Design and performance of two-phase flow pervaporation … · Design and performance of two-phase flow pervaporation and hybrid distallation processes ... flow pervaporation and hybrid

Design and performance of two-phase flow pervaporationand hybrid distallation processesFontalvo, J.

DOI:10.6100/IR602790

Published: 01/01/2006

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Citation for published version (APA):Fontalvo, J. (2006). Design and performance of two-phase flow pervaporation and hybrid distallation processesEindhoven: Technische Universiteit Eindhoven DOI: 10.6100/IR602790

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https://doi.org/10.6100/IR602790

https://research.tue.nl/en/publications/design-and-performance-of-twophase-flow-pervaporation-and-hybrid-distallation-processes(2b859659-7752-4958-ac32-98b89a92fac9).html

DESIGN AND PERFORMANCE OF TWO-PHASE FLOW PERVAPORATION AND

HYBRID DISTILLATION PROCESSES

PROEFSCHRIFT

ter verkrijging van de graad van doctor aan de

Technische Universiteit Eindhoven, op gezag van de Rector Magnificus, prof.dr.ir. C.J. van Duijn, voor een

commissie aangewezen door het College voor Promoties in het openbaar te verdedigen op woensdag 8 februari 2006 om 16.00 uur

door

Javier Fontalvo Alzate

geboren te Bogotá, Colombia

Dit proefschrift is goedgekeurd door de promotor: prof.dr.ir. J.T.F. Keurentjes Copromotor: ir. J.G. Wijers

CIP-DATA LIBRARY TECHNISCHE UNIVERSITEIT EINDHOVEN

Fontalvo, Javier Design and performance of two-phase flow pervaporation and hybrid distillation processes / by Javier Fontalvo Alzate. – Eindhoven : Technische Universiteit Eindhoven, 2006. Proefschrift. – ISBN 10: 90-386-3007-7. – ISBN 13: 978-90-386-3007-6 NUR 913 Subject headings: ceramic membranes / pervaporation / hybrid processes / distillation / two-phase flow / Maxwell-Stefan / simulation of processes / dehydration Druk: JWL boekproducties Copyright 2006, J. Fontalvo

Por Paula, mis padres, mi hermano y Miguel

The financial support by the Programme Office on Economy, Ecology and Technology from the Dutch Ministries of Economic affairs, Education and Environmental affairs under contract EETK20046 is acknowledged.

Table of contents

SUMMARY Chapter 1 INTRODUCTION AND OUTLINE OF THE THESIS. 1 1.1. Introduction 2 1.2. Concentration and temperature polarization 3 1.3. Temperature drop 5 1.4. Hybrid distillation-pervaporation processes 7 1.5. Scope of this thesis 8 1.6. Outline of this thesis 9 1.7. Reference List 11 Chapter 2 SYNTHESIS, PERFORMANCE AND STABILITY OF SILICA MEMBRANES FOR GAS PERMEATION AND PERVAPORATION

13

2.1. Introduction 14 2.2. Theory 14

2.2.1. Gas permeation 14 2.2.2. Pervaporation 15

2.3. Experimental 16 2.3.1. Hollow fibre membrane preparation 16 2.3.2. Tubular membranes 17 2.3.3. Membrane characterization 17 2.3.4. Gas permeation 17

2.4. Pervaporation 17 2.5. Results 18

2.5.1. Membrane characterization 18 2.5.2. Gas permeation 20 2.5.3. Pervaporation 22

2.6. Conclusions 25 2.7. Notation 25 2.8. Reference list 26

Chapter 3 STUDY OF THE HYDRODYNAMICS IN A PERVAPORATION MODULE AND IMPLICATIONS FOR THE DESIGN OF MULTI-TUBULAR SYSTEMS

29

3.1. Introduction 30 3.2. Setup 31 3.3. Ultrasound reconstructions and calibration 32 3.4. Numerical simulations 34 3.5. Comparison of experimental and calculated data 35

3.5.1. Pervaporation of pure water 35 3.5.2. Pervaporation of ethanol-water mixtures 38

3.6. Hydrodynamics during pervaporation of several organic-water mixtures 38 3.7. Implications for the design of multi-tubular pervaporation modules 40 3.8. Conclusions 42 3.9. Reference list 42 Chapter 4 HEAT SUPPLY AND REDUCTION OF POLARIZATION EFFECTS IN PERVAPORATION BY TWO-PHASE FEED

45

4.1. Introduction 46 4.2. Theory 48

4.2.1. Pervaporation 48 4.2.2. Two-phase flow 48 4.2.3. Mass and heat transfer coefficients in slug flow 50

4.3. Experimental 51 4.3.1. Lab scale setup; air as secondary phase 51 4.3.2. Bench scale setup; vapor as secondary phase 53

4.4. Results 54 4.4.1. Dioxane-water system; gas –liquid pervaporation 54 4.4.2. Experimental values on slow flow and theoretical prediction of polarization

56

4.4.3. IPA-water system; vapor-liquid pervaporation 58 4.4.4. Water-IPA system; experimental comparison of two-phase pervaporation using air and vapor

60

4.5. Conclusions 61 4.6. Notation 61 4.7. Appendix 63 4.8. Reference list 65

Chapter 5 SEPARATION OF ORGANIC-WATER MIXTURES BY CO-CURRENT VAPOR-LIQUID PERVAPORATION WITH TRANSVERSAL HOLLOW FIBRE MEMBRANES

69

5.1. Introduction 70 5.2. Description of the pervaporation module and the simulations 72 5.3. Solid-liquid mass transfer measurements and gas void fractions 74 5.4. Results 76

5.4.1. Gas void fractions 76 5.4.2. Mass transfer coefficients 77

5.5. Comparing single phase and two-phase flow pervaporation units 79 5.5.1. Single phase pervaporation units with inter-stage heating 79 5.5.2. Liquid-vapor two-phase pervaporation unit 80

5.6. Conclusions 82 5.7. Notation 82 5.8. Reference list 84 Chapter 6 COMPARING PERVAPORATION AND VAPOR PERMEATION HYBRID DISTILLATION PROCESSES

87

6.1. Introduction 88 6.2. Process modeling 88 6.3. Conventional distillation based processes 91 6.4. Membrane-distillation hybrid processes 93

6.4.1. Distillation as the final dewatering step 93 6.4.2. Membrane as a final dewatering step 95

6.5. Guidelines for selecting hybrid pervaporation – distillation or hybrid vapor permeation-distillation processes

98

6.6. Economical evaluation 100 6.7. Conclusions 102 6.8. Notation 103 6.9. Appendix 106 6.10. Reference list 108

Chapter 7 SEPARATION OF MULTICOMPONENT MIXTURES USING AN INTEGRATED DISTILLATION – PERVAPORATION COLUMN

111

7.1. Introduction 112 7.2. Theory and simulations 113 7.3. Results 116

7.3.1. Binary systems 117 7.3.2. Multicomponent systems 120

7.4. Conclusions 123 7.5. Nomenclature 124 7.6. Reference list 124 Chapter 8 CONCLUSIONS AND PERSPECTIVES 129 8.1. Introduction 130 8.2. Major findings 130

8.2.1. Membrane stability 130 8.2.2. Hydrodynamics, mass and heat transfer 131 8.2.3. Performance of pervaporation units in two-phase flow in stand alone and hybrid applications

132

8.3. Perspectives 133 8.3.1. Industrial application of two-phase flow 133 8.3.2. Further developments in pervaporation processes 134

8.4. Conclusions 135

Summary Pervaporation has been successfully applied for the separation of mixtures in

combination with distillation. In particular, this is because the energy consumption of distillation for the separation of azeotropic mixtures or mixtures with low relative volatility is high. Pervaporation is attractive because the energy required for the separation is strongly reduced due to the selective vaporization of the permeating components only. However, pervaporation involves higher capital cost than distillation processes. Thus, a hybrid process in principle combines the advantages of both separation techniques while the disadvantages can be minimized.

