comparative study of water-gas shift reactors with and
TRANSCRIPT
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Comparative Study of Water-Gas Shift Reactors
with and without Selective Membranes
Wesam Adieb Abbas Hassan
B.Sc.(Honors), in Chemical Engineering
University of Khartoum (2009)
A Dissertation
Submitted to the University of Gezira in Partial Fulfillment of the
Requirements for the Award of the Degree of Master of Science
in
Chemical Engineering
Department of Chemical Engineering and Chemical Technology
Faculty of Engineering and Technology
April, 2018
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Comparative Study of Water-Gas Shift Reactors
with and without Selective Membranes
Wesam Adieb Abbas Hassan
Supervision Committee:
Name Position Signature
Prof. Kamil Mohammed -Elhassan Main Supervisor …………
Dr. Fathelrahman Abbas Elshikh Co-supervisor …………
Date: April , 2018
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Comparative Study of Water-Gas Shift Reactors
with and without Selective Membranes
Wesam Adieb Abbas Hassan
Examination Committee:
Name Position Signature
Prof. Kamil Mohammed -Elhassan Chair person …………
Prof. Babiker Karama Abdalla
Dr. Magdi Ali Osman
External Examiner
Internal Examiner
................
…………
Date of Examination: 14/ April/ 2018
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Dedication
To My Parents:
Adeib and Manal
To My Brothers And My Sister.
TO My Teachers And My Colleagues.
To All My Friends.
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Acknowledgement
Above all, I render my thanks to the merciful “Allah” who offered me the health and the patience
to accomplish this research.
I would like to express my gratitude and thanks to my research supervisor Prof. Kamil
Mohammed –Elhassan for his support, help and constructive advice during planning and
development of this research. Thanks and appreciations are also extended to my Co-Supervisor
Dr. Fathelrahman Abbas Elshikh.
Deep thanks are extended to Dr. Imadeldeen Abdelmoniem Mahjoub the head of the
Department of Chemical Engineering and Chemical Technology, University of Gezira, and Dr.
Alaedin Mohamed Elhassan the head of the Department of Mining Engineering, University of
Khartoum, for their help and support.
Finally, sincere appreciation with thanks to my family and friends for their support and
encouragements.
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Abstract
The water-gas shift reaction (WGSR) is a well-known step for upgrading Carbon monoxide to
carbon dioxide in the production of purified Hydrogen gas. For more than 90 years since its first
industrial application, many issues with respect to the catalyst, process configuration, reactor
design, reaction mechanisms and kinetics have been investigated. More recently, a renewed
interest in the (WGSR) carried out in hydrogen selective membrane reactors has been observed,
because of the growing use of polymeric electrolyte membrane (PEM) fuel cells, that operate using
high-purity hydrogen. Moreover, membrane reactors are viewed as an interesting technology in
order to overcome the equilibrium conversion limitations in non-membrane reactors (traditional
packed bed reactors). The objective of this study is to investigate the usage of membrane reactor
technology in the (WGSR) to examine the improvement extent of carbon monoxide conversion
that possible through the use of this technology, to overcome the thermodynamic limitations for
this reversible reaction. Furthermore, an assessment of the optimum parameters for the maximum
conversion of carbon monoxide was carried out. The study was divided into two parts. In the first
part the modeling and simulation of non-membrane reactor was carried out. The obtained results
were validated against the experimental results of Sing and Saraf (1980). In the second part of the
study, the membrane reactor was simulated by employing a hydrogen-selective Palladium (Pd)
membrane. A one dimensional steady state model was developed; mass and heat balances were
solved simultaneously using MATLAB. It was found that the conversion of carbon monoxide
reaches a maximum of 29% for the packed-bed reactor and 90% for the membrane reactor, which
represents an improvement of 210%. It was found that the optimum operating conditions for
membrane reactor were: temperature of 740 K, pressure of 10 atm and a reactor length and
diameter of 115 cm and 25 cm respectively. It is recommended that a complete economic analysis
should be carried out to assess the commercial viability for the usage of the membrane reactor in
gaseous reversible reactions.
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ملخص الدراسة
غاز الور ون في إأتاج( هو خطوة معروفة لترقية أول أكسييييل الور ون إلي ياأي أكسييييل WGSRإن تفاعل تحول الغاز المائي )
عاما منذ أول تطبيق صيييييناعيا تح البح في العل ل من القفيييييا ا فيما تعتق و يييييائ تحفي 90الهيلروجين النقي. لأكثر من
التفاعلا وتوو ن أظح العمتيةا وتصييييييميح المفاعلا والياا التفاعل الحركية. وفي الأوأة الأخيرةا لوما اهتماد متفلع تفاعل
الغشيييييائية المائي في المفاعلاا الغشيييييائية ائأتقائية لتهيلروجينا وسلب سيييييب ائ يييييتخلاد المت ا ل لخلا ا الوقوع تحول الغاز
التي تعمل ا تخلاد هيلروجين عالي النقاوة. وعلاوة عتى سلبا نظر إلى المفاعلاا الغشائية عتى أأها ةالبوليمر الإلوتروليتية
ائت ان في المفاعلاا التقتيل ة. الهلف من هذه اللرا يية هو التحقيق في ا ييتخلاد التغت عتى قيوعتقنية مثيرة للاهتماد من أجل
تونولوجيا المفاعل الغشيييائي في تفاعل تحول الغاز المائي للرا ييية ملن تحسيييين أسيييبة تحول أول أكسييييل الور ون الممونة من
ميوية الحرار ة لهذا التفاعل العوسيييييي و علاوة عتى سلب تقييح روف خلال ا يييييتخلاد هذه التونولوجياا لتتغت عتى القيوع الل نا
أمذجة ومحاكاة التفاعل المثتى لأقصيييى قلر من تحول أول أكسييييل الور ون. قسيييم اللرا ييية إلى ج ئينلأ في الف الأول تم
(. في 1980) Sarafو Singل النتائب التفر بية مفاعل مشييواا تقتيلم من عون غشييا و تح التحقق من صييحة النتائب مقا تت
لتهيلروجينلأ طور أموسج كحالة يا ا تخلاد غشا البلاع ود ائأتقائي الف الثاأي من اللرا ةا تح محاكاة المفاعل الغشائي
ساا عل وامل ومت موازأاا الوتتة والحرارة في وق وامل ا يييييتخلاد رأامب الماتلاد. وجل أن أسيييييبة تحول أول أكسييييييل
. وقل وجل أن ٪ 210لتمفاعل الغشائيا وهو ما مثل تحسنا نسبة ٪90كحل أقصى لتمفاعل التقتيلم و ٪29ر ون صل إلى الو
ضيييغ جوم وطول مفاعل 10عرجة مطتقةا ضيييغ 740 روف التشيييغيل المثتى لمفاعل الغشيييا كاأ عنل عرجة مرارة
اقتصييياعم كامل لتقييح الفلون التفار ة ئ يييتخلاد المفاعل يييح .توصيييي اللرا ييية يجرا تحتيل 25وعرض مفاعل يييح115
الغشائي في التفاعلاا الغاز ة العوسية.
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Table of Contents
Title Page No.
Dedication IV
Acknowledgement V
Abstract VI
Arabic Abstract VII
Tablet of Contents VIII
List of Tables X
List of Figures XI
Nomenclature XII
Chapter One: 1. Introduction
1.1 Background 1
1.1.1 Membrane 1
1.1.2 Catalytic Membrane Reactor 2
1.1.3 Water-Gas Shift Reaction 2
1.2 Problem Statement 2
1.3 Objectives 2
Chapter Two: 2. Literature Review
2.1 Membrane 3
2.1.1 Historical Development of Membranes 3
2.1.2 Types of Membranes 5
2.1.2.1 Microporous Membranes 5
2.1.2.2 Noneporous, Dense Membranes 5
2.1.2.3 Electrically Charged Membranes 6
2.1.2.4 Antistropic Membranes 6
2.1.2.5 Ceramic, Metal And Liquid Membranes 7
2.1.3 Using Membranes As Reactors 7
2.2 Membrane Reactors 11
2.2.1 Definition 11
2.2.2 Types of Membrane Reactor 11
2.2.2.1 Inert Membrane Reactor 11
2.2.2.2 Catalytic Membrane Reactor (CMR) 11
2.2.3 Application of Membrane Reactors 12
2.2.3.1 Cell Culture And Fermentation Processes 12
2.2.3.2 Light Hydrocarbon Gas-Phase Catalytic Reactions 15
2.2.4 Future Directions of Membrane Reactors 16
2.3 Water- Gas Shift Reaction 18
2.3.1 Back ground 18
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2.3.2 Reaction Kinetics 21
2.3.3 Description of Water –Gas Shift Reaction In Traditional Reactor 23
2.3.4 Selection of Water-Gas Shift Reaction Catalyst for Membrane
Reactor Applications
25
2.3.5 The Water- Gas Shift Reaction in Membrane Reactors 26
Chapter Three: 3. Materials and Methods
3.1 Packed Bed Reactor (TR) 29
3.1.1 Rate Equation 29
3.1.2 Description of Mathematical Model 31
3.1.3 Catalyst Specification 33
3.14 Operating Conditions 33
3.2 Membrane Reactor 34
Chapter Four: 4. Results and Discussions
4.1 Model Validation 35
4.2 Comparing the Performance of both Traditional and Membrane
Reactors
36
4.3 Sensitivity Analysis 38
4.3.1 Temperature Effect 38
4.3.2 Pressure Effect 39
4.3.3 Reactor Length Effect 41
4.3.4 Transport Coefficient Effect 42
Chapter Five : 5. Conclusions and Recommendations
5.1 Conclusions 43
5.2 Recommendations 43
References 44
Appendices 46
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List of Tables
Table Title Page No.
