structured catalytic reactors can be divided into two categories. The first involvesa structured catalyst, whereas the second one involves “normal” catalyst particlesarranged in a nonrandom way. In the first category, the catalyst and the reactor areessentially identical entities.

Because of their low pressure drop, structured reactors in practice dominatethe field for treating tail gases. Figure 2 presents the major types of reactor. Themonolithic reactor represents the class of “real” structured catalytic reactors,whereas the parallel-passage reactor and the lateral-flow reactor are based on astructured arrangement of packings with “normal” catalyst particles.

Structuring is possible at all length scales. In structured reactors the level isconsidered above that of a single particle. Structuring can be done based on dedi-cated structured catalyst shapes in such a way that the catalyst is an integrated part

FIGURE 2 Low-pressure-drop reactors used for tail-gas treating.


of the reactor shape. An alternative is to arrange catalyst particles in such a waythat a structured reactor is the result. Table 1 gives the subdivision together withtypical examples. By inspection of Figure 2, most of them can be recognized.

2.1. Monolithic Catalysts and Reactors

Monoliths usually are made from ceramics, but metals are also used. They can beproduced by extrusion of support material (often cordierite is used, but varioustypes of clays or typical catalyst carrier materials, such as alumina and titania, arealso used), a paste containing catalyst particles (e.g., zeolites, V-based catalysts)or a precursor for the final product (e.g., polymers for carbon monoliths).Alternatively, catalysts, supports, or their precursors can be coated onto a mono-lithic support structure (“washcoating”). Zeolites have been coated by growingthem directly on the support during the synthesis (7). The coating literature andpatents represent a large field, and, in principle, a variety of preparation proced-ures are available. All major catalyst support materials, ceramic and polymeric,have been extruded as monolith (4,8). Metallic support structures are used forautomotive applications (9). The choice for a certain catalyst type will stronglydepend on the balance between maximizing the catalyst inventory and catalysteffectiveness. For slow reactions, a high catalyst loading is desired and the pure

TABLE 1 Subdivision and Typical Examples of Structured Reactors

Structuring on the level of the catalyst Structuring exclusively on and the reactor the level of the reactor

Monoliths Three-levels-of-porosity (TLP) Extruded parallel-channel systems reactors

(honeycombs), usually ceramic Bead-string reactorsEmission reduction for cars Membrane-enclosed catalyticOzone decomposition in airplanes reactorsSelective catalytic reduction of NOx

Arrays of corrugated platesArrays of fibersGauzes

AgMethanol → formaldehyde

Pt/RhNO production from ammoniaHCN production from methane

FoamsCatalytic membranes


catalyst-type monolith is desired; for fast reactions or if diffusion is slow, a thincoating with a maximum geometric area is preferred.

Monoliths are the dominant catalyst structures for three-way catalysts incars (10–13), selective catalytic reduction catalysts in power stations (14–17) andfor ozone destruction in airplanes. What causes this popularity? The catalyst con-sists of one piece, so no attrition due to moving particles in a vibrating caseoccurs. The large open frontal area and straight channels result in an extremelylow pressure drop, essential for end-of-pipe solutions like exhaust pipes and stackgases. The straight channels prevent the accumulation of dust. In all these appli-cations the reaction system is relatively simple; a single fluid phase (gas) has tobe treated at reasonable conditions. More demanding applications of monolithsare now being investigated, fast reactions at high temperatures such as steamreforming, partial oxidation of hydrocarbons to syngas, and oxidative dehydro-genation (18–20). These examples are limited to single-phase applications. Aswill be discussed later, monoliths for multiphase applications have already proventheir value.

Monoliths are industrially produced in large quantities by extrusion. Thisleads to the attractive situation that, although they are sophisticated structures,they are commercially available at reasonable cost. Of course, monolithic cata-lysts have disadvantages. They share with packed-bed catalysts the requirementof sufficient stability or in any case good regenerability. With respect to masstransfer and heat transfer characteristics, the major limitations are the laminarflow through the channels, no interconnectivity between the channels, and a poorradial heat conductivity. The latter two properties are much better for the foam-type monoliths, but with a trade-off in a higher pressure drop and/or lower cata-lyst loading (sites/m3). In principle, a laminar flow velocity profile is associatedwith low mass transfer rates and a wide residence time distribution. Fortunately,for gases, due to the small channel size and high diffusivity, this radial transportin the channels is sufficiently fast. Typical time scales for diffusion are given inTable 2. In liquid phases the diffusivity is three orders of magnitude smaller,which is one of the reasons that monoliths do not enjoy a high popularity in liquid-phase operations. It will be shown that this is based on a misconception.

TABLE 2 Diffusion Time Scales in Catalytic Reactors (�lD2/2D)

D (m2/s) lD for 1 mm lD for 0.1 mm lD for 1 �m

Gas 10�5 50 ms 0.5 ms 50 nsLiquid 10�9 500 s 5 s 500 nsLiquid in cat pore 10�10 5000 s 50 s 5 msLiquid in zeolite pore 10�11 50,000 s 500 s 50 ms


The honeycomb-type monoliths are characterized by a very high geometricsurface area. Dependent on the cell density, this can exceed 3000 m2/m3

reactor!Figure 3 shows the values for three different cell densities, i.e., for 200, 400, and600 cpsi (cells per square inch). These examples are quite realistic. At present thenormal monolith for cars is a 400-cpsi monolith. The values for the geometricalsurface amount to 3440 m2/m3

reactor. In packed beds this value is much lower inorder to avoid unrealistic pressure drops. It is to be expected that future monolithswill exhibit even larger geometric surface areas. That alone makes them highlyuseful for process intensification programs.

Metal monoliths can be shaped rather freely. A good example is given inFigure 4 (9), where it can be seen that in these parallel-channel systems the struc-ture of the channels is such that the turbulence increases. The reasoning behindthat is the wish to counteract the low mass transfer rates associated with laminarflow in the thin channels of the monolith.

2.2. Gauzes

The appearance of gauzes is illustrated by Figures 5 and 6 (9). The use of noblemetal gauzes goes back to the beginning of the 20th century for the oxidation ofammonia into NO. This work followed up work of Ostwald, who applied plat-inized asbestos and later a roll of corrugated strip of Pt. Probably, this was the firstapplication of a structured reactor.

FIGURE 3 Geometric surface areas for three different cell densities.


2.3. Structured Packings

It is often suggested to coat structured packings of the type given in Figure 7 with catalysts, in the same way as monoliths. However, in normal applicationsthis does not lead to a satisfactory reactor. The geometric surface area is orders ofmagnitude smaller than packed beds and monoliths. Of course, this problem canbe solved by packing the channels with catalyst particles. Also, they can beapplied as mixing device to be used for a good inlet distribution in the case of

FIGURE 4 Shaped channels in metal monoliths in order to increase masstransfer in gas-phase applications.


multiphase reactors. Later it will be seen that combinations of static mixers andmonolithic catalysts have a high potential in process intensification.

2.4. Foams

Foams are to some extent the negative images of packed beds. They can be usedwhen turbulent mixing is important. Figure 8 gives an example of a foam that isused as a carrier for a molten salt catalyst in diesel soot trapping and combustion.The openness and the mixing characteristics of foams have stimulated research inthe potential application in soot trapping. An advantage is the robustness of the

FIGURE 5 Pt/Rh gauze for the production of NO from ammonia.


FIGURE 6 Details of a Pt/Rh gauze before (A) and after (B) use in the oxida-tion of ammonia into NO.


system: Plugging does not occur. A disadvantage is the relatively low trappingefficiency. In combustion foams result in stable combustion behavior.

2.5. Arranged Catalysts—Three-Levels-of-Porosity (TLP) Reactors

Three-levels-of-porosity (TLP) reactors (21–23) are alternatives for monolithreactors in certain applications. Conventional catalyst particles can be arranged inany geometric configuration. In such arrays, three levels of porosity can be dis-tinguished: the pore space within the particles, the intraparticle space, and thespace between the arrays. An example of such a TLP reactor is the parallel-passage reactor (PPR); see Figure 9 (24). The catalyst particles are confinedbetween wire gauze screens that divide the reactor into a large number of catalystlayers with empty passages in between.

The gas flows along the catalysts layers instead of through the bed as in atraditional fixed-bed reactor. Because the gas flows through straight channels

FIGURE 7 Structured packing (Sulzer Katapak-K).


(ca. 10 mm wide), the pressure drop over the PPR is much lower than over afixed-bed reactor. Reactants are transferred from the gas to the catalyst inside thegauzes, mainly by diffusion. The PPR is very suitable for treating dust-containinggases, e.g., flue gases from power plants, because dust will not be collected on thecatalyst particles as a result of the straightness of the gas passages.

Bead-string reactors represent the limit of parallel-passage reactors: Theycontain single-catalyst-particle subunits. Figure 10 gives a schematic representa-tion (25).

Bead-string reactors have the advantages of lateral-flow reactors but not thedisadvantage of the low mass transfer rates in the units of the lateral-flow reactors.

FIGURE 8 Foam as a catalyst support (alumina impregnated with Cs2SO4.V2O5).


The disadvantage, of course, is the high cost, due to the difficult, labor-intensiveproduction. In a recent patent of ABB, an arrangement is claimed that combinesthe advantages of the bead-string reactor with an easily produced arranged catalystconfiguration: Monolith channels are packed with catalyst particles, resulting instrings of particles. This was described as a structured packed bed (26).

The reactor internals, consisting of structured packings packed with catalystparticles, are also examples of arranged catalysts, e.g., the Sulzers Katapak-S type.

Multifunctional reactors often are also structured reactors. A good exampleis the membrane reactor (27,28). Two types can be distinguished, those based ona catalytic membrane and those in which the membrane only provides a selectiveseparation function without being catalytically active itself. The former is anexample of a structured catalyst, while the latter belongs to the category ofarranged catalysts. The reactor containing a nonactive membrane is referred to asa membrane-enclosed catalytic reactor (MECR). In the following, an example ofa MECR is described.

2.6. Membrane-Enclosed Catalytic Reactor (MECR)

Catalytic membrane reactors are not yet commercial. In fact, this is not surpris-ing. When catalysis is coupled with separation in one vessel, compared to separ-ate pieces of equipment, degrees of freedom are lost. The MECR is in that respectmore promising for the short term. Examples are the dehydrogenation of alkanesin order to shift the equilibrium and the methane steam reforming for hydrogenproduction (29,30). An enzyme-based example is the hydrolysis of fats describedin the following.

FIGURE 9 Example of the TLP reactor: parallel-passage reactor (PPR). (Adaptedfrom Ref. 24.)


2.6.1. Production of Fatty Acids by Fat Hydrolysis in a MembraneReactor

A recent development at laboratory scale is the application of an enzyme (lipase)to catalyze the hydrolysis: Water and fat are mixed at low temperature (300 K) ina continuous stirred-tank reactor (CSTR). The water phase contains the enzyme.A much purer glycerol solution is obtained than in the conventional process. Thedisadvantage is that the equilibrium is not favorable.

An elegant solution has been proposed based on a membrane reactor con-sisting of a module with hollow cellulose fibers [see Fig. 11 (31)]. The enzyme isplaced at the inner side of the fibers, to which the fat is fed. Water passes at the

FIGURE 10 Schematic (a) and reactor configuration (b) of the bead-stringreactor.


outside and diffuses through the membrane to react at the fat/lipase interface. Thefatty acid formed stays in the oil phase, whereas the glycerol formed is trans-ported through the membrane into the water phase. Laboratory studies shownearly complete conversions.

3. GAS-PHASE REACTIONS

It is fair to state that by and large the most important application of structuredreactors is in environmental catalysis. The major applications are in automotiveemission reduction. For diesel exhaust gases a complication is that it is overalloxidizing and contains soot. The three-way catalyst does not work under the con-ditions of the diesel exhaust gas. The cleaning of exhaust gas from stationarysources is also done in structured catalytic reactors. Important areas are reductionof NOx from power plants and the oxidation of volatile organic compounds(VOCs). Structured reactors also suggest themselves in synthesis gas production,for instance, in catalytic partial oxidation (CPO) of methane.

3.1. Environmental Catalysis

Converters for cars are usually ceramic monoliths and occasionally metal based.Without much exaggeration, they can be claimed to be one of the major successesof recent decades in the area of chemical engineering and catalysis. In the begin-ning, the catalytic converter was placed underbody, where sufficient space wasavailable and where the temperature was expected to be mild. There was no need

FIGURE 11 Membrane reactor for the production of fatty acids. (Adaptedfrom Ref. 31.)


for process intensification! Later, performance became critical, mainly because ofthe observation that under steady-state conditions the converter works well butcold start emission is relatively large. Options that are offered include a catalystclose to the engine, electrically heated catalysts, and combinations of catalyst andadsorbent.

Initially, packed beds were also used. They, however, were no success, andat present monoliths are applied exclusively. This should not be misunderstood.Monolith means literally “a single stone.” However, metal-based analogues arealso included in the definition of monolith. In fact, for catalytic converters in cars,in addition to ceramics, metal-based monoliths have been and still are used. Amajor advantage of metal was the thin wall thickness that could be achieved.Later, industry succeeded in manufacturing ceramic structures of comparablewall thickness. In view of their higher resistance against corrosion, ceramicmonoliths are now more generally applied than metal ones.

Structured catalysts are also essential in diesel exhaust gas purification.State-of-the-art solutions are marketed by PSA and by Johnson Matthey. The truckmarket is dominated by diesel engines. In that application, space requirement is a major issue, and intensification is badly needed. Space velocities exceeding100,000 h�1 are demanded. Reactive structured filters are the way to go.

In the wake of the spectacular application of monoliths in the treatment ofautomobile exhaust gas, the potential of monoliths in other applications was stud-ied. Gas-phase reactions were the major area. Catalytic oxidation has received alot of attention. Low-NOx burners based on monoliths were designed, catalyticoxidation of VOCs also benefits from structured catalysts, basically because ofthe low pressure drop and the resistance against dust.

Originally, packed-bed reactors were applied in selective catalytic reduc-tion (SCR). They could be used only in low-dust applications (15). They weresuccessfully replaced by several types of structured catalysts, viz., honeycombs,plate-type catalysts, and parallel-flow systems. Also, this technology is withoutdoubt successful. Volatile organic compounds are destroyed by combustion instructured catalysts usually containing Pt or Pd. Compared to automotive appli-cations, the size of the reactors is large. Figure 12 explains the engineering for thedestruction of VOCs of the large gas flows in industry (9).

3.2. Production of Syngas

In the production of syngas, the following reactions are usually undesired. Thedesired reaction is the production of CO/H2 mixtures according to

2CH4 � O2 → 2CO � 4H2

whereas sequential oxidation giving CO2 and H2O is not desired. This calls forshort residence times (ms), short diffusion length (as small as possible a diameter


FIGURE 12 Industrial unit for SRC containing metal monolith units.


of the catalyst particle), and the absence of nonuniformity with respect to temp-erature and concentrations. Moreover, the pressure drop in this type of applica-tion should be minimal. It is obvious that a structured reactor should be used. Thisleaves three candidates:

MonolithGauzeFoam

Foams have the highest turbulence and the highest pressure drop. Prob-ably because of the first phenomenon they are the most suited, although gauzesalso might be good. Monoliths have the advantage of being well defined, but the absence of radial heat transport will lead to scale-up problems: When a cata-lyst in a channel would “die,” the temperature will drop, and so will the vis-cosity, leading to a “leak” in the reactor. It is clear that radial heat transfer is a keyissue (as it is in packed beds). Advanced designs have been described in the liter-ature. By adapting the geometry, turbulence can be enhanced. Figure 4 illustratesthis.

Gauzes are the state of the art for many millisecond-reactions performed inindustry. The best-known examples are the oxidation of ammonia to NO for theproduction of nitric acid and the Andrussov process, in which HCN is producedfrom methane and ammonia (32):

NH3 � CH4 � 1.5O2 → HCN � 3H2O

The temperature in this process is quite high, 1100–1200�C. It is not surprisingthat under these severe conditions extended reorganization of the alloys takesplace. This is shown in Figure 6 for the oxidation of ammonia into NO.

Many more options are imaginable. A good example is the crosscurrentmonolith (Figure 13). In theory such a system allows ideal heat exchangebetween adjacent channels. Such an elaborate structure might look improbable.However, Corning recently filed a patent claiming the direct extrusion of cross-current structures. So these advanced reactor types might be applied in practicein the future. Naphtha cracking, as a large-scale endothermic reaction, might bea good case for such a reactor.

3.3. Scale-Up

Scale-up of structured reactors is usually easier than for packed-bed reactors. Themajor point is that the hydrodynamics are independent of the scale of the reactor(assuming a good inlet device). When the radial temperature profile is also inde-pendent of the scale, scale-up is straightforward. This is the case for millisecondreactors. In these reactors, rates are very high; as a consequence, in exothermicreactions they operate adiabatically. So they scale easily.


4. MULTIPHASE REACTIONS

Various types of reactors are being used commercially for multiphase applica-tions, the major ones being the slurry reactor, the bubble-column reactor, and thetrickle-bed reactor (5). Figure 14 gives a schematic of these three types of reac-tor. Each reactor has its own advantages and disadvantages. Slurry catalysts aresmall (typically 50 �m), while trickle-bed particles are larger (millimeter scale),in view of the allowable pressure drop over the bed. The particle size is a crucialparameter. In general it can be stated that larger particles are less efficient and,even more important, are less selective in those reactions where the desired prod-uct is subject to the following undesired reaction (A → B → C, with B as the

FIGURE 13 Crosscurrent monolithic structure.


desired product). In that often-encountered case, the slurry reactor is more select-ive than the trickle bed reactor. In terms of process intensification, a mechanicallystirred-tank reactor often is not a good choice. In practice it is no exception thatgas–liquid mass transfer is rate determining. This implies that only the part of thespace close to the tip of the stirrer(s) is well used. A large part of the reactor doesnot contribute much to the productivity and, depending on the kinetics, will leadto low selectivity. Moreover, the major disadvantages of the slurry reactor are theseparation of product and catalyst and catalyst attrition. The trickle-bed reactor ismuch more convenient, but large particle sizes are unavoidable. An importantlimitation of trickle-bed reactors is that, in practice, they are nearly always oper-ated cocurrently, to avoid liquid entrainment by the gas (“flooding”). Someimportant commercial applications, however, would benefit from a countercur-rent operation, especially for equilibrium-limited reactions and in the case ofstrong product inhibition (33). Examples are hydrotreating processes likehydrodesulfurization (HDS), hydrodenitrogenation (HDN), and hydrocracking.Only for large particles or low flow rates could this operational mode be achievedin a packed bed [Synsat process (34)]. Deep desulfiding is a good example of areaction where the concentration profile in countercurrent operation is more opti-mal from a reaction kinetics point of view (2,35). Also, more active catalysts(e.g., noble metals) can be used in the last part of the reactor (“catalyst profiling”)that are more susceptible to H2S poisoning and, as a consequence, are not suit-able for cocurrent operation. Overall, countercurrent operation leads to deeperdesulfurization with smaller catalyst units or to larger throughputs (21). In all

FIGURE 14 Schematic representations of three basic gas–liquid–solid reactorsystems.


three reactor types, in principle a runaway is possible because when hot spotswould be formed, large amounts of reactants can reach the hot spot, leading to aclassical runaway. It will be shown that structured reactors based on monoliths areimaginable that do not possess the unfavorable properties mentioned. In currentpractice, one application is known. AKZO-NOBEL (previously EKA-NOBEL)operates five plants based on the anthraquinone process, in which the reductionstep is carried out in monolith reactors (36,37). Many multiphase reactions havebeen carried out at laboratory scale, and in industry interest is also increasing, asis apparent from patents (26,38–46). Recently, monolith structures have beentried in photocatalysis. This might well be an important application in the future.

4.1. Hydrodynamics and Mass Transfer in Monoliths

4.1.1. Cocurrent Operation

For cocurrent gas–liquid flow, several flow regimes can occur. The preferred oneis usually the so-called Taylor, or slug, flow (47–49). This type consists of gasbubbles and liquid slugs flowing consecutively through the small monolith chan-nels. The gas bubble fills up the whole space of the channel, and only a thin liq-uid film separates the gas from the catalyst (Figure 15). For two reasons, the rateof mass transfer is large. First, the liquid layer between bubble and catalyst coat-ing is thin, increasing mass transfer. Second, the liquid slugs show an internalrecirculation during their travel through a channel. Because of this, radial trans-fer of mass is increased. Moreover, the gas bubbles push the liquid slugs forwardas a piston, and a type of plug flow is created. Compare this with single-phase liq-uid flow through the channels. Because of the low channel diameter, the flow willbe laminar and, as a consequence, the radial transport will be extremely slow,leading to very poor reactor performance: Rates are slow and the reactor exhibitsstrong nonplug-flow behavior. For multiphase operation under slug-flow condi-tions, the mass transfer increase is an order of magnitude larger than for single-phase liquid flow, whereas the increase in friction—that is, pressure drop—ismuch less [Figure 16 (50)]. A fortunate finding is that Taylor flow conditions areeasily realised under practical conditions.

Ideally, in contrast to packed beds, scale-up of monolithic reactors is verysimple. When we know the behavior of one channel, we should be able to predictthe whole reactor. Is this really true? Compared to a packed bed, a monolithic reac-tor differs in radial transport. When the initial distribution of liquid in the radialdirection is nonideal, going down through the reactor, this unfavorable distributiondoes not change. In a packed-bed reactor this happens to a certain degree. Thereforein scale-up, the reactor inlet system has to be designed well so that the distributionof the liquid in the top of the reactor is ideal. We found that if a bubble emulsion ontop of the monolith is present, a satisfactory distribution seems to be guaranteed, asfound for trickle-bed reactor operation. We carried out a large experimental pro-gram and defined the conditions where this happens to be the case. It appeared that


the flow rate has to be above a specified minimal value. Stacking of monolith pieceson top of each other or with some spacing in between, to allow some radial mixing,does not seem to have a negative impact on the flow characteristics.

So the flow rates have to be sufficiently high (linear velocities > 0.1 m/s) inorder to guarantee a good distribution of liquid over the cross section of the reac-tor. One might wonder if upflow of gas and liquid is not to be preferred, becauselower flow rates might be applied. This appeared not to be the case. Again, highflow rates are needed to establish a good gas–liquid flow distribution. It might beworthwhile to investigate whether systems can be developed or conditions estab-lished that allow low flow rates. Combinations of monolithic catalyst packageswith the Sulzer type of contactors are being conceptually investigated in ourgroup. They might increase the window of operability toward lower flow rates.

FIGURE 15 Taylor flow through a single tube. Left: picture of air–water flow;middle: schematic representation of the gas and liquid slugs; right: CFDvelocity pattern in a liquid slug showing the liquid recirculation.


Moreover, they might lead to flexibility, allowing more compact reactor systems.The first results are promising.

4.1.2. Mass Transfer

Mass transfer was studied experimentally in various ways. Nonreactive studiesinvolved the uptake or release of oxygen by the liquid for the measurement of gas–liquid transfer (51–53), while in reactive studies the overall gas–solid orliquid–solid transfer could be determined. As an example of the performance, amonolith in the hydrogenation of �-methylstyrene was compared with a trickle-bed reactor under identical reaction conditions in cocurrent mode. Per unit reac-tor volume, the washcoated 400-cpsi monolith yielded a hydrogenation rate morethan four times higher. For a reaction that is mass transfer controlled, this stressesthe better mass transfer in the monolith. Overall, the Ni was used 40 times moreefficiently in the monolith than in the trickle-bed reactor, even in spite of the useof an eggshell catalyst in both cases. In spite of the high rates observed, it was feltthat not all the Ni in the washcoat layer was optimally used (5,54). In subsequentwork a more eggshell type of coating was realized and the rates observed were anorder of magnitude higher. Mass transfer is usually expressed as the factor kla, themass transfer coefficient times the exposed surface area per unit volume a. Valuesof kla depend strongly on the gas and liquid properties, but for many systems val-ues of 0.5 s�1 are found, and values even much larger than 1 s�1 possibly apply.

FIGURE 16 Relative increase of friction and mass transfer due to gas–liquidTaylor flow, compared to developed laminar flow in small tubes. � repre-sents the dimensionless length of a liquid slug, Re the Reynolds numberbased on the liquid.


This is one order of magnitude higher than in conventional reactor types (1),which underlines the process intensification potential of monolithic reactors.

In Table 3 the three common reactor types are compared. Obviously, themonolithic reactor in the Taylor-flow regime leads to a high degree of processintensification. When these numbers are recalculated into production rates, valuesof 40 mol/m3

reactor-s were found. Figure 17 illustrates the high value in relation tothe “Weisz window of reality.” This demonstrates the attractiveness of usingmonoliths in fast catalyzed gas–liquid–solid reactions.

4.1.3. Countercurrent Operation in Monoliths and Arranged Packings

Under practical conditions, countercurrent operation in a packed bed reactor isnot feasible, because flooding occurs (55,56). The reason is that in the small inter-stitial space, extended momentum transfer takes place between the liquid flowingdown and the gas flowing upward. At velocities used in industry this would imply

TABLE 3 Comparison of Gas-to-LiquidMass Transfer in Three Common Three-Phase Reactor Types

Reactor type kla (s�1)

Trickle bed 0.05–0.2Slurry 0.1–0.3Monolith �1

FIGURE 17 High productivity of multiphase monolith reactors.


that the particle size has to be increased by an order of magnitude. This leads tounacceptable internal diffusion limitations. Clearly, momentum transfer has to bedecreased while maintaining high rates. This can be done by structuring the cata-lyst or by clever arranging of the catalyst particles in the reactor.

Various arranged catalyst structures are used or can be envisaged. Figure 18gives an overview of the most important ones. The principle of these structures isthat relatively large channels are present, leaving space for countercurrent flowwithout extended momentum transfer. In catalytic distillation a lot of experiencehas already been gained in packings based on particles arranged in bales (2).From an extensive study, it appeared that in structured reactors as well, counter-current operation is possible at industrially relevant conditions. The breakthroughwas the design of optimal monolithic structures and dedicated inlet and outletsystems. For example, good results were obtained by cutting the monolith underan angle of 70� as the optimal value (57) or by a special outlet construction, guid-ing the liquid away from the exit (Figure 19). Finned tubes exhibit outlet flood-ing mainly only, whereas the unfinned tube also exhibits inlet flooding. Theunfinned tube has a larger hydraulic diameter due to the absence of the fins andhence a wider flooding-free region. Injecting the gas via a capillary and guidingaway the liquid through quartz wool plugs even enlarges this region for the finnedchannel. This graph illustrates that current operating region for trickle bed reac-tors (HDS) is well covered by the finned monoliths (58).

FIGURE 18 Various arranged particle packings.


4.1.4. Monolith Reactors

The catalyst to be used in a reactor operation can be coated as a thin layer on thechannel walls, and, hence, the reactor can be described as a “frozen slurry reac-tor.” The diffusion length is small and well controllable.

The catalyst loading often is relatively small, but using thicker coatings orusing a monolith extruded from the catalyst support, e.g., an all-alumina mono-lith, can increase it. The high cell density of the monoliths creates a high geo-metric surface area. Using a packed bed, unrealistically small particles would beneeded to achieve this. Catalyst separation and handling are as convenient as in acommon packed bed.

Scale-up is in principle straightforward. Larger channel geometries (e.g., inthe internally finned monolith channels) allow countercurrent operation of gasand liquid. Monolith reactors are intrinsically safer. The monolith channels haveno radial communication in terms of mass transport, and the development of run-away by local hot spots in a trickle-bed reactor cannot occur. Moreover, when thefeed of liquid or gas is stopped, the channels are quickly emptied.

From the foregoing it should be evident that monolithic reactors (and otherstructured reactors) in many respects are superior to classical reactors. Indeed, for

FIGURE 19 Flow map for countercurrent gas–liquid flow (n-decane/air) throughfinned channels with different outlet geometries. Indicated are the floodinglimits for single tubes. Finned tubes exhibit mainly only outlet flooding,whereas the unfinned tube also exhibits inlet flooding. The unfinned tube hasa larger hydraulic diameter due to the absence of the fins and hence a widerflooding-free region. Injecting the gas via a capillary and guiding away the liquid through quartz wool plugs even enlarges this region for the finnedchannel. This graph illustrates that current operating region for trickle-bedreactors (HDS) is well covered by the finned monoliths (From Ref. 58.)


several reactions, monolithic catalysts have been reported, although, except forone case, only at the bench or pilot scale. The interesting points are to demon-strate that the “theoretically” outlined advantages are indeed present. Comparedto classical reactors they in fact boil down to a larger reactor productivity, a better selectivity control, and a higher efficiency. The first point also implies abetter catalyst utilization. Obvious is the fact the catalyst is fixed in a reactor andpressure drops are low.

Highly exothermal reactions can be applied by external heat exchange(1,39). If a CSTR-type reactor is not desired, the horizontal reactor with inter-stage cooling is an attractive alternative.

The hydrogenation step in the anthraquinone process of AKZO-Nobel is anindustrial realization of a monolithic reactor and includes a lot of pioneering workfrom the Anderson group (59–63). More examples of the use of monoliths can befound in Refs. 5 and 64.

In our own group, in cooperation with a chemical industry, we have studiedthe selective hydrogenation of pyrolysis gasoline, a by-product of the naphthacracking that can be upgraded to gasoline by selectively removing gum-formingdienes and styrene-like molecules, leaving intact the internal alkenes. This study(65) demonstrated the plug-flow behavior needed for such a selective conversionand the efficient use of the active phase, which was at least a factor of 3–4 betterthan in a trickle-bed operation. The hydrogenation of �-methylstyrene, men-tioned earlier, is an even more appealing example of better active-phase utiliza-tion and confirms the good mass transfer properties.

An attractive property of monolithic reactors is their flexibility of applica-tion in multiphase reactions. These can be classified according to operation in(semi)batch or continuous mode and as plug-flow or stirred-tank reactor or,according to the contacting mode, as co-, counter-, and crosscurrent. In view ofthe relatively high flow rates and fast responses in the monolith, transient oper-ations also are among the possibilities.

The cocurrent monolith reactor, with its plug-flow characteristics, can inprinciple be used in downflow, upflow, and horizontal-flow modes, provided agood gas–liquid distribution is secured (66). The last mode might solve a majorproblem in practical applications of monoliths: Because, for hydrodynamic rea-sons, high flow rates are needed, the reactor length tends to be very large. Theprocess intensification potential of horizontal configuration, the so-called in-linemonolithic reactor (ILMR), has recently been demonstrated by Stankiewicz (67)for one of the large-scale hydrogenation processes of DSM. It has been shownthat the conventional reactor system, consisting of a stirred-tank reactor and apacked-bed reactor, could be replaced by an ILMR ca. 30–100 times smaller(depending on the type and thickness of the washcoat) (Figure 20). Research withrespect to this type of reactor is in progress. An important outcome of the researchmight be that coupling of monolithic elements, mixing units and heat exchangers,


FIGURE 20 Process intensification in the in-line monolithic reactor, ILMR.


leads to flexible cascade reactor setups, enabling multistep synthesis in one pass.Further extension of the in-line monolithic reactor concept on other unit oper-ations could possibly lead in the future to much more compact, safer, and envir-onmentally friendly chemical plants, in which pipelines would not only serve forsending gases and liquids, but be made functional and used for reactions or sep-arations (68).

The best studied mode is cocurrent downflow. It can be envisaged in twoways, with either a controlled flow of gas or a free recirculation due to entrain-ment by the liquid at the entrance of the monolith (Figure 21). This reactor is analternative to the bubble-column reactor often used in biotechnological applica-tions. Since high reactor types are being used and large gas-flow rates arerequired, the energy input to introduce and compress the gas for injection at thebottom is relatively high. In the downflow monolith reactor, this gas injection isautomatically achieved. The cocurrent reactor type can easily be used as a stirredreactor type by a large recirculation flow without extremely high energy input dueto the low pressure drop. An external heat exchanger can be scaled independentof the reactor to deliver the required heat duty (1,5).

Of course, monoliths have disadvantages. They are at this moment moreexpensive than particle catalysts. In fixed-bed operation, they will have to exhibit

FIGURE 21 Configuration of a cocurrent downflow monolith reactor with freegas recirculation. Only liquid is recirculated, and an external heat exchangercan be scaled independent of the reactor to deliver the required heat duty.


a sufficiently long lifetime. In quickly (irreversibly) deactivating reactions, theywill not be used. Of extreme importance is that the inlet distribution should besecured. In cocurrent flow, both gas and liquid have to be in contact evenly withthe catalyst at the monolith walls.

Countercurrent operation is appealing in many respects and is already exe-cuted in practice. An example is the desulfurization process of Syn Tech, whereco- and countercurrent operation are combined (see Figure 22).

Apart from the GLS-type, LLS- and GLLS-type catalytic reactions are alsopossible using monoliths. The attractive property is to bring reactants efficientlyin contact with the solid catalyst. But there is more. They can be applied in stripping,extraction, evaporation, drying, and distillation, in co- as well as in countercurrentmodes. Monoliths are then used as low-pressure-drop and low-energy-consumingcontacting devices. The combination with catalysis is then obvious to arrive at amultifunctional reactor system in which reaction and controlled reactant additionor product removal is achieved. These applications are not restricted to gaseousor liquid phases, but also work in solid phases. The straight channels are ideal forfixed- or moving-bed applications (Figure 23), the former to combine an opti-mized catalyst inventory and liquid holdup while still having a relatively low flowresistance of a single pellet-string reactor. Moreover, existing catalysts can beapplied. The use of finned channels gives even more freedom. This could be

FIGURE 22 Cocurrent/countercurrent Syn Technology process scheme.


considered a “structured trickle-bed reactor,” where longer residence times can beachieved than in a straight-channel monolith. Blocks of monoliths filled with par-ticles may find application in catalytic distillation or three-levels-of-porosityreactors (26), replacing the catalyst “bales” (5).

Channels filled with a single particle string have much better solid flow char-acteristics than a packed bed, so application of monoliths as the moving-bed reac-tor internal is seducing. This opens a wide range of applications, coveringmoving-bed adsorption processes, moving-bed applications for deactivating cata-lysts (reforming, hydrodemetallization, dehydrogenation), solid trickle-flow reac-tors, and regenerative processes where a moving catalyst is alternatingly subjectedto different atmospheres and transport reaction intermediates and/or heat (FCC,butane to maleic anhydride oxidation). The channel structure also works as a flowstraightener, providing better plug-flow characteristics in large-diameter entrained-flow reactors, which suffer from back-mixing of catalyst at the reactor wall.

Evaluating the properties of catalytic reactors, there are three importantaspects that strongly determine the overall performance: the amount of catalyst andintrinsic kinetics, the transport phenomena (diffusion inside and outside the catalyst),and the hydrodynamics in the reactor. In classical reactors these are strongly inter-related and cannot be defined and designed independently. As an example, for fast

FIGURE 23 Monolith structure packed with catalyst particles (“catalyst bale”)for structured fixed-bed or moving-bed applications.


reactions small catalyst particles are desired from the point of view of catalyst effect-iveness, but a packed bed with small particles will result in an unacceptable pressuredrop. Therefore an optimum has to be sought for the particle size. The elegance ofstructured reactors is that these three aspects can be designed and optimized fairlyindependently, resulting in an optimized reactor performance. Figure 24 shows thesituation for a monolith channel in a gas–liquid reaction. The zeolite catalyst shouldbe very small to take advantage of its high activity. It is embedded in a washcoatlayer on the wall of the monolith channel of a thickness that yields the required cata-lyst effectiveness and selectivity. The channel diameter determines the type of flow,in this case Taylor flow, which optimizes the mass transfer from the gas and liquidphases to the solid catalyst. The straight monolith channels already ensure a lowpressure drop across the structure. This is a structured system covering about 10orders of magnitude, from nanometers to several meters.

If the aim of the catalytic process is to optimize yield and selectivity, onecan distinguish two extremes: fast reactions and slow reactions (Figure 25). Inslow reactions, the intrinsic reaction kinetics control the process, so the catalystinventory should be as high as possible. Increasing the wall thickness of a mono-lith can have the desired effect. In fact the degree of variation in this way is vir-tually from 10–90 volume %, whereas a packed bed will always yield aninventory of around 60% or lower if hollow catalyst particles are used.

In fast reactions, mass transfer or intraparticle diffusion becomes control-ling. Thinner catalyst coatings, Taylor flow, etc. can be applied to optimize these

FIGURE 24 Schematic representation of the operation of a monolith channel,washcoated with a zeolite catalyst, under Taylor-flow conditions.


requirements. If mass transfer is controlling, the productivity is proportional tothe geometric surface area of the monolithic structure. Increasing cell densitiesare recommended, without yielding unacceptable pressure drops. These examplesexemplify the potential power of the application of monolithic structures in cata-lytic reactors.

5. CONCLUSIONS

Monolithic and other structured catalysts exhibit favorable properties with respectto practical convenience, high rates, high selectivity, and low energy consump-tion. From an engineering point of view, the easy scale-up and the potential ofhigh safety are also appealing. This is not limited to single-phase processes, butthey are also well placed for multiphase processing.

Monoliths exhibit a large flexibility in operation. They are well suited foroptimal semibatch, batch, continuous, and transient processing. Catalytic conver-sion can be combined with in situ separation, catalytic reactions can be combined,heat integration is possible, and all lead to process intensification. In the short term,catalytic monoliths will be applied to replace trickle-bed reactor and slurry-phase

FIGURE 25 Aspects controlling the performance of a three-phase catalyticreactor, indicating the flexibility of the use of monolithic catalysts.


operations in view of the better overall conversion and selectivity performance.Monoliths allow efficient use of small catalyst particles, e.g., zeolites, and have asubstantial flexibility with respect to catalyst inventory in a reactor. Multifunc-tional reactor operations like reactive stripping and distillation are challengingapplications that are not too far away. Several options exist for applications in theoil refinery and the chemical process industry.

The essence of the use of a structured reactor is that it allows the decoup-ling of intrinsic reaction kinetics, transport phenomena, and hydrodynamics. Inthis way those phenomena that control the overall behavior of a catalytic reactorcan be optimized independently, giving rise to excellent reactor performance.

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7

Inline and High-Intensity Mixers

Andrew Green

BHR Group Limited, Cranfield, England

1. INTRODUCTION

High-intensity inline devices are often used to mix fluids in the process industries.Such devices include simple pipes, baffled pipes, tees, motionless mixers, dynamicmixers, centrifugal pumps, ejectors, and rotor/stator mixers. In addition to theirtraditional application in physical processes such as mixing and dispersion, suchdevices can provide very effective environments for mass transfer and chemicalreaction to take place. Furthermore, combining effective inline mixing with heattransfer is the basis of combined heat exchanger reactors (HEX reactors).

The chapter provides insight on the importance of mixing and how it relatesto process intensification using inline mixers. Design information for inlinedevices such as motionless mixers, T mixers, ejectors, and HEX reactors is provided.This should assist the reader to: (a) understand the advantages and disadvantagesof these devices as process tools for single-phase, gas–liquid, and liquid–liquidapplications, (b) evaluate manufacturers bids, and (c) identify opportunities forintensifying processes, as either a retrofit for existing plant or as a new process.

1.1. Why Is Mixing Important?

Consider a simple chemical reaction, where two reactants A and B come togetherand produce a product R:

A B R� →


The reaction will have an intrinsic kinetic rate, usually dependent on the localconcentrations of A and B. Often it will produce heat (exothermic reaction) orrequire heat input (endothermic reaction). If this is not removed (or supplied) fastenough, the temperature will rise (or fall), possibly by tens or even hundreds ofdegrees. Clearly this could have disastrous consequences, particularly becausethe rate of reaction will increase with temperature, potentially leading to a run-away reaction. For the reaction to take place, A and B need to be brought together;the reactor must be mixed. This is usually not a problem in a chemist’s beaker,where mixing can be very rapid. However, if it scaled up to a batch stirred vessel,mixing inevitably becomes slower and may take several minutes in a typical production-scale vessel. If this mixing time is slower than the reaction time, thereaction will be artificially slowed down. It becomes mixing, rather than kinetic,limited. In other words, process inefficiency is built in. For highly exothermicreactions, matters become even worse. As a vessel is scaled up, the ratio of heattransfer area to volume reduces, so its ability to remove heat reduces. To cope, aprocess design chemist will alter conditions to slow the reaction down. This mightinvolve running at lower concentrations (i.e., more solvent) or operating semi-batch—feeding B in slowly over many minutes or hours so that the system cancope with the heat release. If a chemical reaction that would naturally take place in a few seconds is slowed down to take 12 hours or more, it is clearly inefficient.

Reactions are rarely as simple as this. There will often be other reactionscompeting with the desired reaction; for example,

In other words, the desired product R reacts with reactant B to form by-product S.If the second reaction is much slower than the first, there should not be too muchS formed. However if mixing is slow, the first reaction can be artificially sloweddown, which will then tend to favor production of S—and yield will reduce. Theflow pattern in the reactor will also influence the production of S. For “back-mixed” flow, as occurs in a stirred vessel, the product stream from the reactionzone will be continually recirculated back into contact with the reactant stream,exposing R to fresh B. In a “plug flow” reactor, reactants are brought together inthe reaction zone and then removed, reducing the likelihood of the formation ofS. To summarize, production of R will be optimized by ensuring that mixing isfaster than the desired reaction step and that the reactor operates in plug flow.

The reactor is the nucleus of the process. Getting the fluid dynamics rightin the reactor means improved safety, productivity, and selectivity, which in turninfluences upstream (reduced raw material costs) and downstream (reduced sep-aration and waste treatment costs); see Figure 1.

A B RR B S

��

→→


1.2. Process Intensification

Process intensification (PI) has various definitions, but from the point of view ofthis chapter it is considered to be a design philosophy in which the fluid dynam-ics of the plant are designed to meet the chemical and physical requirements ofprocess so that it can proceed at its optimal rate. As such, it integrates chemistryand chemical engineering approaches. This can be illustrated by the generalized“S curve” shown in Figure 2. If “plant performance” is poor (e.g., the mixing rateis much lower than the natural speed of the desired reaction), then so is “processperformance” (e.g., selectivity). As plant performance improves (e.g., the mixingrate is increased), so does process performance, up to the point where it becomeschemistry limited. An optimum PI design will be one where the chemistry is

FIGURE 1 Importance of mixing on reactor design.

FIGURE 2 S curve of plant and process performance.


designed to give (in the absence of plant restrictions) the desired process per-formance; the plant is then designed to operate at the point at the top of the S curveas it “flattens out.” Moving further to the right means overdesign and increasedcapital and/or running costs.

In summary, PI aims to match:

Mixing rate to reaction rateHeat transfer performance to heat generationResidence time to reaction timeFlow pattern to reaction scheme

1.3. Motionless (Static) Mixers

A wide range of motionless (static) mixers is available on the market (Figure 3).They are pipe inserts that generate radial mixing (i.e., across the pipe) and (for multi-phase systems) interfacial surface area (e.g., to produce fine bubbles or droplets).The energy for mixing is extracted from the mean flow; as such, an extra pumping

FIGURE 3 Motionless mixers: Chemineer HEV; Sulzer/Koch SMV and SMX;Chemineer Kenics.


duty is incurred. Originally designed for laminar-flow applications, they find wideuse in all flow regimes. The number of elements required for any application isdependent on the difficulty of the mixing duty, more elements being necessary fordifficult tasks.

Motionless mixers (and inline mixers in general) present an alternative tothe more traditional agitated vessel. These devices are particularly useful for thecontinuous processing of chemicals but are also incorporated as part of a batchsystem in pump-around loops.

1.3.1. Attributes and Benefits

Mixing in a motionless mixer is rapid and is achieved by the action of splittingand twisting of the flow by the mixer elements. Energy dissipation rates are high,with typical values between 10 and 1000 W/kg, compared to an upper limit ofaround 5 W/kg in conventional equipment, such as stirred tanks. These large dis-sipation rates give rise to much higher mixing rates for intensified mixers whencompared to stirred tanks.

When two phases are mixed together (gas–liquid, immiscible liquid–liquid),a fine dispersion of bubbles or drops and a high specific interfacial area are pro-duced because of the intensive turbulence and shear. For this reason, resistance tointerphase mass transfer is considerably smaller than in conventional equipment.In addition, a wide range of gas–liquid flow ratios can be handled, whereas in stirredtanks the gas-flow rate is often limited by the onset of flooding. Mass transfercoefficients (kLa) can be 10–100 times higher than in a stirred tank.

The flow pattern in a motionless mixer is approximately plug flow; i.e., dif-ferent elements of fluid spend similar time periods in the mixer. Residence timeis usually short. The combination of rapid mixing and uniform, short, residencetimes is specifically favorable for carrying out reactions with fast kinetics.

Motionless mixers are compact, thus requiring a small site and a lower cap-ital expenditure (CAPEX). Inherent safety is improved due to a smaller reactinginventory. In addition, since there are no moving parts, sealing problems arereduced and maintenance is minimized.

1.3.2. Limitations

High-intensity mixers are not suited to slow reactions (i.e., reaction times greaterthan a few minutes) where long residence times are required. However, it shouldalways be questioned whether the reaction is intrinsically slow or whether it hasbeen artificially slowed to operate safely in a stirred tank. Because these reactorsare almost by definition designed to meet the needs of a specific reaction, therecan be a lack of flexibility if a multiproduct plant is required. This has, however,been addressed by BHR Group’s FlexReactor (Figure 4), which combines motion-less mixers in a highly reconfigurable package.


2. MIXING CONCEPTS

2.1. Reynolds Number

The Reynolds number is the ratio of inertial to viscous forces in a flow. For a pipe:

(1)

The value of Re indicates the flow regime for a specific system. A particularregime is a property of the flow field, not the fluid, which is why the Reynoldsnumber is useful. Re is an important parameter for mixing considerations becausethe flow regime determines the mixing mechanisms of the flow field. At high Re,inertial forces dominate. Energy input is required to sustain turbulent eddies, whichare active at different length scales; a degree of “self-mixing” exists. At low Re,viscous forces dominate. External energy input is required to stretch, chop, andfold fluid and accelerate molecular diffusion. Ultimately, all energy input is dis-sipated to heat.

Repp pu d

��

�

FIGURE 4 FlexReactor (from BHR Group).


2.2. Hydraulic Diameter

The empty-pipe Reynolds number is based on the inner pipe diameter and super-ficial fluid velocity [Eq. (1)]. If the pipe contains a motionless mixer, Eq. (1) needsto be modified to take into account the “metal” in the mixer, which reduces theeffective diameter but increases the fluid velocity (because it blocks part of thecross section). The theoretically sound characteristic dimension for a motionlessmixer is the hydraulic diameter, given by

(2)

The mixer velocity is the superficial velocity divided by the mixer voidage (�),giving

(3)

For a motionless mixer, ReH � Rep, and values of dH /dp vary significantly frommixer to mixer, as shown in Table 1.

2.3. Pressure Drop

In motionless mixers the energy input for mixing is provided by the pressure lossfrom the mean flow. All manufacturers can provide pressure drop data. These areusually given as a friction factor or as a multiplier for the empty-pipe pressuredrop. Values range from 30 to 1000 times the empty-pipe friction factor.

2.3.1. Friction Factors

Care must be taken when comparing the pressure drop in motionless mixers,because three definitions exist. In this chapter, Moody’s friction factor is adopted,

ReHp Hu d

��

��

dH � �4Area open to flow

Wetted perimeter

TABLE 1 Ratios of Hydraulic Diameter to PipeDiameter for Motionless Mixers

Manufacturer Mixer type dH /dp (%)

Sulzer/Koch SMV 7–25SMX 33SMXL 48

Chemineer Kenics KMS 48HEV 86

T1


which is the ratio of the pressure loss in one diameter’s length of pipe to the meanvelocity pressure:

(4)

For fully turbulent flow (generally ReH � 10,000), fM is roughly constant. Forlaminar flow (ReH � 1000), the product fM Re is a constant. Between these val-ues (transitional flow), fM is a function of Re.

Approximate fully turbulent friction factors for motionless mixers are givenin Table 2. (Note: These figures are approximate and for comparison purposesonly; they should not be used for design. The true friction factors vary slightlywith Re and scale.)

The other friction factors in common use are the Newton number andFanning’s friction factor. The relationship between the three is:

(5)Moody’s Newton number Fanning’s

2.3.2. Pressure Drop and Energy Dissipation: Turbulent Flow

For motionless mixers, energy is extracted from mean flow. Data can be correl-ated using an analogy with a rough pipe:

(6)

This can then be used to determine the total energy dissipation rate in the mixer(in W/kg):

(7)��

� �Q p

V

f u

dm

M

� �

3

2

Rep

p pu d�

�

�

f fM F� �2 4Ne

fp d L

uMp m

p

�2

2

� ( / )

�

TABLE 2 Approximate Friction Factors for Motionless Mixers

fMManufacturer Mixer type (approx.)

Sulzer/Koch SMV 6SMXL 2.5

Chemineer Kenics (KMS) 2HEV 0.4Empty pipe 0.001–0.03


Equation (7) illustrates how critical the pipe diameter is to the energy dis-sipation rate: At fixed throughput,

(8)

The energy dissipation rate is a critical parameter in motionless mixers, becauseit affects the rate of mixing. However, not all of the energy dissipated is useful formixing; in particular, laminar dissipation due to shear at the pipe wall or the mixerdoes not contribute to mixing. Total energy dissipation can be split into dissipa-tive losses (ED) and turbulent energy dissipation ():

(9)

2.4. Turbulent-Mixing Length Scales

Turbulent mixing is a complex phenomenon that takes place at a number of scales.Three scales of mixing can be defined (macromixing, mesomixing, and micromix-ing). Macromixing, or blending, is the spreading of an additive by convective flowpatterns and turbulent dispersion. It occurs at scales of typically 10�2 –10�3 m.Following this dispersion, the largest turbulent eddies are broken down into thesmallest turbulent eddies; this is the process of mesomixing, which occurs atscales of typically 10�3–10�4 m. Below the size of the smallest eddy, viscousforces dominate; this is the scale of micromixing. Various processes occur at thisscale, starting with folding and wrapping (“engulfment,” at scales of 10�4–10�5 m),followed by stretching of small eddies with diffusion (10�6–10�7 m).

3. MIXING AND REACTION

For a reactive process, the reactants must be brought into contact by mixingbefore a reaction can occur. In a motionless mixer in turbulent flow, the pressuredrop defines the turbulent energy dissipation rate, which then determines themacro-, meso-, and micromixing rates.

3.1. Slow Reactions

For a “slow” reaction, the mixing rates are all much faster than the inherent kinetics;in this case the mixing and reaction processes are decoupled. For a motionless-mixer system, there must be sufficient residence time downstream of the mixer forthe reaction to go to completion. For long reaction times, a stirred tank can be usedto give the required residence time. The process then becomes:

t tm r1 2 1 2/ /

Component mixing reaction products ⇒ ⇒

� � � �ED

� ∝I

d 7


The term slow needs to be used with care in PI applications. With mixingtimes typically well below 1 second, reactions of only a few seconds’ duration canbe considered slow for a motionless mixer.

3.2. Fast Reactions

For a “fast” reaction, the time scales for mixing and reaction are of the same order(i.e., ), so mixing and reaction are no longer consecutive processes butsimultaneous:

For fast reactions, the mixing rate can limit the product rate of formationand, as described in the next section, product quality/yield.

3.3. Multiple Fast Reactions

Chemical processes often involve multiple, competing reactions. A common situ-ation is that of a competitive-consecutive reaction, such as that described inSection 1.1, where reactant A and the desired product R are competing for react-ant B. The selectively for waste product S can be defined as

(10)

With very fast mixing ( ), the distribution of products is deter-mined by the relative kinetics of the two reactions: If the desired reaction is muchfaster than the undesired reaction, Xs will tend toward zero. However, slow mix-ing compared to the fastest reaction (i.e., ) slows down the desiredreaction, leading to high waste selectivities, Xs → 1. It should be noted that in thisanalysis, is the maximum of the three mixing time scales (macro, meso, micro).Dependent on the mixing conditions, geometry, and chemistry, any one mixingtime scale can be rate determining.

4. MIXING PERFORMANCE OF INLINE MIXERS

4.1. Macromixing (or Blending) Performance

Measurements of macromixing by, for example, a motionless mixer are based on the coefficient of variation (CoV), which is a statistical measure of radialhomogeneity at the macroscale. It is defined as the standard deviation of con-centration measurements made at the exit of a mixer divided by the mean concentration:

t m1 2/

t tm1 2 1 2/ />>

t tm1 2 1 2/ /<<

Xc

c cSS

R S

��

2

2

t tm r1 2 1 2/ /,Component mixing and reaction products⇒

t tm r1 2 1 2/ /≈


(11)

n is the number of measurements (e.g., conductivity probes or sampling positions)used over the pipe cross section (usually �9). ci is the time-averaged concentrationof the ith probe. CoV characterizes the degree of blending achieved between anadditive and the bulk stream. The lower the CoV, the better the streams are mixed.

A CoV of 5%, or 0.05, is often used as the benchmark. The physical mean-ing of this value is that there will be a 95% probability that all samples taken willbe �2CoV (i.e., �10%) of the mean mixed concentration. Correlations for CoVare usually expressed in terms of the ratio between the coefficient of variationdownstream to the coefficient of variation at the inlet to the mixer [i.e.,CoV/(CoV)0].

4.2. Blending Correlations

4.2.1. Empty Pipe: Turbulent Flow

Blending performance for an empty pipe is critically dependent on where andhow the fluid is injected [Eqs. (12) and (13) and Ref. 1].

(12)

(13)

It can be shown that

(14)

For example, for q/Q � 104, Rep � 105, fD � 0.02, and CoV � 0.05, L/d for cen-terline injection is 78, whereas for wall injection it is 234.

It should be noted that when additive flow has significant momentum, muchmore rapid blending is possible (so-called T mixer). An optimum value of momen-tum ratio between main flow and additive can be found (see Ref. 2 for details).

4.2.2. Turbulent and Transitional Flow Mixing in Motionless Mixers

Motionless-mixer manufacturers usually have experimentally based correlationsto predict macromixing performance in turbulent flow. These often use slightlydifferent bases, so care has to be taken when comparing performance.

For the Koch/Sulzer SMV in turbulent flow (Rep � 2000), significant mix-ing continues to be achieved for several diameters downstream of the mixer.