Several problems can be identified in pervaporation processes, including

concentration and temperature polarization, and a temperature drop at the retentate side. These effects will increase the required membrane area, the number of auxiliary equipment and consequently the total cost for a specific separation duty. This thesis explores alternatives for reducing these problems by using a gas-liquid or a vapor-liquid mixture as the feed in pervaporation units, and studies the resulting implications for the design and performance of hybrid distillation processes. For this, two types of hybrid processes are considered: the conventional ones where the pervaporation unit is externally connected to the distillation column and one where distillation and pervaporation are combined in one single column.

Hollow fibre silica membranes have been analyzed using SEM, SNMS, single gas

permeance and pervaporation with dimethylformamide (DMF)-water mixtures. In addition, aging experiments have been performed on tubular silica membranes for the removal of water. The dehydration experiments with DMF –water and alcohol -water mixtures show that the water flux strongly decreases with time. This reduction in flux is due to interactions between the organic compound in the mixture and the silica layer, i.e. adsorption on and reaction with silanol groups. Nevertheless, it is experimentally shown that this interaction is at least partly reversible.

The hydrodynamics and flux in a pervaporation module have been studied using

CFD and ultrasound computer tomography for several density ratios between the permeating and the non-permeating component. Calculated density and temperature profiles are similar to those reconstructed from measurements. An inversion point has been found on the surface of tubular membranes above which the fluid moves upwards and

below which the fluid moves downwards. The inversion point suggests that for the dewatering of mixtures with a low density ratio a multi-tubular pervaporation module with triangular configuration is convenient. For dewatering of mixtures with a high density ratio a squared configuration is preferred.

The performance of pervaporation modules with co-current vapor-liquid and gas-

liquid flow inside of tubular membranes has been studied. A lab scale and a bench scale setup have been used for the dehydration of 1,4-dioxane–water and isopropyl alcohol (IPA)-water mixtures, respectively. Small amounts of air and vapor increase the turbulence in the liquid phase, thereby reducing concentration and temperature polarization. The flux and selectivity have been increased more than twofold as compared to single phase flow in the laminar regime. Moreover, by using vapor the energy required for the pervaporation process is supplied to the liquid by condensation, avoiding the use of inter-stage heat exchangers and reducing the required membrane area. Good predictions of the total flux for air-liquid flow in the slug regime have been obtained based on experimental data of bubble size, liquid slug size and bubble rise velocity. The performance of a two-phase pervaporation module at low liquid flow rates is close to the performance in single phase at turbulent conditions.

In addition, pervaporation modules operating with co-current two-phase flow

around hollow fibre membranes have been analyzed. Gas void fractions and mass transfer coefficients between the liquid and the membrane surface have been measured in two-phase flow using an electrochemical technique in a transversal unit with hollow fibre membranes. The mass transfer coefficient can be four times higher than the one obtained in single phase. The resulting dimensionless correlation has been included in a rate-based model for designing a pervaporation module for dewatering of an IPA-water mixture. A reduction of the membrane area of about 45% can be obtained as compared with pervaporation modules with an economically optimal number of inter-stage heat exchangers. The lower membrane area required and the avoidance of inter-stage heat exchangers results in a strong reduction of capital cost.

Conventional hybrid processes, which consist of a membrane separation unit that is

externally connected to a distillation column, have been studied. Guidelines to decide whether pervaporation or vapor permeation is more convenient for a specific application are presented. The positive influence of relatively low-selective membranes on the total cost of the process is demonstrated for the dewatering of acetonitrile (ACN)-water mixtures. Also, recycling a fraction of the retentate into the permeate side leads to a strong reduction of the required membrane area and thus the total separation cost. For a pressure sensitive azeotropic mixture such as ACN-water, it appears effective to use a relatively low pressure in the distillation column and a higher pressure in the pervaporation

unit. As compared to distillation-based processes, a reduction between 25% and 60% in the total separation cost can be achieved.

Hybrid distillation processes that combine a pervaporation unit operating in

vapor-liquid flow have been analyzed. The pervaporation unit has been integrated with distillation in one single column (DPSU). The pervaporation zone, containing hollow fibre pervaporation membranes, replaces a section of packing or trays in a distillation column. The proposed hybrid system is evaluated for the dewatering of binary mixtures such as ethylene diamine-water and IPA-water and for removal of methanol from a methyl-tert-butyl ether (MTBE)-butene-methanol mixture involved in the production of MTBE. The DPSU can overcome the azeotropic composition and the distillation boundaries. The hybrid distillation pervaporation process in a single unit (DPSU) is compared with a conventional hybrid process where the pervaporation unit is externally connected to the distillation column (DPEC). The performance of the DPSU has been calculated using a modified rate-based model where the interface between the liquid and the membrane has been included. It has been found that a hybrid DPSU is more convenient for the removal of one component from a multicomponent mixture when the component can not be obtained from the distillation column as a pure distillate or bottom product. Thus, for purification of binary mixtures a DPEC process is more convenient while the DPSU is recommended for the removal of methanol from the multicomponent mixture. For the multicomponent system, the difference in performance between a DPEC and a DPSU lies in the higher methanol concentration that occurs in the pervaporation section in a DPSU.

In general, two-phase flow pervaporation units can expand the application window

of pervaporation in stand alone and hybrid applications. Industrial applications are already viable using the multi-tubular pervaporation modules available in the market. The development of pervaporation modules with hollow fibre membranes will take some more time and it requires more information about the mechanical resistance and the wetting properties, since the wetted fraction will probably affect more strongly the energy consumption of the separation process than the required membrane area for a specific separation duty. A distillation-pervaporation process in one single column opens the possibility to separate multicomponent mixtures in one unit. However, it is necessary to explore the dynamic behavior of these systems and the existence of multiple steady-states which are important for industrial operation and start up.