(2-1) Petrochemical reactions being considered as application for
membrane reactor
15
(2-2) Example of Kinetic Expression for Water-Gas Shift Reaction 22
(4-1) Experimental data and calculated result for packed bed reactor 35
(4-2) Comparison between packed bed reactor and membrane reactor 36
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List of Figures
Figure Title Page No.
(2.1) Examples of the three types of membrane reactor 9
(2.2) Examples of two types of chemical reaction and separation 10
(2.3) Continuous recycle fermenter membrane reactor 13
(2.4) A hollow fiber membrane reactor. 14
(2.5) Methylcyclohexane conversion to toluene as a function of reactor
temperature in a membrane and a non-membrane reactor
16
(2.6) Summarized sketch of the Haber–Bosch process 19
(2.7) Conventional two-stage process diagram of the WGS reaction unit 24
(2.8) Hydrogen production and purification based on the WGS MR unit 28
(4.1) Packed bed reactor profile 37
(4.2) Membrane reactor profile 37
(4.3) Carbon monoxide (fractional) VS Temperature 38
(4.4) Carbon monoxide (fractional) VS Pressure 40
(4.5) Carbon monoxide (fractional) VS Reactor Length 41
(4.6) Carbon monoxide (fractional) VS Transport Coefficient
42
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Nomenclature:
𝑨 Membrane surface area ( 𝑐𝑚2)
𝑨𝒈𝒇 Aging factor which accounts for loss in activity of the catalyst with usage
𝐶𝑝𝒊 Heat capacity of component i
D Reactor diameter (cm)
𝑬𝒂 Apparent activation energy
𝑬𝒇𝒇 effectiveness factor which accounts for intrapellet diffusional resistance.
G Volumetric flow rate (𝑐𝑚3 ℎ⁄ ).
𝑯 Heat of reaction at T (𝑐𝑎𝑙 𝑚𝑜𝑙⁄ )
𝑱𝑯 The hydrogen flux through membrane ( 𝑚𝑜𝑙
𝑐𝑚2.ℎ)
𝒌𝒆𝒒 Equilibrium constant
L Reactor length (cm)
𝐏 Total pressure
𝑷𝑷 Hydrogen partial pressure on the permeate sides
𝑷𝒇 accounts for the effect of pressure on the rate of reaction
𝑷𝒆 Permeability
𝑷𝑹 Hydrogen partial pressure on the retentate and
R Gas constant,(𝑐𝑎𝑙
𝑔.𝑚𝑜𝑙 𝐾).
𝒓 Rate of reaction,( cm3of H2
h.g of catyst).
T Temperature, K
𝑻𝒓𝒆𝒇𝒇 Reference Temperature (K)
𝒘 Mass of catalyst (𝑔)
𝒙𝒊 Mole fraction of component i
𝒙𝒄𝒐∗ Carbon monoxide mole fraction at equilibrium condition.
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Greek letters
𝜹 Membrane thickness.
𝝉 Age of the catalyst (ℎ)
Abbreviation
CMR Catalytic membrane reactor
IMRCF Inert membrane reactor with catalyst on the feed side
MR Membrane reactor
TR Traditional reactor
WGSR Water gas shift reaction
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Chapter One
1. Introduction
1.1 Background
1.1.1 Membranes
Membranes and membrane processes are considered as old inventions and are being increasingly
improve upon in recent years. Synthetic membranes have rapidly gained substantial importance
due to their large number of their practical applications. Nowadays, membranes are used on a large
industrial scale such as production of potable water from the sea, to clean industrial effluent and
the recovery of valuable constituents, to concentrate, purify or fractionate macromolecular
mixtures in the foods and drugs industries and to separate gases and vapors. They are also the key
components in energy conversion systems, artificial organs and drugs delivery devices. The
discovery of new membranes materials was the key factor for increasing the applications of the
membranes in the catalysis field. The significant progress in this area is reflected by the increasing
numbers of the scientific publications, which have grown exponentially over the last few years
(Babiker et al., 2016).
Synthetic membrane, is a synthetically created membrane which is usually intended for separation
purposes in laboratory or industry, it has been successfully used for small and large-scale industrial
processes since the middle of the twentieth century. There is variety of synthetic membranes, they
can be produced from organic materials such as polymers, as well as inorganic materials. The
membrane process is known as a filtration process. The chemical and physical properties of
synthetic membranes and separated particle as well as a choice of driving force define a particular
membrane separation processes (Babiker et al., 2016).
1.1.2 Catalytic Membrane Reactors
Catalytic membrane reactor (CMR) can be define a combination of a heterogeneous catalyst
and a perm-selective membrane, which is a thin film or layer that allows at least one component
of a mixture to selectively permeate through it (Richard, 2004).
A catalytic membrane reactor usually operates at higher yields, better reaction selectivity (as
opposed to separation selectivity) or lower cost than the separate catalytic reactor and downstream
separation units. Though the use of catalytic membrane reactors is not widespread at present, the
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development of new membranes particularly porous ceramic and zeolite membranes creates the
potential to significantly improve yields of many catalytic processes.
1.1.3 Water-Gas Shift Reaction
The Water-Gas Shift Reaction (WGS R) equation below:
CO + H2O CO2 +H2
is an old industrial process discovered in (1780) by (Felic Fotona) which water in the form of
steam is mixed with carbon monoxide to obtain hydrogen and carbon dioxide. It is used in
manufacture of ammonia, hydrogen, methanol, hydrocarbons and in conjunction of steam
reforming. It found new significance in fuel processing, in conjunction with fuel cells; It provides
a method for extracting the energy from the toxic carbon monoxide by converting it to usable
hydrogen along with carbon dioxide which can be tolerated by the fuel cell (Caitlin, 2006).
1.2 Problem statement
The Water-Gas Shift Reaction is reversible and exothermic (ΔH= -41.2 KJ/mol), therefore the
equilibrium constant increases with decreasing temperature, but the reaction rate decreases with
decreasing temperature, for this reason Water-Gas Shift Reaction is thermodynamically
unfavorable at elevated temperature, and equilibrium conversion limitations occurs in the Packed
Bed Reactor.
1.3 Objectives
The main objective of this research is to study the membrane in the Membrane Reactor ability to
increase the yield of product compared to that of the Packed Bed Reactor.
The specific objectives are:
To investigate the reactor membrane technology in the Water-Gas Shift Reaction .
To see the extent of the possible carbon monoxide conversion improvement through the use
of this technology.
To evaluate the optimum parameters for the maximum yield of products.
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Chapter Two
2. Literature Review
2.1 Membranes
2.1.1 Historical Development of Membranes
In chemical technology membranes have gained an important place and they are used in a broad
range of applications. The key property that is the ability of a membrane to control the permeation
rate of a chemical species through the membrane (Richard, 2004).
We can traced Systematic studies of membrane phenomena to the eighteenth century philosopher
scientists. For example, Abb´e Nolet coined the word ‘osmosis’to describe permeation of water
through a diaphragm in 1748. Through the nineteenth and early twentieth centuries, membranes
had no industrial or commercial uses, but were used as laboratory tools to develop
physical/chemical theories. For example, the measurements of solution osmotic pressure made
with membranes by Traube and Pfeffer were used by van’t Hoff in 1887 to develop his limit law,
which explains the behavior of ideal dilute solutions; this work led directly to the van’t Hoff
equation. At about the same time, the concept of a perfectly selective semi permeable membrane
was used by Maxwell and others in developing the kinetic theory of gases (Richard,2004).
Bachmann and Ferry improved on Bechhold’s technique, and by the early 1930s microporous
collodion membranes were commercially available. During the next 20 years, this early
microfiltration membrane technology was expanded to other polymers, notably cellulose acetate.
Membranes found their first significant application in the testing of drinking water at the end of
World War II. Drinking water supplies serving large communities in Germany and where in
Europe had broken down, and filters to test for water safety were needed urgently. The research
effort to develop these filters, sponsored by the US Army, was later exploited by the Millipore
Corporation, the first and still the largest US microfiltration membrane producer (Richard,2004).
