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CPD NR 3258Conceptual Process Design

Process Systems Engineering

DelftChemTech - Faculty of Applied SciencesDelft University of Technology

Subject

Dehydration of ethanol

Authors Telephone

M. de Jong

P.F.A. van Rooijen

V. VerboomT. Winkels

0180-419065

015-3108063

015-21328400174-295760

Keywords

Ethanol, dehydration, water, separation, distillation,

membranes

Assignment issued : 09-03-2001Report issued : 12-06-2001Appraisal : 26-06-2001

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Summary

Ethanol is one of the main base chemicals in the world and is often used as a fueladditive to produce gasohol. Before the ethanol can be used as additive it needs to bedehydrated and purified up to 99.8 vol%. This report provides a comparison between

four conceptual designs of ethanol dehydration units, which process a feedstream of88.8 vol% ethanol to the required purity of at least 99.8 vol%. During the processimpurities like ethyl acetate, acetaldehyde, isobutyl alcohol and isopentyl alcohol are

partly removed. The alternatives, azeotropic distillation by toluene, extractivedistillation by gasoline, extractive distillation by polyacrylic acid (PAA) and normaldistillation followed by membrane purification, are chosen out of a wide range ofoptions found in literature. After comparison of the four alternatives on the basis ofvalidity of thermodynamic data, safety, environmental impact and economy, thenormal distillation followed by membrane purification appears to be the mostinteresting option for further design.

The ethanol dehydration plant is intended to use a specific feed stream originatingfrom a nearby ethanol fermentation plant. Therefore the production capacity of thedehydration plant is fixed to 8.6 kton (10.9 kliters) of ethanol per year. This is incomparison with the production level in Europe of 4.7 billion liters per year, a verysmall amount. Therefore the sales price for ethanol produced in the dehydration plantwill be dependent on the world market price. The dehydration plant is a continuousoperated plant, which is 8,620 hours per year on stream (stream factor 0.984). Thisimplies that 150 hours per year are accounted for unexpected shut-downs. The plantlife is 15 years, which includes 2 years of construction and 1 year of deconstruction.

The options azeotropic distillation by toluene, extractive distillation by gasoline and

normal distillation followed by membrane purification are already operated at fullscale and are patented. The extractive distillation by PAA is based on laboratoryexperiments. During the design some doubts arose for the thermodynamic propertiesof this last option and therefore the extractive distillation by PAA is rejected in theselection procedure for the final design.

The total investment costs are the highest for the azeotropic distillation by toluenewith 1.8 million EURO. The costs of extractive distillation by PAA and normaldistillation followed by membrane purification are modest with respectively 1.2million EURO and 0.9 million EURO. The option extractive distillation by gasolinehas the lowest investment of 0.7 million EURO. Calculation of the economic criteriashows that the azeotropic distillation by toluene is not profitable. This option has anegative cash flow of 0.9 million EURO per year and is therefore rejected as a finaldesign option. In reviewing the variable costs it appeared that at least 58 % of theoperating costs are caused by the purchase of the raw materials. Therefore thesensitivity of the designs for changes of 5 % in the purchase costs of the raw materialswere investigated. It appeared that the economic criteria of the extractive distillation

by gasoline are extremely sensitive. This causes a motivation to reject this option asfinal design.

This implies that on basis of several criteria mentioned above, the normal distillationfollowed by membrane purification appeared to be the most robust design. Thisdesign has a Net Cash Flow of 272 kEURO, a Pay-Out Time of 3.4 years and a Rateon Return of 29.4% before tax. The Discounted Cash Flow Rate on Return before taxamounts to 24.6%.

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Table of contents

1 Introduction 1

1.1 Motive 1

1.2 Purpose of the report 2

1.3 Description of report structure 2

2 Process options & selection 5

2.1 Process options 5

2.2 Criteria and selection 6

3 Basis of design 10

3.1 Description of the design 10

3.2 Process definition 103.3 Basic assumptions 12

3.4 Economic Margin 17

4 Thermodynamic properties 20

4.1 Operating window 20

4.2 Non-ideal equations for distillation 20

4.3 Choice of thermodynamic model 21

5 Process structure & description 30

5.1 Design criteria 30

5.2 Unit operations 30

5.3 Process chemicals 34

5.4 Utilities 35

5.5 Final process conditions 36

5.6 Process Flow Schemes 37

5.7 Process performance 39

6 Process control 42

6.1 General considerations 42

6.2 Control of a distillation column 42

6.3 Control of heaters and coolers 44

6.4 Control of a decanter 44

6.5 Control of the ultrafiltration unit in the extractive distillation by PAA 45

6.6 Control of the membrane unit in the normal distillation followed by membrane purification 45

7 Mass and heat balances 46

7.1 Mass and heat balances of the azeotropic distillation by toluene 46

7.2 Mass and heat balances of the extractive distillation by gasoline 47

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7.3 Mass and heat balances of the extractive distillation by PAA 48

7.4 Mass and heat balances of the normal distillation followed by membrane purification 48

8 Process and equipment design 50

8.1 Integration by process simulation 50

8.2 Equipment selection and design 51

8.3 Special issues 60

8.4 Equipment data sheets 61

9 Wastes 62

9.1 Identification of wastes 62

9.2 Biological treatment of waste water 63

9.3 Influence of process on wastes 64

10 Process safety 67

10.1 The Dow Fire and Explosion Index 67

10.2 Hazard and Operability Studies 69

11 Economy 71

11.1 Capital investment 71

11.2 Operating costs 73

11.3 Income 75

11.4 Cash flow 7611.5 Economic criteria 77

11.6 Cost review 79

11.7 Sensitivities 79

11.8 Negative cash flows 81

12 Comparison and conclusions 82

12.1 Data validity 82

12.2 Purity and recovery 82

12.3 Process yields 83

12.4 Wastes 83

12.5 Process Safety 84

12.6 Economy 84

12.7 Selection for recommendation for further design 86

13 Recommendations 87

13.1 General recommendations 87

13.2 Azeotropic distillation by toluene 87

13.3 Extractive distillation by gasoline 88

13.4 Extractive distillation by PAA 88

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13.5 Normal distillation followed by membrane purification 88

Literature 90

Text symbols 92

Appendices:

Appendix 1 Pure component properties, toxicological and heat data, structure ofmain components

Appendix 2 Block schemes of the designsAppendix 3 AssignmentAppendix 4 Comparison of the different processesAppendix 5 VLE data of regression for PAA design optionAppendix 6 Utility costsAppendix 7 Utility summariesAppendix 8 Process flow schemesAppendix 9 Description of the Aspen Plus 10 filesAppendix 10 Process stream summariesAppendix 11 Process yieldsAppendix 12 Heat and mass balancesAppendix 13 Example calculationsAppendix 14 Calculations of the columns, reboilers, condensers, coolers and heatersAppendix 15 Equipment summary & specification sheets of the azeotropic

distillation by tolueneAppendix 16 Equipment summary & specification sheets of the extractive distillationby gasoline

Appendix 17 Equipment summary & specification sheets of the extractive distillationby PAA

Appendix 18 Equipment summary & specification sheets of the normal distillationfollowed by membrane purification

Appendix 19 Determination of the bundle diameter of the reboilerAppendix 20 Dows Fire and Explosion IndexAppendix 21 Hazard and Operability studiesAppendix 22 Investment and production costs

Appendix 23 Column and tray layout

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1 Introduction

1.1 Motive

Ethanol is one of the basic components of the chemical industry. Ethanol is a clear,colourless, flammable, oxygenated hydrocarbon with the chemical formula C2H5OH.It is miscible in all proportions with water and also with ether, acetone, benzene, andsome other organic solvents. The binary mixture ethanol / water contains an

azeotrope. The azeotropic vapour-liquid equilibrium mixture, which occurs at 78.1Cand 1 bar, contains 95.57 w% ethanol and 4.43 w% water (ref. 37). The chemical

properties of ethanol are dominated by the functional - OH group, which can undergomany industrially important chemical reactions, like, dehydration, halogenation,estrification and oxidation.

For the production of ethanol various feedstocks and hence methods are used.Fermentation alcohol can be produced from grain, molasses, fruit, wine, whey,cellulose and numerous of other organic sources. More than 90 % of the world

production of ethanol is based on biological feedstocks. Synthetic alcohol may beproduced from crude oil, gas or coal, but plays a minor role in the world ethanolproduction with a share of only 7 %.

The commonly known application of ethanol is within the field of alcoholicbeverages, but in the industry ethanol is also suitable for many other applications, likeindustrial solvent and antiseptic. Ethanol is also used as raw material for the

preparation of many industrial organic chemicals, like acetaldehyde, butadiene,

diethyl ether, ethyl acetate, ethyl amines, ethylene, glycol ethers and vinegar.

Since 1970 a new application for ethanol has been found as fuel additive. In 1970 itwas realised that the petroleum stocks are limited and a search for alternative fuelsources began. One of the alternatives can be found in gasohol, which is a fuel source

based on the use of ethanol obtained from natural sources. To produce gasohol anextender is made from a mixture of gasoline (90 w%) and ethanol (10 w%). Gasoholhas a higher octane number and burns more slowly, coolly and completely thangasoline, resulting in reduced emission of some pollutants. On the other hand ethanol-

based gasohol is often expensive and energy intensive to produce. Nowadays the useof ethanol as an additive to petrol is an important application, for example in Brazil.