Samenvatting Pervaporatie gecombineerd met destillatie wordt met succes toegepast voor de

scheiding van mengsels. Deze combinatie is zeer geschikt voor de scheiding van azeotropische mengsels of voor mengsels met een lage relatieve vluchtigheid waarbij het energieverbruik hoog is bij toepassing van uitsluitend destillatie. Pervaporatie is dan aantrekkelijk omdat de energie benodigd voor de scheiding sterk verlaagd wordt doordat alleen de permeërende component verdampt. De investeringskosten voor pervaporatie zijn echter hoger dan voor destillatieprocessen. In een hybride proces kunnen de voordelen van beide technieken gecombineerd worden terwijl de nadelen zoveel mogelijk vermeden worden.

In pervaporatieprocessen kunnen concentratie- en temperatuurpolarisatie

optreden, evenals een daling van de temperatuur aan de retentaatzijde. Hierdoor wordt het benodigde membraanoppervlak en de benodigde randapparatuur vergroot en daardoor de totale kosten voor een gegeven scheiding. In dit proefschrift wordt aangegeven hoe deze nadelen van pervaporatieprocessen verminderd kunnen worden door toepassing van een gas-vloeistof of een damp-vloeistof mengsel als voeding voor de pervaporatiemodule. Tevens wordt aangegeven hoe deze werkwijze het ontwerp en de werking van hybride processen beïnvloedt. Daarbij zijn zowel conventionele hybride processen in beschouwing genomen, waarbij de pervaporatie-eenheid extern verbonden is met een destillatiekolom als processen waarin destillatie en pervaporatie in een enkele kolom gecombineerd worden.

Holle vezel silica membranen zijn gekarakteriseerd door gebruik te maken van SEM,

SNMS, permeatie van zuivere gassen en door pervaporatie van dimethylformamide (DMF) – water mengsels. Verouderingsexperimenten aan buisvormige silicamembranen tonen aan dat de waterflux, bij de verwijdering van water uit DMF-water en alcohol-water mengsels, sterk vermindert in de tijd. Deze afname van de flux wordt toegeschreven aan de interactie van de organische component in het mengsel met de silicalaag via absorptie en reactie met silanolgroepen. Experimenteel is aangetoond dat deze interactie in elk geval gedeeltelijk omkeerbaar is.

De hydrodynamica en de flux in een pervaporatiemodule zijn bestudeerd met

behulp van CFD en ultrasone computertomografie als functie van de relatieve dichtheid van de permeërende en de niet-permeërende component. De berekende dichtheids- en

temperatuurprofielen zijn vergelijkbaar met de profielen gereconstrueerd uit de metingen. Op het oppervlak van de membranen wordt een inversiepunt gevonden waarboven de vloeistof naar boven beweegt en daaronder naar beneden. Dit inversiepunt geeft aan dat voor de ontwatering van een mengsel met een lage relatieve dichtheid van de te scheiden componenten, een meerpijps pervaporatiemodule met een driehoekige pijpsteek het meest geschikt is. Voor ontwatering van mengsels met een grote relatieve dichtheid wordt een vierkante pijpsteek aanbevolen.

De werking van buisvormige pervaporatiemembranen is bestudeerd waaraan

inwendig, in gelijkstroom, een twee-fasenstroming wordt toegevoerd. Daarbij is gebruik gemaakt van een opstelling op laboratoriumschaal voor de ontwatering van 1,4-dioxaan – water mengsels en van een opstelling op semi-technische schaal voor de ontwatering van isopropylalcohol (IPA) - water mengsels. Kleine hoeveelheden lucht of damp vergroten de turbulentie in de vloeistoffase waardoor concentratie- en temperatuurpolarisatie verminderd worden. De flux en de selectiviteit nemen daardoor toe met een factor twee vergeleken met laminaire, één-fase, vloeistofstroming. Wanneer gebruik gemaakt wordt van damp wordt door condensatie energie aan de vloeistof geleverd. Hierdoor neemt het vereiste membraanoppervlak af en zijn er geen warmtewisselaars benodigd tussen de pervaporatiemodules. Een goede berekening van de totale flux door het membraan, voor lucht-vloeistofstroming in slug flow, is verkregen met modelberekeningen gebaseerd op experimenteel vastgestelde belgrootten, lengtes van de vloeistofslug en stijgsnelheid van de bellen. De werking van een twee-fasen pervaporatiemodule is vergelijkbaar met de werking van een één -fase module onder turbulente omstandigheden.

Daarnaast zijn pervaporatiemodules onderzocht opgebouwd uit holle vezels die aan

de buitenzijde worden omstroomd met een twee-fasen mengsel. De gasfractie en de stofoverdrachtscoëfficiënt aan het membraanoppervlak zijn gemeten in een transversale opstelling met een elektrochemische methode. De stoverdrachtscoëfficiënt kan tot viermaal hoger zijn dan bij één-fase stroming. De experimenteel bepaalde dimensieloze stofoverdrachtsrelatie is opgenomen in een model waarmee een pervaporatiemodule is ontworpen voor het ontwateren van een IPA-water mengsel. Het benodigde membraanoppervlak kan met ongeveer 45% verkleind worden vergeleken met pervaporatiemodules met een economisch optimaal aantal tussenwarmtewisselaars. Dit kleinere oppervlak, zonder tussenwarmtewisselaars, veroorzaakt een sterke daling van de investeringskosten.

Voor conventionele hybride processen, bestaande uit een membraanscheiding die

extern gekoppeld is aan een destillatiekolom, wordt aangegeven of pervaporatie dan wel damppermeatie het meest geschikt is voor een specifieke toepassing. Het positieve

effect van relatief laag-selectieve membranen op de totale proceskosten wordt aangetoond voor de ontwatering van acetonitril (ACN) – water mengsels. Ook de terugvoer van een gedeelte van de retentaatstroom naar de permeaatzijde leidt tot een sterke vermindering van het benodigde membraanoppervlak en dus van de totale scheidingskosten. Voor een mengsel waarvan de ligging van de azeotroop drukafhankelijk is, zoals ACN-water, blijkt het effectief te zijn om een relatief lage druk in de destillatiekolom toe te passen en een hogere druk in de pervaporatie-eenheid. Vergeleken met destillatieprocessen kan met hybride processen 25 tot 60% op de scheidingskosten worden bespaard.

Naast conventionele hybride processen zijn processen geanalyseerd waarin de

pervaporatie-eenheid in een twee-fasen systeem wordt gebruikt, door deze eenheid te integreren met een destillatieproces binnen een enkele kolom (DPSU). De pervaporatiezone, bestaande uit holle vezel membranen, vervangt daarbij een sectie pakking of schotels in een destillatiekolom. Dit hybride systeem is bestudeerd voor de ontwatering van binaire mengsels zoals ethyleendiamine-water en IPA-water en voor de verwijdering van methanol uit een mengsel van methyl-tert.butylether (MTBE)- buteen – water zoals aanwezig bij de productie van MTBE. Met de DPSU kunnen producten verkregen worden die zuiverder zijn dan met destillatie vanwege de azeotroop of lage relatieve vluchtigheden. Hybride scheidingen in een enkele kolom (DPSU) zijn vergeleken met conventionele hybride processen waarbij de pervaporatie-eenheid extern met de destillatiekolom is verbonden (DPEC). De werking van de DPSU is berekend met een op stofoverdracht gebaseerd model, waarin het grensvlak tussen vloeistof en membraan is opgenomen. Het blijkt dat een hybride DPSU beter geschikt is voor de verwijdering van een component uit een multicomponenten-mengsel wanneer deze component niet als een zuiver top- of bodemproduct uit de destillatiekolom verkregen kan worden. Voor de zuivering van de binaire mengsels is een DPEC meer geschikt en een DPSU wordt aanbevolen voor de verwijdering van methanol uit het multicomponenten-mengsel. Voor het multicomponentensysteem is het verschil tussen een DPEC en een DPSU een hogere methanolconcentratie die optreedt in de pervaporatiesectie van de DPSU.