(CoV)01 2 ( / ) /� Q q

Wall injection (low velocity): CoV/(CoV) 2 exp( . / )01 20 25 /� � f L dD

Centerline injection: CoV/(CoV) 2 exp( . / )01 20 75 /� � f L dD

CoV

( ) ] / ( )

�

� �

��

c c n

c c

i

i

n

2

1

1∑

�


Usually mixing is measured/defined two diameters downstream of the mixer, andperformance is given by (3)

(15)

where

L � Lm � 2dp (mixing length)

Lm � nLe (mixer length)

Equation 15 applies for viscosity ratios µB/µA � 100 and shows that the mixinglength to achieve a given CoV is 13 times shorter than for an empty pipe. To takeadvantage of the mixing downstream of the mixer, SMV elements are oftenspaced out in pairs in turbulent flow. The length of one element depends on themixer diameter: For mixer diameters � 100 mm, the length of an element (Le) isequal to one pipe diameter, but for mixer diameters � 100 mm, Le � 0.5dp.

For the Chemineer Kenics and HEV mixers, a correlation has been devel-oped that covers both mixers (4). For fully developed turbulent flow(ReH � 8700):

(16)

For 1000 � ReH � 8700:

(17)

Equations (16) and (17) are for measurements three pipe diameters downstreamof the mixer and are valid for viscosity ratios µB/µA � 100.

The macromixing length (say, to give CoV � 0.05) is insensitive to Reunder fully turbulent conditions. So for a higher velocity, though the mixer lengthremains constant, the time for mixing will be shorter. Put another way, macromix-ing time is inversely proportional to pipe velocity.

For transitional flow, precise correlations are not available; but for bothSMV and Chemineer Kenics mixers, extra elements are required to achieve a cer-tain degree of mixing. The SMV does not achieve significant mixing downstreamof the mixer as in turbulent flow, so elements are not spaced out. The HEV is notrecommended for transitional flow.

4.2.3. Mixer Rankings for Turbulent-Flow Blending Applications

The ranking of mixers for a blending application will depend on what the user istrying to achieve. If blending efficiency is most critical (i.e., achieving therequired mixing for minimum pressure drop/energy use), the most efficient mixeris, in fact, an empty pipe. After that the ranking is:

Empty pipe � HEV � Kenics or SMV � SMX

log [ ) ] . ( . . )10 0 0 27 0 0879 0 7363CoV/(CoV Re0.24�� H n

log [ ) ] . ( . . )10 0 1 65 0 0879 0 7363CoV/(CoV Re0.043�� H n

CoV/(CoV)0 2 1 5exp( . / )� � L dp


However, if the rate of mixing is important (i.e., the need to achieve the most mix-ing in the shortest length), the ranking becomes:

SMV � HEV � Kenics � SMX � Empty pipe

The performance of the HEV should be noted: It was designed to be hydro-dynamically efficient and to use flow vortices, rather than “metal,” to achievemixing. Despite a low pressure drop, the HEV can achieve mixing more effi-ciently and in a shorter length than the Kenics [as shown in Eqs. (16) and (17),it has the same design correlation with respect to number of elements, but ele-ment spacing is shorter]. However, care needs to be taken in its installation,because performance can be significantly degraded if there is an uneven flow dis-tribution at its inlet (e.g., if it directly follows a bend). The poor performance ofthe SMX is to be expected because it has been designed specifically for laminar-flow applications and is not recommended for turbulent-flow applications by itsmanufacturers.

4.2.4. Axial Dispersion

The coefficient of variation is a measure of mixing across the pipe cross section(“radial dispersion”). In pipe-flow and motionless mixers, mixing along the lengthof the mixer (axial dispersion) occurs. This can be described in terms of the resi-dence time distribution (or RTD), which is a measure of age distributions for fluidelements passing through the mixer. An empty pipe has relatively high axial dis-persion, primarily caused by the flow profile that is established in a pipe (i.e., thefluid in the center of the pipe flows faster than that near the walls). Motionlessmixers tend to have a much tighter RTD.

4.2.5. Blending Correlations: Laminar Flow

Despite the wide use of the “striation thickness” concept in the early commercialliterature, the CoV is now the most widely used mixing index. The following cor-relations are valid for viscosity ratios 0.01 � µB/µA � 100, feeding into the centerof the pipe and with CoV measured 2dp downstream for Sulzer mixers, 3dp down-stream for Kenics mixer.

SMX (Rep � 200, Le/dp � 1.0):

(18)

SMXL (Rep � 200, Le/dp � 3.3):

(19)

Kenics (Rem � 200, Le/dp � 1.5):

(20)log ( ) ) . . /10 0 0 098 0 067CoV/(CoV � �� n L dm p

log ( ) ) . /10 0 26 0 078CoV/(CoV � �� n L dm p

log ( ) ) . . /10 0 0 19 0 19CoV/(CoV � �� n L dm p


where

n � number of motionless-mixer elements

Le � length of one mixer element

These data are consistent between the manufacturers’ and independent investigators.

4.2.6. Mixer Rankings for Laminar-Flow Applications

As with turbulent flow, ranking depends on the requirements of the application.Based on various literature data, where at least two datasets exist, the followingrankings can be made.

Energy efficiency (most efficient first):

SMXL → Kenics → SMX → Hi-mixer → Komax → Lightnin→ Ross ISG

Mixing rate (most rapid first):

SMX → Ross ISG → Hi-mixer → SMXL → Kenics → Komax→ Lightnin

4.2.7. High Viscosity Ratios

Laminar-flow blending duties involving high viscosity ratios (greater than 1000 : 1)are classified as difficult. The SMX mixer appears to have the best track record inachieving satisfactory results.

4.3. Mixing with Reaction in Inline Mixers

If a mixer is to be used for reactive processes, it should be designed such that thelongest mixing time scale (whether micro-, meso-, or macromixing) is signifi-cantly shorter than the characteristic time scale of the desired chemical reaction.As mentioned in Section 3, any of the time scales can be rate determining.

4.3.1. Micromixing Limited

If micro � meso � macro , then the process is micromixing controlled. Micromixingis a complex phenomenon (Section 2.4), but for most liquids engulfment is thelongest step. In this case, micromixing time is the inverse of engulfment rate (E)and can be estimated by

(21)�

micro � �1 1

0 058

1 2

E .

/

ε


Detailed models have been developed that can predict product distribution forcompetitive reactions with known kinetics under micromixing- (engulfment-) con-trolled conditions (5).

4.3.2. Mesomixing Limited

In motionless mixers, mesomixing time can be estimated from

(22)

Numerical models for mesomixing control with reactions are being devel-oped, although they are more complex than micromixing models and require theinput of empirically determined length scales. Mesomixing limitations give riseto worse process performance than if micromixing alone were limiting, so if pos-sible mesomixing time should be reduced (e.g., by increasing the number of addi-tive feeds or reducing the additive flow rate) to the point that micromixing controls.However, in practice this is often not possible.

4.4. Scale-Up/Scale-Down of Motionless Mixers

(Single Phase)

For systems involving fast reactions where reactor performance has been estab-lished at one scale and equal performance is required at different scales, the cri-teria for scale-up/scale-down are:

The mixing rate of the limiting step (characteristic time scale) should be keptconstant.

Residence time in the mixer should be constant.The limiting mixing mechanism should not change.

The process conditions should remain the same, e.g., reactant concentra-tions, flow rate ratio, mixer type, relative feed position. If the friction factor,mixer voidage, and turbulence-generating efficiency do not vary significantlywith scale, then the following scale-up rules can be applied (where k � Qnew/Qold,the mixer diameter should be rescaled from dold to dnew).

4.4.1. Macromixing or Mesomixing Limitation

For a macro- or mesomixing limitation,

(23)d k dnew old� 1 3/

�

meso � 2 17

1 3

.

/Q

n uB

f ε


The resulting design will have the same number of elements at the new scale(resident time constant), so the length of the mixer will be

(24)

4.4.2. Micromixing Limitation

For a micromixing limitation,

(25)

If residence time is kept constant, fewer elements will be required at the largerscale, and

(26)

5. GAS–LIQUID MIXING

5.1. Introduction

Gas–liquid reactions form an integral part of the production of many bulk andspecialty chemicals, such as the dissolution of gases for oxidations, chlorin-ations, sulfonations, nitrations, and hydrogenations. When the gaseous reactantmust be transferred to the liquid phase, mass transfer can become the rate-limitingstep. In this case, the use of high-intensity mixers (motionless mixers or ejectors)can increase the reaction rate. Conversely, for slow reactions a coarse dispersionof gas, as produced by a bubble column, will suffice. Because a large variety ofequipment is available (bubble columns, sieve trays, stirred tanks, motionlessmixers, ejectors, loop reactors, etc.), a criterion for equipment selection can be established and is dictated by the required rate of mass transfer between the phases.

5.2. Mixer Types

5.2.1. Motionless Mixers

When a gas stream is introduced into a turbulent liquid flow in a motionlessmixer, the gas is broken up into bubbles. The breakup is due mainly to the turbulentshear force of the liquid but also partly to the collision between gas and the lead-ing edge of an element.

There are two basic operating modes:

Continuous, as a stand-alone mixerLoop operation (either semibatch or continuous).

ld

dlnew

new

oldold�

1 3/

d k dnew old� 3 7/

ld

dlnew

new

oldold�


A stand-alone mixer requires the mass transfer/reaction to be completedwithin the mixer. If the gas flow rate matches the stoichiometry of the liquid phase,all the gas should be dissolved and reacted at the end of the mixer. This generallyinvolves very high volumetric ratios between gas and liquid. If there is excess gas,there will be some gas at the mixer outlet, which needs to be separated.

5.2.2. Gas–Liquid Ejectors

Ejectors consist of four main sections (Figure 5):

Spinner—orients and stabilizes the flow.Nozzle—provides a high-velocity jet of fluid.Gas chamber—the high-velocity jet creates suction in the gas chamber,

entraining gas into the ejector.

FIGURE 5 Gas–liquid ejector.


Mixing tube—on leaving the gas chamber, the liquid jet attaches itself tothe mixing tube wall, resulting in a rapid dissipation of kinetic energy,creating an intensive mixing zone known as the mixing shock region.High turbulence in this region breaks up the gas, producing a fine dis-persion of bubbles and consequently a large interfacial area for masstransfer.

By and large, ejectors and motionless mixers have similar mass transferperformance at a given gas-to-liquid flow ratio and energy input. However, eject-ors have a number of benefits and drawbacks compared to a motionless mixer. Onthe positive side, the ejector suction means that a pressurized gas supply is notrequired. The unrestricted mixing tube means that solid formation due to reactionis not problematic. Against this, the operation is sensitive to changes in thegas–liquid flow ratio and diameter/length ratio. Gas-to-liquid flow ratios are alsomore limited in ejectors.

5.3. Loop Reactors

Motionless mixers and ejectors are useful for applications requiring short resi-dence times (on the order of seconds or less). If long residence times are required,e.g., if the reaction is relatively slow, the use of a motionless mixer alone wouldlead to a very long mixer, which may not be practical. One way to overcome thisproblem is to use a loop reactor, which combines a high-intensity mixer, such asa motionless mixer or ejector, with a separation tank.

5.4. Guide to Equipment Selection

Tables 3 and 4 summarize where different mixers/configurations are most appro-priate.

TABLE 3 Application of Mixer Types and Configurations

Motionless-mixer Motionless- Ejector Ejector stand-alone mixer loop stand-alone loop

Solids present � �Slow reaction � �Fast reaction � �Energy � �

efficiencyimportant

Low gasPressure � �


5.5. Mixing Concepts

5.5.1. Rate of Mass Transfer

The rate of mass transfer for motionless mixers and ejectors can be described by

NA � KLa �CLM (27)

where NA is the amount of transferred species per unit time per unit dispersionvolume, KL is the overall mass transfer coefficient, a is the specific surface areafor mass transfer, and �CLM is the log mean concentration driving force. 1/KL,the overall mass transfer resistance, is usually dominated by the resistance in the liquid phase, 1/kL. Consequently, the gas-phase resistance can be neglectedand KL � kL. However, it is imperative that this assumption be checked, becauseit does not always hold for very soluble gases or when kLa is enhanced by reaction (6).

The large levels of turbulent energy dissipation produced in high-intensitymixers act to reduce the bubble size, typically from 0.5 to 2.0 mm in high-intensitymixers, compared to 1.0 to 5.0 mm in stirred tanks and bubble columns. In addition,much higher gas-to-liquid ratios can be achieved, and turbulence enhances kL,leading to overall mass transfer coefficients (kLa) 10–100 times greater than for astirred tank.

5.5.2. Reaction Regime

The relative speed of kinetics to mixing is described by the Hatta number,

(28)Ha �t

tMT

R

0 5.

TABLE 4 Application of Chemineer Kenics andKoch/Sulzer SMV Mixers

Mixer type

Situation Kenics SMV

�p constraint � �Space constraint � �Solids present � �Energy efficiency important � �Need for heat removal � �


where tMT and tR are the time constants for mass transfer and reaction, respectively.The lower the value of Ha, the faster the mixing relative to the intrinsic reactionrate. Ha can be calculated from

(29)

Reactions are often classified into four categories:

Ha � 0.02 slow reaction0.02 � Ha � 2 moderately fast reactionHa � 2 very fast or instantaneous reaction

For Ha � 0.02, there is a considerable scope for process intensification. If a reac-tion is intrinsically fast (a large reaction rate constant) the design aim is to provide sufficiently intense mixing to move it into the slow reaction regime(Ha � 0.02) such that the reaction is limited by the intrinsic reaction rate ratherthan the mass transfer rate.

In order to establish the reaction regime and to design equipment, the fol-lowing need to be known:

Flow patternMass transfer coefficientBubble sizes

These can be determined from gas–liquid flow rates, the energy dissipation rate(driven by the pressure drop), and the physical properties of the fluids.

5.6. Design Guidelines and Correlations

5.6.1. Flow Patterns

Gas–liquid flows are much more complicated than single-phase flows, due to theexistence of the gas–liquid interface. The phases can be present in a range of pos-sible flow regimes (flow patterns), which are dependent upon the physical prop-erties of both phases, the flow rates, and the equipment size and orientation. Themost commonly noted flow patterns are (7):

Annular flow—a liquid film on the walls and a continuous gas phase, con-taining a mist of liquid droplets, in the core

Intermediate slug flow—large gas voids containing liquid dropletsBubble flow—continuous liquid flow with a dispersion of gas bubbles

Figure 6 shows the various flow patterns in horizontal flow, and similar pat-terns can be seen in vertical upflow or downflow. In general, bubble flow developsunder high liquid-flow rates and low gas-flow rates; annular flow develops under

Ha ��

�D

kk C C

nA

LAB AL

nBLn

A

A B2

1

0 52

1( )

.


low liquid-flow rates and high gas-flow rates; stratified flow develops in low gas-and liquid-flow rates.

Bubble flow is generally more desirable for liquid-film controlled masstransfer processes because of the high turbulence level in the liquid phase, while anannular flow is more desirable for gas-film controlled processes, where the turbulence level in the gas phase is high. However, in reactive systems, stoichio-metry will often define the gas- and liquid-flow rates, leaving no choice for the flowpattern. Having said this, motionless mixers and ejectors can maintain the bubbleflow regime even at high gas-to-liquid flow ratios, where flooding of the impellerwould occur in a stirred tank or annular flow develop in empty pipes (Figure 7 (8)).

5.6.2. Pressure drop

Pressure drop is a critical parameter, in that it determines pumping requirementsand enables the power input to the mixer to be calculated.

Motionless Mixers. Major mixer manufacturers agree that the Lockhartand Martinelli parameters for two-phase flow in pipes (9) can also be applied tomotionless mixers. To estimate the pressure drop, the single-phase liquid and gaspressure drops are first calculated. The Lockhart and Martinelli para-meter X isfound from

(30)Xp

pL

G

��

�

FIGURE 6 Gas–liquid flow patterns in horizontal flow.


FIGURE 7 Flow pattern maps for cocurrent air–water upflow through motion-less mixers. (From Ref. 8.)


and the two-phase pressure drop can be calculated from

(31)

where �L and �G are functions of X for the liquid and gas phases, respectively,obtained from the Lockhart and Martinelli charts (Figure 8, Ref. 9) or from thefollowing empirical equations:

(32)

�G � X2�L (33)

Ejector. Gas chamber pressure needs to be known in order to calculate ejector power input. A semiempirical equation was developed byHenzler (10) that related the entrainment ratio to other system variables:

(34)

Gas chamber pressure, ps, can be calculated from this equation through iteration.Factor B depends on the mixing tube/nozzle diameter ratio for a given ejectortype and needs to be determined experimentally.

Q

QB

D

D

p

p

p p

uG

L

m

n

L

G s

s

L j

� � ��

�

1 0 38 12

0 09

2

1 6

22.

( ). /

�

� �

�L X X� � �� ( . . . ). . .4 6 12 5 0 651 78 0 68 0 5

� � �p p pL L G G� �� 2 2

FIGURE 8 Lockhart and Martinelli parameters for pressure drop in multi-phase flow.


The nozzle pressure, which determines the selection of the liquid pump, isgiven by

(35)

5.6.3. Power Input

The power input, required for calculation of the mass transfer coefficient, is cal-culated from Eqs. (36) and (37). For a motionless mixer, the power comes fromthe gas and liquid phases; for the ejector, power comes from the liquid only.

In a motionless mixer:

(36)

In an ejector:

(37)

Generally, kLa values at the same power input are similar between the two de-vices (11).

5.6.4. Mass Transfer Coefficients

The amount of gas transferred is proportional to the product of the mass transfercoefficient (kL) and the specific area (a). Because most measurement techniquesmeasure this product, many correlations for kLa appear in the literature. However,caution is advised, because they can give different predictions for the same oper-ating conditions. Equations (38) and (39) are two examples from independentinvestigators for motionless mixers, from Refs. 12 and 8, respectively:

(38)

(39)

The majority of reported correlations for ejectors are for loop-type config-urations, e.g., (13):

(40)

with very little reported on the stand-alone configuration.

k aP

VL � 0 0440 76

..

k aP

mL G� 0 640 75

..

k aP

VL � � �1 74 10 40 8

..

P Q p Q pp

pQ u uL G

GL L L� � � �� av ln

21

22

21

2

( )�

P Q p Q pp

pQ u uL G L L L� � � �� av ln 1

21

22

21

2

( )�

p p uD

Ds L jn

12

1

41

21� � ��


Comprehensive correlations for motionless mixers and ejectors have beendeveloped as part of BHR Group’s HILINE Consortium, but these are availableonly to members.

5.6.5. Bubble-Size Calculations

When a gas stream is introduced into a turbulent liquid flow in a motionlessmixer, the gas is broken up into bubbles. The breakup is due mainly due to theturbulent shear force of the liquid but, for motionless mixers, also partly to thecollision between the gas and the leading edge of an element. Gas dispersion is aphysical process and involves bubble breakup and coalescence, which can bothtake place in the same mixer/reactor.

Bubble breakup and coalescence are both complex processes. In a turbulent-flow field, bubbles are broken up mainly due to the turbulent shear force, and theeventual bubble size is a balance between this force and the surface tension force.For a given gas–liquid system and flow field, a maximum bubble size exists. Anybubbles larger than this size will be broken up. According to theory (14), this max-imum bubble size relates to gas–liquid physical properties and flow characteristics:

(41)

We�crit is the modified critical Weber number, which is close to 1.Coalescence occurs when two bubbles approach each other, collide, and

become one bigger bubble. Two important factors are:

Frequency of collisionEfficiency of coalescence

The frequency of collision relates to the flow pattern and gas volume fraction: Themore random the flow pattern or the higher the gas volume fraction, the higherthe frequency. The efficiency of coalescence relates to physical properties of thegas–liquid system. Some systems, such as air–water, have a high efficiency ofcoalescence and are often called “coalescing systems.” Other systems, such asgas–alcohol or gas–salt solution, have a low efficiency of coalescence and arecalled “noncoalescing systems.”

5.6.6. Characterization of Bubble-Size Distribution

It is useful to define an appropriate average to characterize bubble-size distribu-tion. For heat and mass transfer, the Sauter mean diameter (d32) is generally used:

(42)dn d

n di i

i i32

3

2��

�

dc d

max

. .

..� �We

2crit′

( )

0 6 0 6

2 0 20 4�

� �ε


Hesketh et al. (2) proposed that the general equation of Hinze (14) was valid forthe turbulent dispersion of a gas using any motionless mixer or an empty pipe:

(43)

where Cn� 0.6 and We�c� 0.6–1.6.The Sauter mean diameter (d32) is related to the interfacial area per unit

volume (a) and dispersed phase volume fraction by

(44)

where � � QG/(QG � Q1).If kLa and a are known, kL can be estimated and, provided the reaction

kinetics are known, Ha can be calculated from Eq. (29) and the reaction regimededuced.

6. LIQUID–LIQUID DISPERSIONS

6.1. Introduction

Motionless mixers are highly effective for producing dispersions of immis-cible liquids. Applications can be physical (e.g., for liquid–liquid extraction) or chemical (e.g., many nitration reactions). As with gas–liquid mixing, the most relevant parameter to measure for such applications is the Sauter meandiameter.

6.2. Turbulent-Flow Correlations

A range of correlations is available from the literature, usually relating the Sautermean diameter to the Weber number, which is the ratio of shear forces to surfacetension forces:

(45)

Most correlations show that d32 is proportional to the Weber number raisedto the power of �0.6, which is consistent with the theory of drop breakup by tur-bulent shear forces. Strictly, these correlations should be applied only where thedrop size is in the inertial subrange of turbulence, i.e.,

�k � d32 � dp/4 (46)

We� � �c p pu d2 /

ad

� 632

�

d Cnc

L G

32

0 6 0 6

2 20 4

2� �We′

( )

. ..�

� �ε


where �k is the Kolmogorov length scale

�k � (v3/)1/4 (47)

6.2.1. Kenics Mixer

Two very similar correlations have been developed for Kenics mixers (15,16) ofthe form

(48)

where C � 0.45 for the former and 0.49 for the latter. These are valid for fully tur-bulent flow in a pipe (Rep � 12,000), inviscid drops (�d � �c) and were measuredwith water as the continuous phase. The dispersed-phase fraction has little effect upto a value of 0.25. Equilibrium drop size is achieved with only eight mixer elements.

For viscous drops (�d up to 200 mPa) in turbulent flow, Berkman andCalabrese (16) developed the correlation further to give

(49)

where

(50)

For higher viscosity ratios, more elements were needed (24 in these experiments).

6.2.2. Sulzer (Koch) Mixers

Sulzer published correlations in the open literature for drop size. An early correl-ation for the SMV is (3)

(51)

where We� is the Weber number based on hydraulic diameter, i.e.,

(52)

The correlation was developed for five 50-mm SMV elements (dH � 8 mm) andcovered Reynolds numbers (ReH) in the range 200–20,000 and dispersed-phasevolume fraction up to 0.25.

We c′ ��

�

u dp H2

d

dHH

32 0 5 0 150 21� �. . .We Re′

Vi ��

�

�

�

d p c

d

u

1 2/

d

d

d

dp p

32 32

1 3 0 6

0 49 1 1 38� ��. .

/ .

We Vi0.6

d

dC

p

32 � �We 0.6


More recent work (17) has covered viscous drops in turbulent flow, over a wide range of mixer types (SMV, SMX, SMXL, SMR, SMF), diameters (up to80 mm), lengths, and dispersed-phase viscosities (up to 70 mPa):

(53)

where Wec � 1.8 and Vi is the viscosity number [Eq. (50)].

6.3. Comparison Between Motionless Mixers

and Stirred Tanks

Comparison between drop sizes for a motionless mixer and a (well-) stirred tankunder typical operating conditions yields an interesting result that similar dropsizes can be obtained in both despite very different average energy dissipationrates. This is down to the extremely inhomogeneous energy dissipation in a stirredtank, where the peak rate (usually close to the impeller tip) can be similar in mag-nitude to that in the more homogeneous motionless mixer. However, a stirred tankmay take hours of operation to achieve an equilibrium drop size, whereas a motion-less mixer will achieve it with a few elements and in less than a second. This hasimplications for two-phase applications when it is important to rapidly achievemass transfer to complete fast chemical/physical reactions or to minimize by-product formation from complex fast reactions.

6.4. Scale-Up/Scale-Down

When scaling up or scaling down a liquid–liquid process in turbulent flow, theenergy dissipation rate needs to be kept constant, giving

(54)

However, checks should be made that flow is fully turbulent at both scales andthat the drop size remains within the inertial subrange of turbulence [Eq. (46)].As a minimum, residence time should be maintained, i.e.,

(55)

This suggests that fewer elements will be required at a larger scale; however, fora conservative design on scale-up, the same number of elements should be main-tained. On scaling up or scaling down, the same mixer type and feed arrangementshould be maintained.

ld

dlnew

new

oldold�

1 3/

d Q Q dnew new old old� �( / ) /3 7

d kB

dc

c

c

d32

0 6 0 6 0 1

0 40 65 11

2� �

� �. ( )( ) . . .

.��

�

�

�

Vi We

ε


7. COMBINED HEAT EXCHANGER-REACTORS (HEX REACTORS)

7.1. HEX Reactor: Concepts

The mixers discussed in Sections 4–6 are particularly suitable for reactions wherethe required heat input (endothermic reaction) or heat production (exothermicreaction) is modest (i.e., temperature changes on reaction would be only a fewdegrees in the absence of any heat transfer). HEX reactors can be used for rapid,highly exothermic (or endothermic) reactions; not only are the mixing rate andresidence time of a reactor matched to the kinetic rate and reaction time, but heattransfer performance is also matched to heat production (Figure 9).

7.2. HEX Reactor Types

HEX reactors generally fall into three basic types (18).

a. Jacketed motionless-mixer reactors (Figure 10): Motionless mixersprovide a highly effective and efficient mixing environment for rapidreactions. Heat transfer capacity can be provided by utilizing either single mixers in jackets or multiple mixers in a shell-and-tube geometry.The “FlexReactor” (Figure 4) has been designed to package motionlessmixers in a highly reconfigurable unit with effective heat transfer.

b. Compact heat exchangers: There is a wide variety of commerciallyavailable compact heat exchangers available (e.g., enhanced shell andtube, plate and frame, plate–fin, and diffusion-bonded). Such devicesprovide extremely effective heat transfer but have not been optimizedas reactors, compromising their efficiency.

FIGURE 9 HEX reactors: principle of operation.


TABLE 5 HEX Reactor Considerations

Attribute Bespoke HEX Motionless mixer FlexReactor

Typical heat U � 1500 W/m2-K U � 800 W/m2-K U � 600 W/m2-Ktransfercoefficient

HT density 5000 m2/m3 900 m2/m3 900 m2/m3

Material of Limited Wide Wideconstruction

Scale-up Use more units Increase tube Increase tube procedures in parallel length/ length/

diameter diameterCost High Low MediumMixing Low High HighMultipoint feed Yes—but fixed Yes—but Yes—variable

once fixed once constructed constructed

Availability Low High MediumResidence time Low High High and

variableFlexibility Low Low High

FIGURE 10 Jacketed motionless mixer.


c. Bespoke HEX reactors: A good example is the MarbondTM reactor,supplied by Chart Industries, which has been designed specifically as aheat exchanger-reactor, combining high heat transfer with effectivemixing (although its high surface area results in relatively lowmicromixing efficiencies). Passageways are typically of the order of afew millimeters. Its diffusion-bonded construction is very robust, and itcan be constructed to provide an optimized feed arrangement.

The choice of reactor will be very dependent on the requirements of thechemical reaction scheme, the relative importance of mixing and heat transfer, andpractical considerations (e.g., the effect of solids in the process; materials of con-struction; flexibility). A comparison of the typical performance of different designsis given in Table 5. HEX Reactors are discussed in more depth in Chapter 4.

NOMENCLATURE

Symbol Explanation Units

a Interfacial area m�1

ci Concentration of species i mol-m�3

c Mean concentration mol-m�3

CoV Coefficient of variation —Cn Constant � 0.6 —C*AL Equilibrium concentration of A mol-m�3

CBL Bulk liquid concentration of B mol-m�3

D Stirred-tank impeller or rotor diameter mDAL Diffusion coefficient of A in liquid m2-s�1

Dm Molecular diffusivity m2-s�1

Dta, Dax Axial turbulent dispersion coefficient m2-s�1

Dtr Radial turbulent dispersion coefficient m2-s�1

d Droplet diameter mdB Feed pipe diameter mdH Hydraulic diameter of motionless mixer mdmax Maximum stable drop size mdn nozzle diameter mdp Internal pipe diameter md32 Sauter mean diameter mE Engulfment rate coefficient s�1

EB Ratio of viscous to interfacial forces —Ep Rate of direct energy dissipation W-kg�1

fM Moody’s friction factor —g Acceleration due to gravity ms�2

KL Overall mass transfer coefficient ms�1

k Second-order reaction rate constant m3-mol�1-s�1

kL Liquid-side mass transfer coefficient ms�1


Lc Concentration integral length scale mLe Length of one mixer element mLm Length of motionless mixer mLR Length of reaction zone mm mass kgN Impeller rotational speed rpsNA Mass transfer rate mol/sNi Total number of moles of species i moln Number of sampling positions used in

CoV measurement —ne Number of motionless mixer elements —nf Number of additive feed tubes —P Power WP1 Inlet pressure PaP2 Exit pressure PaPs Gas chamber or suction pressure Pa�p Pressure drop over motionless mixer PaQ Volumetric flow rate m3-s�1

T Dimensionless time —Stirred-tank diameter m

t Time sup Superficial pipe velocity m-s�1

v Velocity m s�1

V Liquid volume m3

X Lockhart and Martinelli parameter —Xi Selectivity for product i —z Ratio of mixer friction factor to pipe

friction factor —

Greek Symbols

� Flow rate ratio (QA,C /QB) —� Shear rate s�1

Turbulent energy dissipation rate W-kg�1

� Efficiency of turbulence generation —� Voidage of a motionless mixer —� Turbulence-generating length scale m�k Kolmogorov microscale m� Dynamic viscosity mPa-s� Kinematic viscosity m2-s�1

� Fluid density kg-m�3

� Standard deviation —Interfacial tension N-m�1

Characteristic time scale s� Total energy dissipation rate W-kg�1

� Dispersed phase volume fraction —


Dimensionless Groups

Da Damkohler number (k2-o cB0 /E)Re Reynolds number (updp/v)Sc Schmidt number (v/Dm)We Motionless-mixer Weber number (�cup

2dp /�)Stirred-tank Weber number (�c N2D3/�)

We�crit Critical Weber number (0.6–1.6)Ha Hatta number

Subscripts

1 At inlet2 At exitav Averagec Continuous phased Dispersed phaseE EngulfmentG GasH Based on hydraulic diameterj JetL LiquidMT Mass transferm Motionless mixermacro Macromixingmeso Mesomixingmicro Micromixingmix Slowest mixing stepp Empty pipeQ Competitive-parallel reactionsR RadialR ReactionT Tangential1/2r Reaction half-life

REFERENCES

1. Henzler HJ, Chem Ing Tech, 1980; 52:659–661.2. Hesketh RP, Russel T, Etchells AW. R&D notes. AIChE J 1987; 33(4):663–667.3. Streiff FA. Sulzer Tech Rev 1977; 3.4. Knight CS. Experimental investigation of the effects of a recycle loop/static mixer/

agitated vessel system on fast, competitive-parallel reactions. PhD dissertation, Univer-sity of Arkansas, 1994.

5. Baldyga J, Bourne JR. Principles of micromixing. J Fluid Mechanics 1986; 1:147.6. Middleton JC. In: Harnby N, Edwards MF, Nienow AW, eds. Mixing in the Process

Industries. 2d ed. London: Butterworth Heineman, 1992:Chapter 15.


7. Drahos J, Cermak. Chem Eng Processes 1989; 26:147.8. Roes AWM, Zeeman AJ, Bukkems FHJ. I Chem E Symposium Series, 1984; (87):

231–238.9. Lockhart RW, Martinelli RC. Chem Eng Prog 1949; 45:39.

10. Henzler HJ. Chem Eng Tech 1980; 52:659–661.11. Zhu M. Proc. 1st International Conference on Process Intensification, Antwerp,

Belgium, December 6–8, 1995, organized by BHR Group, Cranfield, UK, 51–59.12. Middleton JC. AIChE 71st Annual Meeting, Miami Beach, Paper 74E, 1978.13. Dutta NN, Raghavan KV. Chem Eng J 1987; 36:111–121.14. Hinze JO. AIChE J 1955; 1:289–295.15. Middleman S. Ind Eng Proc Des Dev 1974; 13:78.16. Berkmann PD, Calabrese RV. AIChE J 1988; 34(4).17. Streiff FA, Mathys P, Fischer TU. New fundamentals for liquid–liquid dispersion

using static mixers. Récents Progrès en Génie des Procédés 1997; 11(51):307–314.18. Green A, Johnson B, Westall S, Bunegar M, Symonds K. Combined chemical reactor/

heat exchangers: validation and application in industrial processes. 4th InternationalConference in Process Intensification, Brugge, Belgium, September 2001.


8

Reactive and Hybrid Separations: Incentives, Applications, Barriers

Andrzej Stankiewicz

DSM Research, Geleen, The Netherlands

1. INTRODUCTION

Integration of various steps/operations presents one of the most promising waysfor intensifying (bio)chemical processes. It can be achieved either by combiningreaction and separation in a single reactive separation step or by combining two(or more) separation techniques in a hybrid separation unit. Such an integrationmay bring a number of advantages to the process under consideration, not just adecrease in the size of equipment.

This chapter provides a general overview of the reactive and hybrid sep-arations and discusses their place in the intensification of (bio)chemical processes.Written from an industrial point of view, it focuses on the application aspects ofthose integrative technologies. Potential application fields are reported, alongwith already existing commercial-scale operations. Special attention is given tothe barriers that hamper a broader introduction of the reactive and hybrid separ-ations into industrial practice and the ways to overcome those barriers. The mod-eling and design aspects of three reactive separation methods (reactivedistillation, reactive absorption, and reactive extraction) are discussed in moredetail in Chapter 9.


2. REACTIVE SEPARATIONS—WHY INTEGRATE?

In the simplest case, integration of reaction and separation may take place on thepurely equipment level, without introducing any new functional interrelationsbetween the operations involved—the reaction does not influence the separation,nor has the separation process any effect upon the reaction. The aimed result ofsuch combination can be:

Lower investment costs (compact plant layout, integral design)Smaller inventory (safety aspects)Improved heat management/energy utilization.

The Urea 2000plus™ technology, developed by Stamicarbon B.V. (1) anddescribed further in Chapter 12, presents a typical example of such a “noninter-relating” integration. The integration resulted here in a considerably smaller andcheaper plant, with much less high-pressure equipment/piping needed and lessenergy consumption. Yet the interrelations between the reaction and other oper-ations remained basically the same as in the conventional technology.

In most cases, however, the reaction and separation are integrated in orderto benefit from the interaction effect between those two, for instance,

To improve yield/selectivity (e.g., via equilibrium shift)To facilitate separation (e.g., azeotrope problems)For other reasons, e.g., to extend the catalyst lifetime

One speaks in those cases about reactive separations or separative reactors.The industrially important reactive separations include:

Reactive distillationMembrane-based reactive separationsReactive adsorptionReactive absorptionReactive extractionReactive crystallization

2.1. Reactive Distillation

In most industrial applications the reactive distillation is used to improve theyield/selectivity of the required product. Figure 1 shows three examples of indus-trial processes, in which combination of reaction and distillation shifts the equil-ibrium of the reaction A � B ↔ C � D in the required direction (2). The lengthof the reacting, distillation, and stripping zones as well as the positioning of thereactant inlets vary in each particular case, depending on the process requirements.On the other hand, in selective hydrogenations of dienes and aromatics (3), reactivedistillation is used to remove the single-hydrogenated product from the reaction


zone, thus preventing its further hydrogenation and increasing its yield. In themethyl acetate technology of Eastman Chemical, integration of reactive distillationwith extractive distillation in a single unit totally solved the azeotrope problem (4).Reactive distillation can also be used as a powerful separation method in case of mixtures containing reactive and inert components with close boiling points.The method is schematically depicted in Figure 2. Here, a reactive entrainer isintroduced to the first reactive distillation column, to form an intermediate prod-uct having a boiling point much more distant from the boiling point of the inertcomponents. In the first column, inert components are therefore easily separated,while the intermediate product is fed to the second reactive distillation step,where the reversed reaction takes place and the original reactive component isrecovered and separated from the entrainer. Stein et al. (5) investigated the appli-cation of this principle to the separation of close-boiling i-butene and n-butene,using methanol as a reactive entrainer.

Obviously, reactive distillation may lead to significant savings on energy.Hydrolysis of methyl acetate presents an industrial example of such energy savings.

FIGURE 1 Examples of industrial processes employing reactive distillation:(a) methyl tert-butyl ether (MTBE) from isobutene and methanol; (b) cumenevia alkylation of benzene with propylene; (c) ethylene glycol via hydration ofethylene oxide.


The heat requirements of the reactive distillation-based process are ca. 50% lowerthan in the conventional technology. In the alkylation of benzene to cumene reac-tive distillation effectively eliminated the hot spots and reduced the oligomeriza-tion of propylene (6).

Table 1 gives an overview of the possible applications of reactive distillationreported in the literature. Very few of them have been realized so far on the com-mercial scale. One of the common factors that hinders a broader application ofreactive distillation is a small feasible operation window. The overlap region in thepressure–temperature domain, in which chemical reaction and separation and appa-ratus design are feasible, is usually quite narrow (see Figure 2 in Chapter 9). Apossible remedy for this limitation is sought in the development of new types ofcatalysts that would allow one to significantly broaden the feasible operation win-dow for chemical reaction.

2.2. Membrane-Based Reactive Separations

Sirkar et al. (64) give an interesting overview of various functions that a mem-brane may play in a chemical reactor. Those functions are schematically shownin Figure 3 and summarized in Table 2.

FIGURE 2 Separation of reactive and inert components with close boilingpoints, facilitated by reactive distillation. (From Ref. 5.)


TABLE 1 Reported Existing and Potential Applications of ReactiveDistillation

Product/process Selected refs.

Butyl acrylate from butanol and acrylic acid 7n-Butyl acetate from n-butanol and acetic acid 8Ethyl acetate from ethanol and acetic acid 9,10Methyl acetate from methanol and acetic acid 11,12Hydrolysis of methyl acetate 132-Methylpropylacetate from 2-methylpropanol 14

and acetic acidAmyl acetate from amyl alcohol and acetic acid 15Ethyl pentenoate from ethanol and pentenoic acid 16Esterification of fatty acids 17Methylal from formaldehyde and methanol 18,19TAME (tertiary amyl ether) 20–22MTBE (methyl tert-butyl ether) 23,24ETBE (ethyl tert-butyl ether) from ethanol and 25,26

isobuteneETBE (ethyl tert-butyl ether) from bioethanol and 27

tert-butylalcoholDiisopropyl ether from propene 28TAA (tert-amyl alcohol) via hydration of isoamylene 29Isopropanol via hydration of propene 30Cyclohexanol via hydration of cyclohexene 31,32Phenol from cumene 33Ethylene glycol via hydration of ethylene oxide 34,35Isobutene via dehydration of tert-butanol 36–38Isoamylenes via dehydration of 2-methyl-1-butanol 39Isophorone from acetone 40MIBK (Methyl iso-butyl ketone) from acetone 41,42Diacetone alcohol (DAA) and mesityl oxide (MO) via 43,44

aldol condensation of acetoneAcetone via dehydrogenation of propanol 45Tetrahydrofuran from butanediol 46Xylenes via toluene disproportionation 47Hydrogenation of unsaturated hydrocarbons 48,49Isomerization of C5-C6 paraffins 50Isobutene via hydroisomerization of C4 alkenes 51Cumene via alkylation of benzene with propylene 52Ethylbenzene via alkylation of benzene with ethylene 53,54Cyclopentane and/or cyclopentene from 55

dicyclopentadienePurification of hydrofluorocarbons 56


Among all the membrane functions listed in Table 2, catalytic membranesprobably attract the most attention. The scientific literature on catalytic membranereactors is exceptionally rich and includes many interesting ideas, such as heat- andmass-integrated combination of hydrogenation and dehydrogenation processes ina single membrane unit. Yet practically no large-scale industrial applications ofcatalytic membrane reactors have been reported so far, perhaps with the excep-tion of the Russian vitamin K technology (65). The primary reason for this is therelatively high price of membrane units, although other factors, such as low per-meability, sealing problems, as well as mechanical and thermal fragileness of themembranes, also play an important role. Further developments in the field ofmaterial engineering will surely change this picture.

Possible application areas of catalytic membrane reactors include:

Dehydrogenations, e.g., ethane to ethene, ethylbenzene to styrene, methanolto formaldehyde

Methane steam reformingWater–gas shift reactionSelective oxidations, e.g., propane to acroleine, butane to maleic anhydride,

ethylene to ethylene oxideOxidative dehydrogenations of hydrocarbonsOxidative coupling of methaneMethane oxidation to syngas

An excellent review of all potential applications of catalytic membrane reactorsstudied so far can be found in the 2002 book by Sanchez Marcano and Tsotsis (66).

On the other hand, membranes are frequently employed in combination witha bioreactor, for instance, in enzymatic pharmaceutical processes. An example

Naphtha desulfurization 57Dihydroxy polyether polyol via alkanolysis of 58

corresponding diestersGlycine from glycinonitryle 59DEC (diethylcarbonate) via carbonylation of ethanol 60

with dimethylcarbonatePolyamides (e.g., Nylon 6) via hydrolytic 61

polymerization of amino nitrilesNylon 66 via polycondensation 62Propylene oxide from propylene chlorohydrin 63

and calcium hydroxide

TABLE 1 (cont.)



FIGURE 3 Membrane functions in chemical reactor. (Reproduced with per-mission from Sirkar et al. (64), copyright (1999) American Chemical Society.)

of such an application of an ultrafiltration membrane-reactor system for the pro-duction of s-ibuprofen is discussed further in Chapter 12.

2.3. Reactive Adsorption

The vast majority of possible applications of reactive adsorption aim at the improve-ment of the product yield by shifting the equilibrium in the required direction. Incontrast to the nonreactive adsorption techniques, such as simulated moving bedsand pressure-swing adsorption, and despite its great potential [for example, a 12-fold higher conversion per pass in oxidative methane coupling (67)], the


TABLE 2 Membrane Functions in Chemical Reactor Systems

Function Examples

Separation of products from In situ product removal from enzymaticthe reaction mixture reactor via a nanofiltration or

ultrafiltration membraneRemoval of selected enantiomer via

a liquid membraneRemoval of water in esterification

reactions via a pervaporation membrane

Hydrogen removal in catalytic dehydrogenation reactions

Separation of a reactant from Separation of oxygen from air for a mixed stream for intro- oxidizing methane to syngasduction into the reactor Separation of hydrogen from

dehydrogenation reaction to oxidizeit with oxygen on permeate side

Separation of organic priority pollutants from wastewater for biological purification

Controlled addition of one Controlled oxygen addition in partial reactant or two reactants oxidation reactions (to increase

selectivity)Controlled air introduction in oxidative

dehydrogenationsNondispersive phase contacting, Emulsion-free enzymatic splitting of

with reaction at the phase fatsinterface or in the bulk phases Bubble-free oxygen/ozone supply in

wastewater treatment via hollow-fiber membranes

Segregation of a catalyst Segregation of enzymes with respect (and cofactor) in a reactor to molecular weight on ultrafiltration

membranesImmobilization of a catalyst Immobilization of enzymes or cells on

in (or on) a membrane polymeric membranesImmobilization of metals (Pd, Pt) on

ceramic membranesMembrane is the catalyst Cation exchange membranes for

esterification reactionsPalladium membranes for

hydrogenation/dehydrogenationreactions


Membrane is the reactor Reactions in flow-through membrane systems (“pore flow-through reactors”)

Solid electrolyte membrane Solid electrolyte membranes such as supports the electrodes, H� and O2� conductors in fuel cellsconducts ions, and achieves the reactions on its surface

Transfer of heat Membranes coupling endo- and exothermic reaction zones (e.g.,hydrogenation–dehydrogenation)

Immobilizing the liquid Supported liquid membranes (SLM) for reaction medium homogeneous catalytic processes

Source: Ref. 64.

TABLE 2 (cont.)

Function Examples

industrial-scale applications of adsorptive reactors remain to be seen. Challengesinvolve materials development of catalysts/adsorbents and matching of processconditions (same temperature) for both reaction and adsorption so that high yields/selectivities can be achieved. Reactive adsorption processes investigated in thebench or pilot scale are numerous, as shown in Table 3.

One of the more promising types of adsorptive reactors is the so-called gas–solid–solid trickle-flow reactor (GSSTFR), in which fine adsorbent trickles throughthe fixed bed of catalyst (Figure 4), removing selectively in situ one or more ofthe products from the reaction zone. In the case of methanol synthesis this led toconversions significantly exceeding the equilibrium conversions under the givenconditions (98). The economics of the methanol process based on the gas–solid–solid trickle-flow reactor was evaluated and compared with the conventional low-pressure Lurgi process (99). For the production scale of 1000 tons per day, the newtechnology offered considerable reductions in cooling water consumption (50%),recirculation energy (70%), raw materials (12%), and catalyst amount (70%).Further improvement of the GSSTFR concept could be seen in applying a mov-ing bed of adsorbent through straight, parallel channels of a monolithic catalyst,similar to the one shown in Figure 23 of Chapter 6.

2.4. Reactive Extraction

Similar to reactive adsorption, the reactive extraction can be applied primarily inmultireaction systems, for improvement in yields and selectivities to desired prod-ucts. The combination of reaction with liquid–liquid extraction can also be used


TABLE 3 Processes Investigated in Reactive Adsorption Systems

Process Selected refs. Reactor type

Esterification of glycerin with acetic acid 68 Simulated moving bed chromatographic reactor (SMBCR)

Methyl tert-butyl ether (MTBE) synthesis 69,70 SMBCRHydrolysis of methyl formate 71 Discontinuous chromatographic reactorOxidative coupling of methane 67,72 SMBCREnzymatic production of L-amino acids 73 Centrifugal partition chromatographic reactorOxidation of phenols 74 Chromatographic reactorEthyl acetate from ethanol and acetic acid 75 Chromatographic reactorEnzymatic inversion of sucrose 76 SMBCR

77 Rotating cylindrical annulus chromatographic reactor (RCACR)

Dehydroisomerization of n-butane 78 Chromatographic pulse reactorto isobutene

Mesitylene hydrogenation 72 SMBCRHydrogenation of 1,3,5-trimethylbenzene 79,80 SMBCRBiosynthesis of dextran polymer 81 Chromatographic pulse reactor

from sucroseDissociation of dicyclopentadiene 82,83 SMBCRDehydrogenation of cyclohexane 84 Chromatographic pulse reactor

85 RCACRAscorbic acid synthesis 86 SMBCRRegioselective enzymatic 87 SMBCR

diol esterification88 Batch and fixed-bed adsorptive

reactors


Oxidation of lean VOC mixtures 89 SMBCRHydrolysis of methyl formate 90 RCACREnzymatic production of lactosucrose 91 SMBCR

from sucrose and lactoseDiethylacetal from ethanol and 92 Fixed-bed adsorptive reactor

acetaldehydeSteam methane reforming 93 Pressure-swing adsorptive reactor (PSAR)Propene metathesis to ethene 94 PSAR

and 2-butene1-Butene dehydrogenation to 95 Rapid PSAR

1,3-butadieneSulfur from H2S (Claus process) 96 Reverse-flow adsorptive reactorHCN from carbon monoxide 96 Reverse-flow adsorptive reactor

and ammonia6-Aminopenicillanic acid from 97 Trickle-flow fluidized-bed reactor

penicillin GMethanol synthesis 98,99 Gas–solid–solid trickle-flow reactor (GSSTFR)

TABLE 3 (cont.)

Process Selected refs. Reactor type


FIGURE 4 Gas–solid–solid trickle-flow reactor. (From Ref. 98.)

for the separation of waste by-products that are hard to separate using conven-tional techniques (100,101). An overview of processes studied in reactive extrac-tion systems is shown in Table 4.

In 2002, an interesting concept was proposed for coupling a CO2-basedsupercritical extraction with air oxidation in order to remove and decompose pol-lutants from gases or liquids (134). An exemplary process scheme according tothis preliminary concept is shown in Figure 5. Possible (future) environmentalapplications of such an integrated supercritical extraction-reaction system includetreatment of liquid effluents, regeneration of catalysts and adsorption materials,and soil decontamination.


2.5. Reactive Crystallization/Precipitation

Reactive crystallization/precipitation plays a role in a number of industrially relevant processes, such as liquid-phase oxidation of para-xylene to producetechnical-grade terephthalic acid, the acidic hydrolysis of sodium salicylate tosalicylic acid, and the absorption of ammonia in aqueous sulfuric acid to formammonium sulfate (135). Reactive crystallization/precipitation is also widelyapplied in the pharmaceutical industry, to facilitate the resolution of the enan-tiomers (diastereomeric crystallization). Here, the racemate is reacted with a spe-cific optically active material (resolving agent) to produce two diastereomericderivatives (usually salts) that are easily separated by crystallization:

( ) ( ) ( ) ( ) ( ) ( )

DL L D L L L- - - - -

racemate resolving agent -salt -salt

-A A A� �� → ⋅ ⋅

n p

TABLE 4 Some Processes Studied in Reactive Extraction Systems


Penicillin G recovery 102–106Downstream separation of 1,3-propanediol 107Separation of lactic acid 108–111Separation of organic acids from the products of 112

partial oxidation of paraffinsSeparation of salicylic acid 113,114Separation of D,L-phenylalanine 113,115Separation of citric acid 116Separation of aspartic acid 117Cephalosporin C recovery 118Separation of metals (e.g., zinc) 119–121Phenolic wastewater treatment 122Separation of dicarboxylic acids (e.g., oxalic, malonic, 123

succinic, adipic acid)Recovery of gallium from coal fly ash 124Recovery of palladium, platinum, rodium from 125

leaching solutionsFractionation of amino acids 126Recovery of 7-ACA (7-aminocephalosporanic acid) 127Recovery of erythromycin 128Removal of toxic heavy metals from wastewater streams 129,130Production of dioxolane from aldehyde 131Recovery of aldehydes and ketones from hydrocarbon 132

mixturesProduction of cyclic ester oligomers from linear polyesters 133


Diastereomeric crystallization is commonly used in the production of a numberof pharmaceuticals, such as ampicillin, ethambutol, chloramphenicol, diltiazem,fosfomycin, and naproxen (136).

Somewhat similar are the so-called adductive crystallization processes,often (wrongly) called extractive crystallization, where reactions of complex/adduct formation are used to separate compounds that are otherwise difficult toseparate. Examples of adductive crystallization include separation of p- and m-cresols (137), separation of o- and p-nitrochlorobenzenes (138), separation ofquinaldine and isoquinoline (139), separation of nonaromatic compounds fromnaphtha-cracking raffinate (140), and separation of p-cresol from 2,6-xylenol(141). Other examples of reactive crystallization/precipitation reported in the lit-erature are listed in Table 5.

Reactive crystallization/precipitation can also be conducted in high-gravity(Higee) fields using rotating equipment. In China this technique has been usedsuccessfully for the production of nano-size particles of CaCO3. Ultrafine particles

FIGURE 5 Countercurrent supercritical extraction coupled with air oxidationto remove and decompose pollutants from gases and liquids. (From Ref.134.)


with the mean size of 15–40 nm and a very narrow size distribution were pro-duced by carbonation of a lime suspension in a rotating packed-bed reactor (RPBR)(158). The reaction times in RPBR were 4- to 10-fold shorter than the correspon-ding reaction times in a conventional stirred-tank unit. A similar technique wasused for the production of nanofibrils of aluminum hydroxide with a diameter of1–10 nm and 50–300 nm long as well as nanoparticles of SrCO3 with a mean sizeof 40 nm (159).

2.6. Reactive Absorption/Stripping

Reactive absorption is very old as a processing technique and has been used forproduction purposes in a number of classical bulk-chemical technologies, such asnitric or sulfuric acid. The Raschig process for the production of hydroxylamine,an important intermediate in classical caprolactam technologies (Stamicarbon,Inventa), is also an example of a multistep reactive absorption process. Here,water, ammonia, and carbon dioxide react together in an absorption column togive a solution of ammonium carbonate, which subsequently forms an alkaline

TABLE 5 Examples of Reactive Crystallization/Precipitation

SelectedProduct refs. Remarks

Calcium carbonate 142,143 Liquid–liquid and gas–liquid reaction systems

Methyl �-methoximino 144acetoacetate

Magnesium hydroxide 145,146Calcium phosphate 147,148Magnesium ammonium 149 Removal of ammonium and

phosphate phosphate ions from wastewater

Lead sulfate 150Magnesium carbonate 151Nickel hydroxide 152Ziprasidone–HCl�H2O 153 Conducted in impinging

fluid jet stream systemBarium carbonate 154 To remove CO2 from waste

gasBoric acid 155 Reaction of borax solution

with solid oxalic acidProcaine benzylpenicillin 156Sulfamic acid 157 From urea and fuming H2SO4


solution of ammonium nitrite by reactive absorption of nitrous oxide at low tem-perature. In a further step, the ammonium nitrite is converted to hydroxylaminedisulfonate with sulfur dioxide. The hydroxylamine disulfonate solution is drawnoff and the salt is hydrolyzed and neutralized to give hydroxylamine sulfate andammonium sulfate as coproduct.

Carbon dioxide removal by reactive absorption in amine solutions is alsoapplied on the commercial scale, for instance, in the treatment of flue gas (seelater in this chapter). Another possible application field of the technique is gasdesulfurization, in which H2S is removed and converted to sulfur by means ofreactive absorption. Aqueous solutions of ferric chelates (160–162) as well astetramethylene sulfone, pyridine, quinoline, and polyglycol ether solutions ofSO2 (163,164) have been proposed as solvents. Reactive absorption can also beused for NOx reduction and removal from flue or exhaust gases (165,166). Theseparation of light olefins and paraffins by means of a reversible chemical com-plexation of olefins with Ag(I) or Cu(I) compounds in aqueous and nonaqueoussolutions is another very interesting example of reactive absorption, one thatcould possibly replace the conventional cryogenic distillation technology (167).

3. HYBRID SEPARATIONS

Generally speaking, hybrid separations can be described as processing methods thatintegrate two or more different separation techniques in a single operation, makinguse of the synergy between them. The industrially most important (or promising)hybrid separations include:

Extractive distillationAdsorptive distillationMembrane distillationMembrane absorption/strippingAdsorptive membranes (membrane chromatography)Membrane extraction

3.1. Extractive Distillation

Extractive distillation is probably the oldest and most widely applied type ofhybrid separation, particularly useful in close-boiling-point problems or in sys-tems in which components form azeotropes. In the method, an extra component(solvent) is added to the system, which does not form azeotropes with feed com-ponents. The solvent alters the relative volatility of original feed components,allowing one to distill overhead. The solvent leaves the column with the bottomproducts and is separated in a binary column. Energy savings represent the mostimportant advantage of extractive distillation over the conventional (nonhybrid)separation methods (168,169).