Today, fuel ethanol accounts for roughly two thirds of the world ethanol production(ref. 5).

This report covers the design of four ethanol dehydration processes, which produceethanol to be used as fuel additive. The (imaginary) contractor is an ethanolfermentation plant in the republic of Lithuania. The fermented ethanol, which has aconcentration of 12 vol%, is not suitable for use as fuel additive until it is dehydratedand purified till 99.8 vol%. To concentrate this polluted ethanol stream thefermentation plant uses a distillation column to reach 88.8 vol% ethanol. The

pollutants in this stream are water, acetaldehyde, ethyl acetate, isobutyl alcohol andisopentyl alcohol. This stream is the feed stream for the ethanol dehydration plant

discussed in this report. The pure component properties of all relevant substances arelisted in Appendix 1.

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clear. The thermodynamic properties and the chosen thermodynamic model aredefined in Chapter 4. Subsequently in Chapter 5 the process structures of the chosenalternatives are extendedly treated. The equipment choice is explained and the streamsand utilities defined. The process control is described in Chapter 6, and the mass- andheat balances are defined in Chapter 7. The design of the process equipment is made

explicit in Chapter 8. In this chapter the exact sizes of the equipment are describedand calculated. In Chapter 9 the wastes of the four process alternatives are mentioned,subsequently in Chapter 10 the safety of the four options is investigated by a Fire &Explosion Index and a Hazard and Operability study. In Chapter 11 all previoussubjects are used to obtain the economy. In Chapter 12 the conclusions are drawn.Finally in Chapter 13 the recommendations are made.

In the frame below the world market situation for ethanol is dealt with. Also theimpact of the designed plant on this market is forecasted. Finally the patent situationof several parts of the four design alternatives is mentioned.

Market situation for products and competitors1The total world ethanol production in 1998 was approximately 31.2 billion litres. This was alittle downturn compared with previous years, due to a decrease of the Brazilian productionand the Asian financial crisis. The European ethanol production is approximate 15 % of theworld production and accounts 4.7 billion litres. Within Europe and especially within theEuropean Union there is a stimulation program to increase the use of ethanol as fuel additiveup from 5.6 % in 1997 to 12 % in 2010. Because the production of bio-fuels is moreexpensive than conventional fuels, the manufacture of them will have to be subsidised. Franceand Germany are the biggest producers of fuel ethanol in the European Union, withrespectively 500 million litres and 390 million litres. However, in Germany the ethanol

production is dominated by synthetic manufacturing, mainly by Hls (177 million litres peryear) and Erdlchemie (75 million litres per year). The third largest ethanol producer inEurope is the United Kingdom with a total production capacity of around 430 million litres

per year of which BP-Amoco alone accounts for 417 million litres divided over two sites.This will change at the end of 2001, when BP-Amoco increases its production to 462 millionlitres at the site of Grangemouth.

A relatively new fuel-ethanol project in the EU, which has come on stream around the turn ofthe century, is Nedalcos plant in Bergen op Zoom. This plant has a production capacity of 30million litres per year. Other European ethanol production plants are Agroetanols facility inSweden (50 million litres per year) and the 100 million litres per year distillery in Cartagena,Spain to be operated by Biocarburantes Espanoles.

In Eastern Europe the production of ethanol is dominated by the manufactures in the RussianFederation. The total capacity in Russia can be estimated at 2.5 billion litres, with beveragealcohol accounting for 60 %. The enormous capacity is hardly surprising, given the fact thatRussians drink almost 2.2 billion litres of pure ethanol per year. However due to thedissolution of the former Soviet Union, restructuring and a huge illicit production, the Russiangovernment introduced a monopoly on the production of alcohol in 1997. After removal ofthe state support the competitiveness of the Russian producers decreased and the import ofalcohol into the country has risen tremendously. Nowadays the illicit production of ethanol isstill increasing in the chaotic industry and trade policy of the Russian government.The total EU and US ethanol exports to Eastern Europe in 1997 amounted 360 million litres,of which over 270 million litres came from the US. The largest countries of destination in1997 were Georgia (188 million), Ukraine (24 million) and Latvia (20 million).

1This frame is based on ref 5.

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It is believed that the political nature of fuel-ethanol production makes it unlikely that therewill be a consistently large international trade in this product. Fuel-ethanol programs have

been put in place to create additional demand for the purchase of feed stocks from domesticfarmers. It would run contrary to this intention if large-scale imports were allowed, as theywould support foreign farmers. A bio-fuel program usually incurs large costs, and it would

become completely unjustifiable if that money was spent on imports. As a result, world tradewill generally be limited to industrial and potable alcohol.

Review of joining the market

The designed ethanol dehydration plant will be connected with the fermentation plant and istherefore fixed in size. In case of joining the market with the proposed ethanol productivity of8.6 million litres per year, little impact will be exposed on the world market. The new plantwill only account for 0.2 of the world market and therefore has no influence at all at themarket price. The profitability of the new plant will be dependent on the current market price.According to ref. 9 the price of ethanol is predicted to decrease over the next few years

because of better technologies and increasing capacity even though the price of the feedstockswill increase. Additionally in many countries one or two companies control the production of

ethanol. These companies could provide rivals with a competitive edge. Therefore it could bedifficult to join the market with a relatively small plant.

Patent Situation

Patented processes, which will be redesigned for a specified feed stream, will increase thetotal cost. Therefore the current patent situation is an important issue. Azeotropic distillation

by several hydrocarbon entrainers, like benzene, cyclohexane and toluene, is industriallyapplied since 1903. As a result the azeotropic distillation by toluene is widely patented on unitoperation, process unit sequence and thermal integration. The use of gasoline in extractivedistillation to dehydrate aqueous ethanol was granted with a United States patent in 1952(U.S. patent 2,591,672). Integrated distillation / membrane pervaporation plants arecommercially operated since the eighties. Both the design of hybrid systems and the use of

hydrophilic zeolite membranes are extensively patented worldwide. The use of polymericentrainers to break the azeotrope of ethanol / water systems is still in the experimental phase.It is a relatively new area of investigation. Because of the experimental character of this

process option several possibilities still exist to patent the commercial operation of extractivedistillation by polyacrylic acid. So the use of PAA as an entrainer at plant scale is a feasible

process option concerning the patent situation. It is the only process option of which nosimilar plants seems to exist yet.

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2 Process options & selection

Designing requires a well-considered choice of the type of process, which can only bemade after a thorough investigation of the goal of the plant and the possiblealternatives to achieve this goal. Therefore the requirements and the process options

are described in the next paragraph.

2.1 Process options

As mentioned in Chapter 1 the ethanol / water feed stream is bought from a nearbyethanol fermentation plant. The upstream section of this plant imposes thespecifications on the feed stream of the plant to be designed. Within the fermentation

plant the ethanol is upgraded from 12 vol% to 88.8 vol%. This ethanol needs to befurther purified in the designed dehydration unit up to 99.8 vol% for the use as fueladditive.

To reach this requirement of purity, the azeotrope of the ethanol / water mixture has tobe broken. The alternatives to achieve the ethanol / water separation found inliterature are listed shortly in Table 2.1. The alternatives have been divided in fourcategories, namely methods to reach the azeotrope, to break the azeotrope, methodsthat can directly achieve the required purity and methods that achieve the required

purity by different techniques.

Table 2.1: List of separation alternatives for ethanol / water mixtures.

Type ofseparation Ethanolw% Process Estimate energyconsumptionkJ/dm3ethanol

Ref.

To azeotrope 89 96 Conventional distillation 2,600 26

To azeotrope 89 96 Multi-effect vacuum 2,000 26

To azeotrope 89 96 Vapour recompression 1,800 26

Azeotropic 96 100 Adsorptive dehydration by molecularsieves 1,500 26, 1

Azeotropic 96 100 Adsorptive dehydration by solid agents 500 26, 24

Azeotropic 96 100 Adsorptive dehydration by zeolites 1,500 32, 19

Azeotropic 96 100 Azeotropic distillation by entrainers 2,600 26

Azeotropic 96 100 Extraction by gasoline, hydrocarbons 2,200 8

Azeotropic 96 100 Extraction by non-volatile components 2,000 37, 21,18, 14

Azeotropic 96 100 Low pressure distillation (< 11,5 kPa) 3,000 6

Azeotropic 96 100 Membrane Technology by pervaporation 1,000 12, 17, 32

Azeotropic 96 100 Membrane Technology by vapourpermeation 1,000 30

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Table 2.1 (continued)

Complete 89 100 Adsorptive dehydration by solid agents 700 24

Complete 89 100 Convential dual distillation withentrainer 5,000 26, 21, 14

Complete 89 100 Extraction with supercritical carbon

dioxide 2,500 26, 37, 15Complete 89 100 Solvent extraction 2,500 37

Complete 89 100 Vacuum distillation 10,000 26

Other 12-100 Membrane Technology 26, 37, 25

Each of these alternatives has advantages and disadvantages, so several criteria arechosen to decide which alternatives will be designed in detail. Some alternatives can

be combined to use the advantages of these technologies. For example a hybridsystem of a distillation column followed by membrane purification. Another veryinteresting option to purify the ethanol is the use of a hydrophobic zeolite membrane

(ref. 25). Such a membrane will let ethanol through as the permeate stream and willlet the water flow by. The inlet stream of the membrane is the aqueous ethanol streamfrom the ethanol fermentation at 12 vol%. In this case no distillation columns arenecessary. This option is not chosen for the time being because it falls outside thechosen specifications and battery limits (see Chapter 3, Basis of Design).