By 1960, the elements of modern membrane science had been developed, but membranes were
used in only a few laboratory and small, specialized industrial applications. No significant
membrane industry existed, and total annual sales of membranes for all industrial applications
probably did not exceed US$20 million in 2003 dollars. Membranes suffered from four problems
that prohibited their widespread use as a separation process: They were too unreliable, too slow,
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too unselective, and too expensive. Solutions to each of these problems have been developed
during the last 30 years, and membrane-based separation processes are now commonplace
(Richard, 2004).
The seminal discovery that transformed membrane separation from a laboratory to an industrial
process was the development, in the early 1960s, of the Loeb–Sourirajan process for making
defect-free, high-flux, anisotropic reverse osmosis membranes . These membranes consist of an
ultrathin, selective surface film on a much thicker but much more permeable micro porous support,
which provides the mechanical strength. The flux of the first Loeb–Sourirajan reverse osmosis
membrane was 10 times higher than that of any membrane then available and made reverse
osmosis a potentially practical method of desalting water. The work of Loeb and Sourirajan, and
the timely infusion of large sums of research and development dollars from the US Department of
Interior, Office of Saline Water (OSW), resulted in the commercialization of reverse osmosis and
was a major factor in the development of ultra filtration and microfiltration. The development of
electro dialysis was also aided by OSW funding. Concurrently with the development of these
industrial applications of membranes was the independent development of membranes for medical
separation processes, in particular, the artificial kidney. W.J. Kolf had demonstrated the first
successful artificial kidney in the Netherlands in 1945. It took almost 20 years to refine the
technology for use on a large scale, but these developments were complete by the early 1960s.
Since then, the use of membranes in artificial organs has become a major life-saving procedure.
More than 800 000 people are now sustained by artificial kidneys and a further million people
undergo open-heart surgery each year, a procedure made possible by development of the
membrane blood oxygenator. The sales of these devices comfortably exceed the total industrial
membrane separation market. Another important medical application of membranes is for
controlled drug delivery systems. A key figure in this area was Alex Zaffaroni, who founded Alza,
a company dedicated to developing these products in 1966. The membrane techniques developed
by Alza andits competitors are widely used in the pharmaceutical industry to improve the
efficiency and safety of drug delivery (Richard, 2004).
The period from 1960 to 1980 produced a significant change in the status of membrane technology.
Building on the original Loeb–Sourirajan technique, other membrane formation processes,
including interfacial polymerization and multilayer composite casting and coating, were developed
for making high performance membranes. Using these processes, membranes with selective layers
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as thin as 0.1 μm or less are now being produced by a number of companies. Methods of packaging
membranes into large-membrane-area spiral-wound, hollow-fine-fiber, capillary, and plate-and-
frame modules were also developed, and advances were made in improving membrane stability.
By 1980, microfiltration, ultra filtration, reverse osmosis and electro dialysis were all established
processes with large plants installed worldwide (Richard, 2004).
2.1.2Types of Membranes
The topic is still large enough to include a wide variety of membranes that differ in chemical and
physical composition and in the way they operate. In essence, a membrane is nothing more than a
discrete, thin interface that moderates the permeation of chemical species in contact with it. This
interface may be molecularly homogeneous, that is, completely uniform in composition and
structure, or it may be chemically or physically heterogeneous, for example, containing holes or
pores of finite dimensions or consisting of some form of layered structure. A normal filter meets
this definition of a membrane, but, by convention, the term filter is usually limited to structures
that separate particulate suspensions larger than 1 to 10 μm. The principal types of membrane are
described briefly below (Richard,2004).
2.1.2.1 Microporous Membranes
A microporous membrane is very similar in structure and function to a conventional filter. It has a
rigid, highly voided structure with randomly distributed, interconnected pores. However, these
pores differ from those in a conventional filter by being extremely small, on the order of 0.01 to
10 μm in diameter. All particles larger than the largest pores are completely rejected by the
membrane. Particles smaller than the largest pores, but larger than the smallest pores are partially
rejected, according to the pore size distribution of the membrane. Particles much smaller than the
smallest pores will pass through the membrane. Thus, separation of solutes by microporous
membranes is mainly a function of molecular size and pore size distribution. In general, only
molecules that differ considerably in size can be separated effectively by microporous membranes,
for example, in ultra filtration and microfiltration.
2.1.2.2 Nonporous, Dense Membranes
Nonporous, dense membranes consist of a dense film through which permeants are transported by
diffusion under the driving force of a pressure, concentration, or electrical potential gradient. The
separation of various components of a mixture is related directly to their relative transport rate
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within the membrane, which is determined by their diffusivity and solubility in the membrane
material.
Thus, nonporous, dense membranes can separate permeants of similar size if their concentration
in the membrane material (that is, their solubility) differs significantly.
Most gas separation, pervaporation, and reverse osmosis membranes use dense membranes to
perform the separation. Usually these membranes have an anisotropic structure to improve the
flux.
2.1.2.3 Electrically Charged Membranes
Electrically charged membranes can be dense or microporous, but are most commonly very finely
microporous, with the pore walls carrying fixed positively or negatively charged ions. A membrane
with fixed positively charged ions is referred to as an anion-exchange membrane because it binds
anions in the surrounding fluid. Similarly, a membrane containing fixed negatively charged ions
is called a cation-exchange membrane. Separation with charged membranes is achieved mainly by
exclusion of ions of the same charge as the fixed ions of the membrane structure, and to a much
lesser extent by the pore size. The separation is affected by the charge and concentration of the
ions in solution. For example, monovalent ions are excluded less effectively than divalent ions
and, in solutions of high ionic strength, selectivity decreases. Electrically charged membranes are
used for processing electrolyte solutions in electrodialysis.
2.1.2.4 Anisotropic Membranes
The transport rate of a species through a membrane is inversely proportional to the membrane
thickness. High transport rates are desirable in membrane separation processes for economic
reasons; therefore, the membrane should be as thin as possible. Conventional film fabrication
technology limits manufacture of mechanically strong, defect-free films to about 20 μm thickness.
The development of novel membrane fabrication techniques to produce anisotropic membrane
structures was one of the major breakthroughs of membrane technology during the past 30 years.
Anisotropic membranes consist of an extremely thin surface layer supported on a much thicker,
porous substructure. The surface layer and its substructure may be formed in a single operation or
separately. In composite membranes, the layers are usually made from different polymers. The
separation properties and permeation rates of the membrane are determined exclusively by the
surface layer; the substructure functions as a mechanical support. The advantages of the higher
20
fluxes provided by anisotropic membranes are so great that ;almost all commercial processes use
such membranes.
2.1.2.5 Ceramic, Metal and Liquid Membranes:
The discussion so far implies that membrane materials are organic polymers but in fact, the vast
majority of membranes used commercially are polymer-based. However, in recent years, interest
in membranes formed from less conventional materials has increased. Ceramic membranes, a
special class of microporous membranes, are being used in ultrafiltration and microfiltration
applications for which solvent resistance and thermal stability are required. Dense metal
membranes, particularly palladium membranes, are being considered for the separation of
hydrogen from gas mixtures, and supported liquid films are being developed for carrier-facilitated
transport processes.
2.1.3 Using Membrane as Reactors
By the early 1980s membrane technology had developed to the point at which a number of
industrial groups began to consider using membranes to control the products of chemical reactions.
Two properties of membranes are used for the types of membrane reactor (Richard,2004).The first
type is the membrane as a contactor illustrated in Figure 2.1(a).The membrane separates the
reaction medium in one chamber from a second chamber containing a catalyst, enzymes or a cell
culture. This type of application has a long history in fermentation processes involving so-called
(Cheryan, 1998).More recently, membrane reactors are being developed for conventional chemical
separations (Kemmere and Keurentjes, 2001). As the reaction medium flows through the first
chamber, membrane reactants diffuse through the membrane, react in the second chamber, and
then diffuse back out to be collected as a product stream. The membrane provides a large exchange
area between the catalytic material and the reaction medium but performs no separation function
(Richard, 2004).
The second type of membrane reactor illustrated in Figure 2.1(b),uses the separative properties of
a membrane. In this example, the membrane shifts the equilibrium of a chemical reaction by
selectively removing one of the components of the reaction. For example the important
dehydrogenation reaction converting n-butane to butadiene and hydrogen (Richard, 2004).
𝐶4𝐻10 ↔ 𝐶4𝐻6 + 2 𝐻2
21
Removing hydrogen from the reaction chamber by permeation through the membrane causes the
chemical equilibrium to shift to the right, and the conversion of to butadiene increases (Rezac et
al.,1998).
A third type of membrane reactor combines the functions of contactor and separator. An example
of this combination membrane reactor is shown in Figure 2.1(c), in which the membrane is a
multilayer composite. The layer facing the organic feed stream is an immobilized organic liquid
membrane; the layer facing the aqueous product solution contains an enzyme catalyst for the
desterification reaction (Richard, 2004).