Originally, extractive distillation was limited to two-component problems.However, recent developments in solvent technology enabled applications of this hybrid separation in multicomponent systems as well. An example of suchapplication is the BTX process of the GTC Technology Corp., shown in Figure 6,in which extractive distillation replaced the conventional liquid–liquid extractionto separate aromatics from catalytic reformate or pyrolysis gasoline. This led to a ca. 25% lower capital cost and a ca. 15% decrease in energy consumption (170).Some other examples of existing and potential applications of the extractive dis-tillations are listed in Table 6.

Solvents used for extractive distillation vary considerably, but in almost allcases solvent selection presents a trade-off between its selectivity and solvency(194). The effectivity of the solvent can sometimes be improved by the additionof a salt (195).

3.2. Adsorptive Distillation

Although considered by some authors a “novel process,” adsorptive distillation is a relatively old hybrid separation, originating in the early 1950s (196). It is a

FIGURE 6 Scheme of aromatics separation via extractive distillation in theBTX process of the GTC Technology Corp.


three-phase mass transfer operation in which distillation is carried in presence ofa solid selective adsorbent. The adsorbent usually consists of a fine powder (parti-cle size in the 10-�m range), fluidized and circulated by an inert carrier. Theprocess is typically carried in two columns: an adsorptive distillation column forincreasing separation ability, and a distillative desorption column for enhancing

TABLE 6 Reported (Potential) Applications of Extractive Distillation

SelectedSeparation process refs. Remarks

High-purity cyclohexane 171 Close-boiling-point problemfrom petroleum

Benzene and toluene 172,173 Close-boiling-point problemfrom nonaromatics

Isopropyl ether from acetone 174 Azeotrope problemMethyl acetate from 175 Azeotrope problem

methanolAnhydrous ethanol from 176 Azeotrope problem

fermentation brothEthyl acetate from 177 Azeotrope problem

ethanol/waterTernary acetate–alcohol–water 178 Azeotrope problem

systems (propyl, butyl, amyl, hexyl)

m-Xylene from o-xylene 179 Close-boiling-point problemMTBE from impurities 180 Close-boiling-point problemBinary mixtures of 181 Close-boiling-point problem

lower-boiling alcoholsBinary mixtures of 182,183 Close-boiling-point problem

phenolic compounds (chlorophenol, phenol, cresol, xylenol)

Acetone from water 184 Azeotrope problemEthanol dehydration 185 Azeotrope problemC2� alcohols from water 186 Azeotrope problemCyclohexane-cyclohexene- 187 Close-boiling-point problem

benzeneMTBE from ethanol 188 Azeotrope problemMethylcyclohexane from 189 Close-boiling-point problem

tolueneAnhydrous ethanol recovery 190 Azeotrope problem

from wastewater streamsPropylene from propane 191 Close-boiling-point problem1-Butene from 1,3-butadiene 192,193 Close-boiling-point problem


the regeneration of the adsorbent. As in the case of extractive distillation, adsorp-tive distillation also can be used for the separation of mixtures containing close-boiling components or to bypass the azeotrope. Another interesting potentialapplication field is the removal of trace impurities in the production of fine chem-icals. The simplest scheme of an adsorptive distillation system for separating abinary mixture of azeotrope-forming components is shown in Figure 7. Here,adsorbent S carried by an inert fluid carrier enters the adsorptive distillation col-umn, selectively adsorbs component B from the feed, and flows to the desorption(stripping) column, in which separation and enrichment of B takes place. Figure 8shows another variant of adsorptive distillation, as proposed in a patent by Shell(197), for improved separation of closely-boiling hydrocarbon mixtures. Here anextra stripping medium (e.g., pentane) is used to remove the adsorbate in thestripping column.

Despite an almost 50-year history, no large-scale commercial processesusing adsorptive distillation have been reported so far. Some potential applicationfields for this hybrid separation are listed in Table 7.

FIGURE 7 Scheme of an adsorptive distillation system for the separation ofazeotrope-forming components A and B (S–adsorbent).


3.3. Membrane Distillation

Membrane distillation is considered a promising separation method applicable pri-marily in environmental technologies. In membrane distillation a microporous andhydrophobic membrane separates aqueous solutions at different temperatures andcompositions, as shown in Figure 9. The temperature difference existing across themembrane results in a vapor pressure difference. The molecules are transportedthrough the pores of the membrane from the high-vapor-pressure side to the low-vapor-pressure side. At least one side of the membrane remains in contact with theliquid phase. Benefits offered by membrane distillation include (202):

100% (theoretical) rejection of ions, macromolecules, colloids, cells, andother nonvolatiles

Lower operating temperatures than conventional distillationLower operating pressures than conventional pressure-driven membrane

separationReduced chemical interaction between membrane and process solutionsLess demanding membrane mechanical property requirementsReduced vapor spaces compared to conventional distillation

FIGURE 8 Simplified scheme of an adsorptive distillation–based separationof closely boiling hydrocarbon mixtures. (From Ref. 197.)


TABLE 7 Potential Application Fields of Adsorptive Distillation Reported in the Literature

SelectedSystem refs. Remarks

Toluene–methylcyclohexane 196 Suitable adsorbents: silica gel,mixtures and other activated carbon, coconut closely boiling charcoal, bauxite, hydrocarbons activated alumina

Naphtha reformate and 197 Silica gel as adsorbentother close-boiling hydrocarbons

o-Xylene–m-xylene mixture 198 Modeling studyEthanol–water and ethyl 199,200 Zeolite (4A molecular sieve)

acetate–water–n-butanol as adsorbent, glycol as mixtures carrier

p-Xylene–m-xylene mixture 201 Zeolite (NaY molecular sieve)as adsorbent, n-decane as carrier

Membrane distillation systems may be classified into four different cat-egories (203):

Direct contact membrane distillation (DCMD), in which the membrane isin direct contact with the liquid phase on both sides

Air-gap membrane distillation (AGMD), in which an air gap is interposedbetween the membrane and the condensation surface

Vacuum membrane distillation (VMD), in which the vapor phase is evacu-ated from the liquid through the membrane and the condensation takesplace in a separate apparatus

Sweeping-gas membrane distillation (SGMD), in which a stripping gas,instead of vacuum, is used as a carrier

Currently, the most important application area for membrane distillation iswater desalination technology. Figure 10 shows one of the water desalinationprocesses developed by a Japanese organization, the Water Re-Use PromotionCenter, in cooperation with Takenaka Corporation and Organo Corporation (204).The process uses solar energy and can therefore be installed at locations withoutan electricity supply. Other application areas for membrane distillation reportedin the literature are summarized in Table 8.

In 2002, the TNO Environment, Energy and Process Innovation institute in the Netherlands developed a membrane-based distillation concept that radicallyimproves the economy and ecology of existing desalination technology for


FIGURE 9 Scheme of a membrane distillation process.

seawater and brackish water. This so-called Memstill® technology (Figure 11a)combines multistage flash and multieffect distillation modes into one membranemodule. Since the Memstill® module houses a continuum of evaporation stagesin an almost ideal countercurrent flow configuration, a very high recovery of theevaporation heat is possible. The economic advantage of the Memstill® technol-ogy, compared to the “classical” desalination techniques, is shown in Figure 11b.An academic-industrial consortium is currently developing and improving theMemstill® process concept and module design (226). The same TNO Institute has also developed a concept of another membrane-based distillation technologyfor fractionation of non-water-based systems (227). The technology, calledMEMFRAC, offers high energy efficiency in compact equipment. The study car-ried out for fractionation of benzene from toluene showed that with MEMFRACtechnology a HETP between 5 and 10 cm could be obtained. Additional advan-tages of the MEMFRAC technology include: lack of entrainment, flooding, foam-ing, or channeling (due to indirect gas–liquid contact), independent gas/liquidcontrol, and the possibility for modular plant design. Such a modular MEMFRACdistillation unit is schematically presented in Figure 12.

On the other hand, a pervaporation membrane can be coupled with a con-ventional distillation column, resulting in a hybrid membrane/distillation process(228,229). Some of the investigated applications of such hybrid pervaporationmembrane/distillation systems are shown in Table 9. In hybrid pervaporation/distillation systems, the membrane units can be installed on the overhead vaporof the distillation column, as shown in Figure 13a for the case of propylene/propane splitting (234), or they can be installed on the feed to the distillation column,


FIGURE 10 Scheme of the demonstration test plant for water desalination using solar energy and membrane distillation.(Courtesy: CADDET, Center for Renewable Energy, Harwell, UK).


TABLE 8 Application Areas of Membrane Distillation

SelectedSystem/process refs. Remarks

Concentration of H2SO4, 205–207H3PO4, NaOH, HNO3, and HCl solutions

Concentration of 2,3-butanediol 208 VMD processfrom fermentation broths

Wastewater treatment in 209 Integrated reverse the textile industry osmosis/membrane

distillation processRadioactive wastewater 210 Integrated reverse

treatment osmosis/membrane distillation process

Removal of benzene traces 211 VMD processfrom water

Concentration of protein solutions 212Removal of halogenated VOCs 213 VMD processConcentration of oil–water 214

emulsionsConcentration of sugar/sucrose 215,216 DCMD and AGMD

solutions processesSeparation of water and glycols 217 DCMD processEthanol–water separation 218 AGMD processAcetone and ethanol removal 219 AGMD process

from aqueous solutionsPropanone removal from 220 AGMD process

aqueous streamsAcetone–butanol–ethanol 221 AGMD process

(ABE) solvent recoveryFermentative ethanol 222,223 Integration of MD in

production fermentation resultedin ca. 2� increase of production rate

Concentration and purification 224of fluosilicic acid

Removal of trichloroethylene 225 VMD process


FIGURE 11 Memstill® technology of seawater desalination developed at the TNO institute: (a) principle of theprocess; (b) cost comparison with other desalination techniques. (From Ref. 226.)


as shown in Figure 13b for the case of aromatic/aliphatic hydrocarbon separation(235). Shortcut design methods for hybrid pervaporation/distillation processescan be found in Ref. 236.

3.4. Membrane Absorption/Stripping

Membrane absorption is one of the processes that Mother Nature had invented longbefore engineers did. Human lungs and intestines present perfect examples ofmembrane absorption systems. In the simplest case a gaseous component is selec-tively transported via a membrane and dissolved in the absorbing liquid, as shownin Figure 14. It is also possible to carry a membrane-based absorption-desorptionprocess, with two liquids on both sides of the membrane (237), or a membranestripping process, in which selected components are removed from the liquid phasethrough the membrane by a stripping gas (238). An important characteristic feature

FIGURE 12 Scheme of a modular MEMFRAC distillation unit for fractionationof non-water-based systems, developed at the TNO Institute. (From Ref. 227.)


of membrane absorption is that it proceeds without creating a real gas–liquid inter-face in the form of bubbles. Such a bubbleless gas–liquid mass transfer can be ofadvantage in certain processes, for instance, in shear-sensitive biological systems.

One of the most important application areas of the membrane absorption isthe capture of CO2 from flue gas. Kværner has recently developed a membraneabsorption–based technology for the removal of CO2 from turbine exhaust gases inoffshore applications (239). The process, based on membrane-facilitated CO2absorption in amine, followed by membrane-facilitated stripping with steam, isschematically shown in Figure 15. The expected cost reduction, in comparison witha conventional amine separation process, ranges between 30 and 40%, for bothinvestment and operating costs. The new membrane-based process also offers a verysignificant reduction in the weight and size of equipment (70–75% and 65%, respec-tively; see Figure 16), a great advantage in the case of offshore technology. Someother possible applications of membrane absorption/stripping are shown in Table 10.

3.5. Membrane Chromatography (Adsorptive Membranes)

Membrane chromatography is a separation technique used almost exclusively inthe downstream processing of proteins. Traditionally, most chromatographicpurification steps in the downstream processing of proteins take place in columnspacked with bead-shaped particles. Membrane chromatography presents a hybridcombination of liquid chromatography and membrane filtration based on micro-porous or macroporous membranes that contain functional ligands attached totheir inner pore structure, which act as adsorbents. The main feature and advan-tage of this technique, compared to the conventional ones, is the absence of pore

TABLE 9 Possible Applications of Hybrid Pervaporation/DistillationSystems

System/process Selected investigated refs. Remarks/effects

Benzene–cyclohexane 230 Combination of extractive distillation separation and one-stage pervaporation;

high-purity (99.2–99.5%) products;estimated cost savings of 20%

Ethanol dehydration 231 Simulation study; 50% cost reduction in comparison with conventional azeotropic distillation

Propylene/propane 232 Pilot-plant studies; 20–50% savings splitting on operating costs

Propylene/propane 233 26–30% savings on capital splitting investment


FIGURE 13 Combined distillation/pervaporation systems for (a) propylene/ propane splitting and (b) aromatic/aliphatichydrocarbon separation. (Part a from Ref. 234; part b from Ref. 235.)


diffusion, which is the main transport resistance in traditional chromatography.Dissolved molecules are carried directly to the adsorptive sites in the membranesby convective flow (Figure 17), which increases the throughput of the process.Membrane chromatography presents a process-intensive option for the protein A,G, or L affinity chromatography (247–253), as well as for metal affinity, ion-exchange, hydrophobic interaction or reversed-phase chromatography (254–258).In recent years some new potential application fields for membrane chromato-graphy have been demonstrated. Those are listed in Table 11.

FIGURE 14 Membrane absorption.

FIGURE 15 Membrane absorption–based technology for removal of CO2from turbine exhaust gases, developed by Kværner Process Systems.(Courtesy: Kværner.)


FIGURE 16 Plant-size reduction in Kværner’s technology for CO2 removalfrom exhaust gases: (a) conventional process; (b) membrane absorptionprocess. (Courtesy: Kværner.)

3.6. Membrane Extraction

In membrane extraction, the treated solution and the extractant/solvent are sep-arated from each other by means of a solid or liquid membrane. The technique isapplied primarily in three areas: wastewater treatment (e.g., removal of pollutantsor recovery of trace components), biotechnology (e.g., removal of products fromfermentation broths or separation of enantiomers), and analytical chemistry (e.g.,online monitoring of pollutant concentrations in wastewater). Figure 18a showsschematically an industrial hollow fiber–based pertraction unit for water treat-ment, according to the TNO technology (263). The unit can be integrated with afilm evaporator to enable the release of pollutants in pure form (Figure 18b).


Other more promising possible applications of membrane extraction, reported inthe literature, are listed in Table 12.

3.7. Other Hybrid Separations

In extractive crystallization (nonadductive), the driving force for the separationprocess is created by altering the solid–liquid phase relationships via the additionof a third component (usually liquid solvent) to the system. The solvent is chosenin such a way that it binds strongly at the crystallization temperature but separateseasily at another temperature, where it is usually regenerated via distillation.Examples of such defined extractive distillations include separation of m- and p-cresols using acetic acid as the solvent (297), separation of o- and p-nitrochloro-benzenes using p-dichlorobenzene (298), separation of lithium sulfate and lithiumformate using n-butanol or 2-propanol (299), and separation of p-xylene from m-xylene using pentene (300).

TABLE 10 Reported Possible Applications of Membrane Absorption/Stripping

System/process Selected investigated refs. Remarks/effects

Ammonia absorption/ 240 Pilot-plant study in a polypropylene desorption from hollow-fiber column; ammonia isammonia water absorbed in diluted sulfuric acid

CO2 and/or SO2 removal 241 Absorption in NaOH, K2CO3,alkanolamines, and Na2SO3 usinghydrophobic microporous hollow-fiber modules

Cyanide recovery from 242 Recovery via a gas-filled membranewastewater (GFM) placed between the waste-

water and a chemical strippingsolution

H2S removal from gas 243 Asymmetric hollow-fiber modules streams coupled with concentrated

alkaline solutionCO2 production for the 244 Possible energy saving of more

horticultural industry than 30% reported(greenhouses)

H2S and SO2 removal 245 Polyvinylidene fluoride (PVDF) hollow fibers and concentratedNaOH solution used

VOC removal from 246 Air stripping process via a wastewater polypropylene hollow-fiber module


Smith, Bryson, and Valsaray (301,302) investigated solvent sublation, anadsorptive bubble process, combining the transport mechanisms of liquid/liquidextraction and gas stripping. The technique exploits the surface-active nature oforganic compounds in their removal from water systems. The mechanism of solventsublation is shown in Figure 19a. Gas bubbles are used to transport adsorbed solutefrom the bulk solution to the solvent layer. Solvent sublation is particularly prom-ising in the removal of (hydrophobic) organic compounds from wastewater streams.An exemplary process scheme is shown in Figure 19b. An important advantage ofthe method is that the intimate contact between the extracting solvent and the waste-water is prevented (no problem of residual solvent in the treated water).

Zeitsch (303) conducted a preliminary research on the removal of aceticacid from the vapor stream of furfural reactors by means of extractive condensa-tion. It is a hybrid vapor-phase extraction process, in which solvent (triethyl-amine, TEA) forms a high-boiling complex with acetic acid. As a result, a “fog”

FIGURE 17 Comparison of transport mechanisms in (a) conventional chro-matography and in (b) membrane chromatography.


fraction is formed that can be separated via a coalescence filter or an electrostaticseparator. The technique is reported to be highly selective.

In 2002, Drioli and coworkers (304) investigated a process for obtainingprotein crystals by means of membrane crystallization, which actually combinesmembrane distillation and crystallization techniques. The solvent evaporates atthe membrane interface, migrates through the pores of the membrane, and con-denses on the opposite side of the membrane. The reported preliminary resultsindicate interesting potentialities of this new method with respect to macromole-cular crystallization.

4. BARRIERS AND FUTURE PROSPECTS

Despite many ongoing research activities in the field and a number of successfulcommercializations, there still exist numerous technical and nontechnical barriersthat hinder a wider introduction of reactive and hybrid separations into industrialpractice. Two workshops held in 1998 by the Center for Waste ReductionTechnologies of AIChE (305) identified some of the barriers for reactive separ-ations and divided them into three categories:

a. Technical gaps, such as lack of simulation and scale-up capability, lackof validated thermodynamic and kinetic data, lack of materials (e.g.,

TABLE 11 Reported New Potential Applications of MembraneChromatography

SelectedSystem/process investigated refs. Remarks/effects

Separation of polynucleotides 259 Supercoiled plasmid DNA investigated as model

Separation of oligonucleotides 260and peptides

Separation of small 260 Benzene, toluene, homologues hydrophobic molecules of 4-hydroxybenzoate

investigatedEnantiomeric separation 261 Racemic mixtures of

tryptophan and thiophenalinvestigated in microfluidic-based membrane chromatography

Separation of trace metals 262 La-Ce-Pr-Nd-Sm separation and Zr-Hf separation investigated


FIGURE 18 Pertraction technology for wastewater treatment from the TNOInstitute: (a) scheme of the hollow-fiber pertraction unit; (b) integration ofpertraction with film evaporation. (Courtesy: TNO.)


R

TABLE 12 Reported Potential Applications of Membrane Extraction

SelectedSystem/process investigated refs. Remarks

Separation of acetic acid from aqueous solution 264 Microporous polypropylene membrane, MIBK as solvent

Separation of (S)-naproxen from racemic 265 Reactive extraction via hollow-fibernaproxen thioesters membrane

Separation of D,L-alanine and D,L-phenylalanine 266,267 Hollow-fiber zeolite membraneracemic mixturesRemoval of sulfanilic acid from wastewater 268 Hollow-fiber modules

Lactic acid purification and concentration 269 Emulsion liquid membranes (ELMs)Enrichment of bisphenol A 270 Liquid membranePhenol recovery from aqueous solutions 271,272 Various liquid and solid membranesZinc(II) recovery from HCl solution 273 Bulk liquid and hollow-fiber membranesHydrogen separation from methane steam 274 Palladium alloy membranes

conversion productsSeparation of liquid olefin/paraffin mixtures 275 Nonporous polymeric membranesRemoval of 2-chlorophenol 276 Liquid membrane from aqueous solutionsEthanol removal from aqueous solutions 277 Microporous polypropylene membraneSeparation of cephalosporin C from 278 Bulk and emulsion liquid membranes

fermentation brothSeparation of penicillin G 279 Supported liquid membrane

from aqueous streams (Amberlite LA-2)Enrichment of amino acids 280 Supported liquid membrane (Aliquat 336)Separation of cephalexin from a mixture of 7-ADCA 281 Supported liquid membrane (Aliquat 336)Separation of butyric acid from fermentation broth 282 Liquid membrane

(continued)


Separation of propionic and acetic acid from 283 Polymeric membranesfermentation broth

Separation of citric acid from fermentation broth 284 Liquid membraneSeparation of lactic acid from fermentation broth 285 Emulsion liquid membraneProduction of acetone, butanol, and ethanol 286 Polypropylene membrane

(ABE) from potato wastesSeparation of long-chain unsaturated fatty acids 287 Microporous membrane, MeCN

and n-heptane as solvents(Heavy) metals recovery from wastewater 288–292 Various membranesRemoval of organic contaminants from wastewater 293–296 Various membranes

TABLE 12 (cont.)

SelectedSystem/process investigated refs. Remarks


FIGURE 19 Solvent sublation: (a) process mechanism and (b) an exemplaryprocess scheme. (From Refs. 301 and 302.)


integrated catalysts/sorbents, membrane materials), and lack of high-level process synthesis methodology

b. Technology transfer barriers, such as lack of multidisciplinary teamapproaches to process integration, lack of commonality of prob-lems (technology is application-specific) and lack of demonstrations/prototypes on a reasonable scale (reactive and hybrid separations arestill regarded more as a science than a technology)

c. General barriers, such as higher standards, to which new technologiesmust be held, compared to conventional technologies, lack of informa-tion on process economics (early economic and process evaluation),and fear of risk in using new technologies.

Most of these barriers also hold for hybrid separations. Two more factorsthat clearly play a hindering role in the commercial application of many reactiveand hybrid separations are: the already-mentioned small feasible operation win-dows and the reduction of the degree of freedom caused by the integration of reac-tion and separation or by the coupling of two separations in one processing unit.Figure 20 shows an example of how the integration of reaction and membraneseparation reduces the degree of freedom in a membrane reactor, resulting indecreased operational flexibility (306).

FIGURE 20 Degree of integration versus degree of freedom, in an example ofa membrane reactor. (From Ref. 306.)


Despite the existing barriers, the coming years are expected to bring a sig-nificant increase in the number of industrially applied reactive and hybrid sep-aration technologies. In particular, progress can be expected in the application ofreactive distillation, reactive adsorption, and membrane-based operations. Inhybrid separations, expansion of research activities on new product/process areashas already been seen. Reactive and hybrid separations have enormous potentialfor process intensification. Making full use of that potential will lead to substan-tially smaller, cleaner, and more energy-efficient chemical and biochemicalplants.

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M, Hidalgo M, Valiente M, eds. Solvent Extraction for the 21st Century. Proceedingsof ISEC ’99, Barcelona. London: Society of Chemical Industry, 2001:177–182.

294. Klaassen R, Janssen AE, Akkerhuis JJ, Bult BA, Oesterhold FIHM, Schenider J. RecProg Genie Procedes 1992; 6:183–188.

295. Klaassen R. Chemie Technik (Sonderaus., Chemie Umwelt Technik) 1999; 27:24–28.296. Livingston A, Ferreira F, Han S, Boam A, Zhang S. In: Preprints 8. Aachener Membran

Kolloquium, 27–29 März 2001, Aachen. Mainz: VDI, 2001:1205–1214.297. Chivate MR, Shah SM. Chem Eng Sci 1956; 5:232–241.298. Dikshit RC, Chivate MR. Chem Eng Sci 1970; 25:311–317.299. Carton A, Bolado S, Marcos MM. Informacion Tecnologica 2000; 11:73–82.300. Rajagopal S, Ng KM, Douglas JM. AIChE J 1991; 37:437–447.301. Smith JS, Valsaraj KT. Chem Eng Prog 1998; 94(5):69–77.302. Bryson BG, Valsaraj KT. J Hazard Mater 2000; 2601:1–11.303. Zeitsch KJ. Ind Eng Chem Res 1999; 38:4123–4124.304. Curcio E, Di Profio G, Drioli E. Desalination 2002; 145:173–177.305. Adler S, Beaver E, Bryan P, Rogers JEL, Robinson S, Russomanno C. Vision 2020:

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306. Tlatlik S, Schembecker G. Process synthesis for reactive separations. In: Proceed-ings of ARS-1, Advances in Reactive Separations 1, University of Dortmund, Germany,October 12, 2000:1–10.


9

Reactive Separations in Fluid Systems

E.Y. Kenig and A. Górak

University of Dortmund, Dortmund, Germany

H.-J. Bart

University of Kaiserslautern, Kaiserslautern, Germany

1. INTRODUCTION: AN OVERVIEW OF

REACTIVE SEPARATIONS

Chemical manufacturing companies produce materials based on chemical reactionsbetween selected feed stocks. In many cases the completion of the chemical reac-tions is limited by the equilibrium between feed and product. The process must theninclude the separation of this equilibrium mixture and recycling of the reactants.The fundamental process steps of bringing material together, causing them to react,and then separating products from reactants are common to many processes.

Conventionally, each unit operation—whether mixing or absorption, distil-lation, evaporation, crystallization, in fact, any of the heat-, mass-, and momentum-transfer operations so familiar to chemical engineers—is typically performed inindividual items of equipment, which, when arranged together in sequence, makeup the complete process plant. As reaction and separation stages are carried outin discrete equipment units, equipment and energy costs are added up from thesemajor steps. However, this historical view of plant design is now being challengedby the combination of two or more unit operations into one plant unit. The poten-tial for capital cost savings is obvious, but there are often many other processadvantages that accrue from such combinations (1).


In recent decades, a combination of separation and reaction inside a singleunit has become more and more popular. This combination has been recognizedby chemical process industries as having favorable economics for carrying outreaction simultaneously with separation for certain classes of reacting systems,and many new processes (called reactive separations) have been invented basedon this technology (2–9). Reactive separation units may also be treated as a kindof multifunctional reactor in which the functionalities of several processes arecombined to generate the new reactor concept (Figure 1).

The most important examples of reactive separation processes (RSPs) arereactive distillation (RD), reactive absorption (RA), and reactive extraction (RE). InRD, reaction and distillation take place within the same zone of a distillation column.Reactants are converted to products, with simultaneous separation of the productsand recycling of unused reactants. The RD process can be efficient in both size andcost of capital equipment and in energy used to achieve a complete conversion ofreactants. Since reactor costs are often less than 10% of the capital investment, thecombination of a relatively cheap reactor with a distillation column offers greatpotential for overall savings. Among suitable RD processes are etherifications, nitra-tions, esterifications, transesterifications, condensations, and alcylations (2).

FIGURE 1 Reactive separation units as multifunctional reactors. (Inspired byRef. 4.)


Similarly, in RA, reactions occur simultaneously with the component trans-port and absorptive separation, in the same column zone. These processes areused predominantly for the production of basic chemicals, e.g., sulphuric or nitricacids, and for the removal of components from gas and liquid streams. This can beeither the cleanup of process gas streams or the removal of toxic or harmful sub-stances in flue gases. Absorbers or scrubbers where RA is performed are oftenconsidered gas–liquid reactors (10). If more attention is paid to the mass trans-port, these apparatuses are instead treated as absorption units.

Reactive extraction uses liquid ion exchangers that promote a selectivereaction or separation. The solutes are very often ionic species (metal ions ororganic/inorganic acids) or intermediates (furfural phenols, etc.), and the extrac-tion chemistry is discussed elsewhere (11–13). Reactive extraction can be used forseparation/ purification or enrichment or conversion of salts (14). A 2001 reviewon reactive phase equilibria, kinetics, and mass transfer and apparative techniquesis given in Ref. 8. Reactive extraction equipment is discussed in detail in Ref. 15,and recent advances are given in Ref. 16.

Reactive absorption, distillation, and extraction have much in common.First of all, they involve at least one liquid phase, and therefore the properties ofthe liquid state become significant. Second, they occur in moving systems; thusthe process hydrodynamics plays an important part. Third, these processes arebased on the contact of at least two phases, and therefore the interfacial transportphenomena have to be considered. Further common features are multicomponentinteractions of mixture components, a tricky interplay of mass transport andchemical reactions, and complex process chemistry and thermodynamics.

On the other hand, RD, RA, and RE have a number of specific features thatshould be considered with care and described by different approaches. Beforegoing into detail, it is worthy to note that the operating window of reactive separ-ations may be somewhat limited, since these operations are feasible only if theyallow for both separation and reaction within the same range of temperature andpressure and, on the other hand, for the safe operation from the constructionalpoint of view (Figure 2).

1.1. Reactive Absorption

The main purposes of absorption processes are the removal of one or more com-ponents from the gas phase, production of particular substances in the liquidphase, and gas mixture separation (3). Industrial absorption operations are usuallyrealized by combining absorption and desorption units.

The example given in Figure 3 illustrates this combination of two processes.In an absorber, one or several gas components are absorbed by a lean solvent,either physically or chemically. A rich solvent, after preheating in heat exchan-gers H1 and H3, is transported to the top of a desorption unit, which usually operates


FIGURE 2 Feasibility of reactive separation, depending on mechanicaldesign, chemical reaction, and separation performance.

under a pressure lower than that in the absorber. Part of the gas absorbed by therich solvent is desorbed due to flashing and heating. The other part has to be de-sorbed in the stripper via countercurrent contact of liquid with the inert gas orsteam. The lean solvent then flows through heat exchanger H1 to recover heat nec-essary for heating the reach solvent, passes through heat exchanger H2 to cooldown to a desired temperature, and finally enters the absorber (3).

Usually a small amount of fresh solvent should be added to the column inorder to equalize the solvent loss due to evaporation in the desorber or to irre-versible chemical reactions occurring in the whole system (3).

Reactive absorption represents a process in which a selective solution ofgaseous species by a liquid solvent phase is combined with chemical reactions.As compared to purely physical absorption, RA does not necessarily require ele-vated pressure and high solubility of absorbed components; because of the chem-ical reaction, the equilibrium state can be shifted favorably, resulting in enhancedsolution capacity (17). Most RA processes involve reactions in the liquid phaseonly; in some of them, both liquid and gas reactions occur (18,19).

Usually the effect of chemical reactions in RA processes is advantageousonly in the region of low gas-phase concentrations, due to limitations stemmingfrom the reaction stoichiometry or equilibrium (20). Further difficulties of RAapplications may be caused by the reaction heat through exothermic reactions andby relatively difficult solvent regeneration (21,22). Most RA processes are


steady-state operations, either homogeneously catalyzed or auto-catalyzed. Someimportant industrial applications of RA are given in Table 1.

Reactive absorption can be realized in a variety of equipment types, e.g., infilm absorbers, plate columns, packed units, or bubble columns. This process is char-acterized by independent flow of both phases, which is different from distillation andpermits both cocurrent (downflow and upflow) and countercurrent regimes.

Reactive absorption is essentially an old process, known since the founda-tion of modern industry. This is a very important process, too, being the basicoperation in many technological chains. More recently, the role of RA as a keyenvironmental protection process has grown up significantly.

Despite the clear importance of RA, its behavior is still not properly under-stood. This can be attributed to a very complex combination of process thermody-namics and kinetics, with intricate reaction schemes including ionic species, reactionrates varying over a wide range, and complex mass transfer and reaction coupling.As compared to distillation, RA is a fully rate-controlled process, and it definitelyoccurs far from the equilibrium state. Therefore, practitioners and theoreticians arehighly interested in establishing a proper rate-based description of this process.

1.2. Reactive Distillation

Reactive distillation is a combination of chemical reactions and distillation(Figure 4b). This operation provides promising process alternatives to traditional

FIGURE 3 Scheme of an absorber–desorber link. (Adapted from Ref. 3.)


sequential operations, shown in Figure 4a. Among potential advantages of RD are:

New, less expensive productsHigher efficiency because of overcoming thermodynamic and kinetic

limitationsBetter selectivity due to suppressing of undesired reactionsHigher raw material conversionAvoiding of hot spotsSavings due to smaller equipmentLess environmental pollution

One can distinguish between homogeneously and heterogeneously cata-lyzed RD; the latter is often called catalytic distillation (CD).

The applicability of the RD process is highly dependent on the propertiesof the chemical system at hand. A classical example for which RD is recom-mended may be the reaction in which the products are generated by a reversiblereaction, e.g., in the production of methyl acetate. This system is very complexbecause of the occurrence of several azeotropes between reactants and products.

TABLE 1 Applications of RA Processes

Aim of the process Example Application area Refs.

Removal of harmful Coke oven Gas purification 23–25substances gas purification,

amine washingRetrieval/regeneration Solvent Gas separation 26

of valuable substances regenerationor nonreacted reactants

Production/preparation Manufacture of Chemical 27–29of particular products sulphuric acid, synthesis

formaldehydepreparation,manufacture of soda ash

Manufacture of Fertilizer 30nitrogenous industryfertilizers

Water removal Water removal Gas drying 31,32from natural gas, air drying

Conditioning of Synthesis gas Gas separation/ 26gas streams conditioning gas purification


A usual solution in this case is a sequence of a reactor and several separation units(Figure 5). Another way—an integrated RD process such as shown in Figure 6—allows for simultaneous formation of methyl acetate in the reaction zone, extract-ive distillation and product enrichment in the upper part of the column, andmethanol separation in the stripping zone. The production of esters such asmethyl acetate, ethyl acetate, and butyl acetate has for years been an interestingRD application.

The most important application of RD today seems to be the production ofethers such as methyl tertiary butyl ether (MTBE), ethyl tertiary butyl ether(ETBE), and tertiary amyl methyl ether (TAME), which are widely used as mod-ern gasoline components. Figure 7, upper part, shows a traditional process forMTBE production, which is a strongly exothermic reaction. The disadvantages ofthat process can be avoided if the reaction and separation take place within thesame zone of the reactor (Figure 7, lower part).

Table 2 gives a short overview of possible RD applications.The design of RD is currently based on expensive and time-consuming

sequences of laboratory and pilot-plant experiments, since there is no commer-cially available software adequately describing all relevant features of reactions(catalyst, kinetics, holdup) and distillation (VLE, thermodynamics, plate and pack-ing behavior) as well as their combination in RD. There is also a need to improvecatalysts and column internals for RD applications (1,51). Figures 8 and 9 showsome examples of catalytic internals, applied for reactive distillation.

1.3. Reactive Extraction

Liquid–liquid extraction is based on partial miscibility of liquids. In the simplestextraction system, two compounds have to be separated. This can be done byextracting with a carefully selected solvent, in which one compound (solute) easilydissolves whereas the other (nonsolute) does not. The solvent has to be recovered

FIGURE 4 Reactive distillation (b) as alternative to the sequential operation (a).


FIGURE 5 Methyl acetate synthesis: conventional scheme. (From Ref. 33.)


FIGURE 6 Methyl acetate synthesis: reactive distillation scheme. (From Ref. 33.)

FIGURE 7 MTBE synthesis: conventional scheme (above) and reactive distil-lation scheme (below).


from the extract for recycling. In countercurrent extraction processes, there is alight phase and a heavy phase, with one phase dispersed in the other. Which phasehas to be dispersed is an important question in the design of the process.

In reactive extraction, the use of liquid ion exchangers is recommended inorder to extract ionic solutes. These exchangers can be applied to manifold extrac-tion processes in the chemical industry (e.g., extraction of furfural, organic, and inor-ganic acids), biochemical and pharmaceutical productions (e.g., penicillin, aminoacids), hydrometallurgy (e.g., mining of metals), and all related environmental appli-cations. These last are especially attractive, since liquid ion exchangers react veryselectively and have an advantageous performance at very low feed concentrations.

For practical purposes, an ion exchanger is usually diluted in a nonaromatic,high-boiling diluent (boiling point about 500 K) that is immiscible with water.This prevents solvent losses and toxic problems and gives the organic phase therequired physical properties (high interfacial tension, low viscosity, low density),since most liquid ion exchangers are highly viscous or even solid. In some casesa modifier, usually a long-chain alcohol, is added to help in the solubilization ofthe solute–ion exchanger complex. At very high solute loadings, a split of theorganic phase in a solvent-rich and a solvent-poor fraction may occur, especiallywhen using aliphatic diluents. The organic phase in RE is thus not a single substance,as in physical extraction systems in which such three-phase liquid systems are notencountered. Re-extraction is usually performed with chemicals, for instance,with strong mineral acids.

All liquid ion exchangers can be mixed together in order to generate syner-gistic effects. As a special case, an equimolar mixture of cation and anionexchangers gives a “mixed” extraction system, which can extract salts or acids. Inthis case the re-extraction occurs by shift of either temperature, aqueous ionicstrength, or acidity/basicidity.

Equilibrium and selectivity constitute important aspects of reactive andnonreactive extraction processes. Another important factor is the reaction kine-tics, which has to be reasonably fast. Most RE processes are close to equilibriumin less than five minutes. Many ion exchangers need reaction times of less thanone minute, and thus diffusion of the solute complex in the organic phase is therate-determining step.

The cation and anion exchangers are amphiphilic substances that are ad-sorbed at the interface. The latter is then rigid and independent of the dropletdiameter, since friction forces are shielded. This is similar to physical extractionsystems, in which an analogous behavior is caused by surfactants in the aqueousfeed accumulated at the interface.

The problems concerning reaction equilibrium and kinetics descriptionbased on chemical potentials rather than on concentrations are extensively dis-cussed in Refs. 54 and 55, using the zinc system. The latter is recommended as areactive liquid–liquid reference extraction test system by the European Federation of


TABLE 2 Applications of RD Processes

Reaction type Synthesis Catalysta Refs.

Esterification Methyl acetate from hom. 33methanol and acetic acid

Methyl acetate from het. 34,35methanol and acetic acid

Ethyl acetate from no data 36ethanol and acetic acid

Butyl acetate from hom. 37butanol and acetic acid

Transesterification Ethyl acetate from hom. 38ethanol and butyl acetate

Diethyl carbonate from het. 39ethanol and dimethyl carbonate

Hydrolysis Acetic acid and het. 40methanol from methyl acetate and water

Etherification MTBE from isobutene het. 41,42and methanol

ETBE from isobutene het. 43and ethanol

TAME from het. 44isoamylene and methanol

Alcylation Cumene from het. 45propylene and benzene

Condensation Diacetone alcohol het. 46from acetone

Bisphenol-A from no data 47phenol and acetone

Dismutation Monosilane from het. 48trichlorsilane

Hydration Mono ethylene glycol hom. 49from ethylene oxide and water

Nitration 4-Nitrochlorobenzene hom. 50from chlorobenzene and nitric acid

a hom.: homogeneously catalyzed, het.: heterogeneously catalyzed.


Chemical Engineering (EFCE) and is thus well documented (http://www.icheme.uk/learning/ or http://www.dechema.de/extraktion).

The selection of the right solvent is the key to successful separation by non-reactive and reactive liquid–liquid extraction. In this respect, different criteriashould be taken into account, e.g.,

SelectivityCapacityRecoverability of solventDensityViscosity and melting pointInsolubility of solventInterfacial tensionToxicity and flammabilityCorrosivityThermal and chemical stabilityAvailability and costsEnvironmental impact.

FIGURE 8 (a) Schematic of an RD column filled with catalytic internals CD TECH(1—catalytic balls, 2—feed, 3—distillate, 4—bottom product, 5—sieve tray) and(b) catalytic structured packing Sulzer Katapak-S. (Part a from Ref. 52.)


http://www.icheme.uk/ learning

http://www.icheme.uk/ learning

http://www.dechema.de/extraktion

Some of these criteria are crucial, while others are desirable properties improvingseparation and/or making it more economical. Solvent selectivity, recoverability, anda large density difference in respect to the raffinate are essential. Some of the require-ments on the solvent can be in conflict, and thus a compromise may be necessary.

Because aromatic diluents are more expensive and more toxic than alipha-tic ones, the latter are preferably used in industrial practice (see earlier). Aromaticdiluents, with equivalent molecular weights comparable to those of aliphaticones, are more polar and thus more water soluble. The degradation of the diluentis usually negligible in comparison with that of the ion exchanger. The latter onecan be chemically and thermally degrading and also can be poisoned by an irre-versibly extracted compound.

Reactive extraction is closely related to the droplet phenomena, and thusmost theoretical models are based on droplet consideration. Their experimentalevaluation can be done using either a rising (falling) droplet apparatus (Figure 10a)for short residence times or a Venturi tube for long contact times (Figure 10b) (56).

FIGURE 9 (a) Catalytic structured packing Montz Multipak and (b) an exam-ple of reactive trays. (Part b from Ref. 53.)


Monodispersed droplets can be produced and in the latter case captured by thecounterflowing continuous phase in the conus of the Venturi tube (see Figure 10b).

The RE process proceeds in three major types of equipment: mixer-settlersystems, column extractors, and centrifugal extractors. Countercurrent columnextractors can be further subdivided into nonagitated nonproprietary columns andagitated proprietary extractors. Agitating the liquid–liquid system breaks updroplets and increases the interfacial area to improve the mass transfer and column efficiency. Various forms of energy input are used, e.g., rotation of pro-pellers, impellers, and discs; pulsation, vibration, and ultrasonic devices; and cen-trifugal devices.

Some examples of mechanically agitated contactors are the rotating-diskcontactor (RDC), Karr, Oldshue–Rushton, Scheibel, and Kühni columns shownin Figure 11.

There are three types of nonproprietary nonagitated types of extractioncolumns (see Figure 12). The spray columns are the simplest type of extractors,containing only distributors for the feed (often through perforated pipes). This

FIGURE 10 (a) Rising-droplet apparatus and (b) Venturi tube for droplet masstransfer experiments. (1—feed storage, 2—metering pump, 3—double-flowvalve, 4, 5—pumps, 6—heat exchanger, 7—collecting funnel, 8—stream toanalysis, 9—valve).


makes them cheap; however, they are limited in use due to significant axial mix-ing in the column and the fact that the phases are not coalesced and redistributed.

This often results in low efficiencies, which are comparable to one or twotheoretical equilibrium stages. Packed columns are much more efficient since thepacking reduces back-mixing and enhances drop reformation. The packing typesthat can be used are the same as those for normal distillation operations (e.g.,rings, saddles, or slightly modified structured packings of corrugated metalsheets). Compared to packed beds, structured packings need a reduced cross-sectional area for liquid flow, resulting in smaller column diameters. Sieve-traycolumns resemble the distillation column design, except that there is no weir. In

FIGURE 11 Agitated extractors. Left to right: RDC, Karr, Oldshue–Rushton,Scheibel, Kühni extraction columns.

FIGURE 12 Nonagitated extractors. Left to right: spray, packed, sieve-tray(light), sieve-tray (heavy) column.


an extraction process, either phase, the light or the heavy one, can be dispersed.This means that there are also two sieve-tray designs in respect to downcomer andtop- or bottom-settler/distributor design.

Generally, the selection of a specific RE contactor is complicated due to thelarge number of types available and the number of design parameters. The prac-tical handling and design of a reactive solvent extraction processes can be foundelsewhere (see, e.g., Refs. 12 and 13).

2. FUNDAMENTALS OF PROCESS MODELING

2.1. General

As already mentioned, all three considered RSPs reveal significant similarity, andhence their modeling methods are based largely on the same framework.

Because of their multicomponent nature, RSPs are affected by a complexthermodynamic and diffusional coupling, which, in turn, is accompanied by simul-taneous chemical reactions (57–59). To describe such phenomena adequately,specially developed mathematical models capable of taking into consideration col-umn hydrodynamics, mass transfer resistances, and reaction kinetics are required.

Homogeneously catalyzed RD, with a liquid catalyst acting as a mixturecomponent, and auto-catalyzed RD present essentially a combination of transportphenomena and reactions taking place in a two-phase system with an interface. Inthis respect they are very similar to RA and RE, and, generally, reaction has to beconsidered both in the bulk and in the film region. For slow reactions, a reactionaccount exclusively in the bulk phase is usually sufficient.

For heterogeneous systems (CD), it is generally necessary to consider add-itionally the phenomena in the solid catalyst phase. In this case, very detailedmodels using intrinsic kinetics and covering mass transport inside the porous cata-lyst arise (see, e.g., Refs. 60–62). However, it is often assumed that all internal(inside the porous medium) and external mass transfer resistances can be lumpedtogether (35,63,64). In this case each catalytically active site is in contact with theliquid bulk, i.e., the catalyst surface is totally exposed to the liquid bulk phase andcan be completely described by the bulk variables (9,64). This results in the so-called pseudo-homogeneous models. If the reaction (either homogeneous or het-erogeneous) is very fast, it does not depend on the reaction kinetics and thus canbe described using the data on chemical equilibrium only.

Modeling of hydrodynamics in gas/vapor/liquid–liquid contactors includesan appropriate description of axial dispersion, liquid holdup, and pressure drop. Thecorrelations giving such a description have been published in numerous papers andare collected in several reviews and textbooks (e.g., Refs. 65 and 66). Nevertheless,there is still a need for a better description of the hydrodynamics in catalytic col-umn internals; this is being reflected by research activities in progress (67).


The description of thermodynamics and chemical properties of the RSP isvery process specific, and hence its general detailed discussion would constitutea separate issue. Therefore, we will give only a brief discussion of these topics inthe context of the following case studies (Section 3). Further related details canbe found in Refs. 68–74.

In order to model large industrial reactive separation units, a proper sub-division of a column apparatus into smaller elements is usually necessary. Theseelements (the so-called stages) are identified with real trays or segments of apacked column. They can be described using different theoretical concepts, witha wide range of physicochemical assumptions and accuracy.

2.2. Equilibrium-Stage Model

In recent decades, the modeling and design of RSPs has usually been based onthe equilibrium-stage model. Since 1893, when the first equilibrium-stage modelwas published by Sorel (75), numerous publications discussing various aspects ofmodel development, application, and solution have appeared in the literature (76).The equilibrium-stage model assumes that each gas/vapor/liquid stream leaving atray is in thermodynamic equilibrium with the corresponding liquid stream leav-ing the same tray. For the packed columns, the idea of the height equivalent to thetheoretical stage (HETS) is used. In case of RSPs, the chemical reaction has to beadditionally taken into account, either via reaction equilibrium equations or viarate expressions integrated into the mass and energy balances.

In this respect, much depends on the relation between the mass transfer andreaction rates in a particular RSP. The definition of the Hatta number represent-ing the reaction rate in reference to that of the mass transfer helps to discriminatebetween very fast, fast, average, and slow chemical reactions (68,77).

If a fast reaction system is considered, the RSP can be satisfactorydescribed assuming a reaction equilibrium. Here, a proper modeling approach isbased on the nonreactive equilibrium-stage model, extended by simultaneouslyusing the chemical equilibrium relationship.

Such descriptions can be appropriate enough for instantaneous reactionsand those close to them. In contrast, if the chemical reaction is slow, the reactionrate dominates the whole process, and therefore, a reaction kinetics expressionhas to be integrated into the mass and energy balances. This concept has beenused in a number of studies, for RA (e.g., Refs. 78 and 79), RD (e.g., Refs. 80 and 81), and RE (e.g., Refs. 8 and 12) process simulations.

In practice, RSPs rarely operate at thermodynamic equilibrium. Therefore,some correlation parameters, such as tray efficiencies or HETS values, have beenintroduced to adjust the equilibrium-based theoretical description to reality. Formulticomponent mixtures, however, this concept often fails, since diffusion interac-tions of several components result in unusual phenomena such as osmotic or reverse


diffusion and mass transfer barrier (82,83). These effects cause a strange behavior ofthe efficiency factors, which are different for each component, vary along the columnheight, and show a strong dependency on the component concentration (57,83,84).

The acceleration of mass transfer due to chemical reactions in the inter-facial region is often accounted for via the so-called enhancement factors(27,68,69). They are either obtained by fitting experimental results or derived theo-retically on the grounds of simplified model assumptions. It is not possible toderive the enhancement factors properly from binary experiments, and significantproblems arise if reversible, parallel, or consecutive reactions take place.

The equilibrium-stage model seems to be suitable for esterification reactionin CD processes (see Refs. 35 and 74). However, it cannot be recommended forall reaction types, especially those with higher reaction rates.

2.3. Rate-Based Approach

A more physically consistent way to describe a column stage is known as the rate-based approach (57,85,86). This approach implies that actual rates of multi-component mass and heat transfer and chemical reactions are taken into accountdirectly.

Considering homogeneous RSPs, mass transfer at the gas/vapor/liquid–liquid interface can be described using different theoretical concepts (57,59). Mostoften the two-film model (87) or the penetration/surface renewal model (27,88) isused, in which the model parameters are estimated via experimental correlations.In this respect the two-film model is advantageous since there is a broad spectrumof correlations available in the literature, for all types of internals and systems.For the penetration/surface renewal model, such a choice is limited.

In the two-film model (Figure 13), it is assumed that all of the resistance tomass transfer is concentrated in thin stagnant films adjacent to the phase interfaceand that transfer occurs within these films by steady-state molecular diffusionalone. Outside the films, in the bulk fluid phases, the level of mixing is so highthat there is no composition gradient at all. This means that in the film region,only one-dimensional diffusion transport normal to the interface takes place.

Multicomponent diffusion in the films is described by the Maxwell–Stefanequations, which can be derived from the kinetic theory of gases (89). TheMaxwell–Stefan equations connect diffusion fluxes of the components with thegradients of their chemical potential. With some modification these equationstake a generalized form in which they can be used for the description of real gasesand liquids (57):

(1)dx N x N

c Di ni

i Lj j

Lt ijj

nLi�

��

�1

1∑ , . . . ,


where di is the generalized driving force:

(2)

Similar equations can be also written for the gas/vapor phase.Thus the gas/vapor/liquid–liquid mass transfer is modeled as a combination

of the two-film model and the Maxwell–Stefan diffusion description. In this stagemodel, the equilibrium state exists only at the interface.

The film thickness represents a model parameter that can be estimatedusing mass transfer coefficient correlations. These correlations reflect the masstransport dependence on physical properties and process hydrodynamics and areavailable from the literature (see, e.g., Refs. 57, 68 and 90).

The two-film model representation can serve as a basis for more compli-cated models used to describe heterogeneously catalyzed RSPs or systems con-taining suspended solids. In these processes a third solid phase is present, andthus the two-film model is combined with the description of this third phase. Thiscan be done using different levels of model complexity, from quasi-homogeneousdescription up to the four-film presentations that provide a very detailed descrip-tion of both vapor/gas/liquid–liquid and solid/liquid interfaces (see, e.g., Refs. 62,68 and 91). A comparative study of the modeling complexity is given in Ref. 64for fuel ether synthesis of MTBE and TAME by CD.

2.4. Computational Fluid Dynamics

Every separation unit operation is governed by the continuum conservation laws,and thus, in principle, everything meaningful to know in the continuum for anyprocess can be determined with computational fluid dynamics (CFD) (92). Inrecent years there have been significant academic and industrial efforts to enable

dx

T

d

dzi ni

i i� �ℜ

�1, . . . ,

FIGURE 13 Two-film model.


the use of CFD for the design, scale-up, and optimal operation of a variety ofchemical process equipment.

Special attention has been given to the CFD modeling of two-phase flows. Themost frequently encountered computational techniques for calculating multiphaseflows are Euler–Lagrange and Euler–Euler methods. Euler–Lagrange models areapplicable to dispersed flows (93). In these models the flow of the so-called “carrierphase” is simulated by solving continuum-flow equations. The motion of individualparticles (or group of particles) of the dispersed phase is tracked through the flowdomain using the calculated carrier-phase flow field as input; afterwards, mass,momentum, and energy transfer between the two phases are computed and applied tothe carrier-phase flow field prediction. This procedure requires several iterations (94).

Euler–Euler models assume interpenetrating continua to derive averagedcontinuum equations for both phases. The probability that a phase exists at a cer-tain position at a certain time is given by a phase indicator function, which, forsteady-state processes, is equivalent to the volume of fraction of the correspond-ent phase (volume-of-fluid technique). The phase-averaging process introducesfurther unknowns into the basic conservation equations; their description requiresempirical and problem-dependent input (94). In principal, Euler–Euler modelsare applicable to all multiphase flows. Advantages and disadvantages of bothmethods are compared, e.g., in Refs. 95 and 96.

The volume-of-fluid technique can be used for a priori determination of themorphology and rise characteristics of single bubbles rising in a liquid (97,98).Considerable progress has been made in CFD modeling of bubbling gas–solid fluidized beds by adoption of the Eulerian framework for both the dilute (bubbles)and dense (emulsion) phases (99–102). The use of CFD models for gas–liquidbubble columns has also aroused significant interest in recent years, and bothEuler–Euler and Euler–Lagrange methods have been employed for the descriptionof the gas and liquid phases (94–96,103–113).

There are also some attempts available in the literature to model tray hydro-dynamics using CFD (114–119).

Despite considerable success in some fields of application, the CFD simu-lations are still not fully mastered, especially where the considered processesreveal clearly nonhomogeneous, segregated fluid flow patterns. The latter are usu-ally the basic phenomenon in packed or filmlike units used in reactive and non-reactive separations.

One of the important issues with RSPs is the development of efficient col-umn internals. Such internals have to enhance both separation and reaction andmaintain a sound balance between them. This is valid for both homogeneouslyand heterogeneously catalyzed processes, being especially important for CD. Anunderstanding of the complex, multiphase flow on the internals interrelated with themass transport and chemical reaction constitutes a very important challenge for thefuture. Some first steps in this respect have been done concerning the performance


of SULZER packings KATAPAK-S® and SULZER BX and OPTIFLOW(67,120–122) as well as trays on which chemical reaction occurs (119).

Recently, a substantial effort has been made to optimize column internalsfor reactive separations and to reduce the number of expensive hydrodynamicsexperiments via the CFD simulations (67,119,122). Such simulations can beregarded as virtual experiments carried out in order to predict the performance ofthe internals by varying geometrical and structural parameters, thus reducing theoptimization time.

Necessary new hydrodynamic models have to be formulated and tested forthree-dimensional description of two-phase flow through the internals. Sinceaccurate resolution of the trickle-flow scale is not feasible at the moment, suchflow details have to be simplified and are subject to the subgrid modeling sup-ported by experimental investigations of small-scale phenomena.

The CFD simulations should be linked with the rate-based process simula-tor, providing important information on the process hydrodynamics in the form ofcorrelations for mass transfer coefficients, specific contact area, liquid holdup,residence time distribution, and pressure drop. An ability to obtain these correlationvia the purely theoretical way rather than by the traditional experimental one shouldbe considered a significant advantage, because this brings a principal opportunity tovirtually prototyping of new optimized internals for reactive separations.