In het algemeen vergroot het toepassen van twee-fasenstroming het

toepassingsgebied van pervaporatie zowel voor stand-alone als hybride toepassingen. Met de meerpijps membraanmodules die momenteel commercieel verkrijgbaar zijn kan dit principe al industrieel uitvoerbaar worden toegepast. De ontwikkeling van pervaporatiemodules gebaseerd op holle vezel membranen vergt nog tijd. Bij toepassing van deze membranen is informatie benodigd over de drukval en de bevochtigingseigenschappen omdat deze van meer invloed kunnen zijn op het energieverbruik dan op het benodigd membraanoppervlak. Destillatie-pervaporatie

gecombineerd in één enkele kolom geeft de mogelijkheid om multicomponenten-mengsels in één apparaat te scheiden. Het dynamische gedrag en het optreden van meerdere werkpunten dient nog nader bestudeerd te worden omdat deze van belang zijn bij het opstarten en het bedrijven van industriële installaties.

1. Introduction and outline of the thesis.

Chapter 1

2

1.1. Introduction Distillation is the most widely used technique to separate liquid mixtures. However,

distillation separation of mixtures with an azeotropic composition or with components with low relative volatility is energetically expensive and auxiliary substances are usually required.

Separation by pervaporation depends on the difference in partial vapor pressure

between the two sides of a membrane and the selective sorption properties of the membrane with respect to the components in the mixture1. The pressure difference is created by applying a lower pressure at the permeate side. Because the separation is not driven by the liquid-vapor equilibrium, separation of azeotropic mixtures is also feasible.

Distillation is a well-known technique with lower capital cost than pervaporation.

However, the energy consumption in pervaporation is lower because it is required only for the vaporization and expansion of the compounds that selectively have been transported through the membrane. This energy is removed from the sensible heat carried by the liquid, inducing a drop in the retentate temperature and, consequently, in the flux2. The retentate temperature drop increases the required membrane area for a specific removal duty. Usually, auxiliary equipment like heat exchangers is necessary. Capital cost in pervaporation is high due to the cost of the membranes, the modules and the auxiliary equipment.

A hybrid process exploits the advantages of pervaporation and distillation while the

negative aspects are minimized. Several processes have been presented in the literature and are applied in the industry such as for the dehydration of alcohols, aprotic solvents, and esters, as well as for the removal of VOCs from aqueous streams3.

The role and application range of pervaporation in stand-alone applications and in

hybrid processes can be expanded if the involved capital cost of the pervaporation unit is reduced. The stability of the pervaporation membranes, the concentration and temperature polarization and the temperature drop that occurs in the liquid are factors that increase the required membrane area, the amount of auxiliary equipment and the related capital and operating cost. In the next sections some of these drawbacks and also some characteristics of hybrid distillation-pervaporation processes are described.

Introduction and outline of the thesis.

3

1.2. Concentration and temperature polarization Due to the separation process achieved by a pervaporation membrane a

concentration gradient is created in a boundary layer on the membrane surface. Within this boundary layer the concentration of the permeating component decreases. This phenomenon is called concentration polarization4. As a result the concentration of the components that are retained will increase within the boundary layer. A sketch of a boundary layer and the occurring concentration profiles is presented in Figure 1.1. Concentration polarization causes components that are enriched in the permeate to be depleted in the boundary layer, and components that are depleted in the permeate to be enriched in the boundary layer. Thus, concentration polarization works against the separation achieved by the membrane, reducing flux and selectivity.

Figure 1.1. Concentration (C) and temperature (T) polarization on a pervaporation membrane. “p” and “np” indicate the permeating and non-permeating component, respectively. Subscripts “b” and “s” are locations in the liquid at the bulk and the membrane surface. The superscript p indicates the permeate side. δ is the thickness of the mass (m) and heat (h) boundary layers.

Unlike other membrane processes, also temperature polarization occurs in pervaporation processes5. The energy transported through the membrane is high due to the vaporization and expansion of the permeating components2. This energy consumption results in a resistance to the heat transport in the boundary layer (Figure 1.1). The lower temperature obtained on the membrane surface as compared to the bulk reduces the driving force for the mass transport. Due to the lower temperature on the membrane surface, changes in the intrinsic properties of the membrane in relation to the

Chapter 1

4

permeating components may also occur, determined by the kind of material of the membrane.

The significant effect of concentration polarization in pervaporation processes is

presented in Figure 1.2. The concentration polarization modulus defined as the ratio of concentrations between the membrane surface and the bulk is presented as a function of the Peclet number for mass transport of the component that preferentially permeates. Several curves are shown, each corresponding to a constant value of the enrichment factor, which is defined as the ratio of concentrations between the permeate and the retentate on the membrane surface. The region where pervaporation processes operates is shown, based on typical values of Peclet number and enrichment factor. As a worst case scenario, the concentration on the membrane surface can be until 100 times lower than in the bulk.

Figure 1.2. Concentration polarization modulus as a function of the Peclet number at several enrichment factors. Gray areas correspond to ranges of values usually obtained in pervaporation processes4.

If the mass transfer coefficients or hydrodynamic conditions of the current pervaporation modules are kept constant, the development of membranes with higher flux and selectivity will increase polarization effects, as both the Peclet number and enrichment factor increase as flux and selectivity increase. High mass transfer coefficients, and as a consequence lower Peclet numbers, have commonly been achieved by using high flow rates, however, the pressure drop and the involved energy consumption is also high. Moreover, high flow rates are not convenient in processes where long residence times are required, e.g. in membrane reactors or for the removal of traces of a compound.


5

1.3. Temperature drop The evaporation of the permeating components is a fundamental step in

pervaporation processes. A heat flux, which is taken from the liquid, is necessary for the phase change between the retentate and the permeate. Thus, temperature gradients develop perpendicular to the membrane surface (temperature polarization) as well as in the direction of the flow6. This temperature drop reduces the driving force for mass transport and modifies the intrinsic properties of the membrane with respect to the permeating components. Consequently, a lower flux and usually lower selectivities are obtained, increasing the required membrane area for a given separation duty.