𝐻2𝑂 + 𝑅 − 𝐶𝑂𝑂𝐻 ↔ 𝑅 − 𝐶𝑂𝑂𝐻 + 𝑅 − 𝑅𝑂𝐻
Organic-soluble ester is brought to the reactor with the organic feed solution and freely permeates
the immobilized organic liquid membrane to reach the catalyst enzyme. The ester is then
hydrolyzed. The alcohol and acid products of hydrolysis are much more polar than the ester and,
as such, are water soluble but relatively organic insoluble. These products diffuse to the aqueous
permeate solution. The membrane both provides an active site for the reaction and separates the
products of reaction from the feed (Matson and Quinn , 1992).
In the membrane reactor shown in (Figure 2.1(c)), the chemical reaction and the separation step
use the same membrane. However, in some processes it is desirable to separate reaction and
separation into two distinct operations. If the net result of the process is to change the products of
the chemical reaction, the process is still classified under the broad heading of membrane reactor.
Two examplesin which chemical reaction and separation are physically separated are shown in
Figure 2.2, (Figure 2.2(a)) shows the use of a pervaporation membrane to shift the equilibrium of
the de-etherification reaction .A portion of the organic solution in the esterification reactor is
continuously circulated past the surface of a water-permeable membrane. Water produced in the
esterification reaction is removed through the membrane. By removing the water, the reaction can
be driven to completion. (Figure 2.2(b)) shows the use of a hydrogen-permeable membrane to shift
the equilibrium of the n-butane dehydrogenation reaction. The catalytic reactor is divided into
steps, and the hydrogen-permeable membrane placed between each step. Because the hydrogen is
removed from the reactor in two discrete steps, some inefficiency results, but separating the
membrane separation step from the catalytic reactor allows the gas to be cooled before being sent
to the membrane separator. Polymeric membranes can then be used for the gas separation operation
22
(Rezac et al., 1998).Such membranes can remove hydrogen very efficiently from the butane–
butadiene/hydrogen mixture but cannot be used at the 400–500 ℃operating temperature of the
catalytic reactor (Richard, 2004).
Figure 2.1: Examples of the three types of membrane reactor (Richard, 2004).
23
Figure 2.2: Examples of the two types of chemical reaction and separation
(Richard, 2004).
24
2.2 Membrane Reactors
2.2.1 Definition
Considering a IUPAC definition, a membrane reactor (MR) is a device for simultaneously
performing a reaction and a membrane based separation in the same physical device. It is not only
plays the role of separator, but also takes place in the reaction itself. It is just a plug flow reactor
that contains an additional cylinder with some porous material within it like the tube of the ( shell
and tube heat exchanger ). This inner porous cylinder is the membrane that gives this reactor it is
name (Babiker et al., 2016).
2.2.2 Types of membrane reactors
Membrane reactors are most commonly used when a reaction involves some form of catalyst and
there are two types of these membrane reactors (Babiker et al., 2016).
2.2.2.1 Inert membrane reactor
It is known as (IMRCF) which stands for inert membrane reactors with catalyst on the feed side.
It allows catalyst pellets to flow with the reactants on the feed side (usually the inside of the
membrane).In this kind the membrane does not participate in the reaction directly; it simply acts
as a barrier to the reactant and some products.
2.2.2.2 Catalytic membrane reactor(CMR)
It has a membrane that has either been coated with or is made a material that contains catalyst,
which means that membrane itself participates in the reaction.
Some of the reaction products (those that are small enough) pass through the membrane and exit
the reactor on the permeate side (Richard, 2004)
A catalytic membrane reactor (CMR) is a combination of a heterogeneous catalyst and a
permselective membrane, which is a thin film or layer that allows at least one component of a
mixture to selectively permeate through it. A CMR usually operates at higher yields, better reaction
selectivity (as opposed to separation selectivity), or lower cost than a separate catalytic reactor and
downstream separation units. Though the use of CMRs is not widespread at present, the
development of new membranes, particularly porous ceramic and zeolite membranes, creates the
potential to significantly improve yields of many catalytic processes. (Hsieh 1989) and (Armor
1989) recently reviewed inorganic membrane reactors (Richard, 2004).
25
(Roth 1990), in a review of the future opportunities in industrial catalysis, indicated that we are at
a threshold of major changes in separation technology (particularly in a shift from distillation to
separation by synthetic membranes), these changes will have substantial impact on chemical
process technology. He stated that membranes will be of importance in the emerging area of
catalytic membrane reactors. He referenced a quote by Gryaznov of the former Soviet Union:
Priority in our country should be given to the use of membranecatalysis in all plants that utilize
the selective hydrogenation of acetylenicalcohols to ethylenic alcohols, nitro compounds to
amines, and the hydrogenation of cyclopentadiene to monomer for synthetic rubber (Richard,
2004).
Roth further pointed out a number of opportunities for CMRs for equilibrium limited reactions,
such as the production of ethylene from ethane, ammonia synthesis, methanol synthesis, and the
water gas shift reaction. Itoh (1990) stated recently, however, that for practical applications of
membrane reactor technology, further developments of technology to manufacture membranes that
possess high selectivity, high permeability, and high temperature durability are necessary
(Richard, 2004).
2.2.3Applications of Membrane Reactors
Membrane reactors are being considered for many processes, and some are already being used on
an industrial scale. The three main application categories are described briefly below (Richard,
2004).
2.2.3.1 Cell Culture and Fermentation Processes
The traditional, and still the most common, fermentation process involves the addition of microbial
cell cultures to the reaction medium in a batch reactor. This type of batch process is inherently
slow, and microbial cells are lost with each batch of product. Recently there has been a great deal
of interest in developing continuous fermentation processes using membrane bioreactors
(Cheryan, 1998), (Cheryan and Mehaia, 1986), (Cheryan andMehaia, 1985). Much of this work
has concentrated on fermentation of ethanol or acetone/butanol from low-grade food processing
waste such as cheese whey, using a recycle membrane reactor design as shown in (Figure2.3) The
principal advantages of the reactor are its continuous operation, the high cell densities that are
maintained, and the lack of build-up of reaction products that inhibit the reaction.
Another type of microbiological reactor is the hollow fiber membrane bioreactor shown in (Figure
2.4). In this device, the microbial cells are trapped on the shell side of a capillary hollow fiber
26
module. The feed solution, containing substrate and the products of microbial reaction, is
circulated down the bore of the fibers (Kanazek et al., 1972), (Hu and Dodge, 1985).This device
has proved particularly useful in producing protein antibodies by genetically engineered
mammalian cells. By manipulating the molecular weight cut-off of the fiber, the flux of molecules
of different molecular weights across the filter can be controlled. Very high cell densities can be
achieved in these hollow fiber cartridges, which have been used to produce monoclonal antibodies.
Figure 2.3: Continuous recycle fermentor membrane reactor (Richard, 2004).
27
Figure 2.4: A hollow fiber membrane reactor (Richard, 2004).
2.2.3.2 Light Hydrocarbon Gas-phase Catalytic Reactions
28
Several important refinery and chemical feedstock reactions appear to be good candidates for
membrane reactor systems; some such reactions are listed in Table (2.1). Because of the high
temperatures involved, developing the appropriate selective membranes is difficult, and this type
of membrane reactor has not moved beyond the laboratory stage. The first four reactions listed in
Table (2-1) are dehydrogenation reactions in which one of the reaction products is hydrogen. By
removing hydrogen, the reaction can be driven to completion, increasing the degree of conversion
the dehydrogenated product significantly. An example of the improvement in that is possible is
shown in Figure (2.5) (Ali and Rippin, 1995).
Table (2-1): Petrochemical reactions being considered as applications for
membrane reactor
Source:(Ali and Rippin, 1995).
29
Figure 2.5: Methylcyclohexane conversion to toluene as a function of reactor
temperature in a membrane and a nonmembrane reactor (Ali and Rippin,
1995).
30
In (Figure 2.5), the fractional conversion of methylcyclohexane to toluene in a simple tube reactor
is compared to that in a reactor with hydrogen-permeable, palladium-silver alloy walls. Without
the membrane, the degree of conversion is limited to the equilibrium value of the reaction. By
removing the hydrogen, higher degrees of conversion can be achieved. (Figure 2.5) also illustrates
the problem that has inhibited widespread use of membrane reactors the high temperature of the
reactions. The reactions listed in Table 2.1 are all normally performed at300–500 ◦C. These
temperatures are far above the normal operating range of polymeric membranes, so hydrogen-
permeable metal membranes, micro porous carbon membranes, or ceramic membranes must be
used.
2.2.4 Future Directions of Membrane reactors
Membrane reactors have the greatest potential to develop into large-scale commercial processes.
This technologies are already used on a small scale in niche applications, and they are being
developed for much larger and more important processes.