The local aspects of liquid–liquid two-phase flow in RE has been the focusof CFD analysis by different research groups (123–126). In principle, all aspectsconcerning single-phase flow phenomena (residence time distribution, impellerdischarge flow rate, etc.) can be tackled, even with complex geometries. However,the two-phase CFD is still a challenge, and the droplet interactions (breakup andcoalescence) and mass transfer are not implemented in commercially availablecodes. Thus these issues constitute an open area for further research and devel-opment (127).

3. CASE STUDIES

3.1. Absorption of NOx

3.1.1. Chemical System

The reactive system considered is a basic one in the production of nitric acid aswell as in some other industrial processes (19). It consists of 10 components,including air (N2, O2), water (H2O), oxyacids of nitrogen (HNO2, HNO3), andnitrogen oxides (NO, NO2, N2O, N2O3, N2O4). The components are involved insimultaneous, parallel, and consecutive reactions occurring in both phases. Thereactions are of high orders and most of them are exothermic.

Reaction kinetics is described by the scheme suggested in Ref. 128 andmodified in Ref. 129. This scheme involves eight reactions and can be regarded


as the most extensive reaction system so far. The gas-phase reactions are gov-erned by the following equations:

whereas the corresponding equations for the liquid phase are

The liquid-phase reactions are valid for nitric acid concentrations below34 wt %. In the case of higher nitric acid concentrations, Reactions (R5) to (R7)become reversible. The oxidation of NO (Reaction (R1)) is the slowest reaction inthis system. Therefore, the total gas-phase holdup in absorbers can be determinedusing the kinetic data for this reaction (130). The other gas-phase reactions areinstantaneous equilibrium reactions.

3.1.2. Process Setup

Measurements of an industrial NOx absorption process, schematically shown inFigure 14, were described in Ref. 131. The absorption plant constitutes a sequenceof four units used for the removal of nitrogen oxides from the waste gas of an adipinacid factory. Each unit is separated by a metal plate into two sections. In fact thereare eight columns joined together as a countercurrent absorption plant. This plantis operated at atmospheric pressure. Columns 1–7 have a pump around for coolingof the liquid. The diameter of each column is 2.2 m; the height is 7 m. The packingheight is 3.2 m. The packing consists of 35 mm INTALOX ceramic saddles.

2 10 72

3 99

5 03

7 17

2 30

0

30

0

NO H O HNO HNO kJ/mol (R5)

N O H O 2HNO kJ/mol (R6)

N O H O HNO HNO kJ/mol (R7)

3HNO HNO H O 2NO kJ/mol (R8)

2 2

2 3 2 2

2 4 2 2

2 3 2

� � ��

� ��

� � ��

� � ��

→

→

→

↔

�

�

�

�

H

H

H

H

R

R

R

R

.

.

.

.

2NO O NO kJ/mol (R1)

NO NO N O kJ/mol (R2)

2NO N O kJ/mol (R3)

3NO H O 2HNO NO kJ/mol (R4)

2

2 2 4

2 2 3

� ��

� ��

��

� � ��

2 20

2 30

0

0

2 114

39 9

57 2

35 4

→

↔

↔

↔

�

�

�

�

H

H

H

H

R

R

R

R

.

.

.


The liquid feeds entering columns 7 and 8 are low-concentration nitricacids. The liquid product has a HNO3 concentration of about 35 wt %. The gasfeed has a concentration of NOx of about 60,000 vppm. A quarter of NOx is NO;the rest is NO2.

3.1.3. Results and Discussion

The sensitivity analysis performed in Ref. 129 shows that the suggested modelprovides concentration profiles that are qualitatively correct. For the simulation ofthe industrial absorption process shown in Figure 14, the following correlationsensuring the most reliable results are selected:

The rate constant of Reaction (R1) (the slowest and hence the most import-ant reaction in the system) according to Ref. 132

The liquid-side mass transfer coefficient according to Ref. 133The gas-side mass transfer coefficient according to Wehmeier (see Ref. 134)

Figures 15 and 16 give an illustration of the model quality. Figure 15 showsa comparison of the simulated and measured gas-phase concentrations of NO andNO2 throughout the whole absorption plant, whereas in Figure 16, experimentaland simulated liquid-phase concentrations of HNO3 and HNO2 are demonstrated.

FIGURE 14 Absorption plant consisting of four units (eight columns).


The zigzag form of the simulated concentration profiles results from switchingdifferent sections of each single column (see Ref. 135). Good agreement betweenexperimental and simulation results can be readily observed, except for the firsttwo columns. Here the larger deviations between experiments and simulatedresults can be attributed to the fact that at high concentration of HNO3 Reactions(R5) to (R7), assumed to be irreversible reactions, convert to reversible ones; thedata on their rate constants are lacking.

3.2. Coke Gas Purification


Coke oven gas consists mainly of a mixture of carbon monoxide, hydrogen,methane, and carbon dioxide. It is contaminated with a variety of organic andinorganic compounds that have to be separated in absorption columns before itsfurther use as a synthesis gas. The selective absorption of coke plant gas contam-ination results from a complex system of parallel liquid-phase reactions.

Instantaneous reversible reactions:

NH H O NH OH R9)

H S H O HS H O R10)

2

2 2 3

3 4� �

� �

� �

� �

↔

↔

(

(

FIGURE 15 Experimental and simulated gas-phase concentrations of NO andNO2 throughout the absorption plant.


Finite-rate reversible reactions:

The reactions including CO2 obey first- and second-order kinetics, where-as the other reversible reactions are based on simple proton transfers and aretherefore regarded as instantaneous by the corresponding mass action law equa-tions. The formation of bicarbonate ions (HCO3

�) takes place via two different

CO OH HCO R14)

CO 2H O HCO H O R15)

CO NH H O H NCOO H O R16)

2 3

2 2 3 3

2 3 2 2 3

�

� �

� � �

� �

� �

� �

↔

↔

↔

(

(

(

HCN H O CN H O R11)

HCO H O CO H O R12)

H O OH H O R13)

2 3

2

3 2

3 32

3

2

� �

� �

�

� �

� � �

� �

↔

↔

↔

(

(

(

FIGURE 16 Experimental and simulated liquid-phase concentrations of HNO3and HNO2 throughout the absorption plant.


mechanisms. The rate of the direct reaction between carbon dioxide and hydroxylions (the most important step) is taken from Ref. 28.

Usually the reaction between CO2 and water is very slow and hardly con-tributes to the total rate of reaction of carbon dioxide. Nevertheless, in this workit was considered of the first order with respect to CO2, since the reaction kine-tics depends on the carbonation ratio (136).

The absorption rate of carbon dioxide increases in the presence of aminesor ammonia. Therefore, the reaction kinetics of NH3 and CO2 has been consid-ered in the model equations, too. The rate constant as a function of the temperaturehas been determined according to Ref. 136. The coefficients for the calculation ofthe chemical equilibrium constants in this system of volatile weak electrolytes aretaken from Ref. 137.

The CO2 absorption is hindered by a slow chemical reaction by which thedissolved carbon dioxide molecules are converted into the more reactive ionicspecies. Therefore, when gases containing H2S, NH3, and CO2 contact water, theH2S and ammonia are absorbed much more rapidly than CO2, and this selectivitycan be accentuated by optimizing the operating conditions (23). Nevertheless, allchemical reactions are coupled by hydronium ions, and additional CO2 absorptionleads to the desorption of hydrogen sulfide and decreases the scrubber efficiency.


Today’s coke plant gas purification processes are mostly carried out under atmos-pheric pressure, employing a circulated ammonia-based absorbent. The con-sumption of the external solvent is reduced via the use of ammonia available inthe coke gas (138). An example of innovative purification processes is the ammo-nia hydrogen sulfide circulation scrubbing (ASCS) (Figure 17), in which theammonia contained in the raw gas dissolves in the NH3 absorber and then theabsorbent saturated with the ammonia passes through the H2S absorber to selec-tively absorb the H2S and HCN components from the coke gas. The next step is thethermal regeneration of the absorbent with the steam in a two-step desorption plant,whereas a part of the deacidified water is fed back into the H2S absorber (25).

Pilot-plant experiments have been carried out at real process conditions inthe coke plant “August Thyssen” (Duisburg, Germany). The DN 100 pilot column(Figure 17) was made of stainless steel and equipped with about 4 m of structuredpacking (Sulzer MELLAPAK® 350Y), three liquid distributors, and a digital con-trol system. Several steady-state experiments have been compared with the sim-ulation results and supported the design optimization of the coke gas purificationprocess (25).


A number of steady-state simulations have been performed with the aim of ana-lyzing the influence of numerical and physicochemical parameters, beginning


FIGURE 17 Ammonia hydrogen sulfide circulation scrubbing process for the coke oven gas purification(right) and H2S absorber (left).


with a single stage and ending with a simulation of a column. Different film andpacking section discretizations, several mass transfer and hydrodynamic corre-lations, and different driving forces and diffusion models have been thoroughlytested (Figure 18).

The most sensitive components appeared to be those involved in finite-ratereactions, especially CO2. Furthermore, the impact of electrical forces enhancesthe absorption of the strong electrolytes H2S and HCN by 3–5%, while the CO2absorption rate is dominated by the reaction in the film (139,140). Significantchanges in the concentration profiles and the component absorption rates due tothe film reaction have been established (141,142).

Single-stage simulations reveal that intermolecular friction forces do notlead to reverse diffusion effects, and thus the molar fluxes calculated with theeffective diffusion approach differ only slightly from those obtained via theMaxwell–Stefan equations without the consideration of generalized drivingforces. This result is as expected for dilute solutions and allows one to reducemodel complexity for the process studied (143).

As a further model simplification, a linearization of the film concentrationprofiles has been studied. This causes no significant changes in the simulationresults and at the same time reduces the total number of equations by half and sta-bilizes the numerical solution (142). The assumption of chemical equilibrium in

FIGURE 18 Absorption rates calculated with different model assumptionsconcerning reaction consideration.


the liquid bulk phase does not change the absorption rates significantly, whichindicates fast conversion. Therefore, neglecting the film reaction unrealisticallyreduces the absorption rates. On the other hand, neglecting the reaction kineticswithin the film results in completely different orders of magnitude for the calcu-lated absorption degree. As a consequence, the reactions of carbon dioxide shouldnot be regarded as instantaneous, although the corresponding Hatta number ofabout 7 characterizes the reaction as very fast (3).

The model optimized with respect to the numerical parameters and physico-chemical properties has been validated against experimental data, whereas theaxial concentration and temperature profiles for both phases demonstrated goodagreement (Figure 19). It has also been found that the simulations of the scrubberbased on the equilibrium-stage model extended by the chemical reaction kineticsyield results completely inconsistent with the experimental studies; namely, theselectivity toward H2S and HCN absorption cannot be reflected (Figure 19). Inthis case, the film reaction represents an essential element of the rate-basedapproach that has to be considered in the model. As a result, the only feasible sim-plification is represented by a linearization of the film concentration profiles,including the implementation of the average reaction kinetics in the liquid filmregion (143).

FIGURE 19 Liquid-phase axial concentration profiles for the H2S scrubber:comparison between experimental and simulation results based on differentmodel approaches.


3.2.4. Dynamic Modeling

Steady-state modeling is not sufficient if one faces various disturbances in RA oper-ations (e.g., feed variation) or tries to optimize the startup and shutdown phases ofthe process. In this case, a knowledge of dynamic process behavior is necessary.Further areas where the dynamic information is crucial are the process control aswell as safety issues and training. Dynamic modeling can also be considered as thenext step toward the deep process analysis that follows the steady-state modelingand is based on its results.

The dynamic formulation of the model equations requires a careful analy-sis of the whole system in order to prevent high-index problems during thenumerical solution (144). As a consequence, a consistent set of initial conditionsfor the dynamic simulations and suitable descriptions of the hydrodynamics haveto be introduced. For instance, pressure drop and liquid holdup must be corre-lated with the gas and liquid flows.

The model optimized based on steady-state analysis allows for a dynamicreal-time simulation of the entire absorption process. Because dynamic behavior isdetermined mainly by process hydraulics, it is necessary to consider those elementsof the column periphery that lead to larger time constants than the column itself.Therefore, major elements of the column periphery, such as distributors, stirredtanks, and pipelines, have been additionally implemented into the dynamic model.

The dynamic behavior of the coke gas purification process has been inves-tigated systematically (139,140,145). For instance, local perturbations of the gasload and its composition have been analyzed. A significant dynamic parameter isrepresented by the liquid holdup. Figure 20 demonstrates the changes of the sol-vent composition after a decrease of the gas-flow rate from 67 m3/h to 36.4 m3/hand a simultaneous small increase in the liquid-flow rate.

The liquid holdup of the packing section decreases, which leads to a lowerconversion of the kinetically controlled reactions of CO2 and a reduction in theCO2 absorption rate. As a consequence, the solvent mole fractions of HCO3

� andcarbamate decreases whereas the relative fraction of HS� increases. The select-ivity of the absorption process toward the H2S and HCN reduction is enhanced byminimizing the liquid holdup of the column. At the same time, a larger interfacialarea improves the performance of the plant. Therefore, modern industrial sour gasscrubbers should be equipped with structured packings.

Figure 21 illustrates the system response after a sudden increase in the gasflow by 20% and its H2S load by 100%. As expected, the H2S load increaseseverywhere along the column height in the gas phase. The change is more signifi-cant in the lower part of the absorber than at the top because some additionalhydrogen sulfide is absorbed. The new steady state is already achieved after 30 minutes, which justifies the implementation of dynamic models for the columnperiphery.


FIGURE 20 Dynamic change of solvent composition after a sudden significant decrease in the gas-flowrate and a simultaneous small increase in the liquid-flow rate.


3.3. Methyl Acetate Synthesis, Batch Distillation


The synthesis of methyl acetate from methanol and acetic acid is a slightlyexothermic equilibrium-limited liquid-phase reaction:

(R17)

The low equilibrium constant and the strongly nonideal behavior that causes theforming of the binary azeotropes methyl acetate/methanol and methyl acetate/water make this reaction system interesting as a possible RD application (33).Therefore, methyl acetate synthesis has been chosen as a test system and investi-gated in a semibatch RD column. Since the process is carried out under atmo-spheric pressure, no side reactions in the liquid phase occur (146).

3.3.2. Process Setup and Operation

The catalytic packing MULTIPAK® (147) applied in this case study consists ofcorrugated wire gauze sheets and catalyst bags of the same material assembled in alternate sequence. Sufficient mass transfer between gas and liquid phase is

CH OH (CH )COOH (CH )COO(CH ) H O

kJ/mol

3 3 3 3 2� �

��

↔

�HR0 4 2.

FIGURE 21 Dynamic change of the H2S gas-phase concentration along thecolumn after a sudden increase in the gas flow and its H2S load.


guaranteed by segmentation of the catalyst bags and numerous contact spots withthe wire gauze sheets. The packing was equipped with the acid ion exchange resinknown as an effective catalyst for esterification processes (34,148).

A batch distillation column with a diameter of 100 mm and a reactive pack-ing height of 2 m (MULTIPAK I®) in the bottom section and an additional meterof conventional packing (ROMBOPAK 6M®) in the top section was used. Theflow sheet of the column is shown in Figure 22.

At first, the distillation still was charged with methanol—the low-boilingreactant—and heated under total reflux until steady-state conditions were achieved.At that moment, acetic acid—the high-boiling reactant—was fed above the reactionzone to the second distributor. After 30 min the reflux ratio was changed frominfinity to 2 and the product withdrawal at the top of the column began. During the

FIGURE 22 Reactive distillation column, batch operation.


column operations, the liquid-phase concentration profiles along the column andthe temperature profiles were measured. For the determination of the liquid-phasecomposition, two methods were applied simultaneously. On the one hand, sampleswere taken and analyzed by gas chromatography. On the other hand, an onlineNIR spectrometer was used to determine the concentration without taking anysamples (149).


Figures 23 and 24 show the liquid-phase compositions for, respectively, thereboiler and condenser as functions of time. After column startup, the concentra-tion of methanol decreases continuously whereas the distillate mole fraction ofmethyl acetate reaches about 90%. A comparison of the rate-based simulation(with the Maxwell–Stefan diffusion equations) and experimental results for theliquid-phase composition at the column top and in the column reboiler demonstratestheir satisfactory agreement (Figures 23 and 24). Figure 25 shows the simulation

FIGURE 23 Liquid mole fractions in the column reboiler: lines, simulations;dots, experiments.


results obtained with different modeling approaches, after an operation time of10,000 s. The reference model employs the rate-based approach and the Maxwell–Stefan diffusion equations. Another rate-based model assumes effective diffusioncoefficients instead of the Maxwell–Stefan equations. The third model used is anequilibrium-based one. Both the reference model and effective-diffusion modelshow similar results. The equilibrium-stage model is only able to describe theprocess behavior qualitatively. This is in contrast to the reactive absorptionprocesses (see Sections 3.1 and 3.2) and can be explained by the low reaction rate,which dominates the whole process kinetics.

3.4. Methyl Acetate Synthesis, Steady-State Distillation


The synthesis of methyl acetate from methanol and acetic acid analyzed in thiscase study is the same as described by Reaction (R17):

CH OH (CH COOH CH COO(CH H O3 3 3 3 2� �) ( ) )↔

FIGURE 24 Liquid mole fractions in the column condenser: lines, simula-tions; dots, experiments.



The column (35) has an inner diameter of 50 mm and a total packing height of 4 m composed of a reactive section of 2 m and two nonreactive sections of 1 meach, below and above the reactive part of the column (Figure 26). The acetic acidfeed is located above the catalytic packing, while methanol is fed to the columnbelow the reactive section. A similar column design was presented in Ref. 150.


A series of experiments have been performed with a stoichiometric feed ratio ofacetic acid and methanol. The reflux ratio was kept constant at a value of 2.0, thefeed flow rate at a value of 3.0 kg/h, while the heat duty to the reboiler was variedover a wide range. A comparison of experimental results and model prediction forthe liquid-phase composition profiles along the column is given in Figure 27 fordifferent reboiler duties (151). The theoretical values are displayed with continu-ous lines and empty symbols, whereas the experimental data measured along the

.

FIGURE 25 Axial concentration profiles for the semibatch column (t � 10,000 s ) .


column are shown by the relevant filled symbols. It can be seen that the processbehavior is reflected by the simulations with high accuracy. A maximum concen-tration of methanol and acetic acid can be observed at the respective feed loca-tions, while methyl acetate is enriched toward the top and water toward thebottom of the column.

3.5. Synthesis of Methyl Tertiary Butyl Ether


The synthesis of methyl tertiary butyl ether (MTBE) is one of the most importantapplications of RD. MTBE is produced via an acid-catalyzed reaction betweenmethanol and isobutylene:

(R18)

This reaction has been extensively studied by several authors, e.g., Refs. 152–154.


MTBE synthesis was investigated both theoretically and experimentally. Here,some results for a pilot-scale RD column at Neste Oy Engineering, Finland, are

CH OH C(CH CH C(CH OCH

kJ/mol

3 3 2 3 3�

��

) )

.

2 3

0 37 7

↔

�HR

FIGURE 26 Reactive distillation column, steady-state operation.


presented (155). The column (used as an example here) has a catalytic section in themiddle part. This catalytic section may consist either of a packed bed of catalytical-ly active rings (91) or of structured catalytic packing (147). The rectifying and strip-ping sections are filled with Intalox Metal Tower Packing. The methanol feed isintroduced just above the catalyst section of the column and the hydrocarbon feedjust below.


Figure 28 demonstrates the simulated and measured concentration profiles for thepilot test made in the column, with the reactive section filled with catalytically

FIGURE 27 Liquid-phase composition profiles along the 50-mm-diametercatalytic distillation column for methyl acetate synthesis at reflux ratio of 2.0and different reboiler duties: (a) 295W (b) 873W (c) 1161W.


active rings. In the simulations four components—methanol, isobutene, MTBE,and 1-butene—were chosen to describe the system under consideration. Here,segment 1 corresponds to the reboiler. A satisfactory agreement between calcu-lated and measured values can be clearly observed. In Figure 29, the simulationresults for the column with different reactive internals, catalytic packing MULTIPAK®, are shown. Here, 16 components were considered. Again, the liquid bulk composition profiles demonstrated in Figure 29 agree well with theexperimental data.

3.6. Reactive Extraction of Zinc


The extraction of zinc with the cation exchanger di(2-ethylhexyl)phosphoric acid,is recommended by the EFCE as a test system for RE. Physical properties,

handling, equilibrium data, etc. are documented on the internet (http:// www.dechema.de/Extraction, http://www.icheme.org/learning).

In brief, the ion exchanger is dimeric, in aliphatic diluents (156),and the overall reaction is then

(R19)

At low concentrations, polynuclear complexes do not exist (157); thus a � 1:

(R20)Zn R H ZnR (RH) H22 2 2

��

�� b b↔ 2 2 2

a b a a b aZn R H Zn R (RH) H22 2 2

��

�� ↔ 2 2 2

R H ,2 2

RH,

FIGURE 28 Calculated and experimental liquid compositions for experimentswith catalytically active rings.


http://www.dechema.de/Extraction

http://www.dechema.de/Extraction

http://www.icheme.org/learning

FIGURE 29 Calculated and experimental liquid compositions for experiments with catalyticstructured packing.


The mass action law on the basis of concentrations then yields

(3)

where DZn denotes the partition coefficient of zinc between the organic and theaqueous phase (bar indicates organic species). In logarithmic form, the stoichio-metry of the complex, b, is determined by slope analysis:

(4)

As can be seen, the partition increases with increasing ion exchanger concentra-tion, and with pH. The determination of the equilibrium constant, Keq, isdiscussed, e.g., in Ref. 158.


Apart from the nuclear industry, the most frequently installed pieces of processequipment are mixer-settler cascades. The advantage is the easy control of eachstage regarding the pH value, selection of the phase to be dispersed, etc. (12,15).The disadvantages are the high investment costs and large solvent inventory.Nowadays, modern designed extractants (with fast kinetics) allow a processdesign with highly efficient extraction columns (159). Applications are found inthe chemical industry, e.g., with sulfonic acid extraction (160). The first applica-tion on a big scale in hydrometallurgy was reported in 1997 at WMC OlympicDam, Australia, where 10 pulsed Batman columns (0.5 to 3 m in diameter and35 m tall) were used for uranium recovery. An increased recovery (from 90% to97%) was found after replacing the formerly used mixer-settler units.

The RE of zinc is reported in detail when using an RDC, including the dis-cussion of the stripping process to regenerate the ion exchanger for cyclic reuse(161–163).


A comparison of predicted and experimental mass transfer coefficients is given inFigure 30. The simulated overall mass transfer coefficient originates from amodel that is a combination of the microkinetic reaction according to Eq. (B9)and eddy diffusion according to Eq. (B11) (see Appendix B). Figure 30 showsthat the mass transfer coefficient at higher concentrations is generally underesti-mated, thus including some safety value for the process design. As discussed indetail in Ref. 56, combinations of microkinetics [Eq. (B9)] with other eddy dif-fusion correlations instead of Eq. (B11) (e.g., Refs. 164 and 165) are also appro-priate to describe the system. In contrast to this, the combination of microkineticswith molecular diffusion concepts fails, and the same is true when the equilibrium

R H2 2 ,

log R H pH logZn 2 2eqD b K� � �log([ ]) 2

K Dbb b

eq 2 22

2 2Zn

2 2

ZnR RH) H

Zn R H

H

R H� ��

�

�

�[ ( ][ ]

[ ][ ]

[ ]

[ ]2

2 2


approach is used neglecting the kinetics rate law. The results of the column sim-ulations are discussed in Refs. 162 and 163, and a special discussion on contam-ination effects is given in Ref. 166.

4. SUMMARY AND OUTLOOK

This chapter concerns the most important reactive separation processes: reactiveabsorption, reactive distillation, and reactive extraction. These operations combin-ing the separation and reaction steps inside a single column are advantageous ascompared to traditional unit operations. The three considered processes are similarand at the same time very different. Therefore, their common modeling basis is dis-cussed and their peculiarities are illustrated with a number of industrially relevantcase studies. The theoretical description is supported by the results of laboratory-,pilot-, and industrial-scale experimental investigations. Both steady-state anddynamic issues are treated; in addition, the design of column internals is addressed.

Reactive absorption, reactive distillation, and reactive extraction occur inmulticomponent multiphase fluid systems, and thus a single modeling frameworkfor these processes is desirable. In this respect, different possible ways to build sucha framework are discussed, and it is advocated that the rate-based approach pro-vides the most rigorous and appropriate way. By this approach, direct consideration

FIGURE 30 Comparison of experimental and simulated overall mass transfercoefficients at different initial zinc concentrations and droplet diameters (1, 2, or 3 mm ) .


of the diffusional and reaction kinetics is realized. Special attention is paid to theapplication of CFD, which could become a powerful theoretical tool to predict theflow behavior for different column units and internals geometries for engineeringapplications. In particular, CFD can play an outstanding role in the development ofthe column internals for reactive separations. Fundamental advances in the under-standing of the underlying physicochemical phenomena when coupled with CFDwould go a long way toward support of reactive separation technology.

The modeling of RA is illustrated by the absorption of NOx and by the coke gas purification process. The first case is modeled by using an analytical treat-ment of the film phenomena, whereas the second one is solved by a purely numer-ical technique. The simulation results are compared with the experimental dataobtained at an industrial absorption plant consisting of eight units with pump around (NOx) and at a pilot column for the ammonia hydrogen sulfide circulation scrubbing process (coke gas purification). For the latter case, bothsteady-state and dynamic conditions are considered. The comparison results, on theone hand, demonstrate a good agreement between the rate-based simulations andexperimental data, and, on the other hand, warn of using the equilibrium approach,which appears completely inappropriate in the case of complex finite-rate reactions.

The modeling of RD processes is illustrated with the heterogenously catalyzedsynthesis of methyl acetate and MTBE. The complex character of reactive distilla-tion processes requires a detailed mathematical description of the interaction of mass transfer and chemical reaction and the dynamic column behavior. The mostdetailed model is based on a rigorous dynamic rate-based approach that takes intoaccount diffusional interactions via the Maxwell–Stefan equations and overall reac-tion kinetics for the determination of the total conversion. All major influences of thecolumn internals and the periphery can be considered by this approach.

As an application example, the dynamic model was used for the simulationof the steady-state and semibatch production of methyl acetate, performed in apacked column with a catalytic packing. For the model validation, several experi-ments were carried out in a pilot-plant column. For the investigated operationrange, the simulation results are in good agreement with the experimental data.

The use of this model for model-based process control calls for suitablemodel reductions without a significant decrease in the predictivity. For the methylacetate process, a simplified description of the mass transfer using effective dif-fusion coefficients and neglecting diffusional interactions seems to be sufficient.On the other hand, a detailed description of the reaction, including the specificphenomena of the heterogeneous catalysis by an adequate consideration of the solidphase, is required for the predictive simulation of even more complex systems,including side and consecutive reactions. Optimal functioning of reactive distil-lation depends on careful process design, with appropriately selected columninternals, feed locations, and catalyst placement. Greater understanding of thegeneral and particular features of the process behavior is equally essential.


The modeling of RE processes is strongly related to the knowledge of thereaction equilibrium and kinetics and mass transfer regime. The latter is decisive forcolumn simulations, whereas eddy diffusivity concepts often have to be used. Theparameters can easily be obtained in small-scale laboratory devices, with a minimumof substance involved. A real challenge is the correct hydrodynamic description ofthe two-phase flow. The assumption of plug flow gives acceptable results with col-umn diameters smaller than 0.1 m. The commonly applied axial dispersion modelor back-mixing model uses one parameter to account for nonideal flow in eachphase. Here the dispersed phase is considered to be pseudo-continuous andmonodisperse. Droplet population models taking into account the dynamic pro-cesses of coalescence and breakup of droplets should give a more realistic pictureand thus a more firm design of a process. The use of CFD calculations in liquid–liquid dispersed-phase flow is limited to single-droplet flow or low column holdup.The simulation of large industrial columns especially is not feasible nowadays.

Some important general aspects of rate-based modeling as well as furtherpeculiarities of the specific process applications and the different solution strat-egies are given in Appendices A and B.

The key reactive separation topics to be addressed in the near future are aproper hydrodynamic modeling for catalytic internals, including residence timedistribution account and scale-up methodology. Further studies on the hydrodynam-ics of catalytic internals are essential for a better understanding of RSP behaviorand the availability of optimally designed catalytic column internals for them. Inthis regard, the methods of computational fluid dynamics appear very helpful.

The development of new methodologies enabling the creation of intelligent,tailor-made column internals and consequent RSP optimization constitutes oneof the burning present-day challenges. Such a development is already in progressin some European research projects.

Despite the recent rapid development of computer technology and numer-ical methods, the rate-based approach in its current realization still often requiresa significant computational effort, with related numerical difficulties. This is oneof the reasons the application of rate-based models to industrial tasks is ratherlimited. Therefore, further work is required in order to bridge this gap and pro-vide chemical engineers with reliable, consistent, robust, and comfortable simu-lation tools for reactive separation processes.

ACKNOWLEDGMENTS

We would like to thank our colleagues at the Chair of Fluid Separation Processes,Dortmund University, and all other project partners who have been involved in the research activities. We are also grateful to the German Research Founda-tion (DFG, Grants No. Schm 808/5-1, Ba 1569/2-1� 2-2, Ba 1569/6-1), theVolkswagen Foundation (Project No. I/70 875, 876, 877), the European


Commission (BRITE-EURAM program, CEC Project No. BE95-1335), theGerman Federal Ministry of Education and Research (BMBF, Project No.03C0306), the Foundation “Rheinland-Pfalz für Innovation” (836-386261/193),as well as BASF AG, Bayer AG, Axiva GmbH, Degussa.

NOMENCLATURE

aI specific gas–liquid interfacial area m2/m3

As column cross section m2

B liquid load m3/(m2s)c molar concentration mol/m3

CIP adjustable parameter, Eq. (B10)dC column diameter mdi generalized driving force for component i 1/mdp droplet diameter m

Maxwell–Stefan diffusion coefficient m2/sDax axial dispersion coefficient m2/sDeff effective diffusion coefficient m2/sDZn partition coefficient of zincE length-specific energy holdup J/mE� dimensionless residence time distributionF Faraday’s constant 9.65 � 104 C/molFC gas capacity factor Pa0.5

G gas molar flow rate mol/sh molar enthalpy J/mol

reaction enthalpy J/molky overall mass transfer coefficients m/sKi distribution coefficientKeq equilibrium constant[K] reaction matrix [Eq. (B1)] 1/sl axial coordinate mL liquid molar flow rate mol/sn number of components of mixtureNi molar flux of component i mol/(m2s)Q heat flux W/m2

R total component reaction rate mol/m3sR column vector with elements Ri mol/m3s� gas constant 8.3144 J /(mol K)Re Reynolds numberSc Schmidt numberSh Sherwood numbert time sT temperature KuL liquid velocity m/sU length-specific molar holdup mol/m

�HR0

D


w terminal velocity m/sxi first fluid-phase (liquid) mole fraction mol/mol

of component ix column vector with elements xi mol/molyi second fluid-phase (gas, vapor, or liquid) mol/mol

mole fraction of component iz film coordinate mzi ionic charge of component i

Greek Letters

� film thickness m� dimensionless film coordinate�� forward-reaction constant m3/2/(mol1/2s)�r backward-reaction constant s�1

� thermal conductivity W/(m K)� chemical potential J/mol�c dynamic viscosity of continuous phase Pa � s�d dynamic viscosity of dispersed phase Pa � s� volumetric holdup m3/m3

electrical potential V

Subscripts

G gas or second fluid phasei, j component/reaction indicesL liquid phaset mixture property

Superscripts

B bulk phaseI phase interface

Abbreviations

ADM axial dispersion modelCD catalytic distillationPDE piston flow model with axial

dispersion and mass exchangeRA reactive absorptionRD reactive distillationRE reactive extraction

di(2-ethylhexyl) phosphoric acidRSP reactive separation processRH


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APPENDIX A. A DETAILED DESCRIPTION OF

RATE-BASED MODELING

A.1. Balance Equations

The mass balance equations of the traditional multicomponent rate-based model(see, e.g., Refs. 57 and 58) are written separately for each phase. In order to givea common description to all three considered RSPs (where it is possible, ofcourse) we will use the notion of two contacting fluid phases. The first one isalways the liquid phase, whereas the second fluid phase represents the gas phasefor RA, the vapor phase for RD and the liquid phase for RE. Considering homo-geneous chemical reactions taking place in the fluid phases, the steady-state bal-ance equations should include the reaction source terms:

(A1)

(A2)

If chemical reactions take place in the (first) liquid phase only (this is valid for most of RD processes), the phase balances for the second fluid phasesimplify to

(A3)

The bulk-phase balances are completed by the summation equation for theliquid and second fluid bulk mole fractions:

(A4)xi

B

i

n

�

�

1

1∑

0 1� � �d

dlGy N a A i ni

BGiB I

s( ) , . . . ,

0 1� � � �d

dlGy N a R A i ni

BGiB I

GiB

G s( ) ( ) , . . . ,�

0 1�� d

dlLx N a R A i ni

BLiB I

LiB

L s( ) ( ) , . . . ,�


(A5)

The volumetric liquid holdup, �L, depends on the gas/vapor and liquidflows and is calculated via empirical correlations (e.g., Ref. 65). For the determi-nation of axial temperature profiles, differential energy balances are formulated,including the product of the liquid molar holdup and the specific enthalpy asenergy capacity. The energy balances written for continuous systems are as follows:

(A6)

(A7)

In the dynamic rate-based stage model, molar holdup terms have to be con-sidered in the mass balance equations, whereas the changes in both the specificmolar component holdup and the total molar holdup are taken into account. Forthe liquid phase, these equations are as follows:

(A8)

(A9)

The gas/vapor holdup can often be neglected due to the low gas-phase density, and the component balance equation reduces to Eq. (A2) (see also Ref. 139).

A.2. Mass Transfer and Reaction Coupling in Fluid Films

The component fluxes entering into Eqs. (A1)–(A3) are determined based onthe mass transport in the film region. Because the key assumptions of the filmmodel result in the one-dimensional mass transport normal to the interface, thedifferential component balance equations including simultaneous mass transferand reaction in the film are as follows:

(A10)

Equations (A10), which are generally valid for both liquid and second fluidphases, represent nothing but differential mass balances for the film region, with the

dN

dzR i nLi

Li� � �0 1, . . . ,

NiB

U x U x c A i nLi iB

Lt iB

L Lt s� � �( ) , . . . ,� 1

∂∂

∂∂t

Ul

Lx N a R A i nLi iB

LiB I

LiB

L s�� ( ) ( ) , . . . ,� 1

0 0� � �d

dlGh Q a R H AG

BGB I

GB

G RG s( ) ( )� �

0 0�� d

dlLh Q a R H AL

BLB I

LB

L RL s( ) ( )� �

yiB

i

n

�

�

1

1∑


account of the source term due to the reaction. To link these balances to processvariables like component concentrations, some additional relationships, often calledconstitutive relations (see Ref. 57), are necessary. For the component fluxes Ni,these constitutive relations result from the multicomponent diffusion description[Eqs. (1) and (2)] and, for the source terms, from the reaction kinetics description.

The latter strongly depends on the specific reaction mechanism, the stoi-chiometry, and the presence or absence of parallel reaction schemes (69). The rateexpressions for Ri usually represent nonlinear dependences on the mixture com-position and temperature. Specifically for the coupled reaction–mass transferproblems, such as Eqs. (A10), it is always essential as to whether or not the reac-tion rate is comparable to that of diffusion (68,77). Equations (A10) should becompleted by the boundary conditions relevant to the film model. These conditionsspecify the values of the mixture composition at both film boundaries. For exam-ple, for the liquid phase:

(A11)

Combining Eqs. (A10) with the boundary conditions (A11) written in vec-tor form and using constitutive relations such as Eqs. (1) and (2), we obtain a vector-type boundary-value problem, which permits the component concentrationprofiles to be obtained as functions of the film coordinate. These concentrationprofiles, in turn, allow one to determine the component fluxes. Thus the boundary-value problem describing the film phenomena has to be solved in conjunctionwith all other model equations.

The composition boundary values entering into Eqs. (A11) represent exter-nal values for Eqs. (A10). With some further assumptions concerning the diffu-sion and reaction terms, this allows an analytical solution of the boundary-valueproblem [Eqs. (A10) and (A11)] in a closed matrix form (see Refs. 58 and 135).On the other hand, the boundary values need to be determined from the total sys-tem of equations describing the process. The bulk values in both phases are foundfrom the balance relations, Eqs. (A1) and (A2). The interfacial liquid-phase con-centrations are related to the relevant concentrations of the second fluid phase,

, by the thermodynamic equilibrium relationships and by the continuity condi-tion for the molar fluxes at the interface (57,135).

Due to the chemical conversion in the liquid film, the molar fluxes at theinterface and at the boundary between the film and the bulk of the phase differ.The system of equations is completed by the conservation equations for the massand energy fluxes at the phase interface and the necessary linking conditionsbetween the bulk and film phases (see Refs. 57, 59, and 84).

Generally all these considerations are also valid for the second fluid filmphase, provided that reactions occur there (135). Both analytical and numericalsolutions of the coupled diffusion-reaction film problem are analyzed at fulllength in Ref. 167; their particular applications are considered in Section 3.

yiI

xiI

x z x x z x i ni iI

i L iB( ) , ( ) , . . . ,� � � � �0 1�


A.3. Nonideal Flow Behavior in Catalytic Column Internals

The mass balances [Eqs. (A1) and (A2)] assume plug-flow behavior for both thegas/vapor and liquid phases. However, real flow behavior is much more complexand constitutes a fundamental issue in multiphase reactor design. It has a stronginfluence on the reactor performance, for example, due to back-mixing of bothphases, which is responsible for significant effects on the reaction rates and prod-uct selectivity. Possible development of stagnant zones results in secondary un-desired reactions. To ensure an optimum model development for CD processes,experimental studies on the nonideal flow behavior in the catalytic packing MULTIPAK® are performed (168).

The experimental results confirm that the fluid flow in MULTIPAK® devi-ates from plug-flow behavior (Figure 31). Calculated axial dispersion coefficientsare about 10�4–10�2 m2/s, which are several orders of magnitude larger than thatfor molecular diffusion (Figure 32). Therefore, in the investigated operatingrange, nonideal mixing effects are caused by hydrodynamic rather than molecu-lar diffusion effects. Calculated Bodenstein numbers are one order of magnitude

FIGURE 31 Comparison between the experimental RTD curve for the cata-lytic packing MULTIPAK® (dC � 0.1 m), the ADM model, and the PDE model.


smaller than those for fixed-bed reactors, which may be caused by two effects: theoccurrence of stagnant zones in the catalyst layer, and liquid bypassing due to thehybrid structure of the catalytic packing (168).

The rate-based models suggested up to now do not take liquid back-mixinginto consideration. The only exception is the nonequilibrium-cell model for multicomponent reactive distillation in tray columns presented in Ref. 169. In thiswork a single distillation tray is treated by a series of cells along the vapor andliquid flow paths, whereas each cell is described by the two-film model (seeSection 2.3). Using different numbers of cells in both flow paths allows one todescribe various flow patterns. However, a consistent experimental determinationof necessary model parameters (e.g., cell film thickness) appears difficult, where-as the complex iterative character of the calculation procedure in the dynamiccase limits the applicability of the nonequilibrium cell model.

A far more promising approach is represented by the so-called differentialmodels, such as the axial dispersion model (ADM) (170) as well as the piston-flow model with axial dispersion and mass exchange (PDE) (171). Experimentalstudies (168) show that the ADM gives an appropriate description of the nonidealflow behavior of the liquid phase in catalytic packings (see Figure 31). Considering

FIGURE 32 Axial dispersion coefficients for the catalytic packing MULTIPAK®(dC � 0.1 m ) , calculated based on the ADM model.


the nonideal flow behavior via the ADM, the dynamic mass balances [cf. Eqs.(A8)] take the following form:

(A12)

A thorough investigation of the influence of the flow nonideality in catalyticpackings on the dynamic process behavior of specific CD processes is an objec-tive of some current studies (172).

Equation (A12) is widely used in RE, but it does not account for the spe-cific interactions of the dispersed phase. In this respect current research is focusedon drop population balance models, which account for the different rising velo-cities of the different-size droplets and their interactions, such as droplet breakupand coalescence (173–180).

APPENDIX B. MODELING PECULIARITIES AND

MODEL PARAMETERS FOR THE CASE STUDIES

B.1. Absorption of NOx

In terms of the concentration vector, Eq. (A10) is a nonlinear differential equa-tion of the second order. The boundary-value problem [Eqs. (A10) and (A11)] isusually solved numerically. However, it is also possible to linearize the reactionterm using the method suggested in Ref. 181:

(B1)

Equation (B1) provides a satisfactory representation for many processes over theentire reaction range and is a good linear approximation for most systems in a suf-ficiently small range (see, e.g., Refs. 68 and 182–184). Equation (B1) has gainedwidespread acceptance in various chemical and reactor engineering areas (185) andis recommended for use in the modeling of reactive separation operations (59,184).

The approximation of Eq. (B1) allows one to reduce Eqs. (A10) and (A11)to a linearized boundary-value problem (183,184,186). The latter can then besolved analytically and yields a compact matrix-form solution for the concentra-tion profiles in the film region [58]. Such a solution gives simple analyticalexpressions for the component fluxes with regard to the homogeneous reaction inthe fluid films (see Ref. 135), which can be of particular value when large indus-trial reactive separation units are considered and designed.

The methods of determination of the reaction matrix [K ] are considered inRefs. 167, 181, 183, 184 and 186. Another important matrix parameter enteringinto the linearized film mass transport equation is the multicomponent diffusionmatrix [D]. The latter results from the transformation of the Maxwell–Stefan Eqs. (1)to the form of the generalized Fick’s law (83). Matrix [D] is generally a function of

R x≅ �[ ]K

∂∂

∂∂

∂∂t

UDu l

Lxl

Lx N a R A i nLiL

iB

iB

LiB I

LiB

L s� � � � �ax2

2 1( ) ( ) ( ) , . . . ,�


the mixture composition and it is assumed constant along the diffusion path (83).The direct expressions for the elements of the diffusion matrix [D] can be found,for example, in Ref. 57.

The linearization of the initial film mass transport equation and its analyticalsolution were applied to simulate the industrial NOx absorption process considered.

In order to calculate the multicomponent diffusion matrices [D], the binarydiffusivities in both phases should be known. The film thickness representing animportant model parameter is estimated via the mass transfer coefficients (57,83).The binary diffusivities and mass transfer coefficients were calculated from thecorrelations summarized in Table 3.

The correlations of Billet (66) and Onda et al. (187) are valid for variousmixtures and packings and cover both absorption and distillation processes. Thecorrelation of Kolev (133) is obtained for RA and certain random packings. Ingeneral, the mass transfer coefficient correlations need to be compared to oneanother and validated using experimental data. This shows, in particular, the waythe mass transfer correlations influence the concentration profiles of the compo-nents and other relevant process characteristics.

Nitric acid is a strong electrolyte. Therefore, the solubilities of nitrogenoxides in water given in Ref. 191 and based on Henry’s law are utilized and fur-ther corrected by using the method of van Krevelen and Hoftijzer (77) for elec-trolyte solutions. The chemical equilibrium is calculated in terms of liquid-phaseactivities. The local composition model of Engels (192), based on the UNIQUACmodel, is used for the calculation of vapor pressures and activity coefficients ofwater and nitric acid. Multicomponent diffusion coefficients in the liquid phaseare corrected for the nonideality, as suggested in Ref. 57.

TABLE 3 Binary Diffusion Coefficients and Mass TransferCoefficients

Mass transfer Binary diffusion coefficient

Phase coefficient correlation

Gas Ref. 188 Ref. 187Ref. 66Wehmeier (see

Ref. 134)Liquid Ref. 189 Ref. 187

Ref. 66Ref. 133Ref. 190


B.2. Coke Gas Purification

In Ref. 139 a purely numerical approach to the solution of the considered com-plex RA problem was suggested. The liquid film is treated as an additional bal-ance region, in which reaction and mass transfer occur simultaneously. Therefore,the reactions are considered both in the liquid-bulk-phase mass balances, Eq. (A1),and in the differential balances for the liquid film, Eq. (A10).

To be able to describe the presence of electrolytes in the system, the elec-trical driving force also needs to be taken into account (57). Therefore, the gradi-ent of the electrical potential � is introduced into the generalized driving forcedi [cf. Eq. (2)]:

(B2)

In dilute electrolyte systems, the diffusional interactions can usually beneglected, and the generalized Maxwell–Stefan equations are reduced to theNernst–Planck equations (B3):

(B3)

where n is the solvent index. The consideration of the electrical potential requiresan additional condition, the electroneutrality, which has to be met in each pointof the liquid phase:

(B4)

Thermodynamic nonidealities are considered both in the transport equations(A10) and in the equilibrium relationships at the phase interface. Because elec-trolytes are present in the system, the liquid-phase diffusion coefficients should becorrected to account for the specific transport properties of electrolyte solutions.

The thermodynamic equilibrium at the gas–liquid interface is described asfollows:

(B5)

where the distribution coefficient Ki comprises fugacities in both phases andactivity coefficients in the liquid phase. For the system considered, the values ofKi, Eq. (B5), are determined from the electrolyte NRTL method (70,71).

The liquid-phase diffusion coefficients are found with the Nernst–Hartleyequation (193), which describes the transport properties in weak electrolyte systems. The gas-phase diffusion coefficients are estimated according to the

y K x i niI

i iI� �1, . . . ,

x zi i

i

n

�

�

1

0∑

Nc D dx

dx z

F

T

d

dx N i nLi

Lt Li

L

ii i i Ln��

, , . . . ,eff

� �

�ℜ

1 1

dx

Tx z

F

T

d

di ni

i

L

ii i

L

� � �ℜ

∂∂ ℜ

1 11

�

�

� �

�, . . . ,


Chapman–Enskog–Wilke–Lee model (72). The correlations for the mass transfercoefficients are taken from Ref. 194.

B.3. Methyl Acetate System, Batch Distillation

The rate-based models usually use the two-film theory and comprise the materi-al and energy balances of a differential element of the two-phase volume in thepacking (148). The classical two-film model shown in Figure 13 is extended hereto consider the catalyst phase (Figure 33). A pseudo-homogeneous approach ischosen for the catalyzed reaction (see also Section 2.1), and the correspondingoverall reaction kinetics is determined by fixed-bed experiments (34). Thismacroscopic kinetics includes the influence of the liquid distribution and masstransfer resistances at the liquid–solid interface as well as diffusional transportphenomena inside the porous catalyst.

For the determination of conversion corresponding to the average residencetime, the reaction kinetics is integrated into the mass balances, and the liquid

FIGURE 33 Film model for a differential packing segment with heteroge-neous catalyst.


holdup, as the accumulation term, is accounted for simultaneously, as in Eqs. (A8)and (A9). Because of low vapor-phase density, the vapor holdup is neglected, andthe vapor-phase-component balance equation reduces to Eq. (A2).

Ranzi et al. [195] found that the full energy balances, including the accu-mulation term, have to be considered in order to predict correct dynamic processbehavior. Therefore, the differential dynamic energy balance for the liquid phaseis applied as follows:

(B6)

where

(B7)

Similar to the mass balance equation, the vapor-phase energy balance simplifiesto Eq. (A7).

Experimental studies were carried out to derive correlations for mass trans-fer coefficients, reaction kinetics, liquid holdup, and pressure drop for the pack-ing MULTIPAK® (35). Suitable correlations for ROMBOPAK 6M® are taken fromRefs. 90 and 196. The nonideal thermodynamic behavior of the investigated mul-ticomponent system was described by the NRTL model for activity coefficientsconcerning nonidealities caused by the dimerisation (see Ref. 72).

Binary diffusion coefficients for the vapor phase and for the liquid phasewere estimated via the method proposed by Fuller et al. and Tyn and Calus,respectively (see Ref. 72). Physical properties such as densities, viscosities, andthermal conductivities were calculated from the methods given in Ref. 72. Heatlosses through the column wall were measured at pilot scale.

B.4. Methyl Acetate System, Steady-State Distillation

The model is based on the film theory and comprises the material and energy bal-ances of a differential element of the two-phase volume in the packing. Each elementconsists of an ideally mixed vapor and liquid bulk phase and a vapor film regionadjacent to the interface, as shown in Figure 33. A first guess of the bulk phasecompositions and temperatures was provided by the solution of an equilibrium-stage model without reactions, as suggested in Ref. 198. The catalyzed reactionis described by the quasi-homogeneous approach of Ref. 197, since the concen-tration of acid sites has been determined as aCat � 4.7 molH�/gCat for dry LewatitK2621, which is close to the data of Ref. 197 given for Amberlyst 15.

E h c AL LB

L Lt s� ( )�

∂∂

∂∂

∑t

El

Lh Q a R H A

Q T T N h i n

L LB

LB I

LB

L RL s

LB L

LLB I

Li

i

Li

��

��

�

( ) ( )

( ) , . . . ,

�

�

�

� 0

1

2

1


In order to determine the model parameters, several experiments were per-formed at laboratory scale. Pressure drop experiments were carried out in glasscolumns, with a total packing height of 1 m at ambient pressure. Air/water was usedas a test system, with a circulating liquid phase set at a constant temperature of 20�C.

The experimental data cover a wide range of possible column loads. The gasload for the column with 100-mm diameter was restricted to 1.7 Pa0.5. Therefore,the liquid load was increased to higher values to reach the flooding region of thecatalytic packing. Two different flow regimes similar to those of conventionalstructured packings can be observed. Flooding of the packing can be observed ata pressure drop above 103 Pa/m. The possible column loads for MULTIPAK® arevery similar to those reported in Refs. 199 and 200 for KATAPAK-S.

The number of theoretical stages per meter of the catalytic packing wasdetermined as a function of the gas capacity factor. For the whole range of columnloads, the separation efficiency is at least four theoretical stages per meter. Moritzand Hasse (200) determined an NTSM value of 3 for the laboratory-scale KATAPAK-S. The separation efficiency remains constant for a wide loadingrange of the packing. For lower column loads, the NTSM value increases to 6, aphenomenon already reported in Ref. 90 for the conventional structured packingMontzpak A3-500. A simple transfer-unit concept assuming all mass transferresistance in the vapor phase was used to determine the vapor-side mass transfercoefficients (201). The mass transfer correlation

(B8)

represents all experimental data with an accuracy of 13%. A comparison withexperimental data is shown in Figure 34.

B.5. Synthesis of Methyl Tertiary Butyl Ether

The mathematical description considered in Section 2.3 and Appendix A wasused as a modeling basis for the specially developed completely rate-based simulator DESIGNER (155). This tool consists of several blocks, includingmodel libraries for physical properties, mass and heat transfer, reaction kinetics,and equilibrium, as well as a specific hybrid solver and thermodynamic package.

DESIGNER also contains different hydrodynamic models (e.g., completelymixed liquid–completely mixed vapor, completely mixed liquid–vapor plug flow,mixed pool model, eddy diffusion model) and a model library of hydrodynamiccorrelations for the mass transfer coefficients, interfacial area, pressure drop,holdup, weeping, and entrainment that cover a number of different column inter-nals and flow conditions.

In DESIGNER, different ways of taking account of heterogeneous reactionkinetics are available, depending on the reaction rate and character. One further pos-sibility is to use a detailed model for the heterogeneous catalyst mass transfer

Sh Re Sc1/3G G� � �0 009 0 92. .


efficiency based on the approach of Ref. 91. When applying this type of kineticmodel, the intrinsic kinetics data are needed. Another way is the pseudo-homoge-neous approach, with effective kinetic expressions, by which the kinetics descriptionis introduced as source terms into the balance equations [cf. Eqs. (A1) and (A2)].

For the system considered here, the reaction is slow as compared to themass transfer rate. For this reason the pseudo-homogeneous approach is used, thereaction being accounted for in the liquid bulk only.

Basically, DESIGNER can use different physical property packages that areeasy to interchange with commercial flowsheet simulators. For the case consid-ered, the vapor–liquid equilibrium description is based on the UNIQUAC model.The liquid-phase binary diffusivities are determined using the method of Tyn andCalus (see Ref. 72) for the diluted mixtures, corrected by the Vignes equation(57), to account for finite concentrations. The vapor-phase diffusion coefficientsare assumed constant. The reaction kinetics parameters taken from Ref. 202 areimplemented directly in the DESIGNER code.

B.6. Reactive Extraction of Zinc

In conventional RE processes, the diffusive resistance is concentrated mainly insidethe droplet, whereas the aqueous-side resistance can be neglected. This has beenproven in Ref. 203 using the laser-induced-fluorescence (LIF) technique. Usuallythe organic phase is more viscous and the diffusion coefficients of the organic com-plexes are larger than those at the aqueous side, which supports this finding.

The mass transfer within a rigid droplet is determined by the Maxwell–Stefan diffusion. The appropriate diffusion coefficients experimentally determined

FIGURE 34 Sherwood number correlation for MULTIPAK®.


for this zinc extraction system in Ref. 204 are presented in Table 4. With nonrigiddroplets, a mass transfer enhancement by internal convection has to be consid-ered. However, with industrial feed solutions there are always impurities presentthat may dampen the mass transfer (8). In contrast, there also might be a masstransfer increase due to Marangoni effects (205,206). Therefore, for a final designof a column, mass transfer measurements are recommended.

The macrokinetics of zinc extraction is discussed in detail in Ref. 8. It is acombination of a reaction kinetics term (55) with the Maxwell–Stefan (54) oreddy diffusion (56). The rate law is as follows:

(B9)

where C1, C2, ��, and �r, are the estimated kinetics parameter (see EFCE test systems discussed earlier).

The rate constant for the backward reaction, �r, can be replaced by the ther-modynamic equilibrium constant:

(B10)

The species concentrations are formulated in activities using the Pitzermodel (207) for the aqueous phase and the Hildebrand–Scott solubility parame-ter (208) for the organic phase.

The effective diffusion coefficient is calculated according to the model ofRef. 209, which accounts for interfacial instabilities. This model includes aHandlos–Baron-like correlation (210) and one adjustable parameter, CIP :

(B11)Dw d

C

p

IPd

c

eff �

�

�

�2048 1�

�

K v

r

eq ��

�

� ��

� ��

�

� � �

�

d

dtv R H R RH

R H C

R H

C R H

r[ ] [ ] [ ] [ ] [ ( )]

[ ] [ ]

[ ]

[ ]

.

.

Zn Zn H Zn

H

2 2� �2 2

2

2

2 2

2

2 2

2 2 2

21 5

1 51

TABLE 4 Ternary Fick Diffusion Coefficients for the System(2), and Diluent (3) at 298.15 K .