2.2 Criteria and selection

The criteria to make a selection between the process options are listed below.

- Modern techniquesMost of the conventional processes are invented in the second half of the 20thcentury. Therefore the conventional processes are already designed in detail andoptimised. The first literature on azeotropic distillation with benzene as entrainercan be found in 1902 and is first patented in 1903. Almost a century of researchand optimisation has resulted in an enormous amount of capable entrainers anddehydration methods, which are patented all. Therefore a challenge can be foundin using alternative techniques, which are only known for the last decades (seewindow Chapter 1) and have potential to improve the ethanol dehydration, forexample in the field of economy, energy use or use of hazardous materials. Themost promising alternatives found in literature will be chosen.

- Hazardous materialsIn modern chemistry it is impossible to design and build a plant when theenvironmental consequences are out of proportion. Within the process a minimumof hazardous materials should be used. Whenever possible a hazardous materialshould be avoided or replaced by a less hazardous material.

- Economical profitabilityThe variable costs are mainly dependent on the energy consumption during theoperation of the plant, although some alternatives have relatively high fixed costs.For the first rough distinguish in costs only the energy costs will take into account.

Estimates of the energy consumption can be found in Table 2.1. The energyconsumption of the dehydration methods are based on literature, but there is great

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variety in methods used, extent of purification and year of design. The mentionednumber can only be seen as a rough estimate, because the proper costs of thisconsumption will largely depend on the extent of optimisation and integration, thecosts of the equipment and entrainers used.

- Number of optionsThe time limit of the CPD-project is set in advance to twelve weeks (480 hours

per person). Within this period only a certain amount of work can be done.Therefore it is inevitable to set a maximum number of options to be worked out.In our opinion the maximum feasible number of options to be designed is four,within the imposed time limit. The number of free options is brought back tothree in practice by the requirement of a base reference case to compare thedesigns mutually.

- Data availability and data processingTo facilitate the process of designing a certain amount of data has to be available.

Literature and experts can for example provide this data. A second importantrequirement is the possibility to apply the accumulated data in the process-simulating program Aspen Plus 10.

- Personal interestsIn the assignment a small number of alternatives is listed, which are preferred bythe (imaginary) contractor (Appendix 3). Also the interests of the companyControlec is an important factor in the choice of alternatives. The students owninterests are also taken into account.

In Appendix 4 a comparison is made between the process options given in Table 2.1,based on the criteria mentioned above. The chosen process options are:1. Azeotropic distillation by toluene (standard case)2. Extractive distillation by gasoline3. Extractive distillation by a polymeric entrainer (polyacrylic acid)4. Normal distillation followed by zeolite membrane purification

Each process is operated continuously, because this is the most economic and easiestway of operation and all equipment used is capable of being operated continuously.

2.2.1

Azeotropic distillation by toluene

In azeotropic distillation a third component, in this case toluene, is added to the feed.This component is called the entrainer or mass separating agent. This componentchanges the vapour-liquid equilibrium of the ethanol-water mixture, by forming a newazeotrope. This ternary azeotrope enables the recovery of pure components by usingthree columns. The top stream of the first column is the ternary azeotrope, while the

bottom product contains a high concentrated ethanol mixture. This concentratedmixture is separated from its impurities in the second column, where the 99.8 vol%

pure ethanol is recovered at the top and the impurities like ethyl acetate, andacetaldehyde are the bottom product. The ternary azeotropic mixture of ethanol /

water / toluene and the redundant water are led to a decanter where the water phase isseparated. The azeotropic mixture is led to a third column where the toluene phase is

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separated from its impurities and recycled to the first column. The block scheme isrepresented in Appendix 2.

Benzene is the most common entrainer used in azeotropic distillation and is alreadyknown for a long time. This entrainer however is not chosen because of its

carcinogenic properties. Other candidate third components are listed and treated inChapter 5. The well-known process with toluene as the entrainer will be used as areference.

2.2.2 Extractive distillation by gasolineAn extractive distillation is performed to separate ethanol-water mixtures by adding athird component. This entrainer changes the vapour-liquid equilibrium of the originalmixture. When such a component is chosen, that breaks the azeotrope, it becomes

possible to dehydrate the ethanol mixture relatively easy. Extractive distillation can be

accomplished by using gasoline as an entrainer. Because the produced ethanol will beused in gasohol, an inventive integration can be made between the ethanoldehydration and a gasoline refinery. By using gasoline as entrainer, the ethanol-waterazeotrope will be broken. The ethanol-water separation can be fulfilled like anordinary extractive process, but large savings can be made. The gasoline entrainer is

part of the product, so it is not necessary to recover and recycle it to the distillationprocess. The column will be operated in such way that the gasoline entrainer togetherwith the ethanol is removed as a bottom product, while the aqueous stream will berecovered over the top. There is no need to separate the gasoline-ethanol mixture,

because gasohol is a mixture between gasoline and ethanol. This reduces theinvestments costs and the variable production costs (energy costs) considerable. Ageneral process is presented in Appendix 2.

2.2.3 Extractive distillation by the polymeric entrainer polyacrylic acid (PAA)Another possibility to accomplish an extractive distillation process is the use of the

polymer entrainer PAA. This entrainer changes the vapour liquid equilibrium bybreaking the ethanol / water azeotrope. By adding PAA in the distillation column thepure ethanol can be obtained overhead. The water-entrainer mixture is fed to anultrafiltration unit, where the water is separated from the entrainer. The entrainer,

dissolved in a small amount of water, is recycled to the distillation column. Thegeneral process is represented in Appendix 2.

2.2.4 Normal distillation followed by membrane purificationMembranes can separate water from water-alcohol mixtures in a much moreeconomical way than by conventional distillation. These membranes can overcomethe azeotropic barrier and so can obtain the required purity of ethanol. Membranes areespecially useful to separate mixtures near the azeotropic composition. Theadvantages of membrane technology are the low operation costs and the breaking of

the azeotrope without the aid of a solvent. A disadvantage is the low flux through themembrane, which illustrates the need for multiple membrane-sections in series. A

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rather new technology is the technology of zeolite membranes. These membranes canselectively remove one component. A big advantage of zeolite membranes is that they

have a larger flux (4 kg/(m2.h)) through the membrane than polymeric membranes

(0.2 kg/(m2.h)). Therefore less units will be needed when zeolite membranes areused. Because of these advantages zeolite membranes will be used in the design,

which is shown in Appendix 2. A normal distillation column will purify the feedstream to or near the azeotrope. The azeotrope comes overhead and water flows overthe bottom. Subsequently the top stream will be purified using the zeolite membranes.The membrane sections separate the water by pervaporation. Another option is thevapour permeation technology.

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3 Basis of design

3.1 Description of the design

Pure ethanol has a wide range of applications in both the chemical and the consumerindustry. In the last decades ethanol is discovered as a valuable component in fuel.Since then ethanol mixed with gasoline to produce gasohol. Before ethanol can beused for most purposes the raw, aqueous ethanol needs to be purified from water. Thisimposes a difficulty because the mixture of ethanol and water contains an azeotrope.This binary azeotrope of the ethanol / water system is situated at 95.57 w% ethanol

and 4.43 w% water at 78.1C and 1 bar. This means that purification by normaldistillation cannot recover pure ethanol. To acquire the pure ethanol the azeotrope inthe ethanol / water system has to be broken. Various dehydration processes canachieve the desired separation (see Chapter 2).

For the production of dehydrated ethanol, which is used as an additive to gasoline, apurity of 99.8 vol% is required. In the upsteam production process ethanol is producedin a stream of about 12 vol% aqueous ethanol. To concentrate this stream a distillationcolumn is used to reach 88.8 vol% ethanol. The ethanol stream leaving this distillationcolumn is the feed stream of the design. In the ethanol feed stream there are someimpurities as acetaldehyde, ethyl acetate, isobutyl alcohol and isopentyl alcohol

present in relatively small amounts, but there are no requirements to keep thisimpurities out of the ethanol stream (see Appendix 3).

3.2 Process definition

A well-considered choice in the type of process can only be made after a thoroughinvestigation of the goal of the plant and the possible alternatives to achieve this goal.

3.2.1 Process concepts chosenTo reach the requirements of purity, the azeotrope of the ethanol / water mixture hasto be broken. There are many alternatives available to achieve the desired ethanol /water separation (see Chapter 2, Table 2.1). All these alternatives have advantages

and disadvantages. To make a choice between the suitable process options a selectionis made between them on the basis of several criteria. The following criteria are usedto decide which alternatives will be designed in detail (see Appendix 4):- Modern techniques- Hazardous materials- Economical profitability- Number of options- Data availability- Personal interests

After a comparison between the process options, based on the criteria mentioned

above. The chosen process options are:

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- Azeotropic distillation by toluene- Extractive distillation by gasoline- Extractive distillation by polyacrylic acid (PAA)- Normal distillation followed by membrane purification

Each process is operated in a continuous mode.The option extractive distillation by gasoline is directly producing gasohol in contrastto the other three options that produce pure ethanol. Because the specification of 99.8vol% pure ethanol does not hold anymore, the following specifications are imposed.The gasohol to be produced must contain 10 w% ethanol. This 10 w% of ethanol ingasohol has to be 99.8 w% pure. This imposes a maximum allowable amount of waterin gasohol of 0.2 w% of the ethanol present in gasohol.