Membrane reactors are currently used only in a few specialized biotech applications. However,
long-term, this type of device will be used in the petrochemical industry, with very different
membranes, for dehydrogenation processes or the partial oxidation of methane. Such applications
could be much more important, but the development of suitable membranes poses a number of
very challenging technical problems (Richard, 2004).
31
2.3 Water Gas Shift Reaction
2.3.1. Background
The blossoming of the water-gas shift (WGS) reaction is related to the production of a combustible
gas(water gas ↔ CO + H2) when steam is passed through a bed of incandescent coke, first
observed by Felice Fontana in 1780.
The following century was marked by notorious discoveries and developments related to alkali
chemical manufacture and gas production by gasification (Ihde, 1984).
Ludwig Mond, one of the greatest chemist-industrialists of all time, focused part of his industrial
chemical technology developments on the synthesis of ammonia from coal and coke. Mond
developed the process for producing the so-called ‘Mond gas’ (the product of the reaction of air
and steam passed through coal/coke – CO2, CO, H2, N2, etc.), which became the basis for future
coal gasification processes. Mond and his assistant Carl Lange were the first to use the term ‘fuel
cells’ while performing experiments with the world’s first working fuel cell. using coal-derived
Mond gas (Hoogers, 2002).
One of the hardest tasks was to feed pure hydrogen to the ‘Mond battery’ due to the large quantities
of carbon monoxide present in Mond gas, which poisoned the Pt electrode. Therefore, Mond
solved this problem by passing the Mond gas mixture and steam over finely divided nickel at 400
◦C, reacting the carbon monoxide and steam to give carbon dioxide and more hydrogen. After CO2
removal by a simple alkaline wash, the H2-rich stream obtained could be successfully fed to the
hydrogen cell .
The WGS reaction equation below :
CO + H2O CO2 +H2
ΔH298 k= -41.2 KJ/mol
was discovered and for the first time reported in the literature by the end of the19th century (Mond
and Langer, 1888).
In 1913, the WGS reaction found industrial application in the production of synthesis gas (or water
gas), as a part of the Haber–Bosch process of ammonia manufacture (Figure2.6) Since the Fe based
catalyst for ammonia synthesis is deactivated by carbon oxides, the WGS reaction became an
important step toward CO upgrade to H2. In terms of process, coal/coke was first blasted with air
in a water-gas generator, exothermically producing carbon dioxide and carbon monoxide, thereby
32
maintaining the high temperature(HT) of the coal (around 1000 ◦C) and raising it into an
incandescent form. The water gas (CO + H2)was then produced via an endothermic reaction (C +
H2O → CO + H2), by passing steam through the fuel bed. After the removal of dust particles, the
mixed gas was added to steam, and CO was catalytically shifted in a converter to produce more
hydrogen and carbon dioxide. Then, the gas was dried and compressed and passed through caustic
scrubbers for CO2removal. By passing the gas mixture counter currently over an ammonia
calcuprous solution, the CO was posteriorly eliminated by absorption (Quirk, 1963).
Figure 2.6: Summarized sketch of the Haber–Bosch process (Mendes et al.,
2009).
33
Industrially, the process integration of the WGS reaction is dependent upon the origin of the
synthesis gas (Mendes et al., 2009).
By the beginning of the 20th century, and because the major source of synthesis gas production
was from coal and coke, the WGS reaction was used as a standalone process. By that time, the
most common and economical design was to conduct the reaction in a single stage, at temperatures
around 450–600 ◦C, and employing as catalyst Fe oxide stabilized in Cr oxide (Quirk, 1963).
The next evolution of the process was the introduction of a second-stage converter at temperatures
around320–360 ◦C using the same catalyst. The two-stage converter systems reduced the CO level
to 3000–4000 ppm compared to the single-stage converters that could not reduce the CO content
to much less than 10 000 ppm(1%) (Ladebeck and Wang , 2003).
With the discovery in the 1960s of Cu-based low-temperature (LT) shift catalysts and
improvements in HT Fe-based shift catalysts, a CO content <0.5% in the reformate stream was
achieved CO content <0.5% in the reformate stream was achieved (Ruettinger and Ilinich, 2006).
From the beginning of the 20th century until today, the use of the WGS reaction followed the
increased industrial demand for hydrogen production. This has been accomplished using natural
gas as feedstock instead of coal, and employing better catalysts that improve the yields and permit
the adjustment of the H2 to CO ratio of the product, mainly for ammonia and methanol synthesis
but also for the Fisher–Tropsch process and in refining operations (desulfuration, hydrogenation
processes, etc.). With the growing concerns about environmental issues, H2 production from
synthesis gas for fuel cell applications has become a huge focus of attention. In this context, new
catalysts were developed in order to obey rigid safety requirements(in addition to fuel cell
requirements) such as lower operation temperature, use of nonpyrophoric material sand high
attrition resistance, improving simultaneously the WGS activity for on-board hydrogen processing.
To obtain high-purity hydrogen from either synthesis gas or from the products of the WGS
reaction, competitive separation processes were investigated with the aim of overcoming the
performance limitations and costs of traditional methods (Mendes et al., 2009).
2.3.2 Reaction Kinetics
34
The kinetic models provide the easiest way to represent the reaction and help the designers in
determining the rate of the reaction and thus design the reactors. Basically the kinetic expressions
can be classified as microkinetic approach and the empirical method. The microkinetic method is
based on the knowledge about the elementary steps that are involved in the reaction and its
energetic. This method explores the detailed chemistry of the reaction. Using this method it is
possible to estimate the surface coverage, reaction order and activation enthalpy. This method
provides the accurate pathway and prediction of the reaction, but is computationally intensive. On
the other hand, the empirical models are based other experimental results and are typically
expressed in the Arrhenius model and provide an easy and computationally lighter way to predict
the rate of reaction. Most of the design works use the empirical models. The WGSR has been
explained through both the microkinetic approach and empirical approach (Smith R J et al., 2010).
Almost similar kinetic expressions have been proposed for both the high temperature and low
temperature catalysts. The kinetic models that have been endorsed by many authors from their
experiments with various catalysts have been the Langmuir Hinshelwood model and the power
law model. The recent literature publications use the kinetic expressions of Keiski et al. (1996),
San et al.(2009) for high temperature and Choi and Stenger (2003) for low temperature water gas
shift reaction (Smith R J et al., 2010).
The difference in opinion on the nature of kinetics for the water gas shift reaction has been
attributed to the presence of impurities, mass transfer limitations, experiments carried out at
atmospheric pressure and the use of integral reactor for kinetic studies rather than the differential
reactor (Levent, 2001).
Since all the experiments were reported at atmospheric pressures, the kinetic models can be
corrected for pressure using the modification recommended by Rase (1977) which also takes care
of the diffusion effects of the catalysts or the pressure correction factor of Singh and Saraf (1977)
(Smith R J et al., 2010).
35
Table 2-2: Examples of Kinetic Expressions for WGSR Source: (Smith R J et al., 2010).
2.3.3Description of the (WGS) Reaction in Traditional Reactors
36
The CO-shift process facilities are economically dependent on the feedstock used for synthesis gas
generation, both in terms of equipment and catalysts. Methane is actually the preferred raw
material, with naphtha preferentially used in areas where natural gas is not available.
Taking into account the exothermic nature of the WGS reaction, higher CO conversions are
favored at lower temperatures. But kinetics are almost instantaneous at very high temperatures.
Because of the inherent temperature increase during the reaction, synthesis gas generation is
typically conducted in adiabatic CO shift reactors to avoid catalyst overheating and improve the
reaction conversion Figure (2.6). As mentioned previously, even though the reaction is
thermodynamically favored at lower temperatures, the reaction kinetics is penalized and a large
volume of catalyst is needed to approach high CO equilibrium conversions. The process is
characterized by low space velocities, <20 000 ℎ−1,currently representing the largest reactor in a
fuel processor (Ruettinger and Ilinich, 2006).
37
Figure 2.7 :Conventional two-stage process diagram of the WGS reaction unit
(Mendes et al., 2009).
38
In a typical industrial operation, Ni-based catalysts are primarily used for the SR of natural gas at
high temperatures (>700 ◦C) using high H2O/CO ratios (at least 3–5). Then, the cooled gas from
the reformer(at about 350–450 ◦C) is driven into a HT CO-shift converter Figure (2.7) commonly
in the temperature range of 320–360 ◦C and at a total pressure between 10and 60 bar, containing
an Fe-based oxide catalyst.
Depending on the feedstock and the performance of the HT CO-shift reactor, an outlet stream with
a CO concentration between 1 and 5% is typically obtained at a temperature around 400–450 ◦C.
This stream is then cooled to about 200 ◦C (by thermal quenching with water or by an inter-cooler
system between stages) (Ladebec and Wang, 2003).