Isododecane (3) Toluene (3)

D11 1.01 � 0.30 � 10�10 m2/s 3.60 � 1.10 � 10�10 m2/sD12 �0.11 � 0.10 � 10�10 m2/s �0.35 � 0.42 � 10�10 m2/sD21 4.11 � 0.80 � 10�10 m2/s 0.16 � 3.50 � 10�10 m2/sD22 2.79 � 0.30 � 10�10 m2/s 7.68 � 0.21 � 10�10 m2/s

Zn (RH) (1),RH2R


10

Multifunctional Reactors: Integration of Reaction and Heat Transfer

David W. Agar

University of Dortmund, Dortmund, Germany

1. INTRODUCTION

Reaction engineers devote a lot of time and ingenuity to enhancing reactor per-formance by attempting to follow an optimal trajectory for the reaction system (1)and by overcoming the limitations imposed by the accompanying heat and masstransfer processes. These objectives are often interrelated: Achieving the concen-tration and temperature profiles required to maximize conversion rates and mini-mize by-product formation, for example, dictates the absence of gradients thatmight lead to local deviations from these values.

Process intensification can be considered to be the use of measures to increasethe volume-specific rates of reaction, heat transfer, and mass transfer and thus toenable the chemical system or catalyst to realize its full potential (2). Catalysisitself is an example of process intensification in its broadest sense. The use ofspecial reaction media, such as ionic liquids or supercritical fluids, high-densityenergy sources, such as microwaves or ultrasonics, the exploitation of centrifugalfields, the use of microstructured reactors with very high specific surface areas,and the periodic reactor operation all fall under this definition of process intensi-fication, and the list given is by no means exhaustive.

Reactor performance is dictated by the inputs, by the contacting pattern,i.e., how and when individual elements pass through the reactor and contact one


another and how long they retain their identity, and by the cumulative kinetics andthermodynamics to which elements are exposed along their reaction trajectory (3).Identifying the most suitable reactor configuration, e.g., an ideal plug flow, is awell-established procedure in chemical reaction engineering. Less appreciatedare the full range of possibilities available in manipulating the local rates of reac-tion by imposing favorable temperature, concentration, and activity profiles alonga catalytic reactor. While the first option has received extensive attention, the lasttwo have been somewhat neglected, although they often offer a more selectiveintervention in the progress of the reaction and complement the more commontailoring of the temperature profile.

In general, temperature and concentration profiles may be externally influ-enced by convective or recuperative and, less commonly, regenerative or reactivestrategies (Figure 1). The “convective” addition or withdrawal of side streamsalong the reactor represents a simple technique for temperature control or forimproving selectivity by restricting availability of one reactant. In a recuperativeprocess, examples of which are provided by the cooled tubular reactor and themembrane reactor, heat or material is exchanged, in the latter case usually in a

FIGURE 1 Basic strategies for manipulating temperature and concentrationprofiles in chemical reactors.


selective manner, via diffusive transport processes, such as heat conduction, withspatially distinct external sources and sinks. Regeneration exploits the storage ofheat and mass, by adsorption, for example, on reactor internals—usually a fixedbed—to yield beneficial temperature and concentration profiles that could notarise under steady-state operation. Regenerative processes are inherently unsteadystate in nature and entail a chronologically separate recharging of the storagecapacity drawn down during the reaction phase. The use of a supplementary reac-tion to supply or consume heat and/or reactants or products, as encountered inoxydehydrogenations, is a method requiring considerable finesse and almost per-fect compatibility between the individual reactions.

2. CONVECTIVE HEAT TRANSFER

Each of these approaches has its pros and cons, as can be illustrated for convec-tive cooling in a cold-shot reactor, employed in ammonia synthesis, for example.Plotting conversion against temperature for a reversible exothermic reaction(Figure 2a) shows that intermittent cooling by the discrete introduction of coldfeed along the reactor enables one to circumvent the equilibrium limitationimposed on adiabatic operation but also that the cooling effect desired is accom-panied by a less welcome loss in conversion. Furthermore, as the slope of thecooling line approaches that of the adiabatic reaction path, the efficacy of coolingand its benefits diminish. A possible solution to this problem of limited coolingcapacity is to employ an inert side stream in place of feedstock (Figure 2b) ascoolant (3). However, this would result in dilution of the reactor product streamand complicate the downstream processing steps. A process intensification tech-nique to overcome the difficulties indicated is, for gaseous systems, to inject inertliquid between adiabatic reactor stages. Exploiting the heat of evaporation meansthat much lower quantities of inert are needed than with a gaseous coolant.

3. RECUPERATIVE HEAT TRANSFER

The best-known recuperative reactor—the multitubular reactor used in the partialoxidation of hydrocarbons, for instance—is a ubiquitous piece of equipment inthe chemical processing industry (4). This should not blind one to the fact that itexhibits several serious shortcomings: The reactors are costly, and, despite the useof up to 25,000 tubes in a single reactor to provide a suitably extensive area forheat transfer (�100 m2/m3), one often observes large temperature excursionsfrom the desired temperature level—so-called hot spots—in both the axial andradial directions (Figure 3). These hot spots arise due to a bottleneck in the heatremoval process, arising from a combination of the locally accelerated reactionrates and poor heat transport through the catalyst bed (�100 W/m2K)—an orderof magnitude smaller than what the coolant side is capable of providing.


FIGURE 2 Interstage “convective” cooling of an exothermic equilibrium reaction through introduction of (a) cold-shotfeed by-pass and (b) cold-shot inert side-stream between adiabatic stages.


Hot spots usually dictate the attainable reactor performance in terms of conver-sion, selectivity, safety issues, operating lifetimes, and materials of construction,since most of the critical processes are confined to the immediate vicinity of thetemperature maximum.

3.1. Catalyst Dilution

The elimination of reactor hot spots has also attracted considerable interest overthe years. Because a panacea remains elusive, a variety of countermeasures havebeen adopted reflecting different compromises between the demands of the reac-tion, heat removal, and pressure drop (Figure 4). Perhaps the simplest procedurefor avoiding the formation of pronounced hot spots is to dilute the catalyst at theendangered locations (5). In this manner, the reaction is spread more evenly overthe length of the reactor, and a better harmonization between the heat productionby the reaction and the heat removal via the reactor wall is realized (Figure 5).The resultant reactors are, of course, larger, and the inclusion of inert packingleads to increased pressure drops. But the technique is reliable and involves noadditional developments. Interestingly, the precise activity profile employed isseldom decisive for the improvement in performance; i.e., simple arrangementssuffice.

FIGURE 3 Schematic of temperature profiles and hot spot formation in a multi-tubular reactor.


3.2. Linde™ Isothermal Reactor

An alternative solution is to pack still more heat exchange surface into the cata-lyst bed and to try to augment the heat transfer coefficients modestly by inducinga greater degree of turbulence for the gas flow over the cooling surface. The so-called Linde™ isothermal reactor (6) inverts the situation in the conventionalmultitubular reactor by operating with the coolant—normally pressurized boilingwater—within the tubes and the catalyst outside (Figure 6). The tubes are also nolonger parallel but assume a convoluted spiral geometry that raises the specificsurface area and increases heat transfer coefficients by an order of 50%. Use ofthe reactor is confined to temperatures below 550 K, since the pressure requiredfor evaporative cooling above this value becomes exorbitant, and molten saltcooling is not an option due to the high pressure drop in the cooling circuit. Inaddition, the removal of the catalyst from such reactors, e.g., for regeneration orreplacement, may present problems owing to “arch” formation between coolingtubes. In economic terms, the greater compactness of the Linde™ isothermal reac-tor must be set against its increased construction complexity.

3.3. Fluidized Beds

An extremely effective means of enhancing heat removal from a reactor is to makeuse of fluidized-bed technology (3). Heat transfer coefficients for gaseous systemsare increased to values of around 600 W/m2K or more by virtue of the very efficientconvective-regenerative particle transport mechanism of heat transfer. Further

FIGURE 4 Various measures for the harmonization of reaction, heat trans-port, and pressure drop in chemical reactors benchmarked against a multi-tubular reactor.


advantages include lower resistances to mass transfer (particle size � 100 �m) andfacile catalyst exchange for regeneration purposes if necessary. Less favorableaspects of fluidized beds include the high degree of back-mixing that occurs, the lim-ited range of hydrodynamic loading, and the uncertainty involved in scale-up.Moreover, the catalyst must be very mechanically resilient, and adapting catalystsfor fluidized-bed operation can be a time-consuming and frustrating exercise.

3.4. Catalytic Microreactors

A measure that has been the subject of extensive publication is that of microre-actors with catalytically coated walls (7,8). A microreactor has been defined as:“a miniaturized reaction vessel with characteristic dimensions in the range10–300 �m which has been fabricated using state-of-the-art high-precision engi-neering” (7). Such reactors exhibit well-defined laminar-flow patterns and permitfacile scale-up by simple “numbering up” of the number of channels and flexible

FIGURE 5 Hot spot reduction using spatially structured catalyst dilution.Selectivity profiles for base case with constant coolant temperature, cocur-rent coolant strategy and axially profiled catalyst activity strategy. The basecase chosen in the calculations is the one in which the coolant temperatureis constant and the activity profile along the length of the reactor is at thelevel unity. (From Ref. 5.)


arrangements of individual modules. Furthermore, they have extremely lowholdups and, as a consequence, comparatively short dynamic response times.Most importantly, they enable one to operate under isothermal conditions witheven the most exothermic reactions. The list of applications for which microre-actors are suitable includes the reaction engineering for fuel cell hydrogen pro-duction, the synthesis of hazardous chemicals, high-throughput screening forchemicals and catalysts, as instruments for obtaining insights into chemical reac-tion mechanisms, and as components of “intelligent” chemical sensors.

The excellent heat transfer characteristics of microreactors result not only from their considerably enhanced specific heat exchange surface areas of30,000 m2/m3—a value roughly 300 times higher than that in a conventional mul-titubular reactor—but also from the rapid lateral heat transfer across the channels.

FIGURE 6 Linde™ isothermal reactor for intensified cooling of strongly exo-thermic reactions. (From Ref. 6.)


For the small microchannel dimensions involved (Figure 7), conductive heattransfer plays an important role, augmenting the heat transfer coefficients by afactor of 5 or more to values in the vicinity of 700 W/m2K. This phenomenon isalso responsible for good transverse mixing, counteracting the negative effects oflaminar-flow profiles.

Less advantageous are the high unit production costs for microreactors andthe complexities involved in their manufacture. The basic material of constructionis usually metal or plastic, which has restricted heterogeneous catalytic applica-tions to metallic catalysts or often unsatisfactory metallic-catalytic composites.

FIGURE 7 Structure and typical dimensions of a microreactor. (From Ref. 7.)


Despite the high specific surface areas, the amount of accessible catalyst remainslow due to the limited thickness of the porous catalytic layer dictated by consid-erations such as the adhesion to the substrate. The susceptibility of the fine chan-nels to blockage with solid impurities or deposits formed in the reaction, togetherwith the problems of integrating connections with the external macroenviron-ments and ensuring uniform gas distribution between the individual channels, aprerequisite for numbering up, represent further questions that have to beresolved for the industrial application of microreactors to become practicable.

3.4.1. Catalytic Millireactors

The drawbacks associated with this particular technique of process intensificationmay be ameliorated by analyzing the characteristic length scales for the compo-nent processes. The chemical activity of a heterogeneous catalyst, for example, isexpressed in its nanostructure, i.e., the pore diameter of 10–100 nm reflected inthe typical specific surface area of a catalyst, say, 100 m2/g. This scale providesan adequate number of active sites for the reaction to proceed without imposingunnecessary limitations on the underlying chemical kinetics. Catalyst systemslacking this nanostructure will seldom provide sufficient activity and do not per-mit the catalyst chemistry to realize its full potential. Even in a microreactor, geo-metric surface areas of catalytic layers fall short of the internal specific poresurface area by a factor of almost 10,000. The diffusive transport through the poresystem to and from the active sites dictates the maximal thickness of porous cat-alytic layers. For the diffusion coefficients in gas-phase systems, the conventionalanalysis based on the Thiele or Weisz modulus suggests that diffusive mass trans-fer is sufficient to maintain the reactant flow to the active catalytic sites over dis-tances of about 1 mm. The corresponding value for liquid systems is two ordersof magnitude lower. The validity of this result is confirmed by observing the typ-ical dimensions of industrial catalysts employed for gas- and liquid-phase reac-tions, respectively.

Considering a typical intrinsic rate of catalytic reaction in an intermediatetemperature range of 1 mmol/kg cat�s and a strongly exothermic reaction enthalpyindicates that the rates of heat generation one needs to master can be of the orderof 500 kW/m3 of catalyst. For tolerable temperature gradients (�2 K) and char-acteristic thermal conductivity values for porous catalysts (1 W/mK) one obtainsthe result that specific heat exchange surface areas of 1000 m2/m3 should be suf-ficient to remove the heat of reaction. This analysis thus implies that while the traditional multitubular reactor with its tube diameters of several centimeters maywell be overtaxed, the microreactor offers an unnecessarily excessive specific surface area, way beyond the demands actually being imposed by the chemistryof the catalytic system. The resulting insight is that, in terms of the heat transferdemands, a reduction of the characteristic dimensions down to the millimeterscale yields an entirely adequate performance (Figure 8) (9).


3.4.2. Ceramic Catalytic Millireactors

Effective process intensification requires that all superfluous bottlenecks beremoved, but only to the extent at which a controlling step—usually the intrinsiccatalytic kinetics—can no longer be accelerated further. The considerations pre-sented earlier suggest an optimal reactor structure to meet this criterion. Since theintrinsic catalytic activity may be expressed only when a suitable nanostructureprovides access to the maximum number of active sites and because the mass ofthe catalyst rather than its external surface area determines the total activity, allmicroreactor concepts based on bulk metal catalysts or on thin catalytic films pro-duced by CVD, sol-gel coating, or anodic oxidation on a metallic support arebasically unsuitable because they provide insufficient catalytic material. In addi-tion, the conflicting demands inherent in producing a catalytic layer with thenanostructure thickness and adhesion to the substrate desired often lead to unsat-isfactory compromises, and new composite structures must be developed for eachspecific catalyst system. Whereas the traditional approach to catalytic microreac-tors has been to coat a metallic micro-heat exchanger structure with catalyst, asuperior technique might be to produce a heat exchanger out of the proven cat-alytic material itself or something similar, i.e., to utilize a nanoporous ceramicsubstrate. For the small dimensions involved, the use of slightly more poorly conducting ceramic leads to no deterioration in performance in comparison tometallic walls and can even be advantageous for thermal efficiency by maintain-ing “axial” temperature gradients without conductive short-circuiting.

FIGURE 8 Analysis of the appropriate characteristic dimensions for specificheat transfer surface requirements in a chemical reactor exhibiting typicalreaction rates. Da � Damkoehler number, NTU � number of thermal trans-fer units, Nu � Nusselt number, a � thermal diffusivity. (From Ref. 9.)


The use of proven catalyst recipes would greatly curtail development times,and the absence of extraneous material avoids unwanted catalytic effects andenhances thermal stability. The fixation of catalyst on ceramic substrates such aswashcoats is a well-known, reliable, and relatively straightforward procedure.The fabrication of complex small-scale ceramic structures is, however, more awk-ward than for metals or plastics, and they exhibit relatively poor mechanicalstrength. Furthermore, the porous ceramic nanostructure must be sealed to pre-vent contact between the reaction medium and coolant.

In the past, the principles described have been implicitly recognized in sev-eral attempts to convert monolithic catalysts into catalytic heat exchangers. Whilethe use of millimeter dimensions and nanoporous ceramic supports meets the pri-mary criteria already mentioned, the parallel channel structure of monoliths is notideally tailored for heat exchanger applications, and complex header structuresare required to uniformly distribute and collect reaction medium and coolant toand from the individual channels (Figure 9). The unsatisfactory interface betweenthe “milli-” and “macroscale” has been a major weakness of such concepts.

FIGURE 9 Monolithic catalyst modified to serve as a heat exchanger.


3.4.3. Ceramic Catalytic-Plate Heat Exchanger

An existing item of chemical processing equipment—the plate heat exchanger—suggests a possible solution to this drawback. The integration of manifolds intothe plate stack arrangement means that one can dispense with special headers forfluid distribution. Furthermore, internal structures on the plates can be used toprovide intermittent mixing and induce secondary flows, thus overcoming theprinciple shortcomings of simple laminar-flow conditions. A variety of tech-niques, including an efficient countercurrent CVD method, have been proposedfor sealing porous ceramic structures to prevent the unwanted diffusive slipbetween reaction medium and coolant (10).

Although the fragility of ceramic structures probably precludes the appli-cation of evaporative cooling or molten salt coolants, the use of oil up to 300°Cand gas for higher temperatures is feasible, since the poor performance of gascoolant can, to some extent anyway, be compensated for using by a favorablearrangement of the coolant-side geometry to cut the pressure drop. The coolantpressure in both cases can be adapted to approximate that of the reaction medi-um, thus reducing to a minimum the mechanical stress on the ceramic plates dueto pressure differentials. One can also select the coolant gas to have good thermalproperties (heat capacity, conductivity) and to be inert, so slight leakages into thereaction medium can, if necessary, be tolerated. Heat could be withdrawn effec-tively from the gas coolant circuit using an external inert fluidized-bed or otherproven techniques.

The resulting piece of equipment can be referred to as a “catalytic plate heatexchanger,” with structure and scale similar to that of a conventional plate heatexchanger. The plates comprise porous ceramic of 1- to 2-mm thickness and sep-arate alternating reaction and cooling chambers (Figure 10). The detailed channelprofile desired may be introduced on to the plain support by means of an inter-mediate laser-engraved separating plate. The individuals plates are stacked andcemented with countercurrent, cocurrent, or crossflow cooling configurations, asspecified. More complex temperature profiles can also be attained by combiningindividual stack modules as required. Special attention must be paid to the uni-form distribution of reaction medium and coolant both within and between theplates, to ensure an approximation to plug-flow behavior despite the usuallyunavoidable laminar-flow conditions. Fortunately, such flows are amenable toprecise modeling using computational fluid dynamics, enabling one to developsuitable geometries. Nevertheless, the delusion that numbering-up is a trivial pro-cedure quickly dissipates when one is confronted with the complexity of guaran-teeing almost identical flow conditions in each of several ten thousand channels!

In addition to the harmonization of the underlying physical and chemicalprocesses, the catalytic-plate heat exchanger offers a cost-effective alternative toboth conventional multitubular reactors and catalytic microreactors for industrial


applications. Because it is based on proven catalyst compositions and state-of-the-artfabrication methods, development costs and times should be low, a feature enhancedby the nonspecific, generally applicable nature of the concept. The successfuldesign of such reactors can be ensured by the use of reliable modeling tools andthe flexibility available in the reactor architecture and operation. Finally, thefacile regeneration, recycling, or disposal of deactivated catalyst is an additionaladvantage over metallic catalyst composites.

4. REACTIVE-RECUPERATIVE HEAT TRANSFER

A second method of process intensification for recuperative reactors is to enhanceperformance by using a reactive coolant or heating medium, since the heat effectsassociated with reactions are usually much larger than those available with phasechanges or simpler heating and cooling procedures. The coupling of an exother-mic auxiliary heat source reaction with the desired endothermic reaction, or vice

FIGURE 10 Schematic layout of a catalytic plate heat exchanger.


versa, through a double catalytically coated wall has also been attempted in amodified monolithic catalyst (Figure 11) for the steam reforming reaction, withthe methane combustion reaction being used to supply the heat required (11). Thework done demonstrated that it is extremely difficult to localize the reactions andregulate the temperatures in the interests of performance—the system tends toassume a state dictated by the intrinsic kinetics and heat effects of the two reac-tions. With more “passive” coolants or heating media, the operator is in a betterposition to influence the reactor behavior and can manipulate temperature profilesto a greater extent.

5. REGENERATIVE HEAT TRANSFER

Regenerative reactors, that is to say, those exploiting heat storage on fixed beds,remain a somewhat neglected option in reaction engineering. Although the prin-ciple of heat regeneration had been utilized previously in the chemical industry,

FIGURE 11 Monolithic catalyst adapted for the thermal coupling of endo- andexothermic reactions. (From Ref. 11.)


for example, in a process for high-temperature thermal for nitrogen fixation (12),and is a common feature in related process industries, such as power generationand steel manufacture, its most recent renaissance was brought about the devel-opment (13) and refinement (14) of the reverse-flow reactor concept (Figure 12),which has now established itself firmly in the niche of oxidative waste-gas treat-ment. Additionally, regenerative heat transfer is an important feature of con-ventional reactor operation, for instance, in the occurrence of the transitional “wrong-way” behavior with temperature levels in excess of the adiabatic valuethat can arise when a fixed-bed reactor is shut down (15).

Regenerative heat exchange in chemical reactors offers clear benefits, suchas simplicity, robustness, low costs, and high efficiencies, against which must beset its inherently unsteady-state operation, the limited potential for an exact reg-ulation of temperature profiles, and the restriction of its use to gaseous reactionmedia.

5.1. Comparison of Regenerative with Convective,

Recuperative, and Reactive Heat Transfer

In the evaluation of the regenerative heat exchange option, it is instructive to con-sider the heat exchange techniques presently employed in the following chemicalprocesses: styrene synthesis, steam reforming, and hydrogen cyanide production(Table 1).

FIGURE 12 Reverse flow reactor concept.


TABLE 1 Heat Exchange Employed for High-Temperature Endothermic Reactions

Ethylbenzene Steam Hydrogen CyanideDehydrogenation Reforming Manufacture

C8H10 ↔ C8H8 � H2 CH4 � H2O ↔ CO � 3H2 CH4 � NH3 ↔ HCN � 3H2T-Profilingtechnique 600�C 900�C 1200�C

Convection Badger/Mobil “adiabatic”process

Recuperation BASF “isothermal” process conventional primary steam Degussa BMAreforming process

Regeneration ?

Reaction autothermal reforming Andrussov ammonoxidation(fuel cell applications) process


The moderately endothermic dehydrogenation of ethylbenzene to styrene iscarried out catalytically at temperatures of around 600°C. The dominant technol-ogy for meeting the heat requirements of this reaction involves the intermediateintroduction of superheated steam as a heating medium between adiabatic reac-tor stages, i.e., a convective process. Several additional features of the reactionfavor the convective approach: Steam can easily be removed from the products bya phase separation following condensation, steam acts as an inert dilutant enhanc-ing the equilibrium conversion, and steam also helps to maintain catalytic activity.An alternative recuperative technology in which combustion gases are used as aspatially segregated heat source in a multitubular suffers from the low heat trans-fer coefficients in gas–gas recuperative heat exchangers and entails a more costlyreactor construction.

The steam reforming of methane to synthesis gas, a strongly endothermicreaction carried out catalytically at around 900°C, primarily utilizes recuperativeheat transfer. The higher reaction temperature makes the convective supply ofheat problematic, particularly at the preferred operating pressures of about 25 bar.The heat required can also be supplied reactively by simultaneously carrying outthe exothermic partial oxidation of methane to carbon monoxide and hydrogen.The main impediment for this so-called “autothermal” reforming was previouslythe rapid deactivation of the catalysts used. In connection with research on the“on-board” generation of hydrogen for mobile fuel cell applications, novel noblemetal catalysts have been developed that maintain an adequate activity. Such cat-alysts are, however, almost certainly too costly for regular industrial purposes.

The prevalent manufacturing process for hydrogen cyanide—the Andrussovprocess—represents the successful industrial application of “reactive” heatexchange. In the catalytic ammonoxidation of methane at 1100°C, the actualendothermic synthesis reaction between methane and ammonia is thermal supported by the oxidation of the hydrogen formed, so no additional heat need be supplied (Figure 13). Reactive heat transfer of this sort is, of course, very efficient, in that it entails virtually no temperature gradients, due to the almostmolecular scale at which it occurs. On the other hand, the inclusion of an extrareactant—oxygen—gives rise to unwanted side reactions, such as the formationof carbon oxides, and encourages undesirable ammonia decomposition. Further-more, because it is uneconomic to use pure oxygen in place of air, the nitrogenintroduced into the process results in the need for much larger reactors and down-stream units to treat the diluted gas streams. The alternative recuperative BMAprocess offers much higher yields, due to the absence of side reactions, and higherproduct concentrations with lower flows. Recuperative heat exchange at suchhigh temperatures necessitates the use of catalytically coated ceramic tubularreactors, which are costly and lack operational robustness. In addition, only abouthalf of the heat supplied in the form of combustion gases can actually be utilizedfor the reaction.


5.2. Regenerative Heat Transfer Process for Hydrogen

Cyanide Manufacture

Although regenerative heat exchange solutions have been proposed for the firsttwo examples presented (16), hydrogen cyanide synthesis is, in fact, a morepromising candidate by virtue of the higher and broader temperature range andthe lower operating pressures. A cursory analysis reveals that a regenerative heat exchange process could in fact combine the most important advantages ofthe reactive and recuperative processes while surmounting their shortcomings(Table 2) (17).

More detailed modeling exposes some weaknesses, for example, the needto use coarse monolithic catalyst structure to achieve reasonable reheat periodswithout excessive pressure drops. The cycle time of approximately 4 minutes issomewhat short for practical purposes. The “cold spot” formation in the reactionphase (Figure 14) and the resultant inability of the reaction to distribute itself overthe catalyst to utilize the stored heat optimally is probably a modeling artefactcaused by too low literature values for the activation energy, possibly reflectingan incorrectly interpreted film transport limitation.

FIGURE 13 Possible thermal coupling mechanisms between the endothermicsynthesis reaction and the exothermic hydrogen oxidation in hydrogencyanide manufacture.


5.3. Desorptive Cooling for Enhanced Regenerative

Heat Transfer

Process intensification measures to extend regenerative cycle times could enablethe reaction engineer to utilize the strengths of this form of heat exchange morefully. The use of the larger enthalpies associated with phase changes rather thansimple specific heat effects to store thermal energy is a technique already exploitedwith evaporative coolants in recuperative processes and in the liquid injectionmentioned earlier in the context of convective processes. The regenerative analogyto this principle can be illustrated using a technique that can be referred to as desorptive cooling (18).

If an inert material is initially adsorbed on the fixed bed comprising an appropriate adsorbent and a catalyst, the heat of adsorption—having the sameorder of magnitude as the latent heat of evaporation—will be released (Figure 15).Since no reaction takes place in this phase, moderate temperature excursions areacceptable, and recycle flows over external heat exchangers or injection of liquidadsorptives may serve as heat sinks. In the subsequent reaction phase, the heat lib-erated by an exothermic reaction on the catalyst is taken up by the desorption ofthe inert from the previously loaded neighboring adsorbent particles. As long asthis desorption occurs, the heat of reaction will not lead to major temperatureincreases. Sooner or later, of course, the adsorbent will be depleted and the tem-peratures will drift upward, at which point the adsorption phase must be repeated.

Such a system yields intensive cooling without the need for an extensiveheat exchange surface within the reactor. The operation of the reactor in the reaction

TABLE 2 Advantages and Disadvantages of Reactive, Recuperative, and Regenerative Heat Exchange in Hydrogen Cyanide Manufacture


phase is adiabatic. One can even customize the heat removal process by dis-tributing the adsorbent and catalyst in different ratios along the length of the fixedbed to maximize cycle duration. Furthermore, if the desorption process exhibitsgreater sensitivity toward temperature changes than the reaction, the process willbecome self-regulating to a certain extent. On the negative side, desorptive cool-ing is, as are all regenerative processes, inherently unsteady state, space timeyields are diminished by the presence of adsorbent in the fixed bed, and the adsor-bent-adsorptive system must be compatible, i.e. inert, with respect to the catalyt-ic reaction being conducted.

To demonstrate the potential available, simulations were carried out for theoxidation of carbon monoxide on a palladium shell catalyst with water desorptionfrom 3A zeolite as a heat sink, based on experimentally validated model param-eters for the individual steps (Figure 16). The calculations indicated that the reac-tion cycle time could be lengthened by a factor of 10, to a total 20 minutes, incomparison to a simple regenerative process with a similar amount of inert mate-rial instead of adsorbent in the fixed bed and for the same threshold for tempera-ture deviation from the initial value.

FIGURE 14 Temperature profiles of gas and catalysts in a regenerative processfor hydrogen cyanide manufacture at the start and finish of the reaction cycle.Cycle duration is 4 minutes and the monolithic catalyst used has the follow-ing dimensions: wall thickness 4 mm, channel width 20 mm, length 2000 mm.


Experimental measurements yielded only a fivefold extension of the reac-tion cycle time, a difference largely caused by heat storage effects in the small-scale equipment used, which disproportionately enhanced the cooling effectobserved in the inert bed control experiment. Despite this less satisfactory result,the desorptive cooling concept would still seem to offer potential for dramaticimprovement in performance for regenerative heat exchange processes.

The desorptive cooling principle is effectively equivalent to the distributedinjection of a cooling liquid along the length of a fixed-bed reactor into a gaseousreaction medium undergoing exothermic reaction. An extension of the idea, theadsorption of a reaction product to both enhance the equilibrium position andprovide some of the heat required for the endothermic reaction, has also been pro-posed (19). More mundanely, the latent heat effects of wax solidification havebeen exploited in temperature-regulating fabrics incorporating microencapsulatedwax particles! A certain analogy can also be drawn with the previously mentioneduse of catalyst bed dilution with inert material to better harmonize recuperative

FIGURE 15 Operating principle of “desorptive” cooling.


heat removal with the rate of heat formation due to reaction, even though thismeasure must be considered a example of process deintensification in the strictestsense.

5.4. Regenerative Heat Transfer in Adsorptive Reactors

The benefits obtained by integrating adsorption into regenerative heat exchangedemonstrate the synergies available between these two related processes. Similaradvantages accrue when heat regeneration is incorporated into adsorptive reac-tors, in which concentration profiles are manipulated to improve reactor per-formance through selective ad- and desorption of components in the reactionmedium.

An example of the potential in this second direction is provided by an inno-vative regenerative reactor for carrying out the Deacon reaction (20), in whichhydrogen chloride is catalytically oxidized to chlorine—an important step inchlorine recycling for the chemical industry. By resolving the gas-phase reactioninto two sequential gas–solid subreactions corresponding to adsorption-desorptionsteps (Figure 17), one can overcome the equilibrium thermodynamics that otherwise

FIGURE 16 Simulated temperature profiles along a reactor with and without“desorptive” cooling at various times for the oxidation of CO on a Pt catalystwith water vapor desorption from 3A zeolite in a fixed bed comprising equalproportions of catalyst and adsorbent. The solid curves give the simpleregenerative behavior and the dotted curves describe the desorptivelycooled case. Initial reactor temperature is 125°C, initial adsorbent loading0.12 kg/kg, inlet CO-concentration 0.2 mol/l, gas loading 6000 h�1.


limit conversion to around 85% and complicate downstream processing of theresultant partially converted reactor output. In the exothermic chlorination step,hydrogen chloride reacts with the oxidized form of the catalyst to yield copperchloride and water vapor. During the following endothermic oxidation stage, thiscopper chloride is converted back to its oxidized form, liberating chlorine in theprocess. In this manner, one can achieve total conversion and considerably sim-plify the subsequent processing necessary. In view of the extremely corrosivenature of the reaction (and catalyst!) system a regenerative transfer of heatbetween the two phases is to be preferred over any recuperative alternative. Theuse of the fixed bed to store both heat and chlorine between the two steps of thereaction cycle leads to contradictions and bottlenecks in catalyst and reactordesign—a common drawback of multifunctionality as a technique for processintensification. For example, the removal of the excess heat of reaction from thereactor favors cocurrent operation of the fixed bed in its function as a heat regen-erator. For optimal utilization of the concentration profiles and maximal chlorinecapacity of the fixed-bed countercurrent, operation of the two cycle phases is tobe preferred. Despite these difficulties, it still proved possible to modify the cat-alyst and reactor design sufficiently to fulfill both the “adsorptive” and the“regenerative” function.

6. ELECTROMAGNETIC HEAT TRANSFER

The use of electromagnetic techniques represents the final method for processintensification of heat transfer that will be dealt with here. Developments in thefuel cell sector hold out the promise of cheap electric power based on surplushydrogen in chemical plants. The liberalization of the energy sector has also cutelectricity prices and enhanced the attractiveness of using electricity in connectionwith chemical reactors. The precise regulation possible with electrical processesand their clean, environmentally friendly nature are further inducements for theirapplication.

FIGURE 17 Operating concept of an unsteady-state Deacon process for chlo-rine recycling.


Once again it must be said that the use of electricity in chemical reactors ishardly new. Apart from numerous electrochemical processes, electrothermalprocesses are employed for the manufacture of phosphorus, hydrogen cyanide,and calcium carbide. Furthermore, syntheses based on electric arcs have beenused in the past for the large-scale production of nitric oxide and acetylene andare still employed today for ozone manufacture or chemical vapor depostion coating procedures. A major advantage of the electrical heating employed in elec-trothermal processes that offsets the higher operating costs is the absence of anylimit on the temperature level at which it can be applied, making it particularlysuitable for processes above 1000°C, where conventional heat transfer becomesdifficult to realize.

6.1. Ohmic Heating

Ohmic heating of catalyst is often used as a simple method of igniting the chem-ical reaction during reactor startup, for instance, in the oxidation of ammonia onplatinum-rhodium gauze catalysts. Another application is the prevention of “cold-start” emissions from automotive catalysts responsible for much of the residualpollution still produced from this source (21). The startup times needed for thecatalyst to attain its operating temperature can be cut by a factor of 5 or more by installing an electrically heated catalyst element with a metallic supportupstream of the main catalyst unit. Direct electrical catalyst heating permits faciletemperature control but requires a well-defined catalyst structure to functioneffectively.

6.2. Dielectric Heating

Much attention has recently also been devoted to dielectric heating of reactorsusing microwaves (! � 1 cm to 1 m, � � 30 GHz to 300 MHz) (22). As in domes-tic applications, the primary attraction lies in the absence of heat exchange sur-face, and thus of fouling, and local overheating. Dielectric heating is especiallysuitable for temperature-sensitive materials for which even slight nonuniformitiesor temperature gradients might prove damaging, which explains its use in themanufacture of ceramics and catalysts and in the sterilization of complex fer-mentation media. It has also been proposed for the local production of limitedquantities of hazardous chemicals, such as hydrogen cyanide (Figure 18) (23).The rapid startup and exact temperature regulation possible, even at varyingthroughputs, together with the low inventories entailed by such production reac-tors compensate for the higher costs of the microwave heating. The speed ofmicrowave heating also makes it a suitable technology for the oxidative regener-ation of diesel particulate filters.

Although the efficiency of microwave heating is high, the generation of micro-waves is by no means free of losses. Specific chemical effects due to microwaves


remain controversial and are probably of little significance for industrial purposes.Questions of scale-up, capacity limits, and materials of construction still need tobe addressed before industrial applications of microwaves in combination withchemical reactors become more widespread.

An interesting application, derived from the use of microwaves for selectivedesorption processes, is the modification of catalyst performance by the imposi-tion of a temperature profile on a catalyst pellet, which is usually dictated by theinteraction between the heat of reaction and the thermal conductivity of the pellet. Microwave heating together with the use of carrier materials of variouspermittivities and conductivities would permit one to regulate the temperatureconditions within the catalyst pellet independently. An extension of this principlewould be the selective thermal activation of one sort of catalyst in a mixed fixed-bed system. The external “switching” of catalytic activity in this manner could beemployed expediently to realize multistep syntheses in a single reactor. To preventthermal short-circuiting it would be necessary to isolate the individual catalystparticles from one another in an insulating matrix.

6.3. Electric Arc Processes

Electric arc processes have been given a new lease on life in the guise of plasmareactors, especially those involving “cold,” or nonthermal plasmas, with electron“temperatures” of 104–105 K and gas temperatures of 102–103 K. Plasmas of thiskind can be used to activate and functionalize inert molecules, but usually withonly poor selectivities and low energy yields (� 0.01 mol/kWh!). The use of cat-alytic additives may offer some potential for improvement, but reactive plasmaprocesses will probably remain restricted to a few specific applications.

FIGURE 18 Yields of hydrogen cyanide in catalytic and noncatalyticmicrowave-heated reactors. (From Ref. 23.)


6.4. Peltier Cooling

A major drawback of the electrical techniques described is that they are, withoutexception, heating processes, whereas the reaction engineer is often more inter-ested in reactor cooling. The active cooling of a catalytic surface using the Peltiereffect enables one to achieve almost perfect temperature control, even in the faceof strongly exothermic reaction behavior, but the low efficiency of the phenome-non means that a large additional amount of heat also has to be dissipated. As aPeltier element only “pumps” heat from one location to another (Figure 19), thequestion of the ultimate heat sink must still be resolved. For these reasons suchtechniques will probably be restricted to laboratory use in the form of specialmicroreactors for the foreseeable future.

7. CATALYST MODIFICATION FOR ENHANCED

HEAT TRANSFER

All of the recuperative, regenerative, and electrical methods described for inten-sifying heat transfer in chemical reactors illustrate the importance of developingappropriate catalyst systems tailored to the behavior being sought. Improving catalyst performance in this way by using “active” multifunctional supports whileleaving the catalytic chemistry unchanged can be referred to as “commensal”catalysis, in analogy to the natural relationship between two species that benefitsone partner (the catalyst) while being neither advantageous nor disadvantageousfor the other (the carrier). Bifunctional catalysts, with spillover diffusion effectsto reduce coke formation, and zeolitic catalysts with their selective access toactive sites represent two conventional examples of commensal catalysis. How-ever, the support properties can also be modified to enhance interfacial areas, heatand mass transport, heat and mass storage, mechanical, thermal, or chemical

FIGURE 19 Active electrical reactor cooling using Peltier elements.


resistance, and electromagnetic behavior (Table 3). The use of such microstruc-tured hybrid catalysts can make an important contribution to process intensifica-tion measures in other areas as well.

8. SUMMARY

The reaction engineer has a variety of tools at his disposal when attempting tointensify heat transfer in chemical reactors, ranging from well-established meth-ods to innovative technologies. For recuperative heat transfer, the most dramaticimprovements can be achieved by using catalytic or catalytically coated heatexchange surfaces and working at the millimeter scale to harmonize the physicaland chemical processes taking place and render the catalytic chemistry the performance-limiting step. These two measures overcome the most serious bot-tlenecks in the traditional multitubular reactor.

The operation of regenerative and reactive processes for the manipulationof temperature profiles in chemical reactors is usually more complex than con-vective and recuperative techniques, due to the inherent dynamics and high sen-sitivities involved. Regenerative and reactive processes can, however, permit

TABLE 3 Performance Enhancement Through the Modification of thePhysical Properties of Catalyst Supports

• Bifunctional Catalystse.g. propene oxidation, Pt-doped zeolites, spillover-oxygen

• Supported Catalystse.g. impregnated catalysts, SLPC, SAPC

• Mass Transporte.g. zeolites, membrane encapsulation, Aerogel

• Heat Transporte.g. graphite carriers, coated-wall reactors, full metal catalysts

• Mass Storagee.g. active C-carriers, hydride-containing catalysts

• Heat Storagee.g. metallic monoliths

• Mechanical Resistancee.g. protective coat, “washcoat”, gauzes

• Thermal Resistancee.g. non-oxide ceramics, doped Ba-hexaaluminate

• Chemical Resistancee.g. non-oxide ceramic, silicon dioxide, Al-phosphate

• Electromagnetic Propertiese.g. dielectrically heated catalysts, magnetic fluidized beds


a more active (self-) regulation of the temperature conditions within the reactor,which can, in some cases, simplify their implementation.

The full potential of hybrid operation, employing more than one of the fun-damental processes for manipulating temperature and concentration profiles,remains to be realized. Special catalysts and reactors must be developed toaccommodate the conflicting demands that often arise in the design process.

REFERENCES

1. Nicol W, Hernier M, Hildebrandt D, Glasser D. The attainable region and processsynthesis: reaction systems with external cooling and heating. Chem Eng Sci 2001;56:173–191.

2. DECHEMA. Trendbericht No. 5: Prozessintensivierung. Frankfurt: DECHEMAGesellschaft für Chemische Technik und Biotechnolgie eV, Press release, 2000.

3. Levenspiel O. The Chemical Reactor Omnibook; Corvallis, Or: OSU Book Stores,1996.

4. Eigenberger G. Fixed-bed reactors. In: Ullmann’s Encyclopedia of IndustrialChemistry. Weinheim, Germany: VCH, 1992, Vol. B4:199–238.

5. Krishna R, Sie ST. Strategies for multiphase reactor selection. Chem. Eng. Sci. 1994;49:4029–4065.

6. Lahne U, Lohmüller R. Schüttschichtreaktoren mit gewickelten Kühlrohren, einekonstruktive Neuentwicklung zur Durchführung exothermer katalytischer Prozesse.Chem Ing Tech 1986; 58:212–215.

7. Ehrfeld W, Hessel V, Löwe H. Microreactors. Weinheim, Germany: Wiley VCH,2000.

8. Gavrilidis A, Angeli P, Cao E, Yeong KK, Yan YSS. Technology and applications ofmicroengineered reactors. Chem. Eng. Res. Des. 2002; 80:3–30.

9. Gerhardt W. BASF AG. Personal communication, 2000.10. Dummann G, Pahlke T, Agar D. Versiegelung microporöser Strukturen mit

Gegendiffusion-CVD. Chem Ing Tech 2002; 74:824–827.11. Frauhammer J, Eigenberger G, Hippel LV, Arntz D. A new reactor concept for

endothermic high-temperature reactions. Chem. Eng. Sci. 1999; 54:2661–3670.12. Hendrikson WG, Daniels F. Fixation of atmospheric nitrogen in a gas-heated fur-

nace. Ind. Eng. Chem 1953; 45:26113–2615.13. Matros YS. Catalytic processes under unsteady-state conditions. In: Studies in

Surface Science and Catalysis. Amsterdam: Elsevier, 1989, Vol. 43.14. Nieken U. Abluftreinigung in Katalytischen Festbettreaktoren bei Periodischer

Strömungsumkehr. VDI-Fortschrittberichte No. 328; Düsseldorf: VDI-Verlag, 1993.15. Pinjala V, Chen YC, Luss D. Wrong-way behavior of packed-bed reactors. II. Impact

of thermal dispersion. AIChE J 1988; 34:1663–1672.16. Kolios G, Eigenberger G. Styrene synthesis in a reverse-flow reactor. Chem. Eng. Sci

1999; 54:2637–2646.17. Agar D. Multifunctional reactors: old preconceptions and new dimensions. Chem.

Eng. Sci. 1999; 54:1299–1305.


18. Franke M. Bewertung desorptiver Kühlung von Festbettreaktoren. Master’s thesis,Chemical Engineering Department, University of Dortmund, Dortmund, Germany,2000.

19. Yongsunthon I, Alpay E. Design of periodic adsorptive reactors for the optimal inte-gration of reaction, separation and heat exchange. Chem. Eng. Sci 1999; 54:2647–2657.

20. Agar D, Watzenberger O, Hagemeyer A. A novel unsteady-state process for HCl oxidation—multifunctional reactor operation with regenerative Cl storage in a fixedbed. In: Proceedings of R ’97 Meeting, Recovery, Recycling, Re-integration. Geneva,Switzerland, 1997: Vol. 4, IV.45–IV.50.

21. Kirchner T, Eigenberger G. Optimization of the cold-start behavior of automotivecatalysts using an electrically heated precatalyst. Chem. Eng. Sci. 1996; 51:2409–2418.

22. Bathen D, Schmidt-Traub H. Alternative für Nischen—Anwendungspotentiale derMikrowellentechnologie in der Verfahrenstechnik. ChemieTechnik 1998; 27:80–83.

23. Lerou J, Ng K. Chemical reaction engineering: a multiscale approach to a multiob-jective task. Chem. Eng. Sci. 1996; 51:1595–1614.


11

Process Synthesis/Integration

Patrick Linke and Antonis Kokossis

University of Surrey, Surrey, England

Henk van den Berg

University of Twente, Enschede, The Netherlands, and Ghent University, Ghent, Belgium

1. INTRODUCTION

Process synthesis, also referred to as process integration, deals with the systematicdevelopment of process flowsheets. The process synthesis activity has beendescribed as “the automatic generation of design alternatives and the selection ofthe better ones based on incomplete information” (1). Design technology is requiredto help the engineer find novel, improved solutions to process design problems inthe context of the incomplete information available. The ultimate aim of chemi-cal process design is to synthesize a process that enables the production of desiredchemicals in the most cost-effective and environmentally benign manner possibleand is flexible as well as easily operated. Ideally, process synthesis tools shouldallow one, out of the set of all feasible alternative structural and operationalprocess design options, to systematically determine the most promising processdesigns, those that approach the performance limits of the system closely andmeet the constraints.

An enormous variety of decisions need to be made in order to solve processdesign problems. These range from the selection of the most promising processchemistry (reaction paths and catalysts) to the optimal exploitation of reaction,


mass, and heat transfer interactions at the process level. In order to arrive at a truly optimal design, all decisions have to be made with respect to the overallobjectives in order to achieve the best balance of process trade-offs. Due to thecomplexity of the overall problem, applicabilities of systematic decision-makingtechnologies are generally confined to closely defined design subproblems. The“onion” model (see e.g., Ref. 2) captures a shared view that the process designproblem is best decomposed into a reactor design (the heart of the process) sub-problem, a separation system design subproblem, an energy system design sub-problem, and a utility system design subproblem. These subproblems appearchallenging for a variety of reasons. Challenges in reactor optimization arisemainly from the high nonlinearities that need to be addressed. In energy systemssuch as heat exchanger networks, on the other hand, the challenge is to search thevast number of combinatorial design options. Separation system design presentsexamples of intermediate complexity in terms of model nonlinearities. Despite thechallenges, the design problem decomposition itself leaves plenty of unansweredquestions to be addressed by future research. Clearly, there is significant scope forimprovement by looking at a bigger picture and systematically exploiting interac-tions between solutions to the classical design subproblems. Recent developmentsin process synthesis technology aim at giving a more global perspective by solv-ing less confined but more conceptual problems in a variety of areas.

Process synthesis and integration tools are generally developed in acade-mia. Successful technology development requires close collaboration between theacademic and the industrialist to ensure applicability to real-life problems. Themost successful example of process synthesis and integration technology devel-opment is pinch analysis (3). After many years of close academia–industry col-laboration, pinch analysis has led to significant energy savings in the chemicalindustries and has become a standard tool employed in most energy system syn-thesis projects worldwide. In contrast to the simpler techniques for energy sys-tems synthesis that have relatively quickly been taking up applications inindustry, the methods for separation and reaction system synthesis have been slowin being widely accepted by industrialists. However, these tools have the poten-tial to significantly help design engineers in their quest to develop new plantsfaster and with fewer resources.

The next section will reflect on common conceptual process design prac-tices for overall flowsheet development. The remainder of this chapter reviewsrecent developments in process synthesis methods for reaction and separationsystems that systematically guide the design decision-making process towardnovel and improved designs.

2. CONVENTIONAL CONCEPTUAL DESIGN PRACTICES

In this section a number of issues of process synthesis and conceptual design willbe reviewed. Based on available textbook knowledge (2,4–6), a synergistic


approach is discussed that uses project organization, decomposition, processbuilding, selection of alternatives, and design rounds. Process functions are takeninto account at a level above the unit operations. Due to the importance of proj-ect organization (6), management techniques also need to be applied to processengineering. This section reflects on experiences with industrial applications aswell as graduate courses on process design. New aspects on the combination andarrangement of activities will also be discussed (7).

2.1. Elements of Conceptual Process Design

This section deals with a number of basic points of process design strategy, cor-nerstones that have to be applied in a synergistic procedure. The application real-izes a synergy of process functional analysis, process synthesis, and projectorganization to generate structure and trace alternatives and decisions. This syn-ergy in conceptual process design uses the elements of heuristics, mathematics,and creativity. Components of conventional design approaches include:

Existing flowsheet and technology analysis—collection of know-how.Flowsheet decomposition into functional sections where tasks prevail over

equipment and unit operations.A black box approach as the endpoint of the analysis; raw materials and

products, overall process-relevant data (e.g., yields, reactor conditions),and boundary conditions are given, leading to a first evaluation of theoverall process; economics are based on raw materials and products.

Formulation of goals for the new design; the black box is the start of therebuilding of the process from outside to inside.

Systematic process rebuilding, while including known and creative tech-nologies and breakthroughs and using heuristics, expert systems, andprocess simulation.

Application of management tools in the form of tree diagrams (interrela-tionship of goals and means) and work diagrams (to structure the projectstep by step), document alternatives and choices.

The approach has similarities with a bow-tie model: One starts broad andcollects all available process information. A structure is used, e.g., a tree diagram,to create a format for the information needed and collected. One should concen-trate on the essentials and collect all relevant factors that influence the process,from the outside to the inside. Setting the boundaries of the process early reducesrework later on; however, this clearly limits the degree of design novelty that canbe achieved.

Phasing of process design is a normal issue in the process industries. A processdesign or substantial modification is done stepwise. Sequential steps in processdesign are carried out after management approvals of the process proposals. Thismeans that project organization is of vital importance to direct activities and to


avoid rework. This aspect is hardly covered in the textbooks for process synthe-sis. Biegler et al. (5) use several levels or loops in their book, but they are notlinked to the phases of process design. The concept of process flowsheet decom-position is used by Douglas (4) and Biegler et al. (5) but needs more emphasisfrom the moment a project is started. Flowsheets have to be translated into func-tional activities so that a critical and systematic analysis can be started early. Theresults can be used during the buildup phase after the black (grey) box modeldevelopment. Conventional technology is easily used in the buildup phasebecause it is mature and has less or no risk, but new technology has to be takeninto consideration to enable step changes in performance improvements and toavoid copy engineering.

All design textbooks mention that important decisions are made in the earlystages of design and that the know-how to make these decisions is developedgradually during the design process. The consequences of choices made in earlystages of design cannot be foreseen at the moment of the decision. It is clear thatminor decisions about the type of a centrifugal pump to be used for the reflux ofa distillation column are less important in the early stages of design than in theperformance of the reactor. Douglas (4) advocates the application of rejection ofless attractive alternatives and proposes to continue the development of all alter-natives that cannot be rejected. The conventional way is to continue with a singlealternative; however, this tends to cause problems later in the project. Biegler et al. (5) and Seider et al. (6) show how to document alternatives and choices.Criteria for rejection are, e.g., process yield and selectivity, costs, safety, ecology,and reliability of equipment. These criteria and the decision process have to bedocumented. It must be possible to return to the design path in case the resultshave to be rejected and early decisions need to be reconsidered, as shown inFigure 1.

A systematic process design cannot be done in one step. It is common prac-tice to develop a draft flowsheet first. This will be a feasible, unoptimized flowsheet,

FIGURE 1 Alternatives and choices.


where the consequences of the choices made are not yet clear. Biegler et al. (5)refer to this as the base-case design and propose four sequential rounds to com-plete the design of the process. The process know-how is extended from round toround, and the consequences of choices from among alternatives become moreand more clear. In the fourth step Biegler et al. (5) use optimization techniques tosearch for design alternatives. Ideally, the project phasing applied in the processindustry and the growing ability of the process engineers to collect and judgealternatives during the project need to be matched. The next section describessuch an approach.

2.2. Step-by-Step Conceptual Process Design

This section gives a short description of a step-by-step design strategy that appliestopics treated in the previous section. Assume that the task of a process is to con-vert raw materials A and B into product C and to minimize by-product or wasteD. The states of materials A, B, C, and D are defined by: mass flow, composition,phase (vapor, liquid, solid), form (e.g., particle size), temperature, and pressure(6). The process design has to consider alternatives for raw materials and processroutes that lead to a given product. This is done in step 0 of the design activities.An early set design basis is often the production capacity (ton/yr). Each processoperation can be viewed as having a role in eliminating one or more of the prop-erty differences between the raw materials and the products. The first step is toeliminate differences in molecular type by chemical reaction. The function of thereactor is usually considered the heart of a chemical process. Raw materials areseldom converted into the desired products at the required purities, so a separa-tion function needs to be defined in Step 2. The need for and task of separation isoften strengthened by the reactor conditions, e.g., excess of a component may berequired inside the reactor. Generally, the reactor outlet needs the separation func-tion to match the product specification. This information is collected in thedecomposition stage of process design. Other functions in the process area thatone needs to take into account are change of phase, phase separation, change oftemperature and pressure, as well as mixing and splitting streams. Each operationcan be viewed as having a role in eliminating one or more of the property differ-ence between raw materials and products. The transition from raw materials toproducts is schematically represented in Figure 2, based on the information givenin Seider et al. (6) and Biegler et al. (5). Alternative arrangements of reactors andseparators are considered as a next design action. Both Douglas (4) and Seider et al. (6) consider the choice between batchwise operation and continuous operationa first choice in process design. Our experience is that a black box evaluation—amass balance and a first economic evaluation—based on raw material and prod-uct flows, yields, and prices can be made without making a choice between batchand continuous operation.


Connections between process units are created in a third step while consid-ering in detail the following aspects:

From raw materials to reactor: purification and conditioning of feeds (purif-ication, heating, pressure)

From separator to product: final purification (finishing)Reactor–separator connections: stream reconditioning (e.g., reduction of

temperature and/or pressure)Connections of separator outlets: reactor performance, excess components,

recoveryRecycles from separators: connections to, e.g., reactor system

Step 4 consists of defining and matching the mass flows, including, e.g., theexcess of one component required in the reactor. The activity starts with buildingpreliminary mass balances before other conservation laws, i.e., heat and impulse,are considered. Other relevant factors that have to be evaluated early in processdesign are controllability and safety. The outcome of Step 4 is an overview ofprocess functions needed to realize the transition of raw materials into products,including alternative flowsheets and unit operations that could be applied. Step 5considers the combination of unit operations or task integration to develop theopportunities for process intensification, viz., leading to smaller, more efficient,and cheaper processes. This step includes screening for options related to heatintegration and reactive separations, among others.

Step 6 consists of the setup of the mass and heat balances, and alternativesfor the reaction and separation functions are considered. The alternatives are doc-umented in a structure such as the one shown in Figure 1 or the synthesis tree pro-posed by Seider et al. (6). Proceeding with the steps, many process alternativesare generated form which the most promising ones need to be selected taking

FIGURE 2 Moving from synthesis steps to process operations.


into account profit, process yield, controllability, safety, and ecological aspects.Criteria for assessing preliminary designs are presented in Ref. 5.