The influence of the temperature drop on the driving force for water is depicted in

Figure 1.3 for several organic-water mixtures. In the figure the water driving force is defined as the ratio of the driving force at a given temperature to the driving force at the bubble temperature. For alcohols, a lost in the driving force of around 10% can be expected for temperatures drops of only 2°C, a value that can easily be obtained as a consequence of temperature polarization7.

Figure 1.3. Effect of the temperature drop on the water driving force, relative to the driving force at saturated conditions at 1.013 bar. Organic mixtures containing 5 wt. % water. Dimethyl formamide (DMF), Ethylene diamine (EDA), Acetonitrile (ACN) and isopropyl alcohol (IPA).

The temperature drop in the liquid is conventionally controlled by using a series of

alternating heat exchangers and pervaporation modules7 as show Figure 1.4. In each heat exchanger the liquid is heated to the feed temperature. An increasing number of

Chapter 1

6

heat exchangers reduce the temperature drop per pervaporation module. If few heat exchangers are used the required membrane area increases, while increasing the number of heat exchanger reduces the membrane area required but the cost of auxiliary equipment rises. Thus, an economical optimal exist8. At this optimum the temperature drop per module is between 6 and 10°C. According to Figure 1.3, this temperature drop per module reduces the driving force for water transport 30 to 40% in alcohol-water mixtures. As a consequence the required membrane area strongly increases.

Figure 1.4. Pervaporation unit consisting of pervaporation modules and heat exchangers connected in series.

Various designs have been presented in the literature to supply energy directly into the pervaporation module, for instance using electrical resistances9 or external heating with steam or another fluid10. Figure 1.5 shows the required membrane area for reducing the water concentration in a DMF-water mixture. Several curves are shown as a function of the heat supply to the liquid. The heat supply corresponds in the figure to a fraction of the energy required for an isothermal operation. As it can be seen from the figure, heat supply strongly increases the performance of the pervaporation unit. At constant membrane area, an isothermal operation can remove about twice the amount of water of an adiabatic operation.


7

Figure 1.5. Membrane area required to reduce the water concentration starting from 5 wt. % for a DMF-water mixture as a function of the fractional heat supply.

1.4. Hybrid distillation-pervaporation processes In most of the hybrid configurations distillation is more economical for the bulk of the

separation while the membrane is used to perform the part of the separation where distillation is difficult or impossible. Several kind of hybrid configurations have been studied by Stephan et al.11, Pettersen et al.12, Pressly et al.13 and more recently based on a formal mathematical methodology by Kookos14. In general the pervaporation module can be used in a hybrid distillation system to remove a specific component from a lateral stream of the distillation column, e.g. to overcome the azeotropic composition or as a final treatment stage.

Some basic hybrid configurations are presented in Figure 1.6. The first configuration

(Figure 1.6a) is used for systems with minimum-boiling azeotrope or when a large number of trays are required in the rectifying section. Similarly, the membrane can be placed in the bottom stream for systems with a maximum-boiling azeotrope or when a considerable amount of trays are required in the stripping section. Figure 1.6b is used for systems with an azeotrope at an intermediate concentration. The last system shown in the figure is used for systems with a low relative volatility in the whole range of concentrations where the number of trays and the reflux ratio can significantly be reduced11.

Chapter 1

8

Figure 1.6. Basic configuration for hybrid distillation-pervaporation processes.

1.5. Scope of this thesis This thesis explores alternatives to improve the performance of hybrid distillation -

pervaporation processes by reducing the required membrane area and the amount of auxiliary equipment by optimizing the interaction between the pervaporation and the distillation operations. These aims have been achieved by studying the effect of the hydrodynamics in multi-tubular pervaporation modules, operating in single and multi-phase flow, on the concentration and temperature polarization and on the flux: Also the influence of the membrane properties and process variables on the performance of hybrid systems has been studied here, both in single and multi-phase flow. Multi-phase flow is referred to a retentate that consists of liquid and vapor or liquid and gas.

The results from this research show that a liquid-vapor feed stream can reduce

concentration and temperature polarization while at the same time the vapor supplies the energy required for the pervaporation process. Flux and selectivity are increased reducing the required membrane area and the amount of auxiliary equipment, e.g. inter-stage heat exchangers.

The information presented in the following chapters leads to the design of a hybrid

process that combines a distillation column with a pervaporation module operating in liquid-vapor flow. The role of the pervaporation unit in the hybrid process can be


9

conceived not only as an externally connected unit to the distillation column but also, for some applications, as a distillation column and a pervaporation module operating within one single unit.

1.6. Outline of this thesis In the next chapter the stability and performance of silica membranes is studied in

pervaporation and gas separation. Stability tests on hollow fibres and tubular membranes by pervaporation have been performed using several solvents and some characteristics of the interaction between the liquid and the silica layer are addressed. Due to the strong interactions between the alcohol and the silica layer, the water flux decreases with time in alcohol-water mixtures. However, it is shown experimentally that the negative interactions are partially reversible. These hollow fibre and tubular membranes are used in the subsequent chapters.

The influence of the hydrodynamics on the design of multi-tubular pervaporation

modules is presented in Chapter 3. CFD simulations have been carried out with an experimental verification of the results by comparing density and temperature profiles inside a pervaporation module. Ultrasound computer tomography, a non-intrusive technique, has been used to measure these profiles. Recommendations for the design of multi-tubular pervaporation modules are given based on the density ratio between the components in the mixture. The flux in the pervaporation modules is hindered by concentration and temperature polarization especially when long residence times are required. The polarization effects in multi-tubular pervaporation modules can be reduced by a two-phase flow and this is the core topic of the next chapters.

In this thesis it has been found that injecting a liquid – vapor or liquid-gas feed in the

pervaporation unit strongly improves the performance of pervaporation modules. This alternative is studied experimentally in Chapter 4 in a lab scale and a bench scale setup with commercially available tubular membranes. Vapor and air reduce concentration and temperature polarization. Simultaneously, vapor supplies energy to the liquid by condensation, thus reducing the liquid temperature drop that affects the performance of pervaporation modules in single phase. For laminar flow conditions the flux is close to the one in the turbulent regime.

Unlike the tubular membranes used in the previous chapter, hollow fibre membranes

offer higher packing densities and lower pressure drops. The application and design of pervaporation modules with hollow fibre membranes in two-phase flow are studied in chapter 5. A design of a two-phase pervaporation module using hollow fibre membranes

Chapter 1

10

is presented. The design is based on experimental data of mass transfer in the two-phase regime with a lab scale module. Dehydration of isopropyl alcohol is shown as an example. Also, a comparison with the performance of a pervaporation unit in single phase flow is shown.

The information collected in the previous chapters leads to the question: Is it

convenient to have a hybrid process where the liquid-vapor pervaporation module is combined with the distillation operation within one single column? In order to arrive to an answer, Chapter 6 studies hybrid distillation processes where a pervaporation unit is externally connected and Chapter 7 analyses advantages and disadvantages of hybrid distillation-pervaporation operations within a single column.