To avoid water condensation in the catalyst pores, and its subsequent deactivation, the inlet
temperature is usually kept at least about 20 ◦C above the dew point of the feed gas (Ruettinger
and Ilinich, 2006).
In some operations, water is injected between the stages to adjust the S/G ratio before entering into
an LT CO-shift reactor, operating in the range of 190–250 ◦C and 10–40 bar. Typical temperature
gradients of 20–30 ◦C in the catalyst bed are observed, and an outlet stream with a CO
concentration less than0.5% is obtained (Ruettinger and Ilinich, 2006).
Further reduction of CO and CO2 to acceptable limits for proton exchange membrane fuel cell
(PEMFC) applications and for ammonia synthesis involves others H2 clean-up systems. In
particular, CO concentration ∼2 ppm for on-board vehicle applications (2007 DOE Guidelines)
poisons the Pt catalyst of the PEMFC (Mendes et al., 2009).
2.3.4Selection of WGS catalysts for membrane reactor applications
It is known that the WGS reaction products (CO2+ H2) inhibit the reaction and lower the reaction
rate over WGS catalysts. This inhibition is dependent on the nature of the catalyst and temperature
range and, therefore, the reaction is clearly more effective in membrane reactors (MRs).(Mendes
et al., 2009). Depending on the properties of the membrane, there are two possible outcomes. In
thecae of a CO2-selective membrane, the concentration of H2 along the catalyst bed will attain high
levels. As referred to above, an excess of H2is negative for a traditional Fe-based catalyst due to
over-reduction of the magnetite active phase (Mendes et al., 2009).
.
The other case is the use of an H2 selective membrane, so CO2 will be present at a higher
concentration in the reaction medium, affecting the reaction rate. Therefore, a new challenge is
39
presented in making highly effective conventional HT catalysts used in traditional WGS processes,
suitable for MR applications. Generally, for large stationary fuel processing plants, Lund (Lund,
2002) proposed that weakening the surface oxygen bond strength would lead to higher catalytic
activity as well as to a reduction of the inhibition by CO2. identified Cu–Ce (La-doped) Ox catalyst
as the best composition for MR applications in the temperature range 300–600 ◦C using simulated
coal gas (Mendes et al., 2009).
Other promising catalysts for WGS MR applications are the Pt-based catalysts, which generally
reveal a lower inhibition effect by the forward WGS reaction products over a wide temperature
range, i.e. relatively low partial orders with respect to both CO2 and H2 (Mendes et al., 2009).
Even though there is not much data available in the literature for Au-based catalysts catalyzing the
WGS reaction using a real reformate feed composition, they are very promising when compared
to the traditional Cu-based catalysts because their performance is less affected by the presence of
reaction products (although they exhibit higher inhibition by CO2 and H2 than Pt-based catalysts).
2.3.5 The Water-Gas Shift Reaction In Membrane Reactors
Conventionally, the WGS reaction is limited by thermodynamic constrains, as previously
mentioned: its conversion may be closer to the thermodynamic predictions, depending on the
suitable choice of catalyst. In other words, the laws of thermodynamics set a rigid limit for the
conversion achievable in TRs in which this reaction proceeds only to partial completion. As a
consequence, the interest of scientists seems quite justified in searching for alternatives to TRs.
Among different technologies, the membrane one seems to be promising (Mendes et al., 2009).
In particular, due to the attractive possibility of realizing both reaction and gas
separation/purification in the same device, MRs are currently considered as good candidates for
replacing TRs. With respect to a classic configuration of a TR consisting of a reactor unit in series
with a separation unit, an MR represents a modern solution having many potential advantages:
reduced capital and downstream separation costs, as well as enhanced yields and selectivities
(Mendes et al., 2009).
From the viewpoint of the WGS process in an MR, a reaction product (e.g. hydrogen, in the case
of Pd membranes) moves to the permeate side, enabling the WGS reaction to proceed toward
completion and so making it possible to achieve the higher conversion than a TR working under
the same operating conditions or the same conversion as a TR, but working under milder operative
conditions. In fact, great interest toward the WGS reaction assisted by MRs has been evidenced in
40
the literature, and many studies are focused on hydrogen recovery from a catalytic shift MR, either
using Pd based or silica membranes (Mendes et al., 2009).
Finally, the combination of low cost and long-lasting hydrogen separation membranes and the
WGS reaction Figure (2.8) has been the focus of attention of several R&D programs. Membranes
and a WGS reactor can be combined in an MR unit with synergistic benefits. For equilibrium-
limited reactions, the continuous removal of a reactant from the reaction medium shifts the reaction
toward the formation of products, thereby giving a higher conversion. Depending on the type of
membrane, either hydrogen or carbon dioxide can be selectively removed in principle, having both
strategic advantages and disadvantages (Mendes et al., 2009).
41
Figure 2.8: Hydrogen production and purification based on the WGS MR unit
(Mendes et al., 2009).
42
Chapter Three
3. Materials and Methods
In this chapter for both types of reactor mathematical modeling were developed and validated
using Sing and Saraf (1980) experimental data. The isothermal simulations were done using
MATLAB software. This simulation is divided in two sections in the first one it was only
considered the reaction where the products remain in the reactor section, it was simulated the
conversion obtained in the TR, in which the maximum attainable reaction conversion is
intrinsically limited by the equilibrium. Following the MR, where besides the chemical reaction it
was considered the permeation of one product of the reaction through the wall (hydrogen-selective
membrane).
3.1 Packed Bed Reactor
In this type of reactor the maximum conversion obtained for the WGS reaction is the equilibrium-
conversion, because the products obtained are not removed.
3.1.1Rate Equation
The water-gas shift reaction over a low-temperature catalyst, similar to that over a high-
temperature catalyst, has been used. This rate equation takes into account the effects of
temperature, pressure, and age of the catalyst on the catalyst activity. It also considers the reduction
in reaction rate due to diffusional resistances. The following equation has been used to represent
the rate of the shift reaction over the catalyst pellets (Singh and Saraf, 1980).
𝒓 = 𝑬𝒇𝒇 × 𝟐. 𝟗𝟓𝟓 × 𝟏𝟎𝟏𝟑𝐞𝐱𝐩 (−𝟐𝟎𝟗𝟔𝟎
𝑹×𝑻) × 𝑨𝒈𝒇 × 𝑷𝒇(𝒙𝒄𝒐 − 𝒙𝒄𝒐
∗) (3.1)
Where:
𝒓 is rate of reaction,( cm3of H2
h.g of catyst).
𝑹 is gas constant,(𝑐𝑎𝑙
𝑔.𝑚𝑜𝑙 𝐾).
𝑻 is temperature, K.
𝑬𝒇𝒇 is the effectiveness factor which accounts for intrapellet diffusional resistance.
𝑨𝒈𝒇 is an aging factor which accounts for loss in activity of the catalyst with usage.
43
This has been correlated with temperature and age from the data reported for the catalyst (Singh
and Saraf,1980) as:
𝐥𝐨𝐠 𝑨𝒈𝒇 = (𝟒. 𝟔𝟔 × 𝟏𝟎−𝟒 − 𝟏. 𝟔 × 𝟏𝟎−𝟔𝑻) × 𝝉 (3.2)
Where:
𝝉 is the age of the catalyst, (days).
𝑷𝒇 is accounts for the effect of pressure on the rate of reaction, the following expression valid for
HT catalysts has been used here :
𝑷𝒇= 𝑷(𝟎.𝟓−(𝑷 𝟐𝟓𝟎⁄ )) (3.3)
𝐏 is total pressure, (atm).
Also:
𝒙co is mole fraction of CO.
𝒙𝒄𝒐∗ is mole fraction of CO in equilibrium condition.
𝒙𝒄𝒐∗ = (𝒙𝑯𝟐
× 𝒙𝒄𝒐𝟐)/(𝒙𝑯𝟐𝒐 × 𝒌𝒆𝒒) (3.4)
Where
𝒙𝑯𝟐is mole fraction of 𝑯𝟐 in equilibrium condition
𝒙𝒄𝒐𝟐 is mole fraction of 𝒄𝒐𝟐 in equilibrium condition
𝒙𝑯𝟐𝒐 is mole fraction of 𝑯𝟐𝑶 in equilibrium condition
𝒌𝒆𝒒 is equilibrium constant
𝒌𝒆𝒒 = 𝐞𝐱𝐩[((𝟗𝟗𝟗𝟖.𝟐𝟐
𝐓) − 𝟏𝟎. 𝟐𝟏𝟑 + 𝟐. 𝟕𝟒𝟔𝟓 × 𝟏𝟎−𝟑𝐓 − 𝟎. 𝟒𝟓𝟑 × 𝟏𝟎−𝟔𝐓𝟐 − 𝟎. 𝟐𝟎𝟏 × 𝐥𝐧𝐓)/𝐑]
(3.5)
44
3.1.2 Description of the Mathematical Model
A general mass balance can be written as follow:
𝑨𝒄𝒄𝒖𝒎𝒖𝒍𝒂𝒕𝒊𝒐𝒏 𝒊=𝑰𝒏𝒊- 𝑶𝒖𝒕 𝒊 − 𝑷𝒆𝒓𝒎𝒂𝒕𝒊𝒐𝒏𝒊+ 𝑮𝒆𝒏𝒓𝒂𝒕𝒊𝒐𝒏𝒊 (3.6)
Where 𝐴𝑐𝑐𝑢𝑚𝑢𝑙𝑎𝑡𝑖𝑜𝑛𝑖 is the rate of accumulation of for species i , 𝐼𝑛𝑖 is the input flow
rate, 𝑂𝑢𝑡 𝑖is the output rate flow, 𝑃𝑒𝑟𝑚𝑎𝑡𝑖𝑜𝑛𝑖is the rate flow of permeation, 𝐺𝑒𝑛𝑟𝑎𝑡𝑖𝑜𝑛𝑖 the rate
of generation by chemical reaction.