Between the steps, interactions will take place and decisions can create theneed for loops to previous steps, e.g., to reconsider the reactions (Step 1) and sep-aration (Step 2) in more detail. Process design is generally not done in a first andsingle round; a next round is required to reconsider choices made, reduce thenumber of alternatives, and select the most promising option—while taking intoaccount best available technology, heuristics, and creative new solutions. A sec-ond round, as proposed by Biegler et al. (5), is carried out to create a more in-depth conceptual design and to extend the process design activities, e.g.,toward process simulation, heat integration, and equipment design. This is donein line with the project phases mentioned.

Various industrial applications of the step-by-step design approach havebeen reported (see, e.g., Ref. 8). A number of these and other confidential projectshave shown the value of the coordinated approach to conceptual process design.

From the foregoing discussion it is clear that the process alternatives arecommonly generated based on intuition and case-based reasoning. This leaves astrong chance that promising design candidates are not arrived at and that noveltyis not automatically accounted for in the design process. Ideally, a systematicapproach should be able to capture all possible design alternatives and screen forthe design that delivers the best possible performance for the specified perform-ance measure. The following sections describe such technology developments forreaction system synthesis, separation system synthesis, and integrated reaction–separation systems.

3. REACTION ENGINEERING

The reactor is undoubtedly the most important ingredient of a chemical processflowsheet, for it is the part where the product value is generated. From a decision-making perspective, the reactor design is more of a difficult process design taskthan the separation and energy systems design tasks. This is because the reaction,heat transfer, and mass transfer phenomena tend to occur simultaneously in thisunit. Most reaction models of commercially relevant systems are highly complex,and the development of both graphical and computational design tools that enablequick decisionmaking so as to obtain high-performance reactor designs is chal-lenging. Such tools are required to provide performance targets and design sug-gestions for mixing and operational policies to give maximum decision supportto the designer.

There is an increasing awareness that the commonly employed textbookknowledge and heuristics (9) is insufficient for the systematic development ofhigh-performance reactor designs. The result is a lack of innovation, quality, andefficiency in many industrial designs. Researchers from various perspectives are


making efforts to develop systematic optimization tools to improve the perform-ance of chemical reactors. At a conceptual level, the area of reactor network syn-thesis has resulted in new methods that focus on a systematic and thoroughconsideration of the available options. In the remainder of this section, we descr-ibe recent developments in reactor network synthesis that provide systematicdecision support for various reactor design aspects. Reactor network synthesisaims at identifying high-performance reactor design candidates that exploit mix-ing, feeding, bypassing, recycling, and temperature effects such that the systemsperformance is maximized with respect to the objective functions employed.

3.1. Single-Phase Systems

Virtually all efforts in reactor network synthesis have addressed single-phase sys-tems. Design approaches can be broadly divided into superstructure optimization-based and graphical synthesis. Achenie and Biegler (10–12) were the first tosynthesize comprehensive reactor superstructures using optimization technology.They developed superstructures using axial dispersion models, recycle-PFR repre-sentations, and environmental reactor models and applied optimization tech-niques in the form of nonlinear programming (NLP) methods to identify the mostpromising design candidates hidden in the them. Kokossis and Floudas (13–15)first introduced the idea of a reactor network superstructure modeled and opti-mized as a mixed-integer nonlinear programming (MINLP) formulation. Thoughgeneral and inclusive, their representation did not follow previous developmentsbut made an effort to facilitate the functionalities of the MINLP technology withthe synthesis objectives. With the primary purpose of scoping, optimizing, andanalyzing the reaction process, Kokossis and Floudas replaced detailed modelswith simple though generic structures, enough to screen for design options andestimate the limiting performance of the reaction system. In the same vein,dynamic components have been replaced with the use of CSTR cascades. Asuperstructure of generic elements (ideal CSTRs and PFRs) was postulated toaccount for all possible interconnections amongst the units (Figure 3). The repre-sentation was modeled and optimized as an MINLP model. Schweiger andFloudas (16) later revisited the approach and optimized superstructures with thePFRs, with side streams being replaced by rigorous DSR representations thatavoid the inaccuracies introduced by the use of CSTR cascades.

Around the same time, Glasser et al. (17) retrieved and extended theinsightful methods of Horn (18) and presented graphical procedures known as theattainable region (AR) method. Their approach requires the graphical construc-tion of the convex hull of the problem and helps to exemplify the need for a sys-tematic and general methodology. In principle, the reactor network withmaximum performance in terms of yield, selectivity, or conversion can be locat-ed on the boundary of the AR in the form of DSR and CSTR cascades with


bypasses. Though useful in two dimensions, in higher dimensions their develop-ments face both graphical and implementation problems. Furthermore, the ARmethods frequently result in complex designs with multiple DSR and CSTR unitsand complex feeding and bypassing strategies. Though fundamental limitationsappear evident, persistent efforts to extend the graphical methods have been pub-lished (19–24). A more promising direction has been pursued by Biegler andcoworkers. The motivation has been to instill better guarantees in the optimiza-tion efforts by exploiting ideas and rules established in the construction of theAR. Applications presented in this area include the work by Balakrishna andBiegler (25,26) and Lakshmanan and Biegler (27–29) and involved mathematicalprogramming applications in the form of NLP and MINLP formulations.Hildebrandt and Biegler (30) presented a review of the attainable regionapproaches and suggested areas for future development of the concept.

In 1999, Marcoulaki and Kokossis (31) presented a different interpretationof the synthesis problem. From a practical viewpoint, the nature of a usefulapproach for reactor network synthesis should primarily account for solid per-formance limits for the reaction system (targeting) and the systematic develop-ment of design candidates that approach this performance (screening andscoping). Targeting is particularly useful for the design decision making becauseit allows design evaluation in light of the ultimate performance possible for agiven system, provided there is enough confidence in the optimization results.The targets can be used for synthesis and retrofit problems because they can pro-vide the incentives to modify the operation and ideas for developing the reactordesign. Because the design equations of chemical reactors feature a significantnumber of nonconvex terms, the importance of confidence assumes a significantplace, and robust optimization technology is called for. On the other hand, the

FIGURE 3 Reactor superstructure. (From Ref. 31.)


borders of the attainable region and the strictly optimal solutions assume a rela-tive meaning in reactor design. The reaction kinetics typically involve significantuncertainties and approximate models that give little justification for a strictemphasis on the optimum. In this context, the development of a CSTR-PFR-CSTR-PFR-CSTR-PFR layout as an “optimal” configuration represents an aca-demic exercise of limited interest. A single, slightly inferior CSTR can featureoperational advantages and prove an even better choice in the form of an indus-trial back-mixed reactor. From a targeting viewpoint, layouts near the targets areconsidered equally important. Therefore, solutions in the interior of the attainableregion make valid options as long as they remain close to the targets (31).Screening and scoping aims at identifying the range of design candidates that per-form reasonably close to the targets so as to provide the design engineer withoptions on which the decision making can be based.

The nonlinear, discontinuous, and discrete nature of the synthesis problemformulation considered here is not suitable to be addressed by mathematical pro-gramming techniques in the form of the commonly employed (mixed-integer)nonlinear programs ((MI)NLPs) because the synthesis aim is to establish per-formance targets in the different synthesis stages. Mathematical programmingsearches for local improvements and terminates at the nearest locally optimalpoint. For a general case, there is no reason to be confident that the obtained solu-tion cannot be substantially improved. The type of information required from thetargeting stage naturally relates to the results one can obtain with the applicationof a stochastic optimization approach to the reactor network superstructure syn-thesis. The application of stochastic optimization gives one confidence in theoptimization results, can yield particularly nonlinear reactor models, and is notrestricted by the dimensionality or the size of the problem. Marcoulaki andKokossis have applied stochastic optimization in the form of the simulatedannealing meta-heuristic to the single-phase reactor network synthesis prob-lem (31). They optimize the rich and inclusive superstructures formulated byKokossis and Floudas (13) to identify performance targets and to extract numer-ous design candidates that approach the targets. The implementation of the sto-chastic search over the superstructure schemes requires venues for the developmentand evolution of states, the assessment and acceptance criteria, and the coolingschedule. The synthesis perspective and the optimization methodology is discussedin detail in Marcoulaki and Kokossis (31). They found the methodology to system-atically converge to the globally optimal domain, i.e., the performance targets of thesystem, and to produce numerous design alternatives with performances close tothese targets for the numerous complex systems studied.

To illustrate the methodology, consider the Denbigh reaction system. Thereaction scheme involves five components and is described by:

2 21 2 3 4A B A D E → → → →, , ,B C B


with the following kinetics:

The data are from Ref. 31. The objective for optimization is the maximization ofthe effluent concentration of component B. The performance limit of the systemis identified with each stochastic run requiring an average of only 120 CPU secon an HP 9000-C100 workstation. Numerous designs are obtained from the sto-chastic search that perform close to the performance target, mostly variations ofseries arrangements of PFRs and CSTRs. A detailed discussion of this and otherstudies is given in Ref. 31.

3.2. Multiphase Systems

After fixed-bed reactors, multiphase reactors are the most widely used reactionsystems in the chemical process Industries. The common industrial practiceemploys conventional designs based on empiricism, past experience, and qualita-tive reasoning on the basis of analogies with similar systems and case studies. Thepresence of multiple fluid phases in the system represents additional degrees offreedom that need to be exploited in process synthesis. As compared to the single-phase reactor network design problem, the multiphase case poses additionalchallenges in the form of a significantly larger number of possible network con-figurations as well as additional modeling complexities arising from the additionalneed to model multiple phases and to address mass transfer and hydrodynamiceffects.

Mehta and Kokossis (32) introduced a systematic methodology for the syn-thesis of multiphase chemical reaction networks that is based on a compact rep-resentation of design options. The approach accounts not only for conventionalindustrial reactors, such as bubble columns, cocurrent and countercurrent beds,and agitated reactors, but also for all possible combinations of compartments thatcan improve or enhance the performance of a multiphase reaction process. Therepresentation is described in the form of a superstructure of generic synthesisunits featuring shadow reactor compartments, and the synthesis scheme providesfunctionalities that are subjected to optimization. The implementation of thesingle-phase reactor network synthesis methodology of Marcoulaki and Kokossis(31,33) enables the development of targets and screening procedures that can helpthe engineer to assess system performance and review promising design options.

The building block of the superstructure representation is the generic reactorunit, which follows the shadow reactor concept (32). This generic unit is illustratedin Figure 4. Each generic unit consists of reactor compartments in each phase of thesystem, and each processes the reaction. The shadow reactor compartment assumesa state from the set of homogeneous reactors. The default units in the set includeCSTRs and PFRs with side streams. The interface between a given pair of

R R ,R ,R ,R1 2 3� � � � � � � � � �[ ] [ , , , ]4 12

2 3 42k x k x k x k xA B A B


compartments in contacting phases establishes mass transfer links through whichmass transfer with its shadows in other phases is accounted for. Mixing is accom-modated for through links with other compartments of the same phase.

Each generic reactor unit can represent all possible combinations of mixingand contacting patterns that can be associated with the ideal representations of theconventional reactor designs. Figure 5 shows the five possible conventional reac-tor designs and their counterparts in generic reactor unit instances for a systemcomprising two phases. The shadow reactor concept is generic and can be appliedto systems with any number of phases. The modeling equations associated withthe generic units are given in Mehta (34).

The shadow reactor superstructure is generated for a specified number ofgeneric reactor units by linking the compartments of a particular phase with astream network realizing complete connectivity. By a selective combination ofthese streams, options related to feed distribution, product removal, bypasses, andrecycles among the generic units in the network can be evaluated for each phase.The shadow reactor superstructure is illustrated in Figure 6 for a system with twophases. The flow directions in the compartments, along with the options availablein the stream network, give rise to conventional as well as novel arrangements.These options can be explored through optimal search of the solution spacedefined by the superstructure.

Nonisothermal systems are accounted for by the introduction of temperature-control units into the generic reactor unit representation. These units consist of ele-ments associated with the manipulation of temperature changes and constitutetemperature profiles (profile-based approach) and heaters/coolers (unit-basedapproach). The assumption of thermal equilibrium between the contacting phasesreduces the need for a single temperature per shadow reactor compartment. Theprofile-based system (PBS) finds the optimum profiles without considering thedetails of heat transfer mechanisms. Because the profiles are imposed rather than

FIGURE 4 Generic multiphase reactor unit. (From Ref. 34.)


developed, the approach adds no further computational difficulties as comparedto the isothermal reactor network synthesis. By biasing the profile complexity ofthe final solutions, the profiles can be controlled effectively. The solutionsobtained are easy to interpret, and thus the approach helps in understanding ofdominant trade-offs in the problems. Results from the unit-based system (UBS)provides the target that can be obtained from a network of adiabatic reactors withconsideration of direct and indirect intermediate heat transfer options. The syn-thesis of nonisothermal homogeneous and multiphase reactor networks is dis-cussed in detail in Mehta and Kokossis (35).

Following a similar reasoning that had led to the choice of an appropriateoptimization technology for the single-phase reactor network synthesis as describedin Section 3.1, stochastic optimization in the form of simulated annealing (SA) isadapted for the synthesis of the multiphase networks as well. The advantages arethe identification of stochastic optima with confidence levels and the provision of avariety of solutions around the targets that can be reviewed as alternative designs.The synthesis framework consists of three stages. The targeting stage calculatesperformance targets and confidence levels using stochastic optimization in the formof SA. Simulated annealing is based on a randomized evolution of states developedthrough stepwise modification. The states are developed using the shadow reactorsuperstructure. In the screening and design stage, the results of the targeting stageare used to develop designs and layouts that feature performances that fall within adesired distance from the targets. These may consist of designs with the same or adifferent number of compartments, layouts with different networks of streams,

FIGURE 5 Conventional reactor designs and generic unit counterparts. (FromRef. 34.)


recycles, and bypasses, or variations on the same structure. The stochastic searchproduces a multitude of solutions with similar performance, and its function isexploited as a major advantage. These designs can be improved further with theapplication of deterministic methods and more accurate models. The solutions devel-oped from the screening stage are functional models based on the shadow reactorsuperstructure model. The analysis and validation stage requires the translation ofthese layouts into practical schemes. In general there are several ways to developpractical schemes from the functional models because one can opt for either phys-ically distinct units or multicompartment reactors.

Consider the reactor design for the production chlorination of butanoic acidas an example to illustrate the technology developments. A full study is given inRef. 34, and only a brief summary of the results is presented here. The chlorina-tion of butanoic acid (BA) involves two reactions in the liquid phase:

1. BA � Cl2 → MBA � HCl

2. BA � 2Cl2 → DBA � 2HCl

where MBA and DBA are abbreviations for monochlorobutanoic acid anddichlorobutanoic acid, respectively. The system involves two phases, a liquidphase where the reactions occur, and a gas phase consisting of chlorine feed andhydrogen chloride product. Solubilities are calculated using Henry’s law, andmass transfer rates are modeled according to film theory. The reaction kineticsand all model parameters and other problem data used in the study are given inRef. 34. The objective of the study was to find those reactor networks with the

FIGURE 6 Shadow reactor superstructure. (From Ref. 34.)


highest possible yield. Initially the study was performed assuming fixed phaseholdups in the reactor units; i.e., hydrodynamic effects are neglected.

Conventional reactor designs have been optimized to create a basis forcomparison. The mechanically agitated vessel achieves the highest yield of74.4%, followed by the bubble column reactor with a yield of 72.9%, and the co-as well as the countercurrent reactors both achieving a yield of 69.5%.

Reactor network optimization of a superstructure comprising three genericunits produces a performance target of 99.6% yield of MBA. The selected designswith performances close to the target value range from simple designs employingonly one reactor unit to designs featuring three units. The simplest solution con-sists of a single reactor with completely back-mixed gas and liquid phases and abypass of fresh chlorine feed. Other, unconventional designs are also found thatconsist of two or three generic units, all approaching the target performanceclosely. Two of the simple designs are illustrated in Figure 7. For the unconven-tional designs, see Ref. 34.

Mehta and Kokossis (36) also demonstrate how existing knowledge in theform of hydrodynamic correlations can be incorporated into the framework whilemaintaining the possibility of achieving design novelty. They also show how todeal with the ranges of application of the different available correlations so as tofind meaningful results. For the example just described, they found that the per-formance target of 99.6% remains unchanged if the known hydrodynamic effectsare considered for the reactor units via common correlations available in the lit-erature. However, changes in the reactor designs from a mechanically agitated

FIGURE 7 Reactor design candidates. (From Ref. 34.)


vessel to a bubble column reactor are observed as a result of the higher interfacialareas available in the bubble column, according to the correlations. The method-ology for the design of nonisothermal multiphase reactor networks is presented indetail in Refs. 35 and 36.

4. COMPLEX DISTILLATION

Despite advances in other separation technologies, distillation is still the mostwidely used operation in chemical processes. Effective screening of separationsystems constitutes a critical stage, for engineers need to review and understandtrade-offs ahead of detailed modeling and simulation. In the separation, it is oftendesired to explore the use of complex rather than simple columns because thecomplex units reduce mixing losses, use available vapor and liquid more effec-tively, and improve the separation efficiency (37). Despite their recognized poten-tial in energy savings, complex distillation applications are limited due to theirdifficult and demanding design-and-synthesis assignment. Synthesis challengesand operability issues that arise from a more complex dynamic behavior have dis-couraged wider acceptance in industry. A prohibitive number of configurationsemerge from different allocations of side-rectifiers, side-strippers, prefractiona-tors, and side-draw columns. Such options are difficult to enumerate and assess.The design alternatives increase rapidly, and the trade-offs are impossible toassess with an exhaustive (implicit or explicit) enumeration of the options.

Previous efforts have focused on the development of shortcut methodsthat had a purpose of evaluating fixed configurations and initializing simulationmodels. Stupin and Lockhart (38) developed the equivalent arrangements ofsimple columns to represent complex configurations. Several other researchers(37, 39–47) extended knowledge from shortcut models as available for simplecolumns to evaluate the performance of complex configurations. Tedder andRudd (48) performed a parametric analysis for complex designs and identifiedoptimality regions as functions of the feed composition and the relative volatili-ties. Glinos and Malone (49) identified dimensionless parameters and proposedguidelines for the selection of complex distillation schemes. With less emphasison fixed layouts, thermodynamic methods (50–53) produced procedures to assessenergy efficiency in the integrated separation.

Mathematical programming approaches promote process novelty with theuse of superstructure development. Sargent and Gaminibandara (54) pioneered aprogressive distillation train that Agrawal (55) later extended with additional con-nections of vapor–liquid streams to include satellite columns. Christiansen et al.(56) added more connections between component states to include structures withtriangular walls in a single shell. Several researchers (57–62) have proposed dif-ferent superstructures and developed mixed-integer nonlinear programming(MINLP) models for the synthesis of distillation systems. These superstructures


can be generalized as the combinations of two extreme representations, the state–task network (STN) of Sargent (63) and the state–equipment network (SEN) ofSmith and Pantelides (59). These developments use superstructures of complicat-ed interconnections with a large number of variables to account for the designoptions. More often than not, however, the optimization technology required toprocess these models is not addressed as a challenge. Researchers simply assumethe technology is readily available and that the models can be solved and opti-mized robustly. Consequently, reported designs contain conventional rather thannovel and unconventional designs. The development of new mathematical for-mulations for synthesis problems should not be detached from the optimizationtechnology required to handle these models. Yeomans and Grossmann (64,65)have made some progress on a solution strategy using generalized disjunctiveprogramming models, but their method is still unable to address major difficultiesof solving industrial-scale distillation synthesis problems.

Papalexandri and Pistikopoulos (66) propose a representation based on asuperstructure of multipurpose heat and mass transfer modules. The authors referto tasks as a more general form of a synthesis unit. They postulate an even moregeneral (and difficult) problem and propose a detailed modeling framework thatrelies on the use of MINLP solvers. In contrast to the work by Papalexandri andPistikopoulos, Shah and Kokossis (67,68) use the concept of a task (69) to reducerather than enlarge the mathematical model and to simplify rather than complicatethe optimization. In all cases the mathematical models are postulated as smallmixed-integer linear programming (MILP) problems one can solve to global opti-mality. The remainder of this section gives a description of these developments.

Shah and Kokossis (67,68) present a simultaneously systematic and rigorousapproach for the automated development of optimal designs. The approach is com-putationally inexpensive, reliable, and very efficient to implement. It provides theengineer with a selected set of optimal designs on which further attention todynamics and operability would yield the final design. The work embraces con-ceptual and engineering knowledge. In the past, conceptual knowledge has beenrestricted to heuristics and evolutionary approaches, both inappropriate for a sys-tematic search. They typically employ a two-step approach where the best simplesequence is first identified. The sequence is next evolved into complex layouts.Although novelties are not excluded, they are left to coincidence and deprived froma systematic framework. The challenge is to develop systematic procedures thatovercome the difficulties encountered with superstructure approaches. In contrast,the new synthesis approach is based on a supertask representation. The approachconceptualizes alternatives using simple tasks and hybrid tasks. The options includeall complex column configurations related to side rectifiers, side strippers, prefrac-tionators, Petlyuk columns, and side-draw columns. Furthermore, the representa-tion accounts for sequences with sloppy separations and enables one to bracketoptimum pressure limits. The approach assumes a number of components to be


separated into a predefined set of products. Given is the basic information about thecomponents with respect to its physical and critical properties. The objective is todetermine the appropriate and cost-effective separation schemes ahead of detailedsimulations. In order to address complex systems effectively, Shah and Kokossis(67,68) replace superstructures with supertasks. The building blocks of the latterare not unit based, but task-based elements. Apart from simple columns, the syn-thesis representation is required to embed elementary complex distillation arrange-ments involving side columns, side-draw columns, and prefractionators. Thesecomplex arrangements are shown in Figure 8. The figure explains the two differenttypes of side columns: the side-rectifier arrangement type (Figure 8a) and the side-stripper arrangement (Figure 8b); two types of side-draw columns: the vapor side-draw column (Figure 8c) and the liquid side-draw column (Figure 8d); and the twoprefractionator arrangements: with and without thermal coupling (Figures 8e and8f). The arrangement shown in Figure 8f is known as a Petlyuk column.

FIGURE 8 Commonly used complex column arrangements. (From Ref. 82.)


Unlike simple columns, these complex configurations produce more thantwo products and feature more than a single light and heavy key component. Atask-based representation of these schemes is accomplished with the ideas ofhybrids and sloppy splits. Additional tasks are made up of different simple tasks.They are subsequently termed hybrid tasks and are defined as an ordered combi-nation of simple distillation tasks, as illustrated in Figure 9.

The sloppy separation improves the separation efficiency by distributingcomponents between the lightest and the heaviest ones. Among the differentoptions to distribute intermediate components (ICs), the maximum efficiency isaccomplished by a general sequence of sloppy splits to separate the lightest andheaviest product at each stage. For an n-product system, a hybrid of order (n � 1)accommodates a sequence of sloppy splits in the supertask representation.However, it processes each product (n � 1) times (for an n-product system) at theexpense of energy and column shells. These factors have an adverse effect on thecapital and operating costs of the separation system. To capitalize on the advan-tages of the sloppy split, hybrids of a second order (over n-product systems) areincluded in arrangements, where sloppy splits are followed by sharp separation.These schemes combine the higher separation efficiency of the sloppy splits whilethe recurrent processing of products in the downstream sequence is minimized bythe sharp separations. The sloppy split arrangements are introduced as an addi-tional transformation on each hybrid.

Instances of a task are replicas of the task operating under different condi-tions. The concept is used to optimize the operating conditions, such as the column pressure, and assumes the development of an operating range and a dis-cretization scheme. Feasible ranges of pressure are identified by the physicalproperties (e.g., critical pressure) of the key components (upper limit) and theavailable utility levels (lower limit). The discretization scheme may be either uni-form or based on the available utilities. The modeler can use a small or large num-ber of discrete levels to capture associated trade-offs.

The synthesis elements of the supertask representation are translated intomodeling terms that combine available physical properties with design aspects.Shah and Kokossis (67,68) exploit modeling accomplishments from a variety ofresearch groups (43,37,70) who have separately addressed the modeling and sim-ulation challenges of complex distillation configurations. The calculations areused to set up the MILP optimization problem. Each task is assessed and evalu-ated independently and separately from the performance of the upstream ordownstream units, because the composition and flow rate of the feed and theproducts streams are known a priori and only depend on the separation functionof the particular task. Design and performance calculations can use any shortcut,semirigorous or rigorous method and any property package without influencingthe constraints and solution strategy of optimization efforts.


FIGURE 9 Illustration of tasks and hybrid tasks. (From Ref. 82.)


The proposed formulation results in a mixed-integer linear programmingproblem that is solved using OSL/GAMS (71). The optimization simultaneouslyyields the optimal sequence, the optimal hybrids, and the optimal transforma-tions. The formulation guarantees the global optimality in all cases because itinvolves only linear expressions of the continuous variables. An industrial appli-cation of the methodology is given in Section 6.

5. REACTION–SEPARATION AND REACTIVE–SEPARATION

SYSTEM SYNTHESIS

Particularly strong and complex interactions prevail among reaction and separa-tion systems that are generally not at all or not fully exploited as a result of theapplication of the available synthesis methods for reactor networks and separa-tion systems in isolation. The lack of generality in the synthesis methods is a tribute to the nonlinear process models required to capture the reaction and separation phenomena as well as to the vast number of feasible process designcandidates. These complexities even make it difficult to synthesize the decom-posed subsystems, which are typically reactor networks, separation systems,reactor–separator–recycle systems, and reactive separation systems. The devel-opment of reliable synthesis tools for these sub-systems is still an active researcharea.

Potentially beneficial interactions between reaction and separation are mani-fold and can result from either separation within the reaction zones of the process(reactive separation systems), appropriate combinations of separate reaction andseparation equipment (reactor–separator systems), or a combination of these. Therealization of reactive separation options requires the introduction of mass sepa-rating options into the reaction equipment and results in additional phases orstates that need to be considered in the process synthesis. Common examples arethe introduction of solvents (reactive extraction, reactive absorption), strippingagents (reactive distillation), solids (reactive crystallization), and diffusion barri-ers (membrane reactors). In general reactor–separator systems, a number of sep-aration tasks exist that enable the generation of intermediates or products of aperformance-enhancing nature in the context of a general reacting system. Theoptimal separation tasks cannot generally be performed by the same unit opera-tion, and a combination of unit operations is likely to be present in optimal-flowschemes.

It is obvious that the simultaneous inclusion of all possible reactor–reactiveseparation–separator design options into an automated design framework quicklyleads to combinatorial explosion. In combination with the nonlinear models usedto describe the reaction, mass, and heat transfer phenomena that occur in the pro-cessing units, the resulting synthesis problem is beyond the scope of existing opti-mization technology, even for relatively small problems. This has led to the


development of synthesis tools, which are constrained to limited reaction and/orseparation options; i.e., the general synthesis problem needs to be reduced in sizesignificantly before any systematic synthesis methods can be applied (e.g., Refs. 14,66, and 72–76). This problem-size reduction poses a major challenge: identifica-tion of the relevant modeling information. Currently, the design engineer needs toanticipate the important phenomena that dictate the performance of the systemwithout the help of systematic screening methods. This demands a thoroughunderstanding of the effects of component concentration and temperature manipu-lations on the system performance on the basis of process knowledge available atthe earliest design stage. A typical example is the effect of component additionand removal to and from the reaction zones and the corresponding introduction ofreactive separation options. Although an experienced designer can spot the rele-vant trade-offs for simple reacting systems and link them with the physical prop-erty differences to be exploited for separation, this task becomes nontrivial forsystems with complex reactions where many components are present. However,the inclusion of all relevant design options is vital for the synthesis of optimalprocess configurations, because any opportunity excluded at this stage willinevitably lead to an inferior design.

In view of the problem under consideration, Linke and Kokossis (77,78)propose a design framework that accounts for two synthesis stages: a screeningstage that allows for identification of the dominant design trends in combinedreaction and separation task systems, followed by a design stage that incorporatesthe insights gained in the screening stage and allows for the generation of optimalreactor–reactive separation–separation configurations.

The screening stage aims at replacing the early problem decompositionbased on intuition with the introduction of simple separation models that enablethe investigation of major trade-offs in the context of reactor design. This is facil-itated by structural optimization of relaxed reactor/separation task network super-structure models. Allowing any separation tasks to be present in the networks, themost important composition manipulations in the process network can be identi-fied, regardless of separation feasibility. By introducing constraints on specificseparation tasks, sensitivity studies allow one to investigate their impact on theperformance targets and process layouts.

Rather than final process layouts, the screening stage generates insights intooptimal component separation, removal and recycle policies, along with reactormixing and contacting patterns. The obtained process layouts feature reactor andseparation task combinations, and only those important tasks are subsequentlyassessed for separation feasibility. The generated process design information isprocessed in a design stage by introducing the additional modeling informationrequired for candidate reactive separation and mass exchange operations into arefined superstructure network model and by excluding infeasible separationtasks. In other words, more modeling rigor is added, to give a more detailed


account of the relevant reactive separation and separation options, while the combinatorial size of the problem is reduced by the irrelevant design options.

The difference in targets between the screening and the design stages reflectpotential savings of separation options termed infeasible upon analysis of thephysical properties and give incentives for the investigation of novel separationtechniques for the system under consideration, such as the development of novelsolvents, membranes, or hybrid separation systems. Both synthesis stages allowfor iterations to incorporate the insight gained during the synthesis process.Reasons for iterations are the identification of additional problem constraints, theremoval of process units without positive impact on the system performance, andthe inclusion of refined process models. Such refinements might result from thelimited validity of the available reaction models in optimal design regions. Inorder to gain confidence about the designs, a wide range of operating conditionsand structural design alternatives needs to be searched in the process synthesisstage. In many cases, the optimized operating conditions do not coincide with theregions of model validity, and the model trade-offs might not reflect the real sys-tem trade-offs. This creates the need for new process models for the operationregions identified as of interest, which may introduce additional model nonlinear-ities or an additional set of reactions and components. The proposed synthesisrepresentation, along with the robust optimization framework, is not limited bythe complexity of such models. The refinement of reaction models demands addi-tional experiments to be carried out. The synthesis methodology is efficient insuggesting the concentration and temperature regions of interest for processdesign, which helps to minimize the experimental efforts.

In both synthesis stages, superstructures of generic units are formulated,and the performance targets as well as a set of design candidates are obtained sub-sequently via robust stochastic optimization techniques.

5.1. Unit Representation

The basic elements of the synthesis representation are the generic reactor/massexchanger (RMX) units and the separation task units (STUs). The underlyingphenomena exploited in chemical process design are reaction, heat exchange, andmass exchange. They are generally exploited simultaneously (nonisothermal mul-tiphase reactors and reactive separators), in combinations (nonisothermal homoge-neous reactors and mass exchangers), or in isolation (isothermal reactors andmass exchangers, heat exchangers), depending on the particular system underinvestigation and the location of the operation within the flowsheet. This workemploys a generic reactor/mass exchanger unit for a flexible and compact syn-thesis representation of all possibilities of phenomena exploitation. The RMXunit follows the shadow compartment concept developed for nonisothermal multi-phase reactor network synthesis (Section 3.2). The unit consists of compartments


in each phase or state present in the system under investigation, and the streamsprocessed in the different compartments of the generic unit can, by definition,exchange mass across a physical boundary, which can be either a phase boundaryor a diffusion barrier. Each compartment features a superset of mutually exclusivemixing patterns through which a compact representation of all possible con-tacting and mixing pattern combinations between streams of different phases canbe realized within a single generic unit. Along with decisions on the existence ofmass transfer links between compartments of the same generic unit and decisionson the consideration of reaction phenomena in the different compartments, a sin-gle RMX unit enables the representation of a reactor, a mass exchanger, a reac-tive mass exchanger, or a combination of these. Each compartment receives inletstreams within its state and corresponds with the compartments in the other statesvia diffusional mass exchange links across the state boundaries. The effluentsfrom the compartments either leave the unit or are recycled within the RMX unit.Recycles can be present within a given state or across the state boundaries if tech-nically possible, e.g., by reboiling, condensing, throttling, or compression.Throughout this work, all states that can receive streams from a reference stateare termed adaptable states to the reference state. Mutually exclusive mixingoptions considered for the compartments include well-mixed and segregated flow.In order to avoid a mathematical model with both differential and algebraic equa-tions, segregated flow behavior is approximated with a serial arrangement ofwell-mixed units of equal volume with the resulting cell model consisting of onlyalgebraic equations. In each compartment, all inlet streams are connected to allmixers prior to the subunits. Each subunit effluent stream is split and connectedto the subsequent subunit as well as to the final product mixer of the compart-ment. The recycle streams from the final product splitter distribute among allwell-mixed units employed in the compartments of the adaptable states.Temperature effects are accounted for in the compartment models in accordancewith the profile-based and the unit-based synthesis approaches (Section 3.2).

In contrast to the rigorous representation of reaction and mass transfer phe-nomena by RMX units, the separation task units (STUs) represent venues forcomposition manipulations of streams without the need for detailed physicalmodels. In accordance with the purpose of any separation system, the separationtask units generate a number of outlet streams of different compositions by dis-tributing the components present in the inlet stream among the outlet streams. Inthe screening stage, the aim is to identify the separation tasks, which have a pos-itive impact on the performance of the reaction–separation system. Only sharpseparations are considered between component lumps; i.e., each lump can bepresent in only a single outlet stream of the separation task unit. Sharp separationsdo not impose a limitation in terms of composition attainability, because any non-sharp separation can be obtained by sharp separation followed by stream mixing.Moreover, potential benefits that might arise from nonsharp separations as compared


to sharp separation followed by mixing cannot be quantified, because a particularseparation process is not specified at this synthesis stage. In the design stage, theseparation task unit is adoptable for the representation of separation tasks that can be performed using particular separation processes. In this case, the possibledistribution policies of component lumps to the outlet streams are constrained bythe separation orders of the separation process. Nonsharp separations arisingfrom operational constraints on the separation tasks can easily be accomplished.The separation task unit performs a set of feasible separation tasks according tothe separation order to define a set of outlet streams. Depending on the order inwhich the tasks are performed, a variety of processing alternatives exist for a sin-gle unit.

5.2. Process Representation

The generic synthesis units presented earlier allow for the representation of allsections of general processes involving reaction and separation, i.e., reaction sec-tions, reactive separation sections, mass exchange sections, and sections per-forming separation tasks. The aim of the superstructure network generations is toprovide for a venue that enables the simultaneous exploration of all functionali-ties of the different synthesis units along with all possible interactions amongthem. The superstructure formulations evolve in the different synthesis stagesaccording to the insights into the process obtained at the previous stages. Noveltyis accounted for in the superstructures because the representations are not con-strained to conventional process configurations but instead include all possiblenovel combinations of the synthesis units. The superstructures feature a numberof generic RMX and separation task units as well as raw material sources andproduct sinks, the interconnections among which are realized by two types ofstream networks: Intraphase streams establish connections between the synthesisunits, products, and raw materials of the same state, whereas interphase streamsestablish those connections across the state boundaries; i.e., the source and sinkof such a stream belong to adaptable states. The compositions of the streamsremain unchanged on state transition. The transfer of material across the stateboundary potentially requires the addition or removal of energy to or from theinterphase streams. In this case, an energy transfer unit such as a total reboiler ora total condenser, a throttle or a compressor is associated with those mixersreceiving a stream from a different state.

Superstructures formulated in the design stage facilitate RMX units for rep-resentation of the relevant reaction, mass exchange, and reactive separationoptions obtained from the screening stage and separation task units associatedwith the different unit operations that allow one to perform the desired separationtasks. This incurs superstructures with additional states resulting from the inclu-sion of separating agents and diffusion barriers. Figure 10 illustrates two process


configurations that can be obtained by eliminating interphase and intraphasestreams, as well as generic units from a superstructure network featuring twoRMX units and two separation task units in two states. A flow scheme featuringa catalytic reactive distillation arrangement is shown in Figure 10a. The designshown in Figure 10b does not feature any interphase connections, and reaction

FIGURE 10 Design instances from the superstructure.


occurs in both RMX units. A separation task unit facilitates the intermediate sep-aration of components in between the RMX units; a plug-flow and a well-mixedRMX compartment are arranged in series, followed by a separation task unit rep-resenting a separation sequence to separate the raw material from the product andthe by-products.

5.3. Optimization

In order to establish a basis for optimization, the reaction and separation super-structure is formulated as a mathematical model that involves the componentmaterial balances for the RMX units, the mixers prior to the separation task units,and the product mixers, as well as the energy balances for the RMX units (for theunit-based nonisothermal representation). The formulation incorporates modelsfor the reaction kinetics, physical property and mass transfer models, shortcutmodels and regression expressions for equipment sizing and costing, and theobjective function. General relationship-modeling terms will introduce nonlinear-ities into the superstructure network model. As for the reactor network synthesisproblems, the reaction and separation superstructures are optimized using sto-chastic search techniques. Linke and Kokossis (77,78) have studied the perform-ance of SA and tabu search (TS) for this type of problem. They found robustperformances for both algorithms and TS to be the more efficient search meta-heuristic. A detailed description of the implementation of the stochastic searchalgorithms is given in Ref. 79.

For illustration purposes, consider the Williams–Otto flowsheeting prob-lem (80). In the conventional design, raw materials A and B are fed to a reactor,where the following reactions occur:

(1) A � B → C

(2) B � C → P � E

(3) P � C → G

The reactor effluents are cooled and fed to a decanter for removal of heavy wasteG, which requires further treatment. From the remaining mixture, the desiredproduct P is removed via distillation. The unreacted raw materials as well asunwanted byproduct E and a fraction of P are partly recycled back to the reactor.As components E and P form an azeotrope, an amount of the desired productequivalent to at least 10% weight fraction of E is lost through the purge, which isused on site as a fuel. The volatilities �i of components i in the system have thefollowing descending order: �P > �E > �C > �B > �A > �G. The reaction rates ofcomponents A, B, C, P, E, and G respectively are functions of the weight fractionsX and given by the vector


R(X) � [�k1XAXB; �k1XAXB �k2XBXC; 2k1XAXB � 2k2XBXC � k3XPXC;

k2XBXC � 0.5k3XPXC; 2k2XBXC; 1.5k3XPXC]

The complete data are given in Ref. 79. The objective of the synthesis exercise isto find the designs that maximize the annual profit of the process for a minimumproduction rate of 400 kg/h of component P.

5.3.1. Screening Stage

The stochastic search of superstructures featuring three RMX and three uncon-strained separation task units yields an absolute profit target of around $618 k/yr.A variety of design alternatives exist that can achieve the targets, featuring one,two, or three reactors and component distributors. The process design require-ments as identified in the screening stage can be stated as follows:

Separation and removal of by-products E and GSeparation and removal of desired product PSeparation of components A, B, and C from the reacting mixture and dis-

tribution among the reactor unitsExcess of component B in the reaction zones minimizes by-product forma-

tion (low concentration of product P)Reaction zones exhibit plug-flow behavior

5.3.2. Design Stage

Based on the insights gained in the first design instance, appropriate separationvenues can be identified and included in the search. Distillation enables separa-tion of mixtures according to the order of volatilities and hence allows separationin support of the raw material and intermediate recovery. To avoid fouling, Gneeds to be decanted prior to the operation, which can be achieved at a low costin a decanter. However, component P forms an azeotropic mixture with compo-nent E, resulting in a loss of desired product to the low-value fuel. A solvent isavailable that allows selective extraction of desired product P from the mixture.The equilibrium relationship and maximum solvent loading are given in Ref. 79.The design stage therefore considers a superstructure of RMX units and STUs tocapture the reaction–extraction–distillation–decanting system.

Stochastic optimization yields a target performance of around $433 k/yr forthe system. Designs with performances close to the target can be grouped intotwo main categories according to their use of the solvent: reactive extractiondesigns and reactor–hybrid separation designs, each achieving performancesclose to the target. Two sample designs are illustrated in Figure 11. Designs uti-lizing the distillation–extraction hybrid achieve higher selectivities in the con-version of the raw materials to product P than do the reactive extraction designs,


in which valuable components A, B, and C are lost to the fuel stream. Bothdesigns achieve a virtually complete recovery of product P. The extraction of thedesired product from the concentrated distillation overheads allows for highersolvent loadings and consequently a significantly reduced solvent flow as com-pared to the reactive extraction case. However, the benefits in terms of higherselectivities and lower solvent flows are offset by the expenditure required for thedistillation operation.

FIGURE 11 Process design candidates from the design stage.


The methodology is illustrated with a further application to the design ofactivated sludge process systems in Section 6 and has seen further applications in multiphase reaction–separation, reactive distillation, and membrane networksynthesis (79).

6. INDUSTRIAL STUDIES AND REAL-LIFE APPLICATIONS

The methods described in Sections 3 through 5 have been tested on a number ofcase studies. This section presents three applications to problems of complexitytypically encountered in industrial practice that have been performed using thesynthesis schemes, partly in collaboration with industrial partners.

6.1. Reactor Design: Ammoxidation of Propylene

to Acrylonitrile

Mehta (34) has carried out a reactor network optimization study to find improveddesigns for the production of acrylonitrile in a collaboration between UMIST andone of its industrial partners. Most industrial installations employ fluidized-bed reac-tors (BP/Sohio process) with a well-mixed reaction zone. Previous process improve-ments have mainly resulted from better catalysts, which have produced an increasein yield from 58% to around 80%. The reaction model employed in the optimizationstudy is taken from Ref. 81 and considers seven reactions and eight components. Air,pure oxygen, and propylene are available as raw material streams. The optimizationstudy assumes negligible pressure drop along the reaction sections, isothermal andisobaric operation, and negligible mass gas–solid transfer effects.

Acrylonitrile yield has been employed as the objective function. Initial tar-geting has shown that the network yield increases with the catalyst volumeemployed and asymptotically approaches a yield of around 90% toward infinitelylarge catalyst volumes. For very large volumes, the well-mixed reactor achievesthe performance targets. Because only smaller catalyst volumes are relevant inindustrial practice, a targeting study is carried out to develop the performance tar-gets as functions of the catalyst volume utilized. For comparison, the targets forthe conventional well-mixed reactor are developed alongside the targets that canbe realized with reactor designs featuring novel unconventional mixing and feed-distribution policies. The targeting curve is shown in Figure 12. The targetingcurve shows that the performance of the CSTR reactor significantly degradestoward smaller catalyst volumes, i.e., toward the industrially more relevantdesigns. Only a slight degradation of reactor yields with catalyst volume isobserved for the novel designs. This suggests that the acrylonitrile yield can bekept high even with significantly reduced catalyst volumes, provided an appro-priate reactor network is selected. When compared to the conventional reactor,the novel designs showed improvements in yield of up to 17% for practical values


of catalyst volume. Mehta (34) observed a change in reactor design patterntoward smaller catalyst volumes, gradually moving single CSTRs to combina-tions of PFRs and CSTRs with distributed feed streams.

The study was carried out over a short period of time and resulted in animpressive enhancement of insights into how an optimal reactor should bedesigned for maximum performance at low catalyst volumes. A commercial reac-tor design has subsequently been developed on the basis of the insights gained inthis study.

6.2. Complex Distillation

This case study (82,67,68) included a feed containing nine components and thathad to be separated into four products: A (C4 fraction separated as lights), B (iso-pentane and some n-pentane), C (C6 fraction), and D (heavier components, C7+).Minimum energy cost is used as the synthesis objective, and the optimizationresults are summarized in Table 1. Designs I and II represent the favorableschemes. Design I separates the lightest product and uses a prefractionator for thedownstream separation in hybrid B/C/D. Design II employs a prefractionator inhybrid AB/C/D to separate plentiful products, C and D, early in the sequence.Design III, the best simple sequence, separates the plentiful product, D, first andfavors difficult separation, B/C, in the end. When simulated rigorously, the rank-ing of the design candidates remains identical to that from the conceptual designlevel. The initial energy costs match very well with the costs from rigorous sim-ulation. An interesting point is that in this example the best simple sequence is closeto an indirect sequence (Design III) but the best complex scheme corresponds to

FIGURE 12 Targeting curve for ammoxidation of propylene to acrylonitrile.


a direct sequence (Design I). The lowest pressures are selected for all the tasks inthe designs. This results in the most favorable separation factors. The complexcolumn configurations require high pressures, and the benefits of thermal cou-pling are counterbalanced by the requirement of higher reflux. Pressure effectswere found to dominate the results and to minimize the thermal coupling.

6.3. Combined Reaction and Separation Systems:

The Activated Sludge Process

The reaction and separation synthesis approach of Section 5 has been adopted tothe problem of activated sludge process design (83). The conventional designs aswell as all novel schemes for combined oxidation/denitrification of wastewaterare explored. The process is optimized using a novel methodology for optimalreaction/separation network synthesis, supplied with a comprehensive kineticmodel (84). The activated sludge process is synthesized using the systematic

TABLE 1 Promising Designs for the Industrial Complex Distillation Case Study

Energy PressureDesign cost Design Column (bar) Trays

I 6.57 1 4.9 30

2 1.8 4

3 1.8 82

II 6.60 1 3.5 4

2 3.5 90

3 4.9 29

III 7.60 1 1.6 23

2 4.9 29

3 1.8 54


reaction/separation network synthesis framework of Section 4 to determine theoptimal biochemical reactor network design along with the sludge separation andrecycle policies. The representation provides for all possible mixing, contacting,and reaction/separation features, and the superstructures are optimized throughapplication of the simulated annealing algorithm. Oxygen mass transfer is mod-eled using film theory. A weighted objective function is formulated so as to min-imize both the effluent COD and total nitrogen content.

An initial optimization study without considering any bounds on the reactorvolumes yields a reference target (lowest value of objective function) of 100%, cor-responding to a 96.8% reduction in COD and 84.8% in N. However, the reactionvolumes observed in the designs that achieve the targets are unpractically large. Forthis reason, volume bounds are introduced in a subsequent study to achieve themean hydraulic residence time of the conventional processes. Surprisingly, the tar-get did not deteriorate significantly, the new target for the objective value being101.4%. COD was reduced by 97.4% and N by 84.9%, a performance much bet-ter than those attained by any conventional process, especially as far as denitrifica-tion is concerned. An inspection of the structures revealed a design pattern: Manystructures did not include an anoxic reactor and yet yielded very good denitrifica-tion. Moreover, the system appeared to seek ways to hinder or control the dissolu-tion of oxygen, in contrast to common practice, which aims at dissolving themaximum amount possible into the oxic reactors. To find out how such good den-itrification could be accomplished without the inclusion of an anoxic reactor, thedetailed oxygen profiles within the aerated reactors (mostly PFRs) were examined.The concentration of oxygen is controlled within extremely low levels (0.1–0.3ppm), an order of magnitude less than those currently used in industrial practice(2–9 ppm). This leads to the conclusion that both organic matter stabilization anddenitrification processes occur simultaneously in these designs, due to the very lowoxygen concentrations—a policy that seems to achieve maximum efficiency whenlow volume is required. A schematic design of the system is illustrated in Figure13. Details on the synthesis study can be found in Ref. 83.

FIGURE 13 Conceptual activated sludge process design.


7. CONCLUSIONS

The drive to improve processes and to create new process routes continuouslyforce process engineers and management to develop new technology. Processsynthesis or integration has proven its value to improve processes with reducedcosts, less pollution, and lower energy consumption. Traditionally, most industrialapplications of process integration techniques have been found in developingenergy-efficient systems. After a brief discussion on common conceptual processdesign practices, this chapter illustrated recently developed process synthesistechnology for process design problems involving reaction and separation. Amajor advantage of the methods is their ability to determine robust performancetargets and to identify a variety of design options with similar close-to-target per-formances. The design tools offer decision support to the design engineer. Thiswill allow the inspection of similarities and differences among high-performancecandidates in order to select the most practical novel designs. However, for suc-cessful technology development and transfer, close collaboration between acade-mia and industry is essential.

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12

Process Intensification in Industrial Practice:Methodology and Application

Remko A. Bakker

DSM Fine Chemicals Austria, Linz, Austria

1. INTRODUCTION

In this chapter a method is explained for how process intensification can be intro-duced in a commercial company. When one would like to introduce process inten-sification, certain steps can be followed that increase the chances of success.

One very important aspect is the awareness of the drivers for introducingprocess intensification for certain processes and companies: the why of processintensification. This is explained in Section 3 in some detail, for without theknowledge of these drivers, one risks introducing process intensification ineffec-tively and inefficiently. After that, a short overview is given of the technologiesavailable for process intensification, serving as a basis for the real process inten-sification study. The method of introducing process intensification is thenexplained in Section 5, which is the main section: What steps should be incorpo-rated in a process intensification study, what a good group of participants is, andin what phase of a chemical process the study should be introduced. The chapterconcludes with two concrete examples of process intensification, one in a bulkchemical process and one in a fine chemical process.

Process intensification can bring many benefits regarding business, legislative,and environmental aspects and is therefore very worthwhile to be implemented.


2. SHORT OVERVIEW OF DEVELOPMENTS

IN PROCESS INTENSIFICATION

As discussed in Chapter 1, process intensification as a way of looking at chemi-cal processes originated in the early 1970s. It was adopted by the ICI NewScience Group in the United Kingdom, where mainly rotational fields were usedfor accelerating chemical process steps.

After the first conferences on process intensification, organized by theBritish Hydromechanics Research (BHR) Group (1), the field became a separatelyidentifiable field in chemical engineering and spread around the world. The gen-eral ideas behind process intensification—improving a chemical process via step-by-step changes—could already be found in various activities of developmentgroups in chemical engineering (in industry and at universities), though at thattime not all of these activities were called process intensification. By grouping theactivities in this field under an umbrella name and by organizing meetings andconferences dedicated to the topic, many new ideas, developments, and applica-tions were discovered, and they are still being discovered. This is one of the rea-sons the field of process intensification developed rapidly at the end of the 20thcentury.

3. WHY WORK ON THE INTENSIFICATION

OF CHEMICAL PROCESSES?

As with all developing technologies, there is always more risk in introducingthem in commercial plants than in using the well-known, traditional technologies.Some conservatism in commercial industry is always present. A very good illus-tration of that can be seen in Figure 1 of Chapter 1, where a 16th century chemicalprocess is compared with contemporary chemical processing. Why then wouldone work on these kinds of developments?

By introducing step-by-step improvements in a process, process intensifi-cation offers a strong possibility to fulfill the current business, legislative, andenvironmental requirements. These aspects are becoming more and more impor-tant in contemporary chemical industry.

1. Business aspects are changing due to the expansion of the globalizationof the world economy: There are more potential competitors in thisenlarged market. The best low-cost position is necessary to obtain andmaintain market share. This is true for the fine chemical and pharma-ceutical industries, because transportation costs are only a smallamount of total production costs and because the total global marketmay in theory be supplied by producers from any place in the world.On the other hand, for the bulk chemical industry as well, small cost


advantages per kilogram of product can make the difference betweenprofit and loss.

2. Legislative and environmental aspects are usually strongly linked.They are getting a lot of attention in discussions on the license to oper-ate (LTO) and sustainable technology (see Chapter 14).

These two aspects are also reflected in the visions developed on this issuein the United States in the so-called Vision 2020 program (see Appendix 1). Onecan see that all aspects of introducing step-by-step changes are mentioned. Andin the European Union many current activities are based on the topic of processintensification, reflected, for example, in the European-Supported ResearchProjects (see Appendix 2). On the industrial side, a lot of effort is being given tothe topic of process intensification: the European Chemical Industry Council(CEFIC) has process intensification as one of its technology focus points in theso-called SUSTECH program (see Appendix 3).

The intensification of chemical processes can help reduce the risks ofchemical processing and lower the energy demand and waste production. Thesepositive possibilities, which are evidently also linked to the business aspects, havecontributed to the growing use of process intensification. This is also expressedin the paper of Elverding (2) at the Defacto conference on “Profit, Planet, andPeople” in combination with process intensification, reflecting the benefits forbusiness and environment.

Many different aspects relating to business, legislation, and environment canbe defined. These are given in Table 1. In this important table one can find manyof the benefits of process intensification summed up. All of these aspects in the endhave to do with business, legislation, or the environment and show how processintensification can have a strong positive influence on these three main drivers andanswers the question of why one would one choose to intensify a process.

The overview in Table 1 holds, in principle, for any type of chemical indus-try. Additionally, different types of chemical industries may have different focuses,depending on their starting position and on their business requirements. Knowl-edge of the perspectives helps in determining the main cost drivers from the listgiven in Table 1 and helps therefore in finding the starting points for processintensification studies. Table 2 presents an overview of some of the main types ofchemical industries and the specific demands they posed on process intensifica-tion when applied in these industries. A distinction can thus be made between:

Large-scale chemical industries, producing—mostly in a continuous way—hundreds of kilotons of product

Fine chemical industries, which can be split into:Multiclient chemicals, producing up to hundreds of tons of chemicals for

various customers, mostly in a batch way


TABLE 1 Positive Results of Process Intensification on Various Aspects

Legislativerequirements

Business (environment Aspect (costs) and safety) Environment Examples

Energy X X X Less heat loss via integration of units (e.g., reactive distillation)

Efficient heat transfer possible (high heat transfer rates)

Space X X X Reduction in floor space neededLess pipingReduced maintenanceFaster grade/product changesEase of dismantling and/or shipping

Number of X X X Fewer unitsprocess Less dangersteps Fewer possibilities for yield losses

Cost X Many aspects together may result in a lower costTheoretical goal: no limitation, other than kinetics,

gives optimal yield and least wasteEmissions X X Fewer possibilities for yield losses

Fewer process steps may result in lowered emissions

Waste X X X Fewer possibilities for yield losses means less wasteTheoretical goal: no limitation, other than kinetics,

gives optimal yield and least wasteLess holdup X X Intrinsically safe design: prevent large volumes, little

intermediate storage


Process X Fewer units, hence less possibility for interruptions malfunctioning of partsand downtime

Fewer stoppages due to lowered maintenanceFaster grade/product changes

Flexible X Easier to adapt quickly to new process conditions feedstock due to smaller equipment, fewer intermediate specifications material or grade changes

Recyclable X X X Easier to dismantle and/or shipEasier to adapt quickly to new processes due to

smaller equipment— X X X Many other aspects can be thought of


TABLE 2 Process Intensification for Different Types of Industries

Fine chemicals

Large-scale chemical Multiclient Custom Pharmaceutical industries chemicals manufacturing industries

Characteristics

Up to hundreds of kilotons Up to hundreds of tons Up to tens of tons Up to tens of tons Continuous Mostly batch, some Mostly batch, some Mostly batch

parts continuous parts continuousTotal profits are 5–30% Total profits are 20–40% Total profits are 30–50% Total profits are 30–70%

turnover turnover turnover of turnoverMaterial and energy costs Material and energy costs Material and energy costs Material and energy costs

are 70–90%, labor costs are 70–90%, labor costs are 50–80%, labor costs are 10–30%,labor costs 10–30% of total costs 10–30% of total costs 20–50% of total costs 70–90% of total costs

Why PI?