Chapter 6 studies conventional or single-phase pervaporation modules in

combination with distillation. The conclusions of this chapter also apply for hybrid systems with externally connected pervaporation modules operating in two-phase flow. Some guidelines are presented to design and select whether pervaporation or an alternative technique like vapor permeation is more convenient for a given application. The important effects of membrane selectivity and product sweep in the performance of the hybrid process are shown. The results show that relatively low membrane selectivities or high product sweeps are required for an economical optimal hybrid process.

Chapter 7 discusses advantages, disadvantages, design and performance of a

distillation column and a two-phase pervaporation module within one single unit for binary mixtures and multicomponent mixtures. The dehydration of ethylene diamine-water and isopropyl alcohol-water mixtures is shown as examples of binary systems. The purification of methyl tert butyl ether (MTBE) - butene – methanol, which is an intermediate stream for the production of MTBE, is presented as example of a multicomponent mixture. The conventional hybrid process, where the membrane is externally connected to the distillation column is more convenient for the separation of binary mixtures. The single hybrid unit is more convenient for the removal of methanol from the multicomponent mixture.


11

1.7. Reference List (1) Huang, R. Y. M. Pervaporation membrane separation processes; Elsevier: New york,

1991.

(2) Karlsson, H. O. E.; Tragardh, G. Heat transfer in pervaporation. J. Membr. Sci. 10-16-1996, 119, 295 - 306

(3) Lipnizki, F.; Field, R. W.; Ten, P. K. Pervaporation based hybrid process: a review of process design, applications and economics. J. Membr. Sci. 1999, 153, 183 - 210

(4) Wijmans, J. G.; Athayde, A. L.; Daniels, R.; Ly, J. H.; Kamaruddin, H. D.; Pinnau, I. The role of boundary layers in the removal of volatile organic compounds from water by pervaporation. J. Membr. Sci. 1-10-1996, 109, 135 - 146

(5) Favre, E. Temperature polarization in Pervaporation. Desalination. 2003, 154,129-138

(6) Ito, A.; Feng, Y.; Sasaki, H. Temperature drop of feed liquid during pervaporation. J. Membr. Sci. 1997, 133, 95 - 102

(7) Rautenbach, R.; Albrecht, R. The separation Potential of pervaporation. Part 2. Process design and economics. J. Membr. Sci. 1985, 25, 25 – 54

(8) Bausa, J.; Marquardt, W. Shortcut design methods for hybrid membranes/distillation processes for the separation of nonideal multicomponent mixtures. Ind. Eng. Chem. Res. 2000, 39, 1658 - 1672

(9) Hilgendorff, W., Wenzlaff, A., Böddeker, K., Kahn, G., and Lührs, G. Einrichtung zur Trennung von Lösungen durch Pervaporation. 1984.

(10) Schleger, M.; Sommer, S.; Melin, T. Module arrangement for solvent dehydration with silica membranes. Desalination. 2004, 163, 281 - 286

(11) Stephan, W.; Noble, R. D.; Koval, C. A. Design methodology for a membrane distillation column hybrid process. J. Membr. Sci. 1995, 99, 259 - 272

(12) Pettersen, T.; Argo, A.; Noble, R. D.; Koval, C. A. Design of combined membrane and distillation processes. Sep. Tech. 1996, 6, 175 - 187

(13) Pressly, T. G.; Ng, K. M. A break - Even analysis of distillation-membrane hybrids. AIChE J. 1998, 44, 93 - 105

(14) Kookos, I. K. Optimal design of membrane distillation column hybrid processes. Ind. Eng. Chem. Res. 2003, 42, 1731 - 1738

Chapter 1

12

________________________________ This chapter is partly based on: T. A. Peters, J. Fontalvo, M. A. G. Vorstman, N. E. Benes, R. A. Van Dam, Z. A. E. P. Vroon, E. L. J. Van Soest-Vercammen, and J. T. F. Keurentjes. J. Membr. Sci. 248 (2005) 73-80.

2. Synthesis, performance and stability of silica membranes for gas

permeation and pervaporation

Abstract

Thin microporous silica membranes were prepared on the outer surface of hollow fibre ceramic substrates. The membranes were analyzed using SEM, SNMS, single gas permeance and pervaporation. High He permeance (1.1-2.9 x 10-6 mol /m2 s Pa), high He/N2 permselectivity (~ 100-1000) and Arrhenius type temperature dependence of gas permeance indicate that the membranes are microporous and posses a low number of defects. In the dehydration of dimethylformamide (DMF) in a tubular and the hollow fibre membranes initially high flux and selectivity were observed. Subsequently, pervaporation performance decreased with time, likely due to interactions of water and DMF with the silica material, i.e., adsorption on and reaction with silanol groups. The strong interactions of the DMF molecules with the silica result in a rather low water flux and selectivity. For the sequential dehydration of ethanol-water, methanol-water and 1,4-dioxane-water mixtures in a tubular membrane also a drop in water permeance was observed for the alcohol mixtures. However, the water permeance was partially recovered when the membrane was in contact with the 1,4-dioxane-water mixture.

Chapter 2

14

2.1. Introduction Compared to their organic counterparts inorganic membrane materials generally

possess superior structural stability, e.g., no swelling and compaction, even in harsh chemical environments and at high temperatures1-5. The majority of inorganic membranes are porous and their selective features are often closely related to their pore size. Amorphous silica is an inorganic material containing exceptionally small pores. Membranes based on this material have an asymmetric structure with the actual selective micro-porous silica positioned on a support structure comprising several α- and γ- alumina layers. Silica membranes were discovered more than a decade ago6-8 and are still subject of extensive study.

Silica membranes reported in literature have either a flat plate or tubular geometry.

The flat plate geometry is advantageous from an academic point of view, but it usually has a small surface area (typically ~ 10-2 m2) due to limitations imposed on the dimensions by the dip-coating technique. The surface area of tubular silica membranes is larger and their geometry is also more compatible with the technology developed in organic membrane science. Consequently, commercially available membranes for pervaporation have a tubular geometry3. Drawbacks of this geometry include a relatively low surface area-to-volume ratio (typically < 500 m2/m3) and high costs associated with tubular ceramic membrane supports.

In this work silica layers are positioned on top of ceramic hollow fibres. In principle

this enables the relatively rapid and inexpensive preparation of large membrane surface area, combined with a high membrane surface-area-to-volume ratio (> 500 m2/m3). Additionally, pervaporation modules prepared with externally coated hollow fibres offer high mass and heat transfer coefficients at expense of low liquid pressure drops. The membranes were analyzed using SEM, SNMS, single gas permeance and pervaporation. Stability measurements have been performed in two tubular membranes for dehydration of DMF in one of the tubes and in the second one for the sequential dehydration of ethanol-water, methanol-water and 1,4-dioxane-water mixtures.

2.2. Theory

2.2.1. Gas permeation Numerous theories for describing transport in microporous media have been

presented in literature9-13. These theories become increasingly complex when the microporous medium is less uniform and when more mobile species are present. A simple

Synthesis, performance and stability of silica membranes for gas permeation and pervaporation

15

phenomenological approach is sufficient in this study for the assessment of membrane quality.