In this study it was considered isothermal operation and the system in steady-state and then the
term that refers to accumulation does not exist Therefore, the general balance for a traditional
packed-bed reactor becomes:
𝑶𝒖𝒕 𝒊 − 𝑰𝒏𝒊 = 𝑮𝒆𝒏𝒓𝒂𝒕𝒊𝒐𝒏𝒊 (3.7)
Obviously, the permeation term only appears in the membrane reactor.
The high temperature shift reactors are single bed adiabatic reactors for which the following
mathematical model has been developed. A differential cross section of the catalyst bed is
considered throughout which temperature and composition are assumed constant. Axial diffusion
of mass and heat has been neglected. The material and energy balance over such a differential
section subject to the above assumptions yield the following equations which describe the
composition and temperature of the reaction system along it (Singh and Saraf,1980).
The system of differential equations :
𝒅𝒙𝐜𝐨
𝒅𝒗= −𝒓 ×
𝒘
𝑮 (3.8)
𝒅𝒙𝑯𝟐𝒐
𝒅𝒗= −𝒓 ×
𝒘
𝑮 (3.9)
𝒅𝒙𝒄𝒐𝟐
𝒅𝒗= 𝒓 ×
𝒘
𝑮 (3.10)
𝒅𝒙𝑯𝟐
𝒅𝒗= 𝒓 ×
𝒘
𝑮 (3.11)
𝒅𝒙𝑪𝑯𝟒
𝒅𝒗= 𝟎 (3.12)
45
𝒅𝒙𝑵𝟐
𝒅𝒗= 𝟎 (3.13)
𝒅𝑻
𝒅𝒗= 𝒓 ×
−𝑯
∑ 𝑪𝒑×
𝒘
𝑮 (3.14)
Where:
𝒘 is mass of catalyst (𝑔) =Density of catalyst (𝑔 𝑐𝑚3)⁄ × volume of catalyst (𝑐𝑚3)
−𝑯 is the Heat of reaction at T(𝑐𝑎𝑙 𝑚𝑜𝑙⁄ )= 𝐻0 + ((𝐶𝑝𝒄𝒐𝟐+ 𝐶𝑝𝑯𝟐
) − (𝐶𝑝𝒄𝒐 + 𝐶𝑝𝑯𝟐𝒐)) × (𝑇 −
𝑇𝑟𝑒𝑓𝑓)
𝑻𝒓𝒆𝒇𝒇 is Reference Temperature (K) = 298.15
∑ 𝑪𝒑 = 𝒙co × 𝐶𝑝𝒄𝒐 + 𝒙𝑯𝟐𝒐 × 𝐶𝑝𝑯𝟐𝒐 + 𝒙𝒄𝒐𝟐
× 𝐶𝑝𝒄𝒐𝟐+ 𝒙𝑯𝟐
× 𝐶𝑝𝑯𝟐+ 𝒙𝑪𝑯𝟒
× 𝐶𝑝𝑪𝑯𝟒+ 𝒙𝑵𝟐
(3.15)
× 𝐶𝑝𝑵𝟐
𝐶𝑝𝒄𝒐 =∫ ((3.376 + (0.557 × 103) × 𝑇 − (0.031 × 10−5) × 𝑇−2) × 𝑅)
𝑇
𝑇𝑟𝑒𝑓𝑓
(𝑇 − 𝑇𝑟𝑒𝑓𝑓)
𝐶𝑝𝑯𝟐𝒐 =∫ ((3.470 + (1.450 × 103) × 𝑇 + (0.121 × 10−5) × 𝑇−2) × 𝑅)
𝑇
𝑇𝑟𝑒𝑓𝑓
(𝑇 − 𝑇𝑟𝑒𝑓𝑓)
𝐶𝑝𝒄𝒐𝟐=
∫ ((5.457 + (1.054 × 103) × 𝑇 − (1.157 × 10−5) × 𝑇−2) × 𝑅)𝑇
𝑇𝑟𝑒𝑓𝑓
(𝑇 − 𝑇𝑟𝑒𝑓𝑓)
𝐶𝑝𝑯𝟐=
∫ ((3.249 + (0.422 × 103) × 𝑇 + (0.083 × 10−5) × 𝑇−2) × 𝑅)𝑇
𝑇𝑟𝑒𝑓𝑓
(𝑇 − 𝑇𝑟𝑒𝑓𝑓)
46
𝐶𝑝𝑪𝑯𝟒=
∫ ((3.280 + (0.593 × 103) × 𝑇 + (0.040 × 10−5) × 𝑇−2) × 𝑅)𝑇
𝑇𝑟𝑒𝑓𝑓
(𝑇 − 𝑇𝑟𝑒𝑓𝑓)
𝐶𝑝𝑵𝟐=
∫ ((1.702 + (9.081 × 103) × 𝑇 − (2.164 × 10−6) × 𝑇−2) × 𝑅)𝑇
𝑇𝑟𝑒𝑓𝑓
(𝑇 − 𝑇𝑟𝑒𝑓𝑓)
feed mole fractions (the initial values):
𝒙𝒄𝒐= 0.011
𝒙𝑯𝟐𝒐 = 0.448
𝒙𝒄𝒐𝟐= 0.094
𝒙𝑯𝟐= 0.343
𝒙𝑪𝑯𝟒= 0.001
𝒙𝑵𝟐= 0.111
3.1.3 Catalyst Specification
For both type of reactor the same catalyst specification is used as follow:
Type : (𝐹𝑒2𝑂3-𝐶𝑟2𝑂).
Volume : 3× 104 (𝑐𝑚3).
Density : 5.09 (𝑔 𝑐𝑚3)⁄ .
3.1.4 Operating conditions
The same operating condition for the both type of reactor are used as follow:
Temperature (T) = 675(k).
Pressure (P) = 1 (atm) at which the reacting gaseous system can be considered ideal.
Volumetric flow rate: (G) = 2533176× 104(𝑐𝑚3 ℎ⁄ ).
Gas constant (R) = 1.987 (𝑐𝑎𝑙
𝑔.𝑚𝑜𝑙 𝐾).
47
Age of the catalyst( 𝝉) = 98 × 24 (ℎ).
Effectiveness factor (𝑬𝒇𝒇) = 1(Pseudo homogenous model).
3.2 Membrane Reactor
The same mathematical modeling equations , operating conditions , catalysts specification and
feed mole fractions that used for Packed Bed Reactor are used for Membrane Reactor the only
difference is in the hydrogen mass balance equation because only hydrogen permeates through the
membrane, the mass balance for membrane reactor can written as:
𝑶𝒖𝒕 𝑯𝟐− 𝑰𝒏𝑯𝟐
= 𝑮𝒆𝒏𝒓𝒂𝒕𝒊𝒐𝒏𝑯𝟐− 𝑷𝒆𝒓𝒎𝒂𝒕𝒊𝒐𝒏𝑯𝟐
𝒅𝒙𝑯𝟐
𝒅𝒗=
(𝒓×𝒘)
𝑮−
𝑨×𝑱𝑯×(𝑹×𝑻
𝑷)
𝑮 (3.15)
𝑱𝑯 is the hydrogen flux through membrane ( 𝑚𝑜𝑙
𝑐𝑚2.ℎ) .
𝑨 is the membrane surface area ( 𝑐𝑚2) .
𝑨 = 𝝅 × 𝑫 × 𝑳
Where 𝑫 = 25 𝑐𝑚 and 𝑳= 100 𝑐𝑚.