Material and energy cost Higher efficiency, labor Higher efficiency, labor Labor cost reduction (also reduction cost reduction cost reduction in cleaning procedures)

Higher efficiency, higher Higher production capacity, Higher production capacity, Higher production production capacity material cost reduction material cost reduction capacity

Space and Energy cost reduction, Optimal price and time Optimal price and time of waste reduction optimal price of delivery delivery

How?

Focus on materials and Focus on labor and Focus on labor and Focus on number of energy (large impact efficiency (efficiency so efficiency (efficiency so steps (labor andon costs) that all steps in the that all steps in the cleaning intensive)

process are well balanced) process are well balanced)


Custom manufacturing, producing up to tens of tons of a chemical for one specific customer, mostly in a batch way

Pharmaceutical industries, producing up to tons of a active pharmaceuticalingredient (API), mostly in a batch way

Table 2 shows that different chemical industries have different cost focus-es. It is important to know the main cost contributors of a certain process, becausethis is a good way to find step-by-step changes in the overall costs of a process.For different types of chemical industries one can find an idea of the correspon-ding main drivers in Table 2 (various examples of these main drivers can also befound in the goals defined in Appendix 1).

Of course, apart from these general features every process is unique andrequires an individual investigation and definition of the cost drivers before onecan start a process intensification study. This will be discussed in more detail inthe next section.

Process Intensification is all about step-by-step changes. At what stage anoptimization in a process is called a step change depends not only on the type ofbusiness but also on the maturity of a product in the lifecycle and the goalsdefined by the three main aspects of process intensification (business, legislativerequirements, and environment). For example, the step change that one wouldlike to achieve in efficiency, hence in optimizing a process without changing thehardware drastically, depends on the lifecycle of a product. This is shown inFigure 1. This last aspect should be taken into account when defining the reasonsfor performing a process intensification study.

4. KEY FEATURES OF PROCESS INTENSIFICATION

TECHNOLOGIES

As stated earlier and in other chapters of this book (e.g., Chapter 1), processintensification is making step-change improvements in processes. Many generalpossibilities are available for reaching that goal. A large number of technologiesthat have the characteristics of process intensification involve one or more of thefollowing three imperatives.

1. Make it small.2. Combine.3. Intensify the driving force.

Examples of these can be found in all chapters throughout the book. For exam-ple, belonging to imperative 1 are microreactors (Chapter 5), minireactors(Chapter 6), and compact heat exchangers (Chapter 4). Imperative 2 represents a large field comprising the combining of reaction and extraction, reaction andseparation (Chapters 8, 9, and 10), reactive extrusion, reactive heat exchangers


(Chapter 4), and many more. The last imperative covers technologies that use cen-trifugal fields for contacting (Chapters 2 and 3), separations, and crystallization orthat involve extremes of pressure and temperature, ultrasound waves, microwaves(for e.g., drying) and electric fields (for e.g., separation or dispersion).

Basically, using these technologies one would like to move forward to thetheoretical optimum of a chemical process, which is that there are no other limi-tations than chemical kinetics. Normally a chemical process is influenced bymore than just kinetics: hydrodynamics (mixing), heat transfer, and mass transferdetermine the quality of the process. Process intensification focuses on removingthese three limitations to reaching the goal of kinetically limited processes. Thisis schematically depicted in Figure 2.

These three basic imperatives can be used to begin thinking about when totry to intensify a process in a chemical plant. They are also part of another areaof chemical engineering: safety aspects of plant designs. In the 1980s, effort wasput into developing strategies for designing so-called “intrinsically safe designs,”which means that the process designs are such that dangers are intrinsically minimized so that the safety of a process does not depend on safety devices.

FIGURE 1 Step changes in efficiency as a function of the lifecycle of a prod-uct (fictitious curve to show the differences).


Examples of strategies for safe designs are: prevent large volumes, intensify con-tact, reduce the need for aids (like solvents), use little intermediate storage. Onecan immediately see the similarities between intrinsically safe design and processintensification. This also means that applying process intensification techniquesis beneficial for process safety.

5. INTRODUCING PROCESS INTENSIFICATION

IN A CHEMICAL PROCESS

This section describes how one may intensify a process, or, said differently, howprocess intensification can be introduced for a process.

When one would like to introduce process intensification for a process it isimportant to consider the development phase a process is in. Roughly said, thefollowing phases can be discerned for a general process:

New ideas for a process → Determining process chemistry → Pilot plants studies → Plant design → Plant startup → Debottlenecking or troubleshooting.

Every phase requires a somewhat different approach, because the boundaryconditions and possibilities for intensifying a process differ for each phase. Whatfollows is a short description of the phases and the approaches.

FIGURE 2 Process Intensification: Approaching the kinetic limit by using thethree imperatives: (1) Make it small; (2) combine; (3) intensify the driving force.


1. New ideas for a process: A product and a chemical route are defined.This is the phase in which the core of the process is determined. This means thatit is already very important to think about the consequences that selecting a cer-tain pathway has on the future final plant and on the plant costs. Choosing a non-optimal pathway may result in a much more expensive final plant design than ifanother pathway (or, e.g., another solvent or catalyst) were to be chosen. This isthe phase in which it is still relatively easy and cheap to make changes that pro-duce the most intensified process. To be able to know the effect of a chemicalroute on the final plant requires a multidisciplinary team to be involved at thisphase, not just chemists, because more aspects play a role than chemistry. Thismultidisciplinary team should consist of at least chemists and technologists.

2. Determining process chemistry: The specific chemical parameters ofthe chosen process route are determined. In this phase, the route itself is more orless frozen. The important parameters of the process are determined, such as therequired temperatures, pressures, and concentrations of substances in the processsteps. Making changes and optimizations in this phase is still relatively inexpen-sive. The effects of this phase on the efficiency of the final process can be large.Applying intensification possibilities in this phase will reduce greatly the costs ofthe final process. In this phase the kinetic limit, as given in Figure 2, is deter-mined, hence this is the last phase where the intrinsically optimal route may befound—the route with the least by-products and the least waste generation (fromauxiliaries such as solvents and from by-products). It is also the last phase inwhich, apart from the optimal kinetics, the optimal process conditions can befound that cause the fewest problems in technical implementation. Therefore, amultidisciplinary team is required to determine the optimal process route, not justchemists. Again, this multidisciplinary team should consist of at least chemistsand technologists.

3. Pilot plants studies: The process is scaled up to pilot-plant size toobserve scale-up effects and to determine parameters for the large-scale plant(though one should take into account that this phase is skipped more and more inprocess designs, to save time and money). In this phase, the basic technologies ofthe process are checked, and the scalability of the process route developed in theprevious phases is checked. Changes in this phase are still relatively easy to makebut more costly than in the previous phases, because pilot-plant studies them-selves cost money, and larger changes often mean a delay in the developmentprocess because the previous phase will also have to be partly redone as well. Inthis phase one can look for the limitations that technical equipment pose for theprocess route. Heat and mass transfer limitations usually result in a poorerprocess performance than the optimal lab conditions gave. In the pilot phase onecan try to optimize the equipment to overcome the limitations. For that purposeone can make use of intensified process technologies, as described throughout thisbook. These new technologies can best be tested during the pilot scale, to collect


experience with using these technologies and for the further scale-up to plant size.Because process technology, chemistry, and future plant equipment play a role, amultidisciplinary team is required consisting of at least chemists, technologists,and process engineers.

4. Plant design: A commercial-scale process is developed and designedin detail. Experience gathered in the previous phases is used for developing thelarge-scale plant. In this phase, it is difficult to make large changes in the process,because then the previous phases would have to be partly redone and the costs andtime spent in the previous phases are lost. It is important that intensified tech-niques are used as much as possible, since normally more limitations in theprocess will arise at plant scale than in the lab or at pilot-plant scale. One wantsto approach the kinetic limit as much as possible.

5. Plant startup: The commercial plant is started. In this phase, no actionsfor process intensification studies can be defined.

6. Debottlenecking or troubleshooting: While running the process, prob-lems that may occur are solved (troubleshooting) or larger optimization projectsstarted (debottlenecking). In this phase, the plant is built and the process is run-ning. One gains experience from the large-scale process. Small changes in theprocess can be made, usually as troubleshooting projects or in creep projects.These small changes, though important, are not the changes meant in processintensification. When one would extend the plant capability with larger steps, thisis usually called debottlenecking, which mostly involves investment in the adap-tations for the bottlenecks in the process. In this phase one also tries to find pos-sibilities to really make step changes in the cost efficiency of the process, whichmay be called process intensification. It again requires a multidisciplinary team,because all aspects play a role, from chemistry to process engineering.

It can be seen that a multidisciplinary team is very important in almostevery phase when the goal is to develop an intensified process. Of course, manyof the phases just described in practice run partly in parallel, to speed up thedevelopment process. This means that it becomes even more important to devel-op the most intensified process using multidisciplinary teams as early as possible.

In all phases of the process development just described one has opportuni-ties to introduce intensification technologies to intensify the process. In phases 1and 2, one should not forget that the choices made are crucial for the processafterwards. Also, in these phases one should already be considering the conse-quences of the choices on the commercial plant and the possible scale-up diffi-culties using specific chemical routes. The steps these teams can take to arrive atnew, intensified solutions are given schematically in Table 3.

Step 1. The first step is to define the goal of the study.

What is the time frame of possible implementation? When the study is astrategic study, the time frame can be many years. The goal then is to


TABLE 3 Steps Needed to Develop New, Intensified Solution for Processes,from Pilot Phase to Debottlenecking Phase

1. Goal definition

Determine the scope The goal of the studyUnit and section of the plant or the

entire plantAll the people involved (also; the

internal customers)

2. Scouting of ideas

Identify the thermodynamics, For existing process: from known data

kinetics, and balances (mass, For new process from literature andenergy) plus the cost factors thermo databanks

Determine the theoretically optimalprocess (no limitations)

Identify the breakthroughs In multidisciplinary teamsMake cost estimates of different options

3. Selection

Select the most feasible option All the people involved (also, theinternal customers)

4. Detailed design

Identify the detailed Databanks and new measurementsthermodynamics and kinetics Concluding the intensification project

Final design of the unit or plant

have new concepts ready for the future. In that case one can also usetechnologies that are still in development at the present date. When thegoal is to achieve direct implementation, the time frame is short and oneshould use the available data and available technologies to optimize aprocess. The budget necessary for the latter study is then also differentfrom that for the former type.

What part of a process should be looked at? This can be one unit, one chemi-cal conversion step, a total process, or even a combination of processes thatare interlinked through streams of intermediate products or energy streams.

Step 2. The second step is the actual study: the scouting of the ideas. Thisphase consists of two subparts. The first part is data collection:

What are the kinetics and material properties in the process?What are the mass and energy balances?


What are the current process steps and their limitations?What is the theoretically optimal process (no limitations)?What are the main cost factors?

Data should also be available to all team members of the new technologies on themarket and in development. This overview should be available before the start ofthe actual idea generation:

The available technologies and technologies under development should beknown. This information should be available before starting process inten-sification studies. It is a one-time effort to collect these data and conse-quently a smaller continuous action to keep the data up to date. Overviewscan of course have various formats and sources. An example of one tech-nology from an extensive list used within DSM is given in Figure 3. Ofcourse, this book is also a good reference for available technologies.

The second part of Step 2 is the identification phase. In this phase ideas arecollected to intensify the process within the boundaries as defined in Step 1. Thisphase requires a well-balanced team of participants from various disciplines, e.g.,engineers, R&D, process technologists, chemical experts. Apart from representa-tive of various disciplines, one should ideally also have participants that know theprocess well and people completely unfamiliar with the process, to get a widevariety of ideas and suggestions. Lastly, it is beneficial to mix older, experiencedmembers with younger members who may be less biased to company-standardprocess solutions.

As a start, the existing process should be described using all the known datagathered in the first part. It is very important that the process also beexplained in functional units, instead of equipment. This means that thefunction of a step is described instead of just the details on how a piece ofexisting equipment works in the actual situation. This helps in finding thedifferences between the goal of a step and the way the equipment actuallyworks and in finding other Intensification methods to make it work better.

Next, free discussion and brainstorming begin. One can use many differentstarting points for discussions and ideas. An overview of available tech-nologies also is important (e.g., Figure 3). Also helpful would be the listgiven in Table 4.

Combine and select ideas. After brainstorming one should of course focuson the main ideas. Often various ideas can be combined to create clus-ters of ideas around main themes.

Design a draft of the new process (alternatives). After having found themain ideas, a flowsheet of possible intensified processes should bedrawn. Such flowcharts help in further evaluations and in gaining clearpictures of the ideas.


FIGURE 3 Example of one technology out of an extensive of technologiesused within DSM.


Based on the selected flowcharts (process flow diagrams) one can make afirst rough estimate of the costs of the new processes. Various standardmethods for this are available, depending also on the company’s systemsfor cost estimation. This part of Step 2 should not be forgotten, becauseit can clearly show the benefits of a new process in terms of costs, whichis one of the main drivers for changes in processes. It also helps in rank-ing ideas.

Step 3. This step involves the selection of the most feasible route, whichshould be done in a large group, including the internal customers.

Step 4. The last step is the engineering or further development of theintensified process route. This step follows the standard engineering and devel-opment processes used for developing conventional processes, although somenew technologies may be involved. Pilot plant testing of these new technologiesis very likely to be necessary.

This method has been applied within DSM for 13 of its existing processesfrom its three main business clusters (polymers and industrial chemicals, life sci-ence products, and performance materials). See Figure 4 for the scheme usedwithin DSM that follows this route. The results showed possibilities for reducingthe cost of the present processes to 60–90% of the current costs. (This was shownearlier, in Figure 5 of Chapter 1.) This shows that the way of performing processintensification studies described in this chapter can result in very economicalresults for various types of businesses (from life sciences to bulk chemicals). Twoconcrete examples are presented in the next section to give a feeling of processintensification.

TABLE 4 Starting Questions That Might Be Used in the BrainstormingPhase of an Intensification Study

Keywords Example

Intensify Make the process smallerSegment Divide into independent partsUse a different aid A different solvent, temporary shielding agentChange conditions Change pressure or temperatureChange in an early stage Prevent the root-cause upstream from where

the problem occurs Combine Combine unit operationsFix Fix one phase to prevent separation problemsAdd or remove Quickly remove or add a product or reactantsPeriodic action Change a continuous system into periodic

actionsPhase transition Use phase transition in the process


6. EXAMPLES OF INTRODUCTIONS

6.1. S-Ibuprofen

The first example of process intensification at DSM is the pilot-scale test of theenzymatic production of S-ibuprofen, a nonsteroidal, anti-inflammatory drug. Themolecular scheme is given in Figure 5. More details can be found in Refs. 3 and 4.

The conventional production consisted of many process steps, typical for afine chemical process. These process steps are given in Figure 6. The selectivityof the enzyme is rather high. The downstream processing is rather laborious,

FIGURE 4 Scheme to arrive at intensified processes as used within DSM.


FIGURE 6 Processing steps for the production of S-ibuprofen.

FIGURE 5 Production of S-ibuprofen with a racemization step and an enzy-matic conversion using the carboxyl esterase enzyme.

because it requires that the product be completely enzyme free. Furthermore, theproduct, S-ibuprofen, causes enzyme deactivation. Altogether, many unit opera-tions are required to obtain the product; due to the deactivation of the enzyme, amaximum conversion of 21% can be reached.

Removing the product as fast as possible from the reaction mixture couldpossibly prevent the enzyme from deactivating. This was experimentally tried onpilot scale using an ultrafiltration unit, parallel to the reactor. In this filtration unitthe product is separated and the unreacted components and the enzyme arereturned to the reactor. An even better option, using a true membrane reactor inwhich the catalyst (enzyme) would remain on one side and the product wouldremain on the other side, was not tested. Both options are given schematically inFigure 7.


FIGURE 8 Processing steps in the production of S-ibuprofen with the ultrafil-tration unit.

FIGURE 9 Conversion of the S-ibuprofen process with and without the ultra-filtration unit.

FIGURE 7 S-Ibuprofen production options: ultrafiltration (left) and a mem-brane reactor (right).


Performing the S-ibuprofen production process with an ultrafiltration unit,6 of the 11 unit production steps shown in Figure 6 could be skipped! The entireprocess scheme was reduced to the one given in Figure 8.

Based on kinetic modeling of the reaction, including the enzyme deactiva-tion, it was found that there would be an additional improvement, apart from skip-ping six unit operations. By removing the product quickly out of the reactingmedium, the enzyme deactivation proceeds more slowly and the conversionincreases. This has been experimentally validated in the pilot-scale system. Theresults are given in Figure 9. It can be seen that the conversion is increased from21% to 50%! This example shows that process intensification can bring a sub-stantial cost benefit to the process.

6.2. Urea

The second example of process intensification at DSM is the urea process (5). Thehistory of the urea process at DSM is rather long, as shown in Table 5. Urea is pro-duced in a two-step process. The first step is the formation of carbamate from NH3and CO2. This reaction is exothermic. The second step is the decomposition of car-bamate into urea and water. This second reaction is slightly endothermic. Bothreactions are equilibrium reactions. The conversion to urea in equilibrium is about60%. This means that substantial recycle flow is necessary to obtain sufficientoverall conversion. In the reaction section the main unit operations are:

The stripper, in which remaining CO2 and NH3 are being stripped with CO2from the product flow

The scrubber, where the reactants in the reactor offgas are stripped

TABLE 5 Development of the Urea Process at DSM

1945 Start of R&D program by DSM1956 First commercial plant (75 mt/d)1966 CO2 stripping process1974 World-scale capacity 1750 mt/d1994 Implementation of first pool condenser1995 200th unit contracted1996 Urea 2000plusTM technology1997 Installation of first pool reactor, largest

single-line urea plant in operation1998 Startup of first Urea 2000plusTM pool

reactor plant


The condensor, in which NH3 and CO2 are condensed into carbamateThe reactor, in which the carbamate decomposes into urea and water

In the old process, the Stamicarbon stripping process, these four units werepositioned above each other. The overall height of the plant was 76 m! After var-ious small and larger improvements, a new concept was developed in which thereactor and the stripper are combined in one unit, the horizontal pool reactor. Aschematic diagram and a photo of the reactor are given in Figure 10.

The layout of the new plant concept with the integrated and intensified poolreactor now has a height of 18 m, a height reduction by a factor of 3. Also, thenumber of units has been decreased. This is illustrated in Figure 11, in which thedevelopment of the urea process is depicted. In this figure it can be seen that evenan old “proven” bulk chemical process can be intensified, resulting in a muchmore compact and economical plant.

FIGURE 10 Urea pool reactor used in the Urea 2000plusTM process.


7. CONCLUSIONS

As shown in this chapter, the methodology for applying process intensification incommercial industries requires a broad interest and cooperation within companies,including a vision of the company management and individuals. Some boundaryconditions need to be met, such as a vision in place that offers the initial time andmoney to set up the necessary infrastructure for applying these new technologies:a basic knowledge of the possibilities and available technologies and of themethodology. Process intensification can then bring substantial benefit in terms ofbusiness, legislation, and environment, as is demonstrated in this chapter as wellas in the entire book, therefore making PI very worthwhile to be implemented.

8. APPENDICES

8.1. Appendix 1: Excerpt from U.S. Vision 2020

This excerpt is taken from Ref. 6.

U.S. chemical companies must innovate and change to keep competitivein the global environment. Chemical Industry Vision 2020 Technology

FIGURE 11 Development of the urea process.


Partnership (Vision 2020) is an initiative to leverage research and devel-opment (R&D) resources through the collaborative efforts of industry,government, and academe and focus them on priority targets identifiedby the chemical industry.New Process Chemistry RoadmapSection: Office of Industrial Technologies (OIT)This roadmap focuses on innovative new process chemistry to supportthe Process Science and Engineering Technology and ChemicalSynthesis technical areas within the New Chemical Science andEngineering Technology section of Vision 2020: The U.S. ChemicalIndustry. However, the roadmap is integrally connected to other sectionsof the Vision and other supporting roadmaps. Process chemistry beginsin the laboratory, but is further developed through reactor design andprocess engineering. Feedstock characteristics, catalytic mechanisms,and downstream processing (e.g., separations) all come into play indeveloping new process chemistries.Goals

Reduce feedstock losses to waste or low-value by-products by 90%.Reduce industry-wide energy intensity (energy per unit product) by 30%.Reduce total emissions and effluents from chemical manufacturing by

30%.Increase usage of C1 and renewable resources to 33% of industry-wide

carbon usage.Reduce cost of production by 25%.Accelerate introduction of new products by 15%.Reduce lead times and time to market for new products and technolo-

gies by 30%.

8.2. Appendix 2: European Union Research Program


The 5th framework of the EU research program runs from 1998 to 2000.One of the four subprograms is called the GROWTH program. In thisGROWTH program, key action 1 is called innovative products, processes,and organization. The description is given here:

KEY ACTION 1: INNOVATIVE PRODUCTS, PROCESSES,AND ORGANIZATION (Budget 731 MEuro)

a. Contributing to modernization of industry and adaptation tochange, through the combined effects of improved industrialcapability and innovation capacity, while introducing more flex-ibility and capability to respond in real time to customer needs.


Research should stimulate cross-sectoral exchanges and partici-pation of SMEs, taking into account their specific needs androles in the supply chain as well as approaches able to create andhold in Europe sufficient jobs to arrest the decline of industrialemployment while improving the overall quality of work.

b. Substantially* improving overall quality within the valuechain (quality is intimately linked to value for and timely sat-isfaction of customer needs at the lowest costs) and conse-quently reducing “inefficiencies” and overall lifecycle productcosts by the same order of magnitude.

c. Minimizing resource consumption (e.g., materials, energy,water) to reduce substantially the overall “lifecycle” impact of“product-service” provision and use. These goals should bedealt with in a synergistic way. They should not be regardedas absolute targets for individual projects but rather as broadindications of the direction toward which the European indus-trial system, supported by improved regulations, shouldevolve.

8.3. Appendix 3. Excerpt from the CEFIC Technology Program

SUSTECH


The SUSTECH program was launched by a consortium of the majorchemical companies in Europe in April 1994. In doing so, they createda framework which would enable them to collaborate effectively in areasof research and development, which would be critical for the long termfuture of the process industries in Europe. The focus for collaborationwould be those areas that while attracting public concern were never-theless not areas in which the companies would normally compete. Suchareas as sustainable development, protection of the environment, andtechnologies for making more efficient use of the earth’s resources wereready targets for collaborative action.

Process Intensification.The objective of the Process Intensification (PI) Cluster is to promote PIand its benefits to a wide industrial audience by:

Exposing technologies critically to industrial needsIdentifying priority industrial applications

* The term substantially means over 20–30% in the shorter term or over 10% per year in the longerterm.


Facilitating partnerships between technology providers, equip-ment suppliers, and end users.

Two sessions were held at SUSTECH 7. In the morning session, processintensification and the future direction of the cluster were discussed ingeneral. In the afternoon, a mini-workshop was held on gas–liquid reac-tors. Total attendance was in excess of 60, including 23 representativesfrom the chemical industry.

An industrial Steering Group has been established to determinethe priority areas for the cluster and the mechanisms required to achieveits objectives. It was proposed to carry out analyses on representativeprocesses to identify the potential business benefits arising from appli-cation of PI technologies.

REFERENCES

1. BHR Group. Process Intensification Conferences. http://www.bhrgroup.co.uk/confsite/pi01home.html.

2. Elverding P. DSM’s Triple Bottom Line. DEFACTO 2001; 15(5):28–32.3. Bakker RA, Stankiewicz AI, Schyns VJAJ. Process intensification within DSM, gen-

eral methodology and concrete examples. In: Proceedings of the 4th InternationalConference on Process Intensification for the Chemical Industry. Cranfield, UK,BHR Group, 2001.

4. Cauwenberg V, Vergossen P, Stankiewicz A, Kierkels H. Integration of reaction andseparation in manufacturing of pharmaceuticals: membrane-mediated production ofS-ibuprofen. Chem Eng Sci 1999; 54:1473–1477.

5. Technical information brochures on the Urea 2000plus™ process can be obtainedfrom Stamicarbon B.V., the licensing company of DSM.

6. U.S. Vision 2020, section of the Office of Industrial Technologies (OIT). http://www.oit.doe.gov/chemicals/visions_new_chemistry.shtml.

7. European Union 5th Framework Research Program 1998–2002. Thematic programGROWTH: http://europa.eu.int/comm/research/growth/index.html and http://europa.eu.int/comm/research/growth/pdf/growth-workprog2000_en.pdf.

8. SUSTECH Technology program, part of the CEFIC organization. http://www.cefic.be.


http://www.bhrgroup.co.uk/confsite/ pi01home.html

http://www.bhrgroup.co.uk/confsite/ pi01home.html

http://www. oit.doe.gov/chemicals/visions_new_chemistry.shtml

http://www. oit.doe.gov/chemicals/visions_new_chemistry.shtml

http://europa.eu.int/comm/research/growth/index.html

http:// europa.eu.int/comm/research/growth/pdf/growth-workprog2000_en.pdf

http:// europa.eu.int/comm/research/growth/pdf/growth-workprog2000_en.pdf

http://www.cefic.be

13

Process Intensification for Safety

Dennis C. Hendershot

Rohm and Haas Company, Bristol,Pennsylvania, USA

1. INTRODUCTION

Process intensification is an important strategy in the development of inherentlysafer chemical processes and plants. By reducing the inventory of hazardous mate-rial or energy in the process, the potential consequence of failure to control thathazardous material or energy is reduced. Rather than relying on add-on safety features such as interlocks, procedures, and consequence mitigation systems, thesafety of the plant is based on reducing the magnitude of the possible damage.While safety devices can be designed to be highly reliable, no safety device is per-fect, and all will have a finite failure probability. If a chemical plant contains alarge amount of hazardous material or energy, the consequences of the failure ofthe add-on safety devices may be large. A smaller plant is safer because we havereduced its inherent capability to cause damage, rather than because we have con-trolled that capability through additional safety devices.

2. INHERENT SAFETY

Inherent is defined as “existing in something as a permanent and inseparable ele-ment, quality, or attribute” (1). A chemical process can be described as inherentlysafer if it reduces or eliminates a hazard when compared to another process alter-native. To understand this definition, it is necessary to understand what is meant


by the term hazard. A hazard is a characteristic of a material or process that resultsin potential harm to people, the environment, or property. A hazard is an inherentproperty of a material or its conditions of use. Some examples of hazards include:

Sulfuric acid is corrosive.Chlorine is toxic by inhalation.Heptane is flammable.Nitroglycerine is unstable and can detonate.Heat transfer oil at 300�C contains a large amount of energy.A vessel full of compressed air at 40 bar contains a large amount of poten-

tial energy.

The hazard cannot be changed, but it is possible to change the material or oper-ating conditions. The magnitude of a potential incident that could result from thefailure to control the hazard can also be reduced. An important way of accom-plishing this objective is to reduce the quantity of material in the process throughprocess intensification.

It is generally more appropriate to describe processes as inherently saferwhen compared to alternatives rather than as inherently safe. All processes havemultiple hazards, and it is not possible to eliminate all hazards. For example, itmay be possible to replace a toxic or flammable solvent for an extraction processwith supercritical carbon dioxide. While the carbon dioxide process is inherentlysafer with respect to flammability and toxicity hazards, it operates at high pressure,which introduces new hazards. Depending on the specific details of the alternativeprocesses—what the flammability properties of the solvent are, how toxic the sol-vent is, what the size of the equipment is, what the operating pressure is—thesupercritical carbon dioxide process may be inherently safer overall. However, itis important to consider all hazards in judging the overall inherent safety of processalternatives and to avoid focusing on a single hazard and forgetting about others.

3. HISTORY OF INHERENTLY SAFER DESIGN

It can be argued that engineers have always attempted to eliminate hazards intheir designs. For example, in 1828 the pioneering railway engineer RobertStevenson argued for simplification and minimization—two of the principles ofinherently safer design—of the newly developed steam locomotive when he dis-cussed “an alteration which I think will considerably reduce the quantity ofmachinery as well as the liability to mismanagement. . . . [I]n their present com-plicated state they cannot be managed by ‘fools’; therefore they must undergosome alteration or amendment” (2).

The concept of inherent safety as a specific set of design strategies to elim-inate or reduce hazards in the chemical industry was first articulated by TrevorKletz in the 1970s. Following the Flixborough explosion in England in 1974,


there was increased concern about safety in chemical plants, from the industryitself, government, and the general public. This resulted in an increased focus oncontrolling hazards in chemical plants. A specific worry was that the magnitudeof potential accidents was now larger because of the large size of a new genera-tion of world-scale petrochemical facilities. In 1977, while working as a safetyadvisor for the ICI Petrochemicals Division, Kletz suggested an alternativeapproach in the annual Jubilee Lecture to the U.K. Society of the ChemicalIndustry (3). He proposed that the industry should direct its attention to eliminat-ing or reducing the hazards of its plants and processes rather than accepting thosehazards and working to control them. Kletz called this approach inherently saferdesign. Furthermore, this hazard elimination or reduction would be accomplishedby means that were inherent in the process and, thus, permanent and inseparablefrom it. Since 1977, concepts and approaches to inherently safer design in thechemical industry have been developed and promoted (4–6), including a brief dis-cussion in Perry’s Chemical Engineers’ Handbook (7).

4. CONCEPT OF LAYER OF PROTECTION

FOR PROCESS SAFETY

The safety of a chemical process relies on multiple layers of protection to protectpeople, the environment, and property from the hazards associated with theprocess. A process designer recognizes that equipment can fail and that peoplewill make mistakes. While we can design more reliable equipment and train andmotivate people to reduce mistakes, we can never completely eliminate these fail-ures and mistakes. Therefore, it is important to provide multiple layers of protec-tion, a defense in depth, to reduce risk (Figure 1). Even so, there will always besome small probability that all of the layers of protection will fail simultaneouslyand an accident will occur. Also, the effectiveness of the layers of protectiondepends on the ongoing maintenance of equipment, the training and performanceof people, and the management systems. If these systems deteriorate, the reliabilityof the protection layers will be reduced and the risk will increase (Figure 2). TheCenter for Chemical Process Safety has published a book that describes the layer-of-protection concept in detail and extends it to a quantitative risk-managementtechnique, layer of protection analysis (LOPA) (8).

If the magnitude of the potential accident is very large, some people maynever feel completely comfortable with the residual risk that it may occur, even ifthis risk is very small and management systems to maintain the protection sys-tems continue to be good and effective. By reducing the magnitude of the potentialaccident, inherently safer design acknowledges the inevitability of the failure ofequipment, people, and management systems, and it bases process safety on reduc-ing the inherent hazard of the process. An inherently safer design reduces theneed for layers of protection for a process, and, if the magnitude of the potential


consequence of an accident can be reduced sufficiently, it may eliminate the needfor protection layers entirely (Figure 3).

In general, the layers of protection applied to a chemical process can beplaced into four categories:

Inherent—Eliminate or reduce the hazard by using materials and conditionsthat are less hazardous or nonhazardous.

Passive—Minimize or control the hazard using design features that reduceeither the frequency or the consequence of incidents without the activefunctioning of any device.

FIGURE 1 Layers of protection for a chemical process.


Active—Control or mitigate incidents using controls, safety interlocks, oremergency shutdown systems to detect hazardous conditions and takeappropriate action to place the plant in a safe condition.

Procedural—Use operating procedures, administrative checks, emergencyresponse, and other management systems to prevent incidents, to detectincidents in time for operators to place the plant in a safe condition, or toreduce the magnitude of the damage resulting from an incident.

These categories are not rigidly defined, and a design may exhibit charac-teristics of more than one category. Safety strategies in the inherent and passive

FIGURE 2 Higher risk resulting from degraded layers of protection.


categories are generally considered to be more robust and reliable. They dependon the physical and chemical properties of the system rather than on the success-ful operation of instruments, safety devices, and procedures. Inherent and passivestrategies are often confused, but they are different. A truly inherent solution to asafety issue will either completely eliminate the hazard or reduce the potentialmagnitude of an incident associated with the hazard sufficiently that it cannotcause significant damage. On the other hand, passive strategies do not eliminate

FIGURE 3 An inherently safer process meets risk requirements with fewer orno layers of protection.


the hazard, but instead prevent injury and damage by eliminating or reducing theexposure of people or property to the hazardous condition without the activefunctioning of any device. A containment dike around a storage tank is an exam-ple of a passive layer of protection. It performs its function of limiting the dam-age in case of a leak from the storage tank simply by being present. There is noneed to detect a leak or for any device or person to perform any function for thecontainment dike to do its job.

5. INHERENT SAFETY STRATEGIES

Approaches to inherently safer process design have been categorized in a numberof ways. The Center for Chemical Process Safety (5) describes four strategies forinherent safety, derived from Kletz’s initial proposals (3,6):

11. Minimize—Use smaller quantities of hazardous substances, includingprocess intensification approaches. Example: Use a small, continuousreaction system for the production of nitroglycerine in place of a largebatch reactor.

12. Substitute—Replace a material with a less hazardous substance.Example: Aqueous latex paints reduce both flammability and toxicityhazards when compared to solvent-based paints, both during manu-facture and for the final consumer.

13. Moderate—Use less hazardous conditions or a less hazardous form ofa substance. Example: Plastic resins can be produced in pellet orgranular form, reducing dust explosion hazards when compared to apowder form of the same material.

14. Simplify—Design processes and facilities that eliminate unnecessarycomplexity and that are tolerant of human error. Example: Designpiping to permit gravity flow of hazardous materials in a plant, elim-inating the need for pumps, which can leak.

Additional discussion and more examples of these strategies can be foundin books by Kletz (6) and CCPS (4,5). The remainder of this discussion will focuson minimization (process intensification) as an inherent safety strategy.

6. PROCESS INTENSIFICATION AS AN INHERENT

SAFETY STRATEGY

6.1. Smaller Is Safer

Reducing the size of chemical processing equipment enhances safety in two ways.The quantity of hazardous material that can be released in case of equipmentleakage or rupture is obviously smaller if the equipment is smaller. In addition,


the potential energy contained in the plant equipment is also smaller if theequipment is smaller. This potential energy can be in many forms, such as hightemperature, high pressure, or heat of reaction from a reactive chemical mixturein the equipment. If this potential energy is released in an uncontrolled manner,an incident such as a fire, explosion, or uncontrolled leak of material from theequipment may occur.

Clearly the potential damage from the uncontrolled release of material orenergy is reduced if equipment can be made smaller. There is also another bene-fit to smaller equipment—it may be more feasible to provide equipment to mitigateor control the consequences of an incident. For example, it may be feasible tototally enclose a small reactor in a blast-proof structure. Doing this for a large reac-tor may not be feasible because the blast-proof enclosure would have to be muchlarger. The enclosure would also have to be much stronger because it would haveto withstand a larger potential explosion from the larger reactor.

6.2. Traditional Approaches to Minimizing Inventory

There are many opportunities to minimize inventory of hazardous material in achemical plant without fundamental changes in process technology. The accidentin Bhopal, India, in 1984 released methyl isocyanate, causing approximately 2000fatalities and injuring tens of thousands of people (9). This is by far the worst acci-dent in the history of the chemical industry. Following the Bhopal accident, mostchemical companies reviewed their operations to identify opportunities to reducethe inventory of toxic and flammable materials. Many significant reductions werereported as a result of this effort, and these reductions were accomplished rela-tively quickly. Obviously, these companies did not rebuild plants using differenttechnology or make dramatic changes to the process equipment in the existingplants in such a short time. So how did chemical plants around the world reducehazardous material inventories? They carefully evaluated existing equipment andoperations and identified changes in operations that would allow the existingplants to operate with a reduced inventory of hazardous materials. The Bhopaltragedy focused the attention of creative engineers on the problem of reducinghazardous material inventory, and they quickly identified ways of accomplishingthis objective, even for existing plants and technologies.

6.2.1. Storage

The single biggest reduction in hazardous material inventory of plants was instorage facilities. This includes raw materials, in-process intermediates, and finalproducts. The terrible consequences of Bhopal caused engineers and managers toquestion the need for storage of large quantities of raw materials and intermedi-ates. In many cases, a large inventory of raw materials makes it easier to operatea plant. The company may have more flexibility in ordering raw materials, to takeadvantage of favorable market conditions to order large amounts of material when


prices are low. The plant is less likely to have to shut down because of trans-portation delays. However, if the raw material is very hazardous, and particularlyif it is very toxic and a large leak has the potential to impact population in a largearea surrounding the plant, the risks associated with storage of a large amount ofthe material may be significant. A better strategy would be to devote the com-pany’s resources to ensuring a reliable and economic supply of material, effectivetransportation systems, and good inventory control systems. Strategic alliances withraw material suppliers and transportation companies, modern inventory manage-ment systems, and improved communication with supplier plants will allow aplant to reduce the quantity of hazardous material that must be stored at the man-ufacturing site. There is also an economic benefit—working capital is reducedbecause the inventory of raw material is reduced.

Often, a plant design includes large storage tanks for hazardous in-processintermediates. This intermediate storage decouples sections of a plant from oneanother. Parts of a plant can continue to operate, either filling or emptying in-process storage tanks, while another unit in the plant is shut down for maintenanceor because of operating problems. In fact, reliability engineers interested in increas-ing the on-stream efficiency of a plant will encourage a plant designer to includeintermediate storage buffers in a design to improve the overall plant reliability. Thiscan be desirable in improving the inherent safety of a plant—a continuous plant isgenerally safer in normal, steady-state operation than while starting up or shuttingdown (10). The buffers will reduce the frequency of starting up and shutting downsections of the plant because of unavailability of feedstock from an upstream unitor a shutdown of a downstream unit resulting in no place to put the product.

However, if the intermediate to be stored in the buffer tank is extremelyhazardous (flammable, toxic, reactive), this benefit must be balanced against therisks inherent in the storage of a large quantity of hazardous material. There maybe other approaches to enhancing the reliability of the units within a large plantthat will eliminate or greatly reduce the need to store large quantities of hazardousintermediates. Perhaps the intermediate buffer can be provided at another point inthe process, allowing storage of a less hazardous material or a less hazardous formof the same material. For example, a reactive material is produced in a continuousgas-phase reaction, then absorbed in water to form an aqueous solution, isolatedas a pure material by distillation, and stored for feeding to downstream processingunits. A buffer between this unit and the downstream units could be provided withintermediate storage of the distilled, purified material. However, it would be inher-ently safer to provide this storage buffer for the aqueous solution of the material,which is less hazardous.

There are other ways to reduce the need for in-process storage of hazardousintermediates. Fully understanding the causes of unreliability in production units isthe key to reducing the need for in-process buffer tanks. Causes of shutdown can beeliminated by using more robust and reliable equipment, improving the process to


make it less sensitive to variations in operating parameters, or providing redun-dancy for critical equipment so that the entire plant does not have to be shut downwhen the critical equipment requires repair. If the reliability of individual units ina plant can be increased, it will not be necessary to store large amounts of hazardousin-process intermediates to maintain the desired production capacity. The designengineer should question the need for all in-process storage. He should ask if the plantcan operate just as efficiently without the intermediate storage tanks if the causes ofplant shutdowns could be understood and eliminated. In many cases, the answeris yes, and in-process storage has been greatly reduced or eliminated. One compa-ny reported a reduction in total inventory of hazardous materials, including chlo-rine and hydrogen cyanide, of over 1 million pounds at a single plant site (11).

One measure of the inherent safety of a process with respect to fire andexplosion risk is the Dow Fire and Explosion Index (12). Table 1 shows examplesof the impact of inventory reduction on the Fire and Explosion Index.

The benefits of reduced inventory can also be quantified by estimating thepotential consequence from a potential incident. For example, one of the hazards ofstorage of a liquified flammable gas such as propylene is a boiling-liquid expandingvapor explosion (BLEVE) and the associated fireball. Figure 4 shows the heat radi-ation intensity as a function of distance from a storage tank for a potential BLEVEfor three sizes of propylene storage vessel, ranging from 500 to 50,000 kg. The mag-nitude of the consequence reduction from reducing propylene inventory is clear.

6.2.2. Piping

When designing piping for hazardous materials, the designer should attempt tominimize inventory, by minimizing both pipe diameter and pipe length. Hazardousmaterial piping should be large enough to transport the quantity of material requiredand no larger. Reducing pipe size from 50-mm diameter to 25-mm diameterreduces the inventory of hazardous material in the pipe by a factor of 4. This has a

TABLE 1 Impact of Inventory on the Fire and Explosion Index

Fire and explosion index

Ethyl acrylate storage inventory907,000 kg (2 million pounds) 15191,000 kg (200,000 pounds) 13023,000 kg (50,000 pounds) 120

Agricultural product with in-process intermediate storage

23 cubic m (6000 gallons) 1850 140

Source: Ref. 13.


major impact on the distance over which hazardous concentrations of material canoccur in the atmosphere if the pipe is broken. Figure 5 shows the footprint of a toxicphosgene cloud for a specific set of conditions, where the only difference is the sizeof the phosgene pipe. The smaller pipe results in a much smaller toxic vapor cloud.

Inventory of hazardous material in pipes can also be minimized by usingthe hazardous material as a gas rather than as a liquid. The Dow Chemical ExposureIndex (14) is a tool that can be used to measure inherent safety with regard to poten-tial toxic exposure risk. Table 2 shows the reduction in the Chemical ExposureIndex that can be realized by handling a number of hazardous materials as a gasrather than as a liquid, assuming that the same-size pipe can deliver the requiredflow rate. Figure 6 shows the decrease in the hazard zone (toxic cloud footprint)that resulted from relocating a chlorine vaporizer from a production building to

FIGURE 4 Radiation intensity from a BLEVE and resulting fireball as a func-tion of distance for propylene storage tanks in three different sizes.


a storage area. Following the modification, the long chlorine transfer line containedgaseous chlorine instead of liquid chlorine.

6.3. Process Intensification

Process intensification refers to a chemical process using significantly smallerequipment. Examples include novel reactors, intense mixing devices, heat andmass transfer devices that provide high surface area per unit of volume, devices

482 Hendershot

FIGURE 5 Cloud footprint to atmospheric concentration of 1 ppm resultingfrom the rupture of three sizes of phosgene pipe. Release conditions: Com-plete rupture of pipe while connected to a large storage tank without shutoff,pipe elevation is 5 m above grade, wind speed is 5 m/sec, atmospheric stabil-ity class D, 1 ppm is the Emergency Response Planning Guideline-3 (ERPG-3)concentration for phosgene, the concentration at which life-threateningeffects might result from exposure for 1 hour.


that combine one or more unit operations in a single piece of equipment, anddevices that use alternate ways of delivering energy to processing equipment—for example, ultrasound, microwaves, laser beams or light, and radiation. Otherchapters in this book discuss these and other process intensification techniques inmore depth. These technologies can greatly increase the rate of physical andchemical processes, allowing a very high productivity from a small volume of in-process inventory. Clearly, this is desirable from an economic perspective—a small, highly efficient plant can be expected to be cheaper and more cost effective.If the material contained in the plant is hazardous, because of either its physicaland toxicological properties or its conditions of use, such as high temperature orpressure, then the risk associated with a small amount of the material will be lessthan for a large quantity.

If the plant is small enough, the maximum possible accident may not pose asignificant hazard to people, the environment, or property. This may result in anadditional reduction in the equipment needed for the plant—it may not require asmuch (or any) safety equipment, emergency alarms and interlocks, or other layersof protection to manage risk. Even if the small plant still requires safety equipment,this equipment will be smaller and cheaper. Installation and ongoing operation ofsafety equipment is often a major expense; if it can be eliminated or reduced in sizeand complexity, there will be cost savings. Safety need not cost money—safer canalso be cheaper if a small, efficient, inherently safer process can be invented.

Some specific examples of process intensification resulting in safer as wellas more economical processes follow. Many more examples, described in moredetail, can be found throughout this book.

TABLE 2 Chemical Exposure Index for Failure of a 2-Inch Pipe

ChemicalMaterial State exposure index

Phosgene Liquid 1000Gas 850

Chlorine Liquid 1000Gas 490

Hydrogen sulfide Liquid 980Gas 310

Sulfur trioxide Liquid 360Gas 110

Source: Ref. 14.


6.3.1. Nitration

Nitration reactions are highly exothermic, can generate high pressure in a closedsystem from noncondensible by-products from undesired side reactions, andoften produce unstable reaction products, such as explosives. Many years ago,products such as nitroglycerine were manufactured in large batch reactors. Asengineers began to understand the physical and chemical processes involved innitration chemistry, they recognized that the chemical reaction actually occursvery rapidly once the reactants come into contact with each other. The large reactor

FIGURE 6 Cloud footprint to an atmospheric concentration of 20 ppm result-ing from the rupture of a 50-mm-diameter chlorine pipe containing eitherchlorine liquid or chlorine vapor. Release conditions: Complete rupture of pipewithout shutoff, pipe elevation is 5 m above grade, wind speed is 5 m/sec,atmospheric stability class D, 20 ppm is the Emergency Response PlanningGuideline-3 (ERPG-3) concentration for chlorine, the concentration at whichlife-threatening effects might result from exposure for 1 hour.


size and long reaction times were actually the result of poor mixing, poor removalof the heat of reaction, and poor mass transfer in two-phase reaction mixtures. Thisknowledge of the physical and chemical processes involved in an industrial-scalenitration process was used to design new types of reactors to efficiently contactthe reactants and liquid phases and remove the heat of reaction. Modern nitrationplants use very small continuous stirred-tank reactors with intense mixing andlarge heat transfer area or jet reactors of various designs to provide intense mix-ing and rapid contacting of reactants (6).

6.3.2. Polymerization

A number of innovative polymerization reactors using loop reactors, plug-flowand static mixer reactors, and continuous stirred-tank reactors have been reported.For example, Wilkinson and Geddes (15) describe a 50-liter reactor that has thesame capacity as a 5000-gallon batch reactor. Extruders, thin-film evaporators,and other devices designed to provide intense mixing for viscous or unstablematerials have also been used as reactors.

6.3.3. Tubular or Jet Reactors

Continuous reactors, including simple plug-flow pipe reactors, tubular reactors con-taining static or other mixing devices, and jet reactors of various types, have beenused to efficiently produce toxic materials for immediate consumption in down-stream processing operations with little or no inventory. Some examples follow.

Methyl isocyanate (MIC), the material that was released at Bhopal, can beproduced and immediately converted to final product in a process thatcontains a total inventory of less than 10 kg of MIC (6).

Caro’s acid, an equilibrium mixture of sulfuric acid, water, and peroxymono-sulfuric acid, is used in the metal-extraction industry. It is manufacturedby reacting concentrated sulfuric acid with hydrogen peroxide. Caro’sacid is a powerful oxidizing agent and decomposes readily. A processwas developed to manufacture 1000 kg/day of Caro’s acid in a tubularreactor with a volume of 20 ml and a residence time of less than one second, with the product immediately mixed with the solution to be treated (16).

A continuous tubular reactor was developed to manufacture phosgene forimmediate consumption by a group of batch-processing buildings (17). Oneplant using the new design contains 70 kg of phosgene gas, compared toa total inventory of 25,000 kg of liquid phosgene in the old plant.Because the new plant is small, it is also feasible to provide a secondarycontainment building for the equipment containing phosgene, providingan additional barrier between a highly toxic material and people and thesurrounding environment (18).


6.3.4. Heat Exchangers

Heat exchange equipment varies widely in efficiency in terms of available heattransfer area per volume of inventory contained in the exchanger. Innovative heatexchanger design can reduce the volume of hazardous material in a heat exchangerby a factor of 10 or more when compared to a standard shell-and-tube exchanger.Some more efficient alternatives include plate, shell-and-finned tube, plate-fin, androtary exchangers. With some of these devices, it is important to consider the poten-tial for additional leak scenarios, such as gasket leaks in plate exchangers or haz-ards associated with seals and moving parts in rotary exchangers. The designermust balance the benefits of reduced inventory against the possibility that leaksmay be more likely. Kletz (6) discusses alternative heat exchanger designs.

6.3.5. Distillation

Innovative design of distillation devices can greatly reduce inventory and can alsoreduce the residence time in the distillation system, which may be important forsafely processing thermally sensitive materials. Wiped film evaporators have beenin use for many years, particularly for distillation of reactive materials, such asmonomers, and for other heat-sensitive compounds. In more traditional distilla-tion equipment, columns and trays can be designed to minimize the inventory ofliquid in the column. Often the largest inventory of hazardous liquid is containedin the bottom of the column and in the reboiler. This is particularly undesirablefor a reactive or thermally sensitive material because this is part of the distillationcolumn that is at the highest temperature. It is possible to design the bottom of acolumn to minimize this inventory. For example, use a conical column bottom ora small-diameter column bottom that still provides sufficient liquid head for bot-toms or reboiler recirculation pumps but reduces the total quantity of hot liquid.Centrifugal distillation equipment such as Higee (19) can be much smaller thanconventional distillation equipment.

6.3.6. Extraction

Extraction columns are often very large devices, and they may contain a largequantity of flammable or toxic solvent. The inventory of hazardous material in anextraction system can be greatly reduced by using a centrifugal extractor. In thiscase, the inherent safety benefits of reduced inventory must be balanced againstthe new hazards introduced by the centrifugal extractor, including operation athigh speed and potential for leakage from seals, to determine the best choice fora particular application.

6.3.7. Combined Unit Operations

Equipment that combines more than one unit operation in a single piece of equip-ment is another approach to process intensification. Reactive distillation is a good


example. In one case, a traditional process for the manufacture of methyl acetate usesa stirred-tank reactor, eight distillation columns, and an extraction column. A new,reactive distillation process uses one reactive distillation column and two additionalcolumns (5). The total inventory of combustible material in the process is reduced.In this example, the reactive distillation system also results in a simpler plant—onewith fewer major vessels. The process is also simpler because there is much less sup-porting equipment—condensers, reboilers, pipes, valves, pumps, instruments. Safetyis enhanced because every valve seal, flange, pump seal, and instrument connectionthat is eliminated from the plant is one less place that the process can leak. Thisreduces fugitive emissions from this equipment, which contribute to plant emissionsand pollution, and also the likelihood of bigger leaks that can cause a fire, personnelexposure or injury, or a significant environmental incident.

6.3.8. Innovative Energy Sources

Energy sources such as laser light, ultraviolet light, microwaves, and ultrasonicenergy can be used to apply energy in a controlled fashion to a chemical reactionor physical unit operation to increase efficiency. For example, technology for thedistributed, small-scale manufacture of hydrogen cyanide using microwave energyto direct heat to the reactor catalyst is under development (20). This would allowhydrogen cyanide to be manufactured in small quantities where it is needed,rather than manufacturing it in a large central facility, storing it, and transportingit to the sites where it is needed.

7. METRICS FOR INHERENT SAFETY

How do you measure inherent safety? This is an important question that must beanswered to effectively promote the development of inherently safer process tech-nologies, including process intensification. Development of new technologyrequires resources, and these resources can be more easily obtained if the costsand potential benefits of the new technology, including economics, safety, andenvironment, can be measured. Also, chemical processes always involve multiplehazards. A process that is inherently safer with respect to one hazard may be lesssafe with respect to a different hazard, or it may introduce new hazards. Forexample, our earlier discussion about extraction points out the benefits of reducedinventory that can be obtained with a centrifugal extractor but acknowledges thenew hazards introduced by the use of high-speed rotating equipment and thepotential for leakage from seals. Some method of quantifying these risks is essen-tial to making the best technology decisions.

The chemical industry is just beginning to develop tools for measuringinherent safety. Some of these tools have been used for risk management and lossprevention for some time, but we are just beginning to recognize their value inunderstanding the inherent safety of processes.


7.1. Accident Consequence Analysis

The analysis of the potential consequences of an accident is a useful way ofunderstanding the relative inherent safety of process alternatives. These conse-quences might consider, for example, the distance to a benchmark level of damageresulting from a fire, explosion, or toxic material release. Accident consequenceanalysis is of particular value in understanding the benefits of minimization, mod-eration, and limitation of effects. This discussion includes several examples of theuse of potential accident consequence analysis as a way of measuring inherentsafety, such as the BLEVE and toxic gas plume model results shown in Figures 4,5, and 6.

There are many commercial and public domain tools available for accidentconsequence modeling. Guidelines for Use of Vapor Cloud Dispersion Models (21)reviews many of the available models for flammable and toxic gas clouds in the atmosphere. Guidelines for Consequence Analysis of Chemical Releases (22)describes modeling techniques for fires and explosions, some of the simpleratmospheric dispersion models, and models for estimating the impact of differenttypes of incidents on people and property. This book also includes a set of spread-sheets for the models discussed. The TNO “Yellow Book” (23) describes a widerange of models for many types of potential accidents in a chemical plant. TheTNO “Green Book” (24) reviews models for estimating the potential impact ofaccidents on people and property.

7.2. Risk Indices

A number of risk indices have been developed over the years as chemical processloss prevention and risk management tools. Many of these are based on the inher-ent characteristics of the processes, and they can be used as measures of processinherent safety. In general, these indices measure a single aspect of inherent safety,and it is necessary to use several indices to obtain a full understanding of the over-all process characteristics.

The Dow Fire and Explosion Index (FEI) (12) and the Dow ChemicalExposure Index (CEI) (14) are two commonly used tools that measure inherentsafety characteristics. Gowland (25) reports on the use of the FEI and CEI in thedevelopment of safety improvements for a urethane plant. Tables 1 and 2 illustratethe application of the FEI and CEI in measuring inherent safety characteristics ofprocess design options. These indices measure the inherent safety characteristicsof processes in only two specific areas—fire and explosion hazards and acutechemical inhalation toxicity hazards. Other indices would be required to evaluateother types of hazards.

Because the material inventory is considered in these indices, process inten-sification will result in a lower value (greater safety) for either index. However,the indices do also consider other factors, such as temperature and pressure, that


might be higher for an alternative process. The index is therefore valuable in help-ing to understand the relative value of a reduced inventory of hazardous materialcompared to other hazard factors that might increase.

7.3. Overall Inherent Safety Index

Some initial work on the development of an overall inherent safety index has beendone at Loughborough University in the United Kingdom (26). Other, similarwork has been done at VTT in Finland (27). Both indices are considered proto-types by the developers, and more work is needed. These proposed inherent safe-ty indices evaluate a number of factors related to inherent safety, including:

InventoryFlammabilityExplosivenessToxicityTemperaturePressure

A single-number overall index characterizing the inherent safety of the overallprocess is generated by both proposed inherent safety indices. Process intensifi-cation will lower the value of the index (indicating an inherently safer process)because it will reduce the penalty for “inventory.” If the alternative process resultsin an increase in the inherent hazard due to other factors, the index will be usefulin understanding the inherent safety characteristics of the different alternatives.The relative contributions of the various components of the index to the totalvalue may also be useful in understanding process safety characteristics. Table 3summarizes the application of this proposed inherent safety index to a number ofalternative routes for the manufacture of methyl methacrylate.