For single gas permeation of permanent gases through amorphous microporous

silica membranes, at sufficiently high temperatures and low pressures, transport is activated and permeance is independent of pressure6-8,14,15. Hence, permeance is described by:

( ) Do o

Q ENP H D expp RT∆

− ≡ =

2.1

where N is the molar flux, oH and oD are pre-exponential factors related to the Henry and

diffusion coefficients, respectively, and R and T have their usual meaning. The overall thermally activated nature of transport arises from the simultaneous occurrence of diffusion (ED) and sorption (Q ).

2.2.2. Pervaporation For dehydration of solvents by pervaporation the performance of a membrane is

usually expressed in terms of water flux, or permeance, and selectivityα . The latter is defined as:

water

water

j

j

yx

yx

α = 2.2

where y and x are the molar fractions in the permeate and retentate streams,

respectively. Permeance ( Γ) is defined as the flux divided by the partial pressure difference over the membrane. The partial pressure of component i at the retentate side

is related to the mole fraction x and activity coefficient iγ in the liquid mixture

* oi i i ip x pγ= 2.3

where and oip is the vapor pressure of the pure component i . When the pressure at the

permeate side is small compared to the equilibrium vapor pressure at the retentate side, permeance can be expressed as2:

Chapter 2

16

oi

wi i i

Nx p

Γγ

= 2.4

2.3. Experimental

2.3.1. Hollow fibre membrane preparation

2.3.1.1. Support

The ceramic hollow fibres membrane supports (CEPAration B.V., The Netherlands) have a porosity of ~30%, pore diameter of either 150 or 300 nm, length in the range of 20-30 cm, and inner and outer diameter of 2.0 and 3.0 mm, respectively.

2.3.1.2. γ-Al2O3 intermediate support preparation

On top of the substrates intermediate mesoporous γ-Al2O3 layers were prepared by sequential dip-coating with a boehmite coating solution. The boehmite solution was made by adding aluminium-tri-sec-butoxide (Aldrich) drop-wise to water at 90 °C under vigorous stirring, and subsequent boiling for 90 minutes to remove the 2-butanol produced during the hydrolysis. A white solution was obtained, which was peptized with 1 mol/L HNO3 (water/alkoxide/acid ratio: 70/1/0.07). The peptization was accompanied by a change in color from white to “nano” blue. After refluxing for 16 hours the resulting solution had a pH of 3.8. Finally, 120 mL polyvinyl alcohol (PVA) solution was added to 180 mL boehmite solution, followed by stirring at room temperature for 30 minutes and subsequently stirring at 90°C for 150 minutes. The PVA solution was prepared by dissolving 8.75 gram PVA (Aldrich, PVA Powder, average Mw 89-98 kD, hydrolysis grade 98%) in 250 ml of 0.05 M HNO3. The dip-coat process was performed at room temperature in a laminar flow cupboard (Interflow, quality class 100) to minimize dust contamination. The substrate speed was 10 mm/s and the dip-time was 25 seconds. The membranes were dried in a climate chamber (Espec 100) at 40 ºC and 60 RH % for at least 120 minutes. After drying, the membranes were sintered at 600 ºC for 180 minutes (heating rate 1 ºC/min). The procedure for dipping, drying and sintering was repeated three times in order to obtain defect-free intermediate γ-Al2O3 membranes.

2.3.1.3. Silica separation layer preparation

The intermediate γ-Al2O3 layers were modified by dip-coating with a polymeric silica sol, prepared via acid catalyzed hydrolysis and subsequent polycondensation of tetraethylorthosilicate (TEOS) (Aldrich, >99%). The polymeric silica solution was prepared by


17

drop-wise addition of water and HNO3 to a TEOS/ethanol solution under vigorous stirring (water/TEOS/acid/ethanol: 6.4/1/0.085/3.8) and refluxing at 60 °C for 180 minutes. The resulting solution was diluted 18-fold with ethanol. Dip-coating was performed in a laminar flow cupboard (Interflow, quality class 100). After dip-coating (substrate speed 10 mm/s, dip-time 5 seconds) the membranes were dried for 30 minutes at 40 °C and 60 RH %. After drying, the membranes were sintered at 350-600 °C for 180 minutes (heating rate 0.5 ºC/min).

2.3.2. Tubular membranes The two tubular pervaporation membranes supplied by Pervatech (The Netherlands)

consisted of an α-alumina support tube, with 7 mm internal diameter and 50 cm length, and a γ-alumina intermediate layer on the internal face of the tube. The silica layer was placed on top of the intermediate layer.

2.3.3. Membrane characterization The thickness and morphology of the different layers of the hollow fibre membranes

were studied by Scanning Electron Microscopy (SEM) using a JEOL 840 microscope. Samples were sputtered with a thin layer of gold. Independently, the thickness of the silica layers was determined by Secondary Neutral Mass Spectrometry (SNMS).

2.3.4. Gas permeation Single gas permeance of a single fibre was measured in a pressure controlled dead-

end set-up. Viton O-rings were used for sealing, limiting the operating temperature to 210 °C. The effective membrane length was 17 cm leading to an effective membrane area of 0.0017 m2. Prior to measurements, the membranes were pre-treated by permeating He at 180 °C for two days. Measurements were performed with He and N2 (>99% pure) at temperatures ranging from 25 °C to 200 °C. The pressure at the permeate side (tube side) was kept slightly above ambient pressure, while the pressure at the feed side (shell-side) was varied in the range 150 – 350 kPa. The flow required to maintain the pressure drop over the membrane was measured by a mass flow indicator.

2.4. Pervaporation The organic solvents used in dehydration experiments were ethanol, methanol, 1,4-

dioxane (pro-analyze, Merck) and dimethylformamide (DMF) (>99.8%, Merck). A typical set-up was used for pervaporation2. A single hollow fibre membrane or a tubular membrane was placed in a tubular module and feed was supplied by a continuous

Chapter 2

18

recycle. The liquid superficial velocity in the membrane module exceeded 3 m/s in order to eliminate polarization effects. On the permeate side vacuum was maintained at 10 mbar by a cascade of liquid nitrogen cold traps and a vacuum pump. For some measurements overnight stops were done during which time the membranes remained in the feed liquid without vacuum applied on the permeate side at the measurement temperature. Subsequent start of an experiment was preceded by a stabilization period of one hour, during which vacuum was applied on the permeate side. For experiments without overnight stops vacuum was applied continuously and the liquid temperature was kept at the measurement temperature. Dehydration of DMF-water mixtures was performed with the hollow fibre membranes and one of the two tubular membranes.

Stability measurements were performed with the second tubular membrane. First the

fresh tubular membrane was tested for dehydration of an ethanol-water mixture for a period of 100 h with vacuum and 100 h without vacuum. Subsequently, the mixture was replaced by a methanol-water mixture and the membrane was tested for 1000 h at four temperatures: 43, 33, 53 and 43 °C, sequentially. Finally, dehydration experiments were performed with a 1,4-dioxane-water mixture for 150 h.