Pd or Pd-alloy membranes show very high hydrogen selectivities (Mendes et al.,2009) therefore
it has been used .The hydrogen flux through Pd and Pd-alloy membranes can be illustrated by the
following equations:
𝑱𝑯 =𝑷𝒆
𝜹 [(𝑷𝑹)𝒏 − (𝑷𝑷)𝒏]
𝑷𝒆 = 𝑷𝒆°𝐞𝐱𝐩 (
−𝑬𝒂
𝑹 × 𝑻)
where 𝑷𝒆is the permeability, 0.5 <n < 1, 𝑷𝒆° is the pre-exponential factor, 𝑬𝒂is the apparent
activation energy, 𝑹 is the ideal gasconstant,𝑻 is the absolute temperature, 𝑷𝑹 and 𝑷𝑷 are the
hydrogen partial pressure on the retentate and permeate sides, respectively, and 𝜹 is the membrane
thickness.
48
Chapter Four
4.0 Results and Discussions
4.1 Model validation:
Sing and Saraf (1980) used FORTRAN IV language to validate their experimental data. In this
chapter MATLAB software is used to validate their experimental data. This step is necessary for
the use of MATLAB to model the proposal membrane reactor.
Table 4.1: Experimental data and calculated results for packed bed reactor
Composition of the gas,
% (dry basis)
Inlet Outlet
(experimental)
Outlet
(FORTRANIV)
Outlet
(MATLAB)
𝐶𝑂 1.9 0.20 0.10 0.22
𝐻2 60.6 61.0 61.30 62.3
𝐶𝑂2 17.1 18.6 18.56 18.8
𝐶𝐻4 0.20 0.19 0.20 0.20
𝑁2 + 𝐴𝑟 20.2 20.0 19.84 20.1
Temperature °C) 193 193 192.64 193
From Table (4.1) it can be seen that the difference in calculated and measured compositions (in
absolute terms) and temperature at the reactor exit is insignificant for both programs. The
agreement between the measured and calculated compositions and temperature at the outlet does
indicate the validity of the model.
49
4.2 Comparing the performance of both Traditional and Membrane reactors :
Table 4.2: Comparison between Packed Bed Reactor and Membrane
Reactor
Packed Bed Reactor Membrane Reactor
Temperature (𝐾°) 675 675
Pressure (atm) 3 3
Feed Flow Rate (𝑐𝑚3
ℎ)
2533176× 104 2533176× 104
Catalyst Type 𝐹𝑒2𝑂3 − 𝑐𝑟2𝑂3 𝐹𝑒2𝑂3 − 𝑐𝑟2𝑂3
Catalyst Weight (g ) 9365× 103 9365× 103
Inlet ( CO) Mole Fraction (𝑥𝑖𝑛) 0.011 0.011
Out let ( CO) Mole Fraction (𝑥𝑜𝑢𝑡) 0.0064 0.0011
(CO) Conversion (𝑋𝐶𝑂)
𝑋𝐶𝑂 =𝑥𝑖𝑛 −𝑥𝑜𝑢𝑡
𝑥𝑖𝑛× 100%
29% 90%
From the table (4.2) it can be seen that the membrane reactor achieves considerable improvement
in conversion as compared with the traditional reactor in similar operating condition. It is worth
noting that the maximum conversion that can theoretically be obtained in the traditional reactor,
imposed by the equilibrium condition, can be overcome with the membrane reactor as a
consequence of the hydrogen withdrawal from the reaction system according to Lechatlier´s
principle.
50
Figure 4.1: Packed Bed Reactor Profile
Figure 4.2: Membrane Reactor Profile
51
4.3 Sensitivity Analysis:
In the following figures for different ranges of operating conditions sensitivity analysis was
developed for both type of reactor.
4.3.1 Tempture Effect :
In Figure (4.3) The following data is used for the both reactor configurations:
Reactor length = 61.15 ( cm) , reactor diameter =25 (cm), pressure = 3 (atm) , feed flow rate =
2533176× 104 (cm3
h). The temperature is varied from 400 K to 1000 K to find out the optimum
temperature for the maximum CO conversion .
Figure 4.3: CO Conversion(fractional) VS Temperature
52
It can be seen that the CO conversion for Packed Bed reactor increases asymptotically with
temperature. The conversion reaches a maximum conversion of 42% at a temperature of 675 k.
Thus the temperature of 675 K is the optimum temperature for the Packed Bed reactor. At this
temperature the conversion for Membrane reactor is 90%.
An increase in temperature beyond it optimum value for the Packed Bed reactor , result in a
decrease in conversion due to the reversibility of the chemical reaction .Therefore equilibrium
limitation is the limiting factor. The conversion approaches zero for Packed Bed reactor at
temperature of 740 K. At this temperature the conversion for the Membrane reactor is 92%.,which
the maximum conversion for this reactor , thus the temperature of 740 K is the optimum
temperature for the reactor operation.
4.3.2 Pressure Effect :
In Figure (4.4) The following data is used for the both reactor configurations:
Reactor length = 61.15 ( cm) , reactor diameter =25 (cm), temperature = 675 (k), feed flow rate =
2533176× 104 (cm3
h). The pressure is varied from 0.1 (atm) to 20 (atm) to find out the optimum
pressure for the maximum CO conversion .
53
Figure 4.4: CO Conversion (fractional) VS Pressure
It can be seen that for the Packed Bed reactor the pressure has a positive effect in CO conversion
up to equilibrium. At equilibrium the conversion reaches a maximum of 42% at a pressure of 3
atm. After this point no change occurs in the Co conversion with pressure. This is because the
total number of moles of react in the gas phase at any axial position in a fixed bed reactor is not
determined by the relative position of the forward and reverse reactions. At equilibrium the CO
conversion is only function of temperature and feed composition.
For Membrane it can be seen that the CO conversion at a pressure of 4 atm increased to 94%
from the 42% of the Packed Bed reactor. At a pressure of 10 atm the CO conversion reaches a
maximum of 98% which is the optimum pressure for the reactor operation.
54
4.3.3 Reactor length effect :
In Figure (4.5) The following data is used for the both reactor configurations:
Reactor diameter =25 (cm) ,pressure = 3 (atm) ,feed flow rate = 2533176× 104 (cm3
h) , temperature
= 675 (k) . The reactor length is varied from 10 (cm) to 150 (cm) to find out the optimum
temperature for the maximum CO conversion .
Figure 4.5 CO Conversion (fractional) VS Reactor Length
It can be seen for the Packed Bed reactor the optimum length is 20 cm at which the CO conversion
reaches the maximum of 43% . There is economically an unjustifiable no increase in conversion
beyond this point.
For the Membrane reactor length has positive effect in CO conversion .is positive .At the length
of 115 cm ,the CO conversion reaches 97% which the optimum length for reactor opreation after
it the Co conversion has not changed.
55
4.3.4 Transport coefficient effect :
In Figure (4.6) The following data is used for the both reactor configurations:
Reactor diameter =25 (cm), pressure = 3 (atm), feed flow rate = 2533176× 104 (cm3
h) ,
temperature = 675 (k) and reactor length = 61.15 (cm) the Transport coefficient (kc) is varied
from 0.1 (1/h) to 19 (1/h) to find out the optimum( kc) value for the maximum CO conversion
.
Figure 4.6: CO Conversion (fractional) VS Transport coefficient
It can be seen that the optimum value of ( kc) is 10, at which the maximum conversion of CO is
95% and after this value of( kc ), the CO conversion does not increase appreciably with ( kc)
increment.
56
Chapter Five
5. Conclusions and Recommendations
5.1 Conclusions
It was found that the membrane reactor has better performance than the packed-bed reactor,
allowing in certain conditions to overcome the thermodynamic equilibrium limitations in
traditional reactors. The conversion of carbon monoxide reaches a maximum of 29% for
the packed-bed reactor and 90% for the membrane reactor, which represents an
improvement of 210% .
The optimum parameters were assessed for both types of the reactors by developing
sensitivity analysis for different ranges of operating conditions. It was found that the
optimum operating conditions for membrane reactor were: temperature of 740 K, pressure
of 10 atm and a reactor length and diameter of 115 cm and 25 cm respectively.
5.2 Recommendations
It is recommended that :
A model validation should be carried out experimentally to confirm the reported findings.
A complete economic analysis should be carried out to assess the commercial viability of
the usage of membrane reactor for the gaseous reversible reactions.
57
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59
Appendix A:Packed Bed Reactor MATLB code
60
61
Figure 4.1: Packed Bed Reactor Profile
62
Appendix B: Membrane Reactor MATLB code
63
64
Figure 4.2: Membrane Reactor Profile
65
Appendix C: (Carbon monoxide conversion(fractional) VS Temperature ) MATLB code
Figure 4.3:Carbon monoxide conversion (fractional) VS Temperature
66
Appendix D:( Carbon monoxide conversion (fractional) VS Pressure) MATLAB code
Figure 4.4: Carbon monoxide conversion (fractional) VS Pressure
67
Appendix E:(Carbon monoxide Conversion (fractional) VS Reactor Length) MATLAB code
68
Appendix F: (CO Conversion (fractional) VS Reactor Length ) MATLAB code
Figure 4.5 Carbon monoxide Conversion (fractional) VS Reactor Length