TABLE 3 Evaluation of Alternate Methyl Methacrylate ProcessesUsing the Edwards et al. Proposed Inherent Safety Index

InherentProcess route safety index

Acetone cyanohydrin �120Ethylene based/propionaldehyde �75Propylene based �70Ethylene based/methyl propionate �50Isobutylene based �50Tertiary butyl alcohol based �50


8. PROCESS INTENSIFICATION BENEFITS FOR PASSIVE

AND ACTIVE LAYERS OF PROTECTION

While a smaller process may not totally eliminate a hazard, it often can have thebenefit of making effective passive layers of protection for the process more feasible and cost effective. Passive protection devices such as containment dikes,blast-resistant enclosures, and containment buildings to prevent the escape oftoxic gas can be smaller. Active safety devices such as rupture disks, flares, andscrubbers will be reduced in size. Smaller processing equipment will also respondmore quickly to other common safety interlock actions, such as shutting off reactantfeeds, increasing cooling to a vessel, adding reaction shortstop agents, and emp-tying a reactor to a reaction quench tank. Some examples of the potential bene-fits of process intensification in the application of other safety features follow.

Unstable materials such as explosives are sometimes manufactured inremotely controlled facilities protected by blast-resistant enclosures or bunkers.In this situation, if an explosion occurs the process equipment may be severelydamaged or destroyed, but nobody will be injured and the environment and otherproperty will be protected. This would be considered to be a passive safety feature—the blast-resistant enclosure does not require the action of any equipment or peopleto perform its function. While this type of blast-resistant enclosure may be practicalfor small equipment, it may become prohibitively expensive for large equipment.Clearly a larger structure would be required for larger equipment. However, theenclosure would also have to be much stronger because the potential explosion inthe large vessel would be of greater magnitude.

Process vessels are often protected from overpressurization by active devicessuch as relief valves and rupture disks. These devices open at a set pressure andallow gas and liquid to escape from the protected vessel to limit the pressure. Manyyears ago, the rupture disk or relief valve would simply discharge to the atmosphere.However, this is often unacceptable today because we cannot allow hazardous mate-rial to escape to the environment, even in an emergency situation. Emergencyrelief devices often must discharge to a collection-and-treatment system to ensurethat no hazardous material is released to the atmosphere. This treatment systemmight include equipment such as vapor liquid separation devices, scrubbers,absorbers, flares, or thermal oxidizers. If the process vessels can be made smaller,the size of the required treatment system for the emergency relief discharge is cor-respondingly reduced. Total containment of the emergency relief device effluent ina large pressure vessel may even be possible, eliminating all discharge to the envi-ronment during the event (but the contents of the containment vessel will have tobe treated and disposed of later on).

For extremely exothermic reactions, emergency quench systems are some-times used to protect against a runaway reaction. If the reactor temperature exceedsa predetermined maximum safe temperature, the reactor contents are rapidly


discharged to another vessel containing a material that will stop the reaction.These systems are more feasible for a small reactor. A reactor of a couple of cubicmeters volume can be emptied to a quench tank in a few seconds through a large-diameter dump valve, but it might take many minutes to discharge the contents ofa large reactor. It may not be possible to empty the large reactor quickly enoughto prevent the runaway reaction.

9. SUMMARY

Safety considerations are an inseparable part of the development of a chemicalprocess and the design and operation of a chemical plant. While risk manage-ment and safety features can be added on to a plant design or to an operatingplant, safety is most reliably and robustly ensured by developing inherently saferprocesses.

Safety strategies can be categorized as inherent, passive, active, and proce-dural. Inherent and passive strategies generally relate to the basic process tech-nology and plant design and are nearly always implemented early in the designlife cycle. They focus on elimination of hazards or minimizing the degree of haz-ard rather than on management of hazards. Process intensification is an importantapproach to the development of inherently safer chemical processes because itreduces the quantity of hazardous material in the process, thereby reducing theinherent risk.

Active and procedural strategies are usually also a part of a chemical processrisk management program—it is not often possible to eliminate all hazards.Process intensification can also make active and procedural safety features moreeffective and economical. Safety equipment can be made smaller and less costly.It may be feasible to use safety devices to protect against the hazards from smallprocessing equipment that are impractical for use in a large plant. The fasterresponse time of small equipment may allow effective automatic or manual inter-vention to detect an incipient problem and take action to prevent it from develop-ing into a serious accident. Chemical process safety cannot be viewed in isolationfrom other process and plant design criteria. The chemical plant must meet manyrequirements for workers (safety, long-term health, employment and wages), own-ers (operating costs, capital investment, profitability), customers (product quality,reliability of supply, cost), neighbors (safety, health, environmental impact, eco-nomic impact), and government (compliance with laws and regulations). All ofthese are important, and they may be in conflict. The chemical process designermust work to select the optimum design that considers all stakeholders. Processintensification is an important approach to minimizing the hazards associatedwith chemical handling and manufacture and will be an important factor in thefuture for designing safe, environmentally friendly, and economically competitivechemical plants.


REFERENCES

1. American College Dictionary. New York: Random House, 1967.2. Rolt LTC. The Railway Revolution: George and Robert Stevenson. New York:

St. Martin’s Press, 1960:147.3. Kletz TA. What you don’t have can’t leak. Chem Ind 1978; May 6, 287–292.4. Center for Chemical Process Safety (CCPS). Guidelines for Engineering Design for

Process Safety. New York: American Institute of Chemical Engineers, 1993:5–52.5. Center for Chemical Process Safety (CCPS). Inherently Safer Chemical Processes:

A Life Cycle Approach. New York: American Institute of Chemical Engineers, 1996.6. Kletz TA. Process Plants: A Handbook for Inherently Safer Design. Bristol, PA:

Taylor and Francis, 1998.7. Perry RH, Green DW, eds. Perry’s Chemical Engineers’ Handbook, 7th ed. New

York: McGraw-Hill, 1997: 26–5, 26–6.8. Center for Chemical Process Safety (CCPS). Layer of Protection Analysis. New

York: American Institute of Chemical Engineers, 2001.9. Lees FP. Loss Prevention in the Process Industries, 2nd ed. Oxford, UK: Butterworth-

Heinemann, 1996:A5/1–A5/11.10. Post RL, Hendershot DC, Kers P. Synergistic design approach to safety and reliabil-

ity yields great benefits. Chem Eng Prog 2002; 98(1):60–66.11. Wade DE. Reduction of risks by reduction of toxic material inventory. In: Woodward

JL, ed. Proceedings of the International Symposium on Preventing Major ChemicalAccidents. February 3–5, 1987, Washington, DC; New York: American Institute ofChemical Engineers, 1987:2.1–2.8.

12. Dow Chemical Company. Dow’s Fire and Explosion Index Hazard ClassificationGuide. 7th ed. New York: American Institute of Chemical Engineers, 1994.

13. Hendershot DC. Process minimization: making plants safer. Chem Eng Prog 2000;96(1):35–40.

14. Dow Chemical Company. Dow’s Chemical Exposure Index Guide. 1st ed. New York:American Institute of Chemical Engineers, 1994.

15. Wilkinson M, Geddes K. An award-winning process. Chem Br 1993; (December):1050–1052.

16. Whiting MJL. The benefits of process intensification for Caro’s acid production.Trans IChemE 1992; 70(March):195–196.

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18. Delseth R. Production industrielle avec le phosgene. Chemia 1998; 52(12):698–701.19. Kelleher T, Fair JB. Distillation studies in a high-gravity contactor. Ind Eng Chem

Res 1996; 35:4646–4655.20. Koch TA, Krause KR, Mehdizadeh M. Improved safety through distributed manu-

facturing of hazardous chemicals. Process Safety Prog 1997; 16(1):23–24.21. Center for Chemical Process Safety (CCPS). Guidelines for Use of Vapor Cloud

Dispersion Models. 2nd ed. New York: American Institute of Chemical Engineers,1996.


22. Center for Chemical Process Safety (CCPS). Guidelines for Consequence Analysisof Chemical Releases. New York: American Institute of Chemical Engineers, 1999.

23. The Netherlands Organization of Applied Scientific Research (TNO). Methods for theCalculation of the Physical Effects of the Escape of Dangerous Materials: Liquids andGases (The Yellow Book). 3rd ed. The Hague, Netherlands: Sdu Uitgevers, 1997.

24. The Netherlands Organization of Applied Scientific Research (TNO). Methods for theDetermination of Possible Damage to People and Objects Resulting from Releasesof Hazardous Materials. The Hague, Netherlands: Sdu Uitgevers, 1992.

25. Gowland RT. Applying inherently safer concepts to a phosgene plant acquisition.Process Safety Prog 1996; 15(1):52–57.

26. Edwards DW, Lawrence, D. Assessing the inherent safety of chemical processroutes: is there a relation between plant costs and inherent safety? Trans IChemE B,1993; 71(November):252–258.

27. Heikkila A. Inherent Safety in Process Plant Design. Espoo, Finland: TechnicalResearch Centre of Finland (VTT), 1999.


14

Process Intensification Contributions to Sustainable Development

G. Jan Harmsen, Gijsbert Korevaar, and

Saul M. Lemkowitz

Delft University of Technology, Delft, The Netherlands

This chapter is meant for developers and designers of processes, in particular ofintensified processes, who want to ensure that their processes contribute to sus-tainable development. To this end problems with the present technological solu-tions in relation to the world-scale environment and society are explained in thefirst section. Then in the second section sustainable development is explained andthe role and criteria for sustainable technology are derived. Finally process inten-sification is assessed on its contribution to sustainable development.

The information presented should be sufficient for a developer or designerof a process to raise awareness of design aspects related to sustainable develop-ment, to yield hints to modify the design in the direction of sustainable develop-ment, and to find references for further detailed information.

1. PROBLEMS LEADING TO SUSTAINABLE DEVELOPMENT

1.1. Environmental Problems

The disturbance of several natural ecological cycles became a problem with thestart of the Industrial Revolution in the 19th century. The emissions to air, water,


and soil increased exponentionally. In the 20th century these changes in the envi-ronment reached global proportions:

Increases of concentrations of “greenhouse” gases in the global atmosphereoccurring at rates and to levels higher than those ever reached during theprevious thousands of years (1).

Massive deforestation occurring faster than at any time in human his-tory (2).

Species extinction and loss of biodiversity at rates higher than 10,000 peryear (3).

The publication of the report Limits to Growth (4) by the Club of Romehad a major impact on thinking about the environmental impact of our culturaldevelopment. Under the assumption that the five basis elements of this study—population, the production of food, industrialization, pollution, and the use of non-renewable resources—will keep increasing exponentially, they showed that, ifunchanged, this would lead to enormous problems, as soon as the 21st century. Thesocial consciousness of the problems caused by unlimited growth of these ele-ments was greatly increased by this report by the Club of Rome.

In the present situation, one can say that natural resources and the environ-ment can no longer be regarded as inexhaustible reservoirs of products and serv-ices. On the contrary, in general it is now recognized that natural resources arelimited, that the natural environment has a limited capacity, and that this musthave consequences for the way we act.

1.2. Socioeconomic Problems

In addition to adverse changes to the world environment, economic differencesbetween rich and poor are increasing (decreasing “equity”), both within nationsand between nations. The richest fifth receives 83% of the total world income (5).Moreover, the gross domestic product (GDP) per capita stays more or less thesame for poor countries. This unfairness is also a part of unsustainability, becauseit blocks possibilities toward worldwide cooperation. It limits the real growth ofthe intrinsic quality of life, which is wanted by every human being thinking abouthis or her children’s future. It also leads to extensive illegal immigration of thepoor to the richer parts of the world.

Also, all parts of society contribute to the problem. For instance, carbondioxide emissions to the atmosphere due to fossil fuel combustion is caused 50%by industry; the other parts are domestics and transport (the electric power con-tribution is divided between industry and domestics) (6,7).

1.3. Present Technology Is Part of the Problem

Since awareness arose of the large-scale environmental, resource, and socioeco-nomic problems, several technological solutions have been implemented. However,


they did not solve the problems and in some cases caused even bigger problems,as is shown in the cases described later.

1.3.1. High Chimneys for Acidic Flue Gases

In the 1960s a large number of cities in western Europe suffered from smog for-mation. One of the major causes was the local emission of acidic flue gases fromlarge industrial areas. Under still weather conditions, these acid components wouldremain in the atmosphere near the earth’s surface and form smog. The technologi-cal solution was to build very high chimneys—over 200 meters high—cuttingthrough the atmospheric inversion layer. The flue gases where dispersed in this wayover a larger area. In the decade that followed, it appeared that in more rural areas,notably Sweden, these acid gases caused acidification of forests and lakes, withnegative effects on their biological quality. Trees and fish died on a large scale.

1.3.2. CFC Use

In the 1960s and ’70s, CFCs were used on a large scale as propellant gases in spraycans and as a refrigerant in refrigerators. These gases were considered inert tohumans and the environment. Hence, release into the atmosphere was consideredacceptable. In 1973, however, Molina discovered that these CFCs broke down theozone layer high up in the atmosphere (8). This ozone layer itself absorbs the ultra-violet radiation of sunlight and in this way prevents ultraviolet radiation fromreaching the earth’s surface and thereby harming humans (skin cancer formation)and other living species (algae in oceans). Due to the slow breakdown of CFCs theeffect on the ozone layer is expected to last for centuries (9).

1.3.3. HFC Replacement of CFC for Refrigeration

Due to the negative effect of CFCs on the ozone layer, a replacement was devel-oped: HFC. This component hardly breaks down the ozone layer. It is now replac-ing CFCs in domestic refrigerators and air conditioners in cars on a very large scale.However, HFC has a global warming gas potential that is a factor of 1000 higherthan carbon dioxide. The present amount of HFC in the atmosphere is estimatedto account for a global warming effect of 0.6�C (10).

1.3.4. MTBE in Gasoline as a Replacement for Lead Components

In gasoline, lead components are used to enhance the octane number. It becameclear that these lead components, once emitted from car exhaust pipes, causebrain damage. California was the first to ban these lead components and to stim-ulate the use of MTBE as a replacement. MTBE both acts as an octane boosterand enhances the gasoline’s combustion, by which less volatile organic carboncomponents are emitted from the car exhaust pipes. However, it appeared thatMTBE entered underground water reservoirs by occasional spillage of gasoline.This made the water unfit for drinking due to its bad smell. In California billions


of gallons of drinking water are already contaminated in this way, and legislationis under way to ban the use of MTBE in gasoline.

1.3.5. Brent Spar Oil Platform in the North Sea

In the 1970s, the awareness of the limited oil resources resulted in a sudden oilprice increase by a factor of 4. It became economical to produce crude oil fromsea bottoms. One of the first oil platforms in the North Sea was the Brent Spar. Inthe 1990s, oil production ceased and the platform had to be demolished. In theBrent Spar design, demolition had not been taken into account, and so the demo-lition appeared to be very difficult. In consultation with the British government itwas decided that dumping in a deep trough in the ocean would be the best eco-nomical and environmental solution.

The environmental activist organization Greenpeace disagreed with thissolution, arguing that the problem was not the dumping of the Brent Spar; the realissue, they held, was the principle of dumping in the ocean in general and that theBrent Spar dumping would be used by other companies (nuclear energy compa-nies and others) to claim the same right of dumping.

Greenpeace started a campaign to prevent this dumping by asking the gen-eral public to boycott Shell gasoline purchases. Especially in Germany this boy-cott had a noticeable effect, and Shell decided to reverse it decision and to inviteany company or organization to propose alternative solutions. In the end, it wasdecided to bring the platform to a fjord and dismantle it there in such a way thatparts could be reused in other applications.

1.3.6. Preliminary Conclusions on Technology Cases

All the technological solutions mentioned already have in common that the prob-lem to be solved was defined too narrowly in space, time, and lifecycle phases.Wider-ranging and longer-term effects on the local and global scales on humans,nature, and ecology were not taken into account.

2. SUSTAINABLE DEVELOPMENT AND REQUIRED

TECHNOLOGY

2.1. General Concept of Sustainable Development

The general opinion about the aforementioned problems and the ad hoc technicalsolutions is that they are by no means sufficient. The essential concept when think-ing about these subjects nowadays is sustainability. It was greatly helped by thedescription and definition by the Brundtland report of the World Commission onEnvironment and Development (WCED), Our Common Future (11):

Sustainable development is not a fixed state of harmony, but a process ofchange in which the exploitation of resources, the direction of investments,


the orientation of technological development and institutional change aremade consistent with future as well as present needs. Sustainable develop-ment is development that meets the needs of the present generation with-out compromising the ability of future generations to meet their own needs.

This definition contains two key concepts:

The concept of needs, in particular the essential needs of the world’s poor,to which overriding priority should be given

The idea of limitations imposed by the state of technology and social organization on the environment’s ability to meet present and futureneeds (12).

The most admirable aspect of this report is the attempt to bring the twolargest world problems—unlimited exploitation of nature and growing inequitywithin and among nations—together in one concept that everyone from govern-ment to citizen can work on.

Two very important remarks have to be made about the definition of theBrundtland Commission.

1. The definition makes clear that development of new technologies, socialstructures, or whatever has to take into account economic and socialissues ( present generations) and long-term and large-scale environmen-tal issues ( future generations). Thus developments that have to lead tosustainability are limited and have to consider the idea that every humanbeing must be able to fulfill his or her needs in a more or less equal way.

2. We have to distinguish between two different kinds of development: onethat leads to technological innovation, as used in the term research anddevelopment, and one that is about improving the welfare of a society,as used in the term developing countries. This distinction is very impor-tant, because it makes clear that sustainable development has differentmeanings in different countries. In a rich country, sustainable develop-ment sets limits to the growth of affluence or the possibilities of new andinnovative technologies. In poor countries, sustainable development hasto do with helping the population to survive and if possible to thrive.

To further understand what the commission meant by this definition it is neces-sary to read the whole book, because a lot of discussion occurred within the com-mission about poverty and affluence, distribution of knowledge and information,the rights of the northern part of the world to live more affluently than the southernpart, the right to develop, etc. Here we use the summary by Installé (13):

Set priorities among the needs.Ever-lasting growth is unsustainable, because large parts of earth’s natural

capital are not substitutable by human-made capital.


The actual world economy rules are unsustainable.The existence of various cultures should be encouraged.

Intergenerational equity results in a maximum well-being for ourselves under theconstraints that the well-being of future generations must be at least equal to ours.

2.1.1. Ideas Complementary to Brundtland

The definition and vision of the Brundtland Report are described more practicallyby three other United Nations Commissions, the International Union for Conser-vation of Nature and Natural Resources (IUCN), the United Nations EnvironmentProgram (UNEP), and the World Wide Fund For Nature (WWF), called Caringfor Earth: A Strategy for Sustainable Living (14). This report defines sustainabledevelopment as: “Improving the quality of human life while living within the carrying capacity of supporting ecosystems.” This report contains nine chapters,with the following titles, that also can be seen as nine issues of sustainable development.

1. Respecting and caring for the community of life2. Improving the quality of human life3. Conserving the earth’s vitality and diversity4. Minimizing the depletion of nonrenewable resources5. Keeping within the earth’s carrying capacity6. Changing personal attitudes and practices7. Enabling communities to care for their own environments8. Providing a national framework for integrating development and con-

servation9. Creating a global alliance

2.1.2. The Three Parts of Sustainable Development

So far the ideas of sustainable development have been described mainly in glob-al political terms. The financial and business world also took hold of the conceptand reformulated it into three parts: social, ecological, and economic (15,16).This in turn was translated into a catch phrase: People, planet, and profit (17). Wewill summarize these parts.

2.1.2.1 Ecological Part (Planet). The ecological part has to do with theimpact of human action on nature. Mainly this means all the known environmen-tal problems and processes that disrupt the ecosystems (ozone depletion, acidifi-cation, greenhouse effect, destruction of species, wastes, etc.). In a sustainableworld all those known problems must be minimized or avoided. In addition, foras far as possible, sustainable development must have the power to avoid newproblems. The precautionary principle is therefore adopted.


Precaution principle: At the United Nations Conference on Environmentand Development (UNCED) in Rio de Janeiro the (nonbinding) Rio Declarationof 1992 was agreed on, with the following text:

Principle 15.In order to protect the environment, the precautionary approach shall bewidely applied by States according to their capabilities. Where there arethreats of serious or irreversible damage, lack of full scientific certaintyshall not be used as a reason for postponing cost-effective measures toprevent environmental degradation.

This is a better-safe-than-sorry principle that advocates the reduction of inputsinto the environment of substances especially where there is reason to believe thatharmful effects are likely to occur (18). Therefore, the precautionary principle hasto play an important role in decisions about new technology.

2.1.2.2 Social Issues (People). The sustainable development of societyconcerns needs. This means that consumer participation is very important andthat the creation of new markets has to be done with caution. Further, concerningthe lasting factor of societies, it is essential that sustainable development be inagreement with local culture.

2.2.2.3 Economic Issues (Profit). It is obvious that if scarce nonrenew-able resources such as fossil fuels and rich metal ores are going to be depleted,prices will become extremely high. Therefore, economics is very important insustainable development. This is also true for the feasibility of sustainability con-cepts: If they are not profitable, then they will not be accepted. However, it is veryimportant to think over the meaning of the term profitability. Indeed, if externalcosts are taken into account, sustainable technology will be more profitable in thelong term than conventional technology.

External costs is an economic concept, which gives the possibility of settingthe prices right. External costs are the cost involved by the rest of the lifecycle ofproducts and by-products and not accounted for in present cost calculations. Forexample, the price of fossil fuels depends not only on exploration costs but also onthe costs involved by the production of CO2 and resultant adverse climatic effects,which will be many trillions of dollars. In an article in Nature the value of theworld ecosystem services is estimated to be worth 33 1012 $ per year (19). Thismeans that for economic activities that destroy part of the ecosystem, the ecoval-ue (service) reduction should be accounted for.

These external cost can be considerable compared to the present calculatedcost. The present global economic production expressed as GNP is estimated at18 1012 $ per year; so the world ecosystems produce a factor 1.8 higher annualvalue (19). This means that the external cost can even exceed the present costwhen the ecosystem is seriously affected by human activity.


2.1.3. Other Views on Sustainable Development

So far the mainstream thoughts about sustainable development have beendescribed. However, there are other voices, which have a different view on sus-tainable development and critisize (part of) the Brundtland description. Theirgeneral comment is that the Brundtland description stems from an anthropocen-tric worldview that places man too much at the center. Two other views are theecocentric and the theocentric Judeo-Christian worldviews. A brief description ofeach follows, together with their main criticism on the Brundtland view.

2.1.3.1. Ecocentric Worldview. In this view nature has a value of itsown. It places nature at the center and man as a part of nature. Man should nothave a large impact on nature. All biological species should be maintained.Technology should fit into and depend on nature. Nature should be kept in asteady state. If this means that human needs are to be restricted, then that isaccepted. To quote Gandhi: “The world provides sufficient for everybody’s needbut not for everybody’s greed.”

People who have this worldview often criticize the Brundtland descrip-tion because it sees nature mainly as a source for goods and services for man-kind (20).

2.1.3.2. Theocentric Judeo-Christian Worldview. In this view, nature isseen as God’s creation. Man is mandated to be God’s steward for this creation. Atypical Bible text in this respect is: “to work it and to take care of it” (Bible,Genesis 2:15). Also, in this worldview people should take care of their neighborsand in particular take care of the weak and the poor. Furthermore, in this viewman should restrain himself in his needs. He should not be greedy. A compre-hensive description is given by Elsdon (21).

Some criticism of people with this worldview on the Brundtland descrip-tion is that it sees nature as man’s possession and that it does not mention arestriction on man’s greed.

2.2. Sustainable Technology Positioning

In the last 30 years, environmental impact has been mainly reduced by end-of-pipe solutions. These solutions always mean that more construction material andenergy is required; hence they reduce the environmental impact for a certain com-ponent but often increase the impact for another component and are always moreexpensive.

Then the next concept arrived: clean production technology. Jackson (1993)defines clean production as “a conceptual and procedural approach to productionthat demands that all phases of the lifecycle of a product or a process should beaddressed with the objective of prevention or minimization of short- and long-term risks to human health and to the environment (18).”


With this type of technology, products and processes are reformulated tomanage emissions and waste “upstream.” These solutions can mean both lessemission and less cost and energy.

In these views of technology, the social part and the economic parts of sus-tainable development are not taken into account. In the next section we will derivea general view on sustainable technology, first by showing all interactions withthe sustainable development aspects and then some specifics for each sustainabledevelopment part.

2.2.1. Social, Economic, and Ecological Interactions with Sustainable Technology

Providing for the needs of society while staying within ecological constraints is notonly a task for technology, but also the result of the interaction of culture, politicalstructure, and technology; (22). Culture determines the size and the nature of theneeds for which fulfillment is justified and the conditions that technology and struc-ture have to satisfy. Structure is the way in which fulfilling the needs is organizedby means of production and consumption. It has to do with organization, economy,and policy. Technology is not only the total of means available for fulfilling theneeds; technology itself also has an influence on the culture. The latter influence isdifficult to predict but can be discovered from the past.

For instance, the development of lightweight, low-cost plastics led to pack-aging of small-portion consumer items. This in turn enhanced the development ofself-service supermarkets, which led to shopping by car rather then walking tonearby small shops (23). This contributed to a more individual culture. Presentlythe whole infrastructure of a developed country is based on auto transport. Theroad is no longer a meeting place of persons but a depersonalized area, againleading to a more individualistic society.

Stainer et al. give some guidelines for the different actors in a technology-based culture (24):

Engineers: Be aware of the implications for society and environment.Nonscientists: Understand applied science.Companies: Sustainable growth is allied to risk management.Politicians: Develop the skills for sustainable policy.All groups: Recognize the need for communication.

A very interesting concept developed by Moser is the principle of invasive-ness in relation to embeddedness of technology (25). This concept makes it clearthat current technologies have a high impact on the environment (invasiveness) butthat their rapid development and the high public concern about their impact are notvery well anchored in society; this means a low embeddedness. New, cleaner tech-nology, and especially sustainable technology, has to be developed so that they willnot need to change very fast and be more strongly embedded in society.


All this is said to make clear that something has to change drastically in thefocus about using technology if we want to speak about sustainable technology.In sustainable technology, technology is still serving humans and their society(otherwise, we need not speak about technology at all!). However, the technologyshould also be the servant of nature, both in an active and a passive way.

2.2.2. Ecological Impacts as a Function of Cultural, Economic, and Technological Activities

The ecological impacts caused by human activities are large. To give more insightinto the effects of economy, culture, and technology on ecological impacts, thefollowing semiquantitative expression is often used. It is the so-called “master”equation. This expresses the environmental impact as a product of population size(P), gross domestic product per person {GDP/P} and environmental impact perunit GDP {EI/GDP}

EI � P * {GDP/P} * {EI/GDP}

Often the following substitution is made:

GDP/P � W (W meaning “wealth”)

EI/GDP � T (T meaning “technology”)

resulting in

EI � P * W * T

This master equation can then be used to calculate the increase in environmentalimpact for the population, or the increase in wealth with the same technology, or tocalculate the required emission reduction per unit GDP by new technology. If, forinstance, wealth worldwide reaches the same level as now in the industrialized world,then W increases by a factor of 4. If the population (P) does not increase and tech-nology (T) stays the same, then EI (e.g., carbon dioxide emissions to the atmosphere)would increase by a factor of 4. To keep EI at the same level, the technology shouldbe improved to give an emission reduction of a factor 4 per unit product or service.

In reality this use of the expression is too simplistic:

Population growth is a function of the level of wealth. In Western societythe natural population growth is less than zero (the birth rate is less than2 per woman). The same could happen in developing countries once acertain level of wealth has been reached.

Replacing GDP/P by wealth is also under debate. All kinds of occurrences,such as car accidents leading to medical care consumption, are adding toGDP, while these things do not add to wealth at all. Corrections to the GDPfor these aspects have already been proposed by (26).

There is a trend that beyond a certain wealth, the environmental impactemission per unit capita is reduced. Jackson shows that on average above


5000 GDP per capita, the hazardous waste per unit GDP is reduced (18).Hence the level of W also has an effect on EI/GDP.

The environmental impact per unit GDP (T) is not only a function of tech-nology but also a function of human behavior and societal culture.Jackson shows that countries with the same GDP per capita show largedifferences in hazardous waste per unit GDP. For instance, Scandinaviancountries, compared to Canada and the United States show, a similarGDP per capita but up to a factor 4 difference in hazardous waste per unitcapita (18, p. 115). This difference is due to differences in culture.Scandinavian countries value nature more than does the United States. Agood source for further reading on this aspect of wealth increase andimpact reduction is von Weizsäcker et al. (27).

2.2.3. Transition to Sustainable Technology System Levels

Now when we talk about the transition to sustainable technology it is helpful tounderstand that technology has to deal with different system levels, as shown inTable 1. It is obvious that changes in societal infrastructure are more complex andrequire more disciplines and more stakeholder’s involvement. These changes areon a very long time scale (decades to centuries). Scenario building is an impor-tant part of the methods to robust long-term developments (28). This field is toobig to elaborate further in this chapter.

Changes to new, sustainable products or new processes, fitting into existinginfrastructures, are less complex and can be achieved on a shorter time scale. Thesystem levels, from process downward, will be further discussed later.

TABLE 1 System Scale Levels and Technology Disciplines

System scale levels Major disciplines involved

World biotic and abiotic Ecology, politics, physics, chemistrySocietal infrastructure Politics, social science, economics, laws,

ecology, civil engineering, city and landscape architecture

Industrial complex Civil engineering, economics, lawProcess Chemical engineeringUnit operation Chemical engineeringEquipment Chemical engineering, mechanical

engineeringCatalyst/dispersed entity Chemistry, physics, chemical

engineeringNano Chemistry, physics, engineeringMolecules Chemistry, physics


2.3. Sustainable Technology Scorecard

To aid the developer or designer of a new technology the author has derived ascorecard to assess the technology on its contribution to sustainable development.It is based on the sustainable development aspects: social, ecology, and economy.The scorecard is shown in Table 2. The basis and use of the scorecard will nowbe explained.

TABLE 2 Sustainable Technology Scorecard

Sustainable development item Comment

Social/peopleProvide for the needs of the poor: Prime goal

water, food, clothing, local energy, etc

Fair distribution of wealth, power, Adaptable, nondisruptive to societyand knowledge

Social acceptance By stakeholder engagementSafe Loss prevention at all conditionsNoise Below legal limitsSmell No obnoxious emissionsOccupational health Long-term effects of exposure to

chemicals should be known andacceptable

Plot area impact Low; important in denselypopulated areas

Skyline impact Low aesthetic and bird friendly

Ecological/planetSensitivity to world-scale nature No emission of components whose

and ecology ultimate environmental fate isunknown

Depletion of abiotic resources Keep air, surface water, and soilhealthy

Depletion of biotic resources; Maintain biodiversitybiodiversity

Dehydration Maintain water reservoirsOzone depletion potential (ODP) No ODP gas emissions Global warming potential Green house gas emissions reduced

by factor 4 Photochemical ozone pollutants Volatile organic component

emissions below expected futurelegal limits


TABLE 2 (cont.)

Sustainable development item Comment

Acidification Volatile organic component emissions below expected futurelegal limits

Human toxicity Volatile organic component emissions below expected futurelegal limits

Ecotoxicity (terrestrial and aquatic) Volatile organic component emissions below expected futurelegal limits

Nutrification (eutrophication) Volatile organic component emissions below expected futurelegal limits

Radiation Volatile organic component emissions below expected futurelegal limits

Thermal pollution Volatile organic component emissions below expected futurelegal limits

Waste Volatile organic component emissions below expected futurelegal limits

Economy/profitScarce resource depletion No full depletion; replacement by

renewablesDrinking water resource depletion No depletionFossile fuel depletion No full depletion; replacement by

renewablesExternal (future) cost Include in present accountingCapital expenditure and

decommisioning costOperational cost long termProfitable over total lifecycle

2.3.1. Social/People

A prime goal of sustainable development on a world scale is to provide for theneeds of the poor. Throughout one should keep in mind that developed economies have different challenges than survival economies. For survival coun-tries the main goal is to provide for the needs of the people in an affordable waywhile not consuming natural capital (29). Hence this is an important item on the


scorecard. One can think of technology providing water, food, clothing, and cleanenergy for cooking and housing while adapting to the local social, economic(resources), and ecological situation (30).

Another item on the scorecard requiring some explanation is social accept-ance. In the previous sections it is shown that sustainable technology is to beembedded in society. This means, in our view, that society should be consideredin the design stage. A discussion with representatives of stakeholders should beheld to clarify the effects of the new technology on the ecology, society, andeconomy and to seek acceptance. Stakeholders can be environmental activists,civilians, government representatives, manufacturing companies, and consumers.

These stakeholders often have different worldviews. In stakeholder discus-sions this is often not recognized. The previous section on worldviews in relationto sustainable development should help to understand the background of certainarguments used in stakeholder discussions.

Scholes reveals that such a stakeholder discussion was carried out by Shellin the design phase and led to changes in the design and to greater acceptance ofthe new process (31).

2.3.2. Ecological Assessment

The precautionary principle should be applied. This means that for all emissionsover the whole lifecycle, the final environmental fate needs to be known. If thisknowledge is not available, the emission of that component should be kept atzero.

All other scorecard items stem directly from lifecycle analysis assessmenttheory (32). For some impact types, quantitative norms can be stated. For stratos-pheric ozone layer depletion components, the emission norm is zero. This meansin practice that components such as CFCs and SF6 should not be used at all, noteven in contained technical applications, because containment over the lifecyclecannot be ensured.

For other emissions, such as potential global-warming gases, the norm willdepend on the time horizon of the application. If the application is foreseen for along time horizon, say, 50 years, then the global-warming gas emissions shouldbe reduced by a factor of 4–10 compared to the present (27).

To be able to assess the environmental impact of a technology over thewhole lifecycle, a lifecycle assessment (LCA) should be performed. A briefdescription of LCA follows.

Life cycle assessment is a method for calculating the environmental impactof a product or service over its whole lifecycle. Steps in a lifecycle include:

ExplorationRefiningManufacturing


UseEnd of product life (i.e., combustion, reuse, or recycle)Process decommission

A process (manufacturing) is therefore part of a product recycle. Often the decom-missioning phase of the process plant is not taken into account in the LCA. As, forinstance, happened in the Brent Spar case and for most nuclear power plants.

An LCA consists of the following stages:

1. Goal definition and scoping2. Inventory analysis3. Impact assessment (classification)4. Valuation5. Improvement

1. Goal definition and scoping defines the purpose of the LCA and itsscope, the latter meaning depth and broadness. This is important because LCAcan be carried out in great detail, requiring great input of time and money. Whenthe goal is comparing alternatives it is sufficient to gather knowledge that is notabsolutely correct but good enough to make relative comparisons. Scope involvesselecting system boundaries—what and what not to include. For example, infra-structure may include the plant that makes a product but also the whole trans-portation system (e.g., roads, their construction and maintenance) to distribute it.Limiting scope is essential for practical execution of LCA studies.

2. The inventory analysis determines the material and energy inputs andoutputs relating to one functional unit of product or service and links these toenvironmental impacts. This stage is simply a material and energy balance, albeita complete one, that is, covering all lifecycles of the product or service being con-sidered and relating to all inputs and outputs of all of these lifecycles, includingeven micro-outputs, such as (when relevant) micro-emissions of dioxins down tomilligrams per year for plants, with total emissions in megatons per year (e.g.,power plant).

3. Impact assessment translates the inventories into recognized environ-mental impacts and calculates a value for each of these impacts. Generally rec-ognized environmental impacts are given in Table 3. Assigning an emission of agiven substance a given environmental impact depends on the properties of thegiven substance. Emitted sulphur dioxide, for example, is converted into sulphuricacid and is therefore classified as an “acidification” impact. Emitted carbon dioxideand nitrous oxide are greenhouse gases and are therefore classified as “globalwarming” substances.

Once emissions have been classified into a given environmental impact,they must be assessed. This means calculating the quantitative value of thisimpact. Environmental impacts are calculated by means of conversion factors, or


so-called equivalency factors. We illustrate this by using the impact category“global warming.”

Conversion factors for global warming potential (GWP) are expressed interms of CO2 equivalency (kg CO2 equivalent), as given in Table 4. Using GWPvalues, emissions to the air of various substances can be converted to an equiva-lent CO2 global-warming effect by means of the following formula:

Example: The GWP of a factory that per year emits 1 megaton of CO2, 10 kilo-tons of CH4, and 1 kiloton of N2O is

3(1 � 1 megaton � 0.01 � 62 � 0.001 � 290)� 1.352 megatons CO2 equivalent per year

Other environmental impacts can be calculated similarly by using theappropriate equivalency factor. Impacts of human toxicity and ecotoxicity are, forexample, determined by factors such as toxicity data (for humans, resp. variousnonhuman life forms), persistence, and bioaccumulation. Equivalency factors for

Global - warming effect (kg) (Emissions to air) (GWP)� ��

ii

n

i1

∑

TABLE 3 Generally Recognized Environmental Impacts

Depletion Pollution Disturbances

Abiotic Depletion of stratospheric Desiccation resources ozone layer (dehydration)

Biotic resources Global warming Physical ecosystemFormation of degradation and

photochemical landscape oxidants and degradationpollutants Human victims

AcidificationHuman toxicityEcotoxicity (terrestrial and

aquatic)Nutrification (eutrophication)RadiationThermal pollution (dispersion

of heat)NoiseSmellOccupational health


a wide range of substances can be found in the literature (32) and for industrialchemicals (33).

4. Valuation. In the valuation step, all impacts are multiplied by norm fac-tors. The resulting figures are added up to yield a single total figure. The normfactors are subjective, for they express the relative importance one gives to totallydifferent environment effects. This subjectivity can be made more acceptable byhaving the norm factors set by a panel with representatives of various worldviewsor political parties.

5. Improvement. One of the great uses of LCA is to pinpoint places in thelifecycle that cause major environmental impacts and thus lead to highly efficientimprovement of the impact spectrum for a given product.

The LCA results can be used for a sustainability assessment of a new tech-nology by comparing the outcome with the existing technology results. In theconceptual phase of the new technology, a relative assessment can be used tohighlight where improvements should be made and where the problem areas are.This is shown in Section 3 on process intensification assessment.

Detailed Information on the LCA method is provided by Hauschild (32).

2.3.3. Economy/Profit

It is obvious that if scarce nonrenewable resources such as fossil fuels and richmetal ores are going to be depleted, prices will become extremely high. Whichmeans that the next generation will experience higher costs for the same goods,which is not fair. If scarce resources are depleted, then developments should alsostart to provide alternatives at a similar price, so-called “strong sustainabilityeconomic development” (2). Obvious scarce resources—water and fossil fuels—are explicitly put on the scorecard.

The future high resource cost and the external cost related to emissionsshould also be taken into account in the total lifecycle cost.

Long-term profit is also placed on the scorecard. If a technology is not prof-itable it will not be accepted by business. However, it is very important to think

TABLE 4 Global Warming Potential (GWP) for a Number of Substances (32)

GWPSubstance Substance formula (kg CO2/kg substance)

Carbon dioxide CO2 1Methane CH4 62Nitrous oxide N2O 290CFC11 CFCl3 5,000

Source: Hauschild, 1998, vol. 1.


over the meaning of the term long-term profitability. It is recommended that theexternal costs, as far as they are known, be included in the cost calculations, forin the long term these external costs will likely become internalized via taxationor otherwise. Some external cost figures can be obtained from the literature. Table 5gives the values for carbon dioxide and volatile organic components.

The estimation of external cost is a scientific field in development. To findmore data, the reader could search the literature on “external cost” or “future cost.”

2.3.4. Use of the Scorecard in the Conceptual Design Phase

In the conceptual design phase a quantitative assessment requiring lots of informa-tion is not desirable. A qualitative assessment highlighting the main improvementsand areas of major concern requiring alternative designs is more suitable. The score-card can be used qualitatively, for instance, by estimating whether the new technol-ogy improves individual scorecard items as compared to existing technologies.

Note: There is a tendency to arrive at a single sustainability score. In theauthor’s opinion little is gained and a lot is lost by the use of a single figure. What is lost is the information about where the problems with the new technolo-gy are located. This in turn means that suggestions for improvements cannot befound. Also, the danger of balancing poor-scoring areas with very positive areasbecomes real. A sustainable technology should score positive on all three legs(social, ecology, economy) and should have no significant negative score in anyimpact area.

3. THE POTENTIAL FOR PROCESS INTENSIFICATION

TO CONTRIBUTE TO SUSTAINABLE DEVELOPMENT

Elaborate descriptions of process intensification will be found elsewhere in thisbook. Here a short description suffices, followed by an assessment on its poten-tial to contribute to sustainable development, based on present industrial casesand general features.

TABLE 5 Estimates of External Cost for Two Components Relevant toChemical Processes

Emission type External cost

Carbon dioxide net emission to atmosphere $46/ton (34)Volatile organic carbon emission to atmosphere $8000/ton (34)

Source: SIKA, 1995.


3.1. Definition of Process Intensification

Process intensification consists of the development of novel apparatuses and techniques that, compared to those commonly used today, are expected to bringdramatic improvements in manufacturing and processing, substantially decreas-ing the ratio of equipment size to production capacity, energy consumption,or waste production and ultimately resulting in cheaper and more sustainabletechnologies (35).

Stankiewicz and Moulijn further divide PI into two areas: equipment andmethods. They give an extensive list of equipment examples. The methods aresubdivided into multifunctional reactors, hybrid separations, alternative energysources, and other methods.

For multifunctional equipment, Siirola provides a general description offunctions and how to combine them into a piece of equipment (36). Functionsoperating in the same temperature and pressure range can be combined in onepiece of equipment. Multifunctional reactors represent one subclass of solutionswith promising features in cost and energy requirements (37).

3.2. Present Status of Process Intensification from Industrial Cases

We now apply the criteria for sustainable technologies derived in Section 2, Table 2,to industrial PI cases. Table 6 shows the results. A short description of the casesfollows. Italic numbers in the table are derived from the references, as is explainedin the following text.

3.2.1. Industrial Cases of Process Intensification: Background for Table 6

3.2.1.1. Case E: Eastman Chemical Methyl Acetate Process (Siirola,1995, 1998). In the commercial operation of the Eastman methyl acetateprocess, many reaction and separation functions are combined in one single largecolumn (a few meters in diameter and 80 meters high (38). The number of piecesof major equipment compared to a conventional design is reduced by a factor of10 (36). The primary energy consumption and capital expenditure is reduced bya factor of 5 (38).

The plot area of this process is probably a factor of 10 smaller than the con-ventional process, because the number of major units has been reduced by a fac-tor of 10.

The major contributions to global warming, acidification, thermal pollu-tion, and fossil fuel depletion are probably all directly related to the energyrequired. Because this is reduced by a factor of 5, it is assumed that these contri-butions reduce by the same factor, 5.


TABLE 6 Sustainable Technology Scorecard Results for Industrial PI Cases

Improvement potential (reduction factor)

Sustainable development item E S SG G SP Range

Social/peopleProvide for the needs of the poor:

water, food, clothing Fair distribution of wealth, power,

and knowledgeSocial acceptanceSafe: lower reactive, dangerous content Yes YesNoiseSmellOccupational health —Size (equipment volume reduction) 3 3Smaller plot area impact 10 4 4–10Smaller skyline impact

Ecological/planetSensitivity to world-scale nature

and ecologyDepletion of abiotic resources

(clean air, etc.)Depletion of biotic resources;

biodiversityDehydrationDepletion of stratospheric ozone layerGlobal warming 5 >1.4 1.4 1.2 1.2–5Formation of photochemical 10 1.2 12–10

pollutantsAcidification 5 >1.4 1.4 1.2 1.2–5Human toxicityEcotoxicity (terrestrial and aquatic)Nutrification (eutrophication) RadiationThermal pollution 5 16 1.4 1.2 1.2–16Waste 3.5 3.5

Economy/profitScarce resource depletionDrinking water resource depletionFossil fuel depletion 5 >1.4 1.2 1.2–5External (future) cost low Capital expenditure 5 0.8–1.6 1.6 0.8–5Operational cost 5 1.4 1.6 1.4–5Profitable over total lifecycle

Derived from process data, assuming other lifecycle steps are hardly affected.


It is likely that the number of flanges and valves has been reduced at leastby the same number as the pieces of equipment. The VOC reduction in theEastman case can therefore be a factor of 10 lower than for the conventionaldesign.

3.2.1.2. Case S: Sulzer Hydrogen Peroxide Distillation System (39).The safety in this new process is drastically improved because of low operatingtemperature, minimal product holdup in the system, reliable safety devices, andproper selection and treatment of the construction materials (39).

The number of units has been reduced by a factor of 4; therefore it isassumed that the plot area is reduced by the same factor, 4.

The energy consumption has been reduced by 30% (reduction factor is 1.4).We have used this factor for global warming, acidification, and fossil fuel deple-tion. The thermal pollution from cooling water is reduced by a factor of 16 andthe operational cost by a factor of 5 (39).

3.2.1.3. Case SG: Shell Global Solutions Natural Gas Dehydration (40).This case involves membrane separation for natural gas conditioning (dehydra-tion). The mass of equipment is reduced by 70% (factor of 3 reduction) comparedto the conventional process. From this we derived that the construction volume isreduced by the same factor, 3. The capital expenditure and operation cost reduc-tion figures are given by Rijkens (40).

3.2.1.4. Case G: GlaxoSmithKline Fine Chemical from Carbonyl Process(41). The fine chemical is produced in a high-heat exchange reactor. The resi-dence time is thereby reduced by a factor of 1800(!) compared to a conventionalbatch reactor. The reactive content is thereby considerably reduced; hence theprocess is safer.

The CO2 emission related to this process is reduced by a factor of 1.4 (from18 to 13 g CO2/mol product). The global warming, acidification, and thermal pol-lution are assumed to be reduced by the same factor.

The waste is reduced by a factor of 3.5.

3.2.1.5. Case SP: Shell International Chemicals—PDC Bulk ChemicalProcess Design (37). This new process design consists of a multifunctionalreactor, which largely combines reaction, heat exchange, and component separa-tion in one piece of equipment. Data on energy and cost element reductions areprovided.

3.2.2. Conclusion on the Industrial PI Cases

It is clear that for all cases, considerable improvements are made compared toconventional processes in some criteria in the social, ecological, and economicareas. It is also clear that for most criteria, no assessment can be made becausedata are not reported.


3.3. Potential Contribution of Process Intensification

to Sustainable Development

From the definition and the examples given it is clear that process intensificationdoes not belong to end-of-pipe technology. Most example cases belong to the cat-egory of clean technology. In the following section we will concentrate on itspotential for contributing to sustainable development.

Table 7 contains a summary of the potential of process intensification for sus-tainable development. Some aspects will be explained in the sections that follow.

3.3.1. Social/People

3.3.1.1. Providing for Needs, Especially of the Poor. Process intensifi-cation is mainly concerned with processes and less with consumer products.Moreover it is only located on the system-level process. So a big overall contri-bution to sustainable development is not expected. However, by miniaturizationof certain processes those processes may become part of a higher system leveland have a larger impact then initially imagined. Take, for instance, the develop-ment of small fuel cells for electricity production in remote villages in developingcountries, fed with hydrogen from biomass waste. These fuel cells may be mass-produced at low cost, by which they enhance the quality of life in developing coun-tries.

3.3.1.2. Social Acceptance. Social acceptance of chemical plants is stillan issue. This is due to the fact that big disasters have occasionally occurred(Bhopal, Seveso) and also because chemical plants smell, due to diffusive emis-sions of volatile components. Via process intensification the amount of chemical-ly hazardous, reactive material can be reduced considerably, by which the size ofan emission in the case of an explosion will be far less and the chance of an explo-sion itself reduced by the lowered hazardous, reactive content.

Often this low reactive content can lead to designing for total containment;i.e., no relief system is required by which further emission reductions are achieved.Moreover, the diffusive emissions (smell) can be reduced by reducing the num-ber of flanges and valves. Because the size of the plant will also be reduced, it islikely that the social status of a chemical plant can be raised by the introductionof process intensification.

3.3.2. Ecology/Planet

3.3.2.1. Dematerialization. By scale, the reduction of material andenergy intensity by process intensification should reduce environmental impact.

3.3.2.2. Global-Warming Impact Reduction. There is a general reasonwhy PI, via by function integration, can result in lower energy requirements andthereby in lower carbon dioxide emissions (given the present way of providing


TABLE 7 Sustainable Technology Scorecard for the Potential Contributionof PI to Sustainable Development

Process intensification Sustainable development item improvement potential

Social/peopleProvide for the needs of the poor: water, Not identified yet

food, clothing Fair distribution of wealth, power, and Not identified yet

knowledgeSocial acceptance Probably yes Safe: lower reactive, dangerous content Yes; factor > 10 possibleNoise Yes; fewer pumps and

pieces of equipmentSmell Yes; fewer flanges, less

diffusive emissionsOccupational health Yes; by less diffusive

emissionsConstruction volume reduction Yes; factor > 4 possiblePlot area impact Yes; factor > 4 possibleSkyline impact Yes; factor > 4 possible

Ecological/planetSensitivity to world-scale nature To be assessed for each case

and ecologyDepletion of abiotic resources Yes; lower emissions

(clean air, etc.)Depletion of biotic resources; biodiversity Not identifiedDehydration Not identifiedDepletion of stratospheric ozone layerGlobal warming Yes; far less energy requiredFormation of photochemical pollutants Yes; less diffusive emissions

VOCAcidification Yes; less energy requiredHuman toxicityEcotoxicity (terrestrial and aquatic)Nutrification (eutrophication) RadiationThermal pollution Yes; less energy requiredWaste Yes

Economy/profitScarce resource depletion To be assessed for each caseDrinking water resource depletion Not identified yetFossil fuel depletion Yes; improvement factor > 4External (future) cost low YesLower capital expenditure Yes; factor > 4Lower operational cost Yes; factor > 4Profitable over total lifecycle Likely


power via fossil fuel combustion) and thereby reducing the impact of the chemi-cal process industry (from CO2 emissions) on global warming.

The basic principle behind high-quality energy carriers (fossil fuels, solarenergy) follows from thermodynamics. All process steps are carried out in a finitetime and are therefore irreversible, by which high-quality energy is converted tolower-quality energy. This is called exergy loss. By reducing the number of highlyirreversible process steps, exergy losses are reduced (42).

This can be achieved by combining process functions into one process step. Itcan also be obtained by reducing the number of units just by selecting the best unitsand their sequence. It can also be achieved by reducing the biggest driving forces byevenly spreading them over the process steps. This will result in general in energysavings and in reductions of carbon dioxide emission (43). For distillations the sav-ings are of the order of 30% (reduction by a factor of 1.4) (43).

Individual pieces of PI equipment, as listed by Stankiewicz and Moulijn(5), requiring alternative energy sources such as microwaves and centrifugalfields will often result in higher energy requirements. Once again, the processstep is made more irreversible (faster by higher driving forces). By assessing theenergy requirements of the whole lifecycle, energy requirement reductions maybe achieved, via, for instance, a higher selectivity, i.e., a reduction in feedstock con-sumption. And feedstock preparation is also associated with energy requirements.

3.3.2.3. Photochemical Pollutant Reduction. Smog is caused by photo-chemical oxidants. The formation of photochemical oxidants is mainly due tovolatile organic compound (VOC) emissions. Seventy to ninety percent of VOCemissions from chemical plants are from leaking flanges, seals, etc. (44). Processintensification can reduce considerably the number of pieces of equipment andthereby the number of flanges and seals. Moreover, the few remaining flangescould be designed such that fugitive emissions are below the limits for smog for-mation, and the additional cost can be limited. This means fewer leaks and lowerfugitive emissions to air, water, and soil. The introduction of process intensifica-tion could therefore make a major contribution to VOC emission reduction.

3.3.2.4. Acidification. The acidification and thermal pollution impactsof a chemical plant in general do not come from the plant itself but are related tothe conversion of fossil fuel to energy (steam, electricity); hence the same emis-sion reduction factor as for carbon dioxide is assumed for the process intensifica-tion potential.

3.3.3. Economy/Profit

3.3.3.1. Scarce Resource Depletion. For any new technology, the long-term and worldwide effects on scarce-resource consumption should be assessedover its entire lifecycle. For instance, even minute amounts of rare earth metals per


application can, in the long run (and when the technology is applied worldwide),deplete that scarce resource. (This is of course not only an economic issue butalso an ethical and ecological one).

The present resources of drinking water and fossil carbon components arealready seen as limited. Careful use and development of alternatives is now beingpursued. Process intensification has the potential of requiring less energy andthereby requiring less fossil carbon as fuel.

3.3.3.2. Economic Cost Reduction Potential. The major part of the costof manufacturing chemical products is the feedstocks. The capital expenditure perunit product is in general between 10% and 30%. The energy cost is in generalbetween 5% and 10%.

With bulk chemicals the feedstock costs are in general close to the theoreti-cal minimum, due to the high selectivity achieved (90% and higher). The introduc-tion of PI in this field can lead to a reduction in capital expenditure and of energyutilities by a factor 5 (38); a variable cost reduction factor of 1.4 is achievable.

In the fine chemicals industry the selectivity figures are in general far lower.Here big feedstock cost savings is achievable by more selective processes. This ismainly achieved by changing from stoichiometric acid and alkaline feed-drivenprocesses to truly catalytic processes (45).

But via, for instance, high-heat exchange reactors (HEX reactors) thatallow a much shorter reaction time and at the optimal temperature, process inten-sification can also reduce the formation of by-products considerably. A by-productreduction by a factor of 4 is achieved via a HEX reactor (46). This means a con-siderable reduction in feedstock cost and waste handling.

3.4. Traps for Process Intensification Leading to

Unsustainable Technologies

Larger emissions or consumption of scarce resource could occur elsewhere in thelifecycle via the large-scale introduction of process intensification. For example,the proposed printed circuit board reactors could in the manufacturing step causelarge emissions of VOCs from cleaning chemicals. And the demolition phase couldalso cause a large waste stream.

4. EPILOGUE

It is clear from the foregoing that presently known PI methods can contribute con-siderably to the process industry in meeting the social, ecological, and economicconstraints of sustainable development. For each new application it remains achallenge to the engineers to identify ways to provide for the needs of people,especially the poor, while meeting all SD constraints.


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