3.2.2 Block schemesThe block schemes for the four process options are provided in Appendix 2. Only thesignificant pieces of equipment are represented in this block scheme. In the blockschemes the total mass streams (ton/annum) and yields (ton/ton product) are given. Inthe separate blocks the process conditions are displayed.

3.2.3 Thermodynamic propertiesTo calculate the exact thermodynamic properties the liquid activity coefficients andvapour fugacities are necessary (See Chapter 4). For an estimation of these parametersseveral methods are in use. To determine the right thermodynamic method, thefeasible models are investigated in Aspen Plus 10. These models are the Wilson, the

NRTL, the UNIQUAC and the UNIFAC model. All these models are applicable to(highly) non-ideal mixtures, Vapour-Liquid Equilibria (VLE) and except for theWilson model they are also applicable to two liquid phases.

For the separation of ethanol and water the models NRTL and UNIQUAC are mostlikely to be used. When the separation is accomplished with the aid of a massseparating agent, for example toluene, the system changes from two phases to three

phases. Also in this case the NRTL and the UNIQUAC models are valid.

To determine the influence of the model both the NRTL and the UNIQUAC methodare evaluated by creatingx,y-diagrams and residue curves for the binary systems andthe ternary system of ethanol, water and toluene (see Chapter 4, Figure 4.1 and 4.2).The differences between the available models mutually and with literature are small.

Nevertheless a choice has to be made between the UNIQUAC and the NRTL model.Both models are appropriate, but here the choice is made for the UNIQUAC model,

because of its wide acceptance in literature and its accuracy in representing VLE datafor a wide range of systems.

Because the thermodynamic model has to be useable for the systems with an entraineras well, the validity of the UNIQUAC model for this systems is investigated (See

Chapter 4). It appears that the UNIQUAC model can also be used as thermodynamic

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Table 3.2: Available utilities at the plant site.

Utility Price

Medium pressure steam at 10 bar 12.5 (Euro/ton)

Electric power 50 (Euro/MWh)

Cooling water 450 (Euro/kton)

Nitrogen 1.35 (Euro/ton)

Also other necessary material, including chemicals, are assumed to be available atmarket prices. The infrastructure is fully developed and electricity, air and seweragefacilities can be easily constructed.

3.3.3 Battery limitsThe ethanol is produced by fermentation in the nearby production plant. During theethanol production some side reactions take place, so small amounts of aldehydes,

higher alcohols and esters are formed. These components will be convertedrespectively to acetaldehyde, ethyl acetate, isobutyl alcohol and isopentyl alcohol. Atthe nearby production site a predistillation already takes place to concentrate theethanol to a value of 88.8 vol%. The preconcentrated ethanol is stored in a tank at 30

C and 1 bar. From this storage tank the feed stream will cross the battery limit. Insidethe battery limits the ethanol is dehydrated by the chosen methods. In each design thedehydration is mainly performed by distillation columns (see Appendix 2). Out of the

battery limit flows respectively a dehydrated ethanol stream or gasohol stream.Furthermore a waste water stream comes out the dehydration plant crossing the

battery limits. The treatment of the waste water will not be included in this design, butthe cost of the treatment on the other hand must be taken into account. In the case of

the azeotropic distillation by toluene (30 C, 1 bar) and extractive distillations bygasoline (20 C, 1 bar) and PAA a stream of entrainer flows into the battery limit froma storage tank. An amount of the entrainer flows out of the system as impurities in theethanol or waste water stream.

3.3.4 Definition in- and outgoing streamsVarious streams enter the battery limits of the four options. In each design an ethanolfeed stream is present and enters the battery limit. Besides this feed stream anentrainer is added in the azeotropic distillation by toluene and in the extractivedistillation by gasoline, respectively a small quantity of toluene and a large quantity ofgasoline. These streams also pass the battery limits. The specifications of the ethanolfeed stream are listed in Table 3.3.In the design all components of the ethanol feedstream have the maximal allowable concentration: the worst case design.

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Table 3.3: Conditions of the ethanol feed stream for the four alternatives.

Stream Name : Ethanol feed

Component Units Specification Additional Information

Available Design Notes

Ethanol vol% 88.8 88.8 (1)

Acetaldehyde mg/m3 300 300 (1)Isobutyl alcohol mg/m3 1250 1250 (1)

Isopentyl alcohol mg/m3 3750 3750 (1)

Ethyl acetate mg/m3 500 500 (1) (1) As worst case scenario.

Process Conditions and Price

Temperature C 30

Pressure Bar 1

Phase V/L/S L

Price ethanol EUR/ton 250

A representative composition for the gasoline entrainer is listed in Table 3.4.

Table 3.4: Composition of the gasoline feed stream.

Stream Name : Gasoline feed

Component Units Composition

Toluene w% 10

1-Hexene w% 6

2-Methyl-2-butene w% 4

Methyl cyclopentane w% 13

Methyl cyclohexane w% 7

n-Pentane w% 15

2-Methylbutane w% 8n-Hexane w% 12

2-Methylpentane w% 12

3-Methylpentane w% 6

n-Heptane w% 7

Process Conditions and Price

Temperature C 20

Pressure bar 1

Phase V/L/S L

Price EUR/ton 1,200

The in- and outgoing streams of feedstocks, products and wastes crossing the batterylimits are summarised for each process in Table 3.5 till Table 3.8. The onlyspecification for the ethanol product stream given, is that it has to contain at least 99.8vol% ethanol. In all four designs there is a waste water stream present.

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Table 3.7: Streams passing battery limits in the extractive distillation by PAA.Stream name: Ethanol feed Ethanol product Waste water

Component Mw kg/s kg/s kg/s

Ethanol 46.07 0.280 0.279 0.000Ethyl acetate 88.11 0.000 0.000 0.000

Acetaldehyde 44.05 0.000 0.000 0.000Water 18.02 0.042 0.000 0.042Isobutyl alcohol 74.12 0.001 0.000 0.001Isopentyl alcohol 88.15 0.002 0.000 0.002PAA 2000 0.000 0.000

Total 0.324 0.280 0.044

Enthalpy kW -2,356 -1,680 -672.2Phase L/V/S L L L

Pressure bar 1.0 1.0 1.0Temperature C 30.0 30.0 40.0

Price EUR/ton 250 550 -125

Table 3.8: Streams passing battery limits in the normal distillation followed by membranepurification.

Stream name : Ethanol feed Ethanol product Waste water

Component Mw kg/s kg/s kg/s

Ethanol 46.07 0.280 0.277 0.003Water 18.02 0.042 0.000 0.042Isopentyl alcohol 88.15 0.002 0.000 0.002Isobutyl alcohol 74.12 0.001 0.000 0.001Ethylacetate 88.11 0.000 0.000 0.000Acetaldehyde 44.05 0.000 0.000 0.000

Total 0.324 0.277 0.047Enthalpy kW -2,356 -1,666 -686.2

Phase L/V/S L L L

Press. bar 1.0 1.0 1

Temp C 30.0 30.0 40

Price EUR/ton 250 550 -224

3.3.5 General assumptionsIn designing thefour process alternatives several general assumptions have been

made. Further design choices are dealt with throughout the report and are not listedhere.

The general assumptions made are:

1. The feed stream of 88.8 vol% ethanol has the conditions of T = 30 C and p = 1bar. These conditions are assumed, because no conditions are given in theassignment (Appendix 3).

2. The ethanol product stream that crosses the battery limits has the conditions T =30 C and p = 1 bar. So the specification for the product purity (99.8 vol%) holdsfor these conditions. (Obviously this assumption doesnt hold for the gasohol

production plant.)

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3. The mass flows of all components in the feed and product stream are calculated byassuming that these streams are ideal liquids. The mass or volume flows arecalculated using the densities of each component separately at the statedconditions, so not one density of the whole mixture (non-ideal). To illustrate thisTable 3.9 is added. The error made by assuming this is very small and therefore

difficult calculations are avoided.4. The cooling water is available at 20 C and 3 bar and it is allowed to be disposed

off at 40 C after using as utility.5. Pressure loss due to pipeline friction is neglected.6. For condensers, reboilers, heaters and coolers a heat loss of 5 % is taken into

account. From there on the needed exchange area and utility amount is calculated.7. All pumps have an outlet flow at 0.50 m above ground level. This is taken into

account to calculate the duties of the pumps that pump liquid up to a certain level.8. The disposal of the condensed utility steam is not taken into account.

Table 3.9: The composition of the feed stream.

Name : Ethanol feed

Component MW kg/h kmol/h m3/h w% mol% vol% kg/m3(ref.28)

Ethanol 46.07 1,006.2 21.841 1.288 0.863 0.720 0.888 781.5

Water 18.02 150.95 8.379 0.152 0.130 0.276 0.105 993.7

Ethyl acetate 88.11 0.725 0.008 0.001 0.001 0.000 0.001 887.7

Acetaldehyde 44.05 0.435 0.010 0.001 0.000 0.000 0.000 767.6

Isobutylalcohol

74.12 1.813 0.024 0.002 0.002 0.001 0.002 790.4

Isopentylalcohol

88.15 5.438 0.062 0.007 0.005 0.002 0.005 803.2

Total 1,165.56 30.324 1.450 1.000 1.000 1.000Phase L/V/S L

Pressure bar 1.0

Temperature C 30.0

The density of isobutyl alcohol is estimated using the values given underneath: B*C/A

Density of 2-methyl-2-propanol at 20 C (kg/m3), A: 788.8 (ref. 28)

Density of 2-methyl-2-propanol at 30 C (kg/m3), B: 777.6 (ref. 28)

Density of isobutyl alcohol at 20 C (kg/m3), C: 801.8 (ref. 20)

3.4 Economic Margin

To determine the maximum allowable investments for the designs the economicmargin is calculated. This margin is the difference between income from sales minusthe costs of the feedstock. According to ref. 10, the market price for ethanol fuel gradeis 550 EUR/ton. The market price of ethanol 88.8 vol% is not available, so the pricehas to be assumed. As the ethanol prices are not proportional to the percentageethanol, because of the efforts to overwin the azeotrope, a price of 88.8 vol% ethanolis assumed at 45 % of the fuel grade price. So the market price of raw ethanol is set at250 EUR/ton. The prices for all the feedstocks, entrainers and products are tabulated

in Table 3.10.

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Table 3.10: Costs raw materials and incomes from product sales.

Raw materials: Unity EUR per unity

Ethanol 88.8 w% 1 ton 250

Gasoline2 ton 1,200

Toluene2 ton 450

Polyacrylic acid1

ton 1,300

Products: Unity EUR per unity

Ethanol ton 550

Gasohol ton 1,2961 Estimation2 ref. 10

The economic margin per year can be defined as:

Margin = Total Product Revenues - Total Feedstock Costs (3.1)

The margin for each process option is summarised in Table 3.11.

Table 3.11: Economic Margin.

Process option Productioncapacity

(t/a)

Productrevenues(kEUR/a)

Feedstock costs(kEUR/a)

Margin(kEUR/a)

Azeotropic distillationby toluene 8,286 4,565 2,562 2,003

Extractive distillationby gasoline 86,737 112,428 96,113 16,315

Extractive distillation

by PAA 8,689 4,616 2,513 2,103Normal distillationfollowed bymembrane purification 8,599 4,406 2,513 1,893

From the table it can be seen that the option extractive distillation by gasoline has thelargest margin due to the high production capacity. The other three options do notvary much from each other. With the economic margin the maximum allowedinvestment is calculated by a discount cash-flow analysis as described in ref. 33,

p.239. This method is used to calculate the present worth of future earnings. This canbe used to determine the maximum allowed investment (equation 3.2). A discounted

cash-flow rate of return (r) of 10 % is assumed and a plant life (n) of 15 years.

'1 1

n t

nn

MarginMaximum Allowable Investment

r

(3.2)

The maximum allowable investments calculated from equation 3.1 for each processoption are summarised in Table 3.12.

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Table 3.12: Maximum allowable investment for the four designs

Process option Maximum allowable investment (kEUR)

Azeotropic distillation by toluene 17,238

Extractive distillation by gasoline 140,408

Extractive distillation by PAA 18,099

Distillation followed by membranepurification 16,291

The results will be compared with the economic evaluation in Chapter 11.1.

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4 Thermodynamic properties

4.1 Operating window

The thermodynamic relations and estimation methods used in the process designsshould be valid for the temperatures and pressures occurring in the equipment.Therefore an operating window is defined for each process and is shown in Table 4.1.

Table 4.1: Operating window for the alternatives.

Process Temperature range (C) Pressure range (bar)

Azeotropic distillation by toluene 30.0 112.7 1 1.5

Extractive distillation by gasoline 20.0 81.6 1 2.4

Extractive distillation by PAA 30.0 92.9 1 3.0

Normal distillation followed bymembrane purification

30.0 94.8 (distillation)3.8 120.0 (membrane)

1 1.2 (distillation)0.0 4.3 (membrane)

4.2 Non-ideal equations for distillation

The main separation of the water from the ethanol takes place in the distillationcolumns in each design. On each stage in the distillation column both vapour andliquid phases are present and are in (thermodynamic) equilibrium. The equilibrium of

component j in the vapour and liquid phase is based on the equality of the fugacity f

in both phases (ref. 34, p.338):

l v

j jf f (4.1)

in which: ijf Fugacity of mixture of component j in phase i

This equation (4.1) accounts under the restrictions of constant temperature andpressure. For component j in the non-ideal vapour phase the fugacity and the fugacitycoefficient are related as follow (ref. 34, p.366):

v vj j jf y P (4.2)

in which:

jy Mole fraction of component j in the vapour phase

vj

Fugacity coefficient of the vapour phase mixture

P Total pressure

For component j in the non-ideal liquid phase a similar relation between fugacity and the

activity coefficient exists (ref. 34, p.368):

l lj j j jf x f (4.3)

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in which:

jx Mole fraction of component j in liquid phase

j Activity coefficient of component j

i

jf Fugacity of component j in phase i.

Using equations (4.1) to (4.3) the vapour-liquid equilibrium constant Kfor component j is:

l

j j j

j vj j

y fK

x P

(4.4)

This quantity determines the relative volatility of two components and thus theseparation of these components.

Other thermodynamic properties like heat capacities and enthalpies for the pure

components can be found in Appendix 1.

4.3 Choice of thermodynamic model

As can be seen in equations (4.1) till (4.4) the liquid activity coefficients and vapourfugacities are necessary to calculate the exact thermodynamic properties. For anestimation of these parameters several methods are in use. To determine the rightthermodynamic model, the feasible models are investigated in Aspen Plus 10. Thesemodels are the Wilson, the NRTL, the UNIQUAC and the UNIFAC model. All thesemodels are applicable to (highly) non-ideal mixtures, vapour-liquid equilibria (VLE)

and except for the Wilson model they are also applicable to two liquid phases.Although a phase separation into two liquids between ethanol and water is notexpected, but is expected between the water or ethanol and one or more entrainers, theWilson method seems to be more inaccurate in this design than the other methods.

The UNIFAC (UNIQUAC functional group activity coefficient) model is valid in thetemperature range from 2 C to 202 C and in a pressure range from 0 to 4 bar. Thismodel is an extended form of the UNIQUAC method and applies different models toestimate unknown thermodynamics properties from the group contributions instead ofmolecular contributions. A disadvantage of this method is the inaccurate values of

parameters. Because this model estimates a large number of parameters it is only usedin cases where almost all parameters are unknown. This is not the case with ethanol /water mixtures, so the UNIFAC method will not be used.

For the separation of ethanol and water the models NRTL and UNIQUAC are mostlikely to be used. When the separation is accomplished with the aid of a massseparating agent, for example toluene, the system changes from two phases (vapourliquid) into three phases (vapour liquid liquid). Also in this case both the models can

be used. To determine the influence of the model both the NRTL and the UNIQUACmethod are evaluated by creatingx,y-diagrams and residue curves for the binarysystems and the ternary system of ethanol, water and toluene (see figures below). In

these diagrams the azeotropic points should be seen at or near the value in literature.In Table 4.2 all relevant azeotropic points found in ref. 20 (p. 6-221, 6-239) are listed.

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Table 4.2: Azeotropic points at atmospheric pressure according to ref. 20.

mixture phases Wtfrac (%) Molfrac (%) T (C)

Ethanol / water LV 96 / 4.0 90 / 10 78.17

Ethanol / toluene LV 68 / 32 81 / 19 76.7

Water / toluene L1L2V * 13.5 / 86.5 44.4 / 55.6 84.1

Ethanol / water /toluene L1L2V * 37 / 12 / 51 40 / 33 / 27 74.4* : The azeotropic mixture is the vapour phase. The two liquid phases are almost immiscible.

4.3.1 Ethanol / water mixtureThe most important azeotrope is that of ethanol / water. For this mixture the VLE-curve obtained from the flowsheet calculation program Aspen Plus10 is comparedwith the theoretical data of Perry (ref. 29, p.13-12) in Figures 4.1 and 4.2.

Figure 4.1: Comparison of thex,y-diagram of the ethanol / water system from Aspen Plus 10

models and Perry.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

liquid molefraction ethanol (-)

vapourmolefractionethanol(-)

Uniquac 1 bar

NRTL 1 bar

Perry

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Figure 4.2: Zoom-in of the azeotropic point in the mixture ethanol-water.

As can be seen from Figure 4.2 the Aspen models are consistent with the literaturevalues given in Table 4.2. Because both the liquid and the vapour phase of the systemare highly non-ideal the NRTL-HOC and the UNIQ-HOC method are also evaluated.

These models use the Hayden-OConnell equation to describe the non-ideal vapourphase. The plot is shown in Figure 4.3.

0.8

0.82

0.84

0.86

0.88

0.9

0.92

0.94

0.96

0.98

1

0.8 0.82 0.84 0.86 0.88 0.9 0.92 0.94 0.96 0.98 1

liquid molefraction ethanol (-)

Vapourfractionethanol(-)

Uniquac

NRTL

Perry

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Figure 4.3: Comparison thermodynamic models UNIQUAC, UNIQUAC-HOC, NRTL andNRTL-HOC.

As can be seen in Figure 4.3, the difference between the available models is small.This is probably the result of the wide acquaintance of the used substances. Therefore

the UNIQUAC-model and the NRTL-model probably will take some non-idealbehaviour of the vapour phase into account. The two HOC-models are especiallysuitable for vapour reaction. Because no such reaction takes place, a choice has to bemade between the UNIQUAC and the NRTL model. Both models are appropriate, butthe UNIQUAC model is used in this situation, because of its wide acceptance in theliterature and its accuracy in representing VLE data for a wide range of systems (ref.2). A small disadvantage of the UNIQUAC (but also of NRTL) method is that the

parameters all inherit a Boltzmann-type Tdependence from the origins of theexpressions for GE, but it is only approximate. (ref. 29, p. 4-23)

4.3.2 Other componentsWithin the four design options three options make use of a third component toseparate the water from the ethanol, namely the entrainers toluene, polyacrylic acidand gasoline. For this third component the thermodynamic model should also bevalid. The UNIQUAC model accounts for liquid-vapour equilibria as well as liquid-liquid-vapour equilibria. Therefore the UNIQUAC model can be used in all fouralternatives. But in each design option the validity of UNIQUAC has to be checked:in Aspen Plus 10 for the interaction between ethanol, water and the third component.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

liquid mole fraction ethanol (-)

vapourmolefractionethanol(-)

UniQuac

NRTL

UniQ-Hoc

NRTL-Hoc

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Azeotropic distillation by toluene

To validate the use of UNIQUAC for the azeotropic distillation by toluene, thethermodynamic properties of ethanol, water and toluene are investigated. This is done

by comparing the equilibrium-curves of the binary mixtures and in a ternary mixture

with literature of Table 4.2 above.

Table 4.3: Azeotropic points at atmospheric pressure according to UNIQUAC-Aspen Plus 10.

mixture Phases molfrac (%)

Ethanol / water LV 90 / 10

Ethanol / toluene LV 81 / 19

Water / toluene L1L2V * 56 / 44

Ethanol / water / toluene L1L2V * 46 / 28 / 26* : The azeotropic mixture is the vapour phase. The two liquid phases are almost immiscible.

Aspen-UNIQUAC simulations are done to analyse all the equilibria of the mixtures of

ethanol, water and toluene. In Table 4.3 all azeotropic points found are summarised.First the equilibrium of ethanol and toluene is simulated. This simulation gives anazeotrope of about 81 mol% ethanol, showed in Figure 4.4. Table 4.2 gives aliterature value for the azeotrope of 81 mol% ethanol. So the UNIQUAC resultcorresponds well with literature.

Figure 4.4:X,y-diagram of the ethanol / toluene system from Aspen-UNIQUAC.

Y-x for ETHANOL/TOLUENE

Liquid Molefrac ETHANOL

0 0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5 0.55 0.6 0.65 0.7 0.75 0.8 0.85 0.9 0.95 1

Vapor

Mo

lefrac

ETHANOL

0.2

0.4

0.6

0.8

1

1.0133E+05 N/sqm

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Figure 4.5:X,y-diagram of the water / toluene system from Aspen-UNIQUAC.

As mentioned in Table 4.2 the mixture of water and toluene is a three phase mixture:two liquid phases and a vapour phase. As can be seen from Figure 4.5 the two liquid

phases are nearly pure water and toluene. This is because the inorganic and organiccomponents are not miscible. The vapour phase contains the azeotropic mixture.Figure 4.5 shows that this azeotrope lies at 55.8 mol% water. In ref. 11 a value of 55.5mol% water is given. A remarkable difference is seen between these two similarvalues and the literature value given in Table 4.2, 55.6 mol% toluene. Other Aspenmodels like NRTL and UNIFAC were consulted and all give about the same results asUNIQUAC. Although it seemed very remarkable that literature can not reachaccordance, it is assumed that (ref. 11) and the several Aspen-simulations are correct.

Finally a residue curve is made for the ternary mixture of ethanol, water and toluene,

as can be seen in Figure 4.6. The literature value stated in Table 4.2 is 40 mol%ethanol, 33 mol% water and 27 mol% toluene. The residue curve shown in Figure 4.6gives an azeotropic point at about 46 mol% ethanol, 28 mol% water and 26 mol%toluene.

This result is quite reasonable. According to the residue curve a greater fractionethanol will go overhead with the azeotropic mixture in a distillation column. Thisimplies that a column, where ethanol is supposed to be received as the bottom stream,

performs better in reality than a distillation column designed in a flowsheet programwith the UNIQUAC parameters.

Y-x for WATER/TOLUENE

Liquid Molefrac WATER

0 0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5 0.55 0.6 0.65 0.7 0.75 0.8 0.85 0.9 0.95 1

VaporMolefracWATER

0.2

0.4

0.6

0.8

1

1.0133E+05 N/sqm

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in which:

ij UNIQUAC binary interaction parameter (-)Aijt/mDij UNIQUAC regression parametersT Temperature (K)

Table 4.4: Regressed Aspen-UNIQUAC parameters of water (i) / ethanol (j) / PAA VLE dataUNIQUAC Aij Aji Bij Bji Cij Cji Dij Dji

Aspen-default(with azeotrope) -2.4936 2.0046 756.95 -728.97 0 0 0 0

Regression(without azeotrope) -3.6339 -143.15 1,519.8 50,000 0 0 0 0

Nevertheless it can be seen from Figure 4.7 that the curve fits quite well through theexperimental data. However, because of the very small amount of experimental data

points at high ethanol mole fractions, which is near the possible azeotrope, theregressed parameters seem not very realistic. Trying to make a feasible design of the

distillation column the regressed equilibrium curve is used.

Figure 4.7: Aspen-UNIQUAC regression of the experimental data of ref. 2.

An additional doubt of the validity when using the regression results arises. Theregressed equilibrium curve is only valid at the pressure of 1 bar, but in the distillationcolumn some pressure drop has to be designed. When using the curve only at 1 bar itreflects the influence of PAA well. But in Figure 4.8 can be seen that the influence of

pressure on the equilibrium curve is large. Here a trade-off situation occurs: shouldpressure drop be designed in this case, in spite of the large influence of theUNIQUAC parameters on pressure dependent VLE data, or should the distillationcolumn be designed without any pressure drop.The choice has been made to design a pressure drop in the column because in practicethere is always some pressure drop in a column. As Figure 4.8 shows the azeotropedoes occur at higher pressures than about 1 bar. Because the purpose of this option isto avoid the azeotrope by adding an entrainer, the pressure in the column design iskept below 1.0 bar. In this way there will be no azeotrope and equilibria curves that

y vs. x

Mole fraction ETHAN-01

0 0.2 0.4 0.6 0.8 1

Molefract

ionETHAN-01

0.2

0.4

0.6

0.8

1

Exp D-1 R-1

Est D-1 R-1

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are extrapolated will be used. Obviously there is doubt of the validity of this design,so further experimental data should be obtained before final design and constructionare realised. To compensate slightly for these uncertainties the column is designed insuch way that the product concentration is higher than the original specification.

Figure 4.8: The influence of pressure on the ethanol / water / PAA equilibrium curve.

Y-x for WATER/ETHAN-01

Liquid Molefrac ETHAN-01

0 0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5 0.55 0.6 0.65 0.7 0.75 0.8 0.85 0.9 0.95 1

VaporMolefracETHAN-01

VaporMolefracETHAN-01

VaporMolefracETHAN-01

0.2

0.4

0.6

0.8

1

0

0.2

0.4

0.6

0.8

1

0

0.2

0.4

0.6

0.8

1

1.1 bar

1.0 bar

0.9 bar

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5 Process structure & description

In this chapter the selection of unit operations, equipment, utilities and process chemicalswill be made clear on the basis of the design criteria mentioned in Paragraph 5.1. Inaccordance to the choice of the equipment and the utilities the final process conditions are

defined. The processes are extendedly drawn in the process flow schemes and in thischapter a short description is available. Finally the process yields are given. The actualdesign details are treated in Chapter 8.

5.1 Design criteria

According to the assignment only one criterion has to be fulfilled, namely the ethanolproduct purity has to be 99.8 vol%. This purity has to be achieved regardless of thespecified variations in the feed stream (see Paragraph 3.3.4). Another criterion, which istried to be fulfilled, is a recovery of at least 99 % of the pure ethanol present in the feedstream. Besides, the costs of the total process will be kept as low as possible. To attainthese design criteria the selection of the equipment, special process condition, utilities and

process chemicals are reviewed in the subsequent paragraphs. The influences of thesechoices will be shown in Paragraph 5.7.

5.2 Unit operations

5.2.1 Distillation columnsIn the chosen dehydration units, a distillation column is the backbone of each design.Distillation columns can be plate columns or packed columns. The criteria for selection

between these two possibilities are given in Table 5.1.

Table 5.1: Selection criteria and their evaluations for plate and packed columns.

Vapour-liquid contact A plate column provides a good vapour-liquid contact and is stage wise,while the vapour-liquid contact in a packed bed column is continuous.However the performance of a packed column is dependent on themaintenance of good liquid and vapour distribution throughout the bed.There is always some doubt that good distribution can be maintainedthroughout a packed column.

Accuracy Plate columns can be designed with more assurance than packed columns,

because in packed columns there is always some doubt about the goodliquid distribution. Because the high requirements of the product there isonly a small operating range. Therefore the accuracy of the design has to bevery good to ensure that the product requirements are reached.

Efficiency A plate column provides a sufficient liquid hold-up. This provides a goodmass transfer and therefore a high efficiency. The efficiency of a plate can

be predicted with more certainty than the equivalent term for packing due tothe liquid distribution.

Pressure drop In general the pressure drop of a packed column can be lower than of a platecolumn. The pressure drop of a plate column can be held within acceptablelimits when a sufficient area and spacing is kept.

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Table 5.1 (continued)

Properties ofchemicals

Ethanol and water are non-corrosive or foaming substances. Therefore is itnot necessary to choose for a packed column on the basis of a foaming orcorrosive mixture.

Economy Plate columns with a small diameter are quite expensive because the plates

are difficult to install. Packing is much cheaper in small diameter columns.

Using the criteria above, a choice is made for plate columns because of the high accuracyof the design. This is a very important factor, because the high requirements of purity of99.8 vol% have to be reached. Because the liquid distribution is an important factor on theaccuracy, efficiency and vapour liquid contact, some fluctuations in the packed columncannot be prevented. Therefore the certainty of a plate column is chosen, in spite of itshigher costs.

Not only the type of column, but also the plate contractor in the column has influence onthe overall performance. For the selection of the plate contactor the choice can be made

between three types:- Sieve plates- Bubble-caps- Valve plates

Considering the costs, sieve plates have been chosen. A sieve plate is approximately 1.5and 3 times cheaper than valves and bubble-caps respectively. Sieve plates operatesatisfactory for most applications. A disadvantage of sieves is that the operating range ofsieves is smaller than the operating ranges for bubble-caps and valves, especially at start-up and shutdown conditions. Because sieves plates rely on the vapour flow through theholes to hold the liquid on the plate, sieves can not be operated at very low vapour rates.

Bubble caps and valves have a positive liquid seal and can operate efficiently at lowvapour rates. So sieves plates are satisfactory on the condition that column weeping ischecked.

5.2.2 CondensersThe top stream of a distillation column can be condensed partially or totally, or totallynot. The type of condenser influences the heat duty and the downstream equipment.Because the downstream equipment operates with liquid phases a total condenser has

been chosen. This implies that one part of the liquid stream is refluxed, and the other part

can be used as liquid feed for the downstream apparatus. For the condenser a fixed tubesheet is chosen, because this is the simplest and cheapest type of exchanger.

5.2.3 ReboilersIn designing a reboiler there can be chosen between different kinds of types:- Forced circulation- Thermosyphon (natural circulation)- Kettle type (no circulation)

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The forced circulation reboiler is especially suitable for viscous and fouling media or isused under high vacuum (< 0.3 bar). Non of these circumstances will occur in the processso there has to be made a choice between the thermosyphon reboiler and the kettlereboiler. In most cases a thermosyphon reboiler is the most economical and most usedreboiler for applications above 0.3 bar. However a disadvantage of the thermosyphon

reboiler is that the column base has to be elevated to provide hydrostatic head required forthe thermosyphon effect (ref. 33). This affects the column supporting structure and willincrease the costs. The kettle reboiler has a lower heat transfer coefficient and there is noliquid circulation. Generally it will be more expensive than an equivalent thermosyphonreboiler, but the costs can be decreased fairly by implementing the reboiler in the base ofthe column. The costs will be in this case competitive or lower than using a thermosyphonreboiler, while no additional column supporting structure is necessary. Therefore a kettlereboiler is chosen in all four design options.

5.2.4 Heat exchangersAn important part of the total plant expenses comes on the account of heating and coolingthe process streams. To utilise these heat streams more efficiently, and to decrease theoverall costs, heat exchangers are used. The most commonly used type of heat-transferequipment is the omnipresent shell and tube exchanger. However there are more optionsavailable as the double pipe exchanger, plate and frame exchangers, plate fin exchangersand air coolers. A distinguish can be made between exchangers between two processstreams and a heating or cooling equipment which requires utilities.

For the process streams exchanger a choice is made for the in chemical industriescommonly used tube and shell exchangers. This is done because of the following usefuladvantages (ref. 33, p. 584):

A large surface area in a small volume

Good mechanical layout for pressure operation

Can be constructed from many materials with well-established fabrication techniques

Can easily be cleaned

Due to the advantages mentioned above, the coolers of the process streams are operatedwith cooling water. However more possibilities are available to cool process streams. Aninteresting option is the use of air-cooled heat exchangers, because air is cheap and easilyavailable and there is no probability of leakage. Air-cooled heat exchangers are mostly

used in areas, where seasonal variations in ambient temperatures are relatively small.Because a research on the climate in Lithuania falls out of the scope of the assignment, afirst choice is made to use cooling water. Moreover cooling water is available accordingto the assignment and quite constant in temperature. However air coolers remain aninteresting possibility and should be investigated (see Chapter 13). For the heating of

product streams medium pressure steam of 10 bar is used. Both the coolers and heaters areshell and tube exchangers.

All heat exchangers are operated counter-currently. This is because a counter-current heatexchanger is more effective than a co-current heat exchanger. The fluids flowing throughthe pipe are continuously changing in temperature. If the two streams are flowing in an

opposite direction, the temperature difference between the shell and tube temperature willshow less variation than in the case of co-current flow. Therefore it is possible for the

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cooling liquid to leave at a higher temperature than the heating liquid, contrary to a co-current operated heat exchanger, where the outlet of the heating fluid must always behigher than that of the cooling fluid. Another advantage of counter-current flow is theextraction of a higher proportion of the heat of the hot fluid.

The factors given in Table 5.2 will determine the allocation of the fluids in the shell or inthe tubes (ref. 33). When several factors contradict the most important one will decide theallocation.

Table 5.2: Factors determining the fluid allocation in heat exchangers.

Factor Rule of thumb

Corrosion The most corrosive fluid should be allocated in the tubes.

Fouling The fluid with the highest fouling-tendency should be allocated in the tubes.

Fluid temperatures The fluid with the highest temperature should be allocated in the tubes.

Operating pressures The fluid with the highest pressure should be allocated in the tubes.

Flow-rates The fluid with the highest flow-rate should be allocated in the tubes.

Process stream heat exchangers are implemented in the design to optimally use theavailable heat capacity of process streams (heat integration). In this way the amount ofheating or cooling heat exchangers and their utilities are minimised.

5.2.5 VesselsIn the four designs several vessels are used: reflux accumulators and decanters. Allvessels are designed as a cylinder because this is the cheapest shape (ref. 33). The

properties of the incoming stream and the purpose of the vessel determine the position of

the vessel. For example decanters are essentially tanks, which have to give sufficientresidence time for the droplets of the dispersed phase to settle readily. For small flowrates, which is the case in both two designs containing a decanter, a vertical cylindricalvessel is more economical than a horizontal one. For great stream rates the decanter will

be cheaper as a horizontal vessel. Furthermore the size of the vessels is based on a chosen(average) residence time.

5.2.6 PumpsFor the selection of the pumps distinction can be made between dynamic pumps and

positive displacement, reciprocating pumps. Positive displacement pumps are normallyused where a high Net Positive Suction Head (NPSH) is required at a low flow rate.Because this is not the case a dynamic pump will be installed. The by far most widelyused type in chemical industry is the centrifugal pump. It is capable of pumping liquidswith very wide-ranging properties and can be constructed from a very wide range of(corrosion resistant) materials. Therefore the centrifugal pump is perfectly capable ofhandling the fluids present in the designs and is cheaper than other types of pumps.For a final design the selection of the pump cannot be separated from the design of thecomplete piping system. The completeNPSHrequired will be the sum of the dynamichead due to friction losses in the piping, fittings, valves and process equipment, and any

static head due to differences in elevation. In this design the friction losses and the designof the piping system are left out of consideration.

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5.4.1 Cooling waterIn all designs cooling water is used for the cooling of the process streams. Another

possibility is the use of air-cooling, but because the uncertainty of the climate in Lithuaniacooling water is preferred. The major users of cooling water are the condensers at the topof the column.

5.4.2 ElectricityIn each design electric pumps are defined. Because the size of most of the pumps is verysmall, electric pumps are the cheapest. Another possibility is to execute the pump onsteam pressure. In the case of a power failure this kind of (essential) pumps are notaffected. However steam pressure pumps are more expensive and therefore in this

conceptual design only electrical pumps are used. In a finite design steam pressure pumpsshould be considered as essential points for maintaining safety.

5.4.3 Medium pressure steamFor the heating in the reboilers of the distillation columns medium pressure steam is used.Other possibilities are furnace-heating or oil-heating. Because the distillation column

bottom streams are too small both alternative possibilities are more expensive than theone using medium pressure steam. Medium pressure steam is also used for the heating of

process streams where necessary.

5.4.4 Liquid nitrogenThis special utility is needed to condense the permeate stream of the pervaporationmembrane-unit. This stream is nearly pure water vapour at 0.008 bar and 3.8 C. Thisstream cannot be cooled with the available cooling water of 20 C. Therefore liquidnitrogen is used as cold utility in accordance with ref. 19 and ref. 36.

5.5 Final process conditions

The main objective for designing the ethanol dehydration plant, is to minimise the costs,without losing touch with environmental and safety issues. This means that whenever

possible the conditions of the process should be at ambient pressure and at lowtemperatures. In most cases these conditions are maintained, but in some cases it is notfeasible to retain this conditions. In the four design options, only the pressure in theoptions extractive distillation by gasoline and extractive distillation by PAA are alteredslightly from ambient pressure.

5.5.1 Extractive distillation by gasoline

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In the case of the extractive distillation by gasoline the separation of ethanol and water inthe distillation column can be excellently achieved at ambient pressure, but then the toptemperature of the column is 30.6 C. This temperature complicates the condensation ofthis top stream. A difference of temperature between the cooling water and the processstream should be at least 10 C to maintain a heat transfer (see Chapter 8). The cooling

water comes in at a temperature of 20 C. To meet the requirement mentioned above, themaximum outlet temperature of the cooling water is 20.6 C. This implies that anenormous amount of cooling water is necessary to cool the process stream and thecapacity of the cooling water is not totally used. The costs to condense the top stream willgrow out of proportion. Therefore the column has to be adjusted to the available utilities.This implies that the column has to be brought under pressure, until the top temperature ishigh enough, to use the cooling water optimal. In the mean time the separation has to staysatisfactory. This point is reached when the column is operated at a pressure of 2.3 bar.The top stream has a temperature of 47.5 C, so cooling water can be used as a utility.

5.5.2 Extractive distillation by PAAIn the case of PAA there is a problem in the thermodynamics of the distillation column,earlier discussed in Chapter 4. The regressed UNIQUAC-parameters are valid at 1 bar,while the pressure in the distillation column changes. Furthermore it appears that thedependence of the UNIQUAC-parameters on the pressure dependent vapour-liquidequilibrium is strong (Figure 4.8 in Chapter 4). In the design the pressure is kept below 1

bar. This is done because the essence of the design is the avoidance of the azeotrope ofethanol and water. The final conditions of the column are 0.84 bar at the top and 1 bar atthe bottom.

5.6 Process Flow Schemes

For each designed option a process flow scheme is added in Appendix 8. In this paragrapheach design is described according the accompanying process flow scheme. The bases forthe process flow schemes are the Aspen Plus 10 simulations. In Appendix 9 a descriptionof the Aspen Plus 10 files are given.

5.6.1 Azeotropic distillation by tolueneThe toluene make up stream (available at T = 30 C and p = 1 bar) is added to theethanol feed stream and is led to the heat exchanger (E01) to heat up before thestream is pumped to distillation column (C01), which is operated at 1 bar. In thedistillation column (C01) the ternary azeotropic mixture of toluene, ethanol and watercomes overhead and a polluted ethanol stream comes over the bottom . Thevapour of the top stream is condensed in the total condenser (E02) and partly refluxed(stream ) to the distillation column (C01). The other part (stream ) is led todecanter (S01), where the toluene and the ethanol / water layer are separated. The toluenestream is fed to distillation column (C03) (1 bar) to purify the toluene fromimpurities like ethyl acetate and acetaldehyde. Although the streams are very small, a

distillation column is necessary to prevent accumulation of the impurities in the process.The top stream of distillation column (C03) which contains waste water is partly

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refluxed and partly mixed with other waste water streams. The bottom stream ofcolumn (C03) contains pure toluene and is recycled to distillation column (C01). Themass flow of toluene that stays in the process by recycling is 560 kg/h.

The bottom stream of distillation column (C01) is led to distillation column (C02)

(1 bar). In this column (C02) the polluted ethanol stream is purified from itsimpurities isobutyl alcohol and isopentyl alcohol. The top stream is partly refluxedand partly led to heat exchanger (E01) to cool down. Further cooling of the productstream , which contains 99.9 w% ethanol, is obtained in heat exchanger (E09). The

bottom stream contains waste water. This stream is mixed with the other wastewater streams and , respectively originated from the decanter (S01) anddistillation column (C02).

For each stream the exact composition, temperature, pressure, phase and enthalpy can befound in the process stream summary in Appendix 10.

5.6.2 Extractive distillation by gasolineAfter an increase in pressure, the ethanol feed and the gasoline feed (available atT = 20 C and p = 1 bar) are led to distillation column (C01). Before entering thedistillation column gasoline stream is split into stream , distillation feed, andstream , used later on to produce gasohol. In the column (C01) the ethanol isseparated from the water and the water comes overhead together with an amount ofgasoline . This vapour is condensed and is partly recycled to the column (C01),stream . The other part, stream is cooled down in heat exchanger (E03) before itis led to decanter (S01). In the decanter (S01) a water stream is separated from agasoline stream . The water stream is led to the water storage. The gasolinestream heated in heat exchanger (E04) before it is recycled to distillation column(C01).

The bottom stream of column (C01) contains a mixture of gasoline and ethanol,which is cooled down in heat exchanger (E04). The mixture is mixed with gasoline stream to produce stream , which is gasohol with the desired 10 w% ethanol.

For each stream the exact composition, temperature, pressure, phase and enthalpy can befound in the process stream summary in Appendix 10.

5.6.3 Extractive distillation by PAAThe ethanol feed is led to pump (P01) before it is led to heat exchangers (E01) and(E02) to heat up. This stream is mixed with the recycle stream of water and PAA. The mixed stream is led to distillation column (C01), where the mixture isseparated in an ethanol stream at the top and a water / PAA stream at the

bottom. The top stream of the distillation column (C01) contains no polyacrylic acid. It iscondensed in heat exchanger (E03) and partly refluxed. The pressure of the ethanol stream is brought to 1 bar in pump (P03). The ethanol is decreased in temperature by heat

exchanging it with the feed stream in (E01) and in heat exchanger (E06). The productstream is led to a storage tank.

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At the bottom of the column (C01) a mixture of water and PAA is increased inpressure by pump (P04) before the stream is cooled down in heat exchanger (E02). Thestream is further cooled down in heat exchanger (E05) and led to the ultrafiltration unit(S01). Therein 95 w% of the water and hydrocarbons is split off the PAA. The wastewater

stream is led to a biological waste water treatment. In stream the mass ratioPAA / water is 39 / 61 (for transporting as recycle stream). It is decreased in pressure andmixed with the feed stream . It is assumed that the polymer is recycled completely soin steady state no PAA will have to be added. The mass flow of PAA that stays in the

process by recycling is 5 kg/h. This is consistent with the 0.45 w% of the azeotropicmixture mentioned in ref. 2.

For each stream the exact composition, temperature, pressure, phase and enthalpy can befound in the process stream summary in Appendix 10.

5.6.4 Normal distillation followed by membrane purificationThe ethanol feed stream is heated in heat exchanger (E01) before the stream is fed todistillation column (C01). In the column (C01) the azeotropic mixture is coming overheadand the redundant water over the bottom. The top stream is condensed and partlyrefluxed. The azeotropic mixture is increase in pressure in pump (P03) and heated in heatexchanger (E04). Further heating to 120 C is obtained by heater (E05) before the mixtureis led into membrane unit (S01). In the membrane unit (S01) water is separatedfrom the ethanol by pervaporation. The ethanol stream is brought back to 120 C byheater (E06) and further purified in membrane unit (S02). The ethanol stream isheated once again to 120 C and its final purification takes place in membrane unit (S03).The purified ethanol stream coming out from membrane unit (S03) is cooled down in heatexchanger (E04). The ethanol stream is cooled down in heat exchanger (E09) and furthercooled down in cooler (E10) before the product stream is stored.

The released water streams and from respectively membrane unit (S02) and(S03) are mixed with water stream from membrane unit (S01) to create stream. This vaporous, vacuum stream is condensed in heat exchanger (E08) with liquidnitrogen. The vacuum pump (P04) is only used during start-up procedures. To increasethe pressure of stream a hydrostatic pressure increase is utilised before the stream isheated in heat exchanger (E09). This is to prevent the production of ice, see Chapter 8.

Subsequently the water stream is increased to ambient pressure in pump (P05).

The water stream originating from the bottom of column (C01) is cooled down inheat exchanger (E01) and then mixed with stream . The mixed stream is led toa biological waste water treatment.

For each stream the exact composition, temperature, pressure, phase and enthalpy can befound in the process stream summary in Appendix 10.

5.7 Process performance

After the considered selections of equipment, entrainers, utilities and process conditionsin Paragraph 5.2, the purities and recoveries of the four options can be calculated. Besides

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a comparison on the purity and recovery a comparison can be made between the processeson the basis of the process yields.

5.7.1 Purity and recoveryIn Table 5.4 an overview is given of the purity and recovery of the four designed

processes.

Table 5.4: Summary of purity and recovery of the designed processes

Process option Purity (vol%) Purity (w%) Recovery(kg/kg)

Azeotropic distillation by toluene 99.95 99.94 0.955

Extractive distillation by gasoline - - 1.000

Extractive distillation by PAA 99.89 99.88 0.999

Normal distillation followed by

membrane purification 99.86 99.84 0.990

As is shown in Table 5.4 all options attain the required minimum purity of 99.8 vol%. Inthe case of extractive distillation by gasoline the required purity cannot be obtained

because the ethanol is directly mixed with the gasoline. Instead of the purity, themaximum allowed amount of water can be compared with the actual amount of water inthe gasohol. This absolute amount is calculated by multiplying the amount of ethanol inthe feed stream with the obtained recovery of ethanol in the gasohol stream and with themaximum percentage of impurities allowed in the ethanol content of gasohol (0.2 w%).Calculated in this ways the maximum allowable amount of water in the gasohol productstream becomes:

8,673 t/a ethanol in the feed stream 100 % recovery 0.2 % = 17.35 t/a = 5.5910-4kg/sH2O. The actual amount of water amount to 1.8

.10-4kg/s. This implies that also thiscriterion is completely satisfied. Furthermore the gasohol produced contains the desired10 w% ethanol.

The other design criterion, to obtain a recovery of 99 %, is achieved in all options exceptfor the azeotropic distillation by toluene. In this option only a recovery of95.5 % is obtained. This is mainly ca