For dehydration of DMF the retentate and permeate compositions were analyzed

using an automated Karl-Fischer titration apparatus (Mettler Toledo DL50 Graphix). For alcohols and 1,4-dioxane the retentate was analyzed using Karl-Fischer, while the permeate composition was analyzed using gas chromatography. Activity coefficients of water with either alcohol or 1,4-dioxane were approximated using the Wilson equation16, for DMF the activity coefficient was estimated using the NRTL model16. Vapor pressures were calculated using the Antoine equation16.

2.5. Results

2.5.1. Membrane characterization In Figure 2.1 SEM-pictures of a hollow fibre substrate (a), coated with intermediate γ-

Al2O3 layers (b), and coated with silica (c) are shown. Figure 2.1b shows that the four γ-Al2O3 layers form a single 3-4 µm thick homogeneous layer on the substrate, providing a sufficiently smooth surface for silica to be deposited on. For thinner intermediate layers the preparation of defect free silica layer appeared unsuccessful. The particle size in the intermediate layer is in the range 30-80 nm (Figure 2.1c), which is comparable to the thickness of the silica layer (20-60 nm). The SNMS depth profile of the silica membrane (Figure 2.2) shows a monotone decrease in Si concentration from the membrane surface (0 nm) towards the intermediate γ-Al2O3 layer. Correspondingly, the Al concentration


19

increases with depth and at approximately 20 nm reaches a practically constant value. The crossover point is at 5 nm.

Figure 2.1. SEM micrographs of the various layers constituting the hollow fibre membranes: (a) cross- section substrate (magnification 100x), (b) cross-section intermediate γ-Al2O3 layers on top of the substrate (magnification 3300x), (c) cross-section silica membrane on the intermediate γ-Al2O3 layers (magnification 75000x).

γ-Al2O3 intermediate layer

SiO2 layer

(a) (b)

(c)

Chapter 2

20

Figure 2.2. SNMS depth profile of a silica membrane from batch TNO-3.

The SNMS profiles indicate that the silica layer on top of the γ-Al2O3 layer has a thickness of at least 20 nm, which is in reasonable agreement with the thickness determined by SEM (20-60 nm). The gradual decrease in Si concentration with depth may be due to the surface roughness of the γ-Al2O3 and partial penetration of silica into this layer.

2.5.2. Gas permeation Single gas permeation was performed using helium and nitrogen at several trans-

membrane pressures. For both gases, permeance appears to be independent of pressure. At a temperature of 200 °C helium and nitrogen flux were 2.2 x 10-6 mol/m2 s Pa and 4.0 x 10-9 mol/m2 s Pa, respectively. The high helium permeance and the high permselectivity of this gas with respect to the larger nitrogen are typical for microporous silica membranes. The low permeance of nitrogen suggests that the pore size of the silica is similar to the molecular dimensions of N2 (kinetic diameter 3.64 Å).

For both helium and nitrogen excellent linear fits of the Arrhenius plots were obtained

(Figure 2.3) confirming that transport is thermally activated. The increase in helium permeance with temperature (apparent energy of activation is -6.6 kJ/mol) indicates that for this inert gas the temperature dependence of diffusion predominates that of sorption. For nitrogen a decrease in permeance is observed with temperature (apparent energy of activation is 3.3 kJ/mol). The less pronounced change with temperature


21

indicates that for nitrogen the temperature dependence of sorption predominates that of diffusion, albeit only slightly.

Figure 2.3. Arrhenius plots of helium and nitrogen with trans-membrane pressure difference 100 kPa on membrane TNO-8a.

Table 2.1 contains permeance data for several batches of silica membranes. Every

batch of membranes consists of 20 membranes from which at least 5 membranes were tested. In total 42 membranes were tested out of which 35 show a helium/nitrogen permselectivity exceeding 150. High values are observed for the permeance of helium (1.1 - 2.9 x 10-6 mol/m2 s Pa, 200 ºC). Differences in helium/nitrogen permselectivity have both been found between the seven batches likely due to differences in silica dipcoat solution and in one batch probably due to differences in substrate quality.

Table 2.1. Helium and nitrogen single gas permeance and He/N2 permselectivities for several silica membranes, 200 °C.

Membrane series code

Manufacturing Period

P(He) (10-6 mol/m2 s Pa)

P(N2) (10-6 mol/m2 s Pa)

Permselectivity He/N2

TNO-3 Aug 2002 0.85 0.003 290 TNO-4 Nov 2002 1.51 0.002 760 TNO-5 Dec 2002 2.07 0.022 100 TNO-6 Apr 2003 1.51 0.003 510 TNO-7 June 2003 1.86 0.002 940 TNO-8 Aug 2003 2.20 0.004 560 TNO-9 Nov 2003 1.86 0.010 190

Chapter 2

22

The performance of the hollow fibres prepared in our study is comparable to that of the flat plate membranes discussed by De Vos and Verweij17. Nair et al.18 measured a helium permeance of approximately one order of magnitude lower (2 x 10-7 mol/m2 s Pa, 135 ºC), with a permselectivity of 1230 comparable to our results. The high permeance value can partly be attributed to the small contribution of support resistance.

2.5.3. Pervaporation

2.5.3.1. Dehydration of dimethylformamide with hollow fibre membranes

Dehydration of dimethylformamide (DMF) has been carried out using several silica membranes and the values of flux, water permeance are given in Table 2.2. A decline of the water permeance is observed with time. Typical data are depicted in Figure 2.4 (TNO-4b, 75 / 98 ºC, 5 wt. % water). For most of the membranes the water permeance finally reaches a steady value of ~1 kg/m2 h bar, only for membranes TNO-3 and TNO-6a the final value was higher: 3 kg/m2 h bar. For all membranes both the water permeance and selectivity are low which is in agreement with the findings of Ten Elshof et al.19. The low water permeances are due to the high interaction between the DMF molecules, with an exceptionally high dipole moment (3.82 D), and the silica material.

Table 2.2. Pervaporation performance of various hollow fibre membranes in the dehydration of DMF, 5 wt. % water.

Tc Tr tpr No Nt Γo Γt αo αt Membrane code (°C) (h) (kg/m2 h) (kg/m2 h bar)

TNO-3c♣ 350 73 24 0.70 0.36 6.2 3.6 17 25 350 86 48 0.35 0.50 3.1 2.0 75 15 TNO-4a 350 75 168 0.37 0.65 0.9 0.9 4 2 350 98 8 1.68 1.62 0.9 0.8 2 2 TNO-4b 350 75 75 0.84 0.51 2.4 1.7 4 5 350 98 85 0.84 0.75 1.3 1.1 7 7 350 75 40 0.45 0.35 1.5 1.5 7 5 TNO-6a 350 75 168 0.30 0.68 3.3 3.7 39 10 350 100 168 1.09 2.08 3.1 3.8 21 10 TNO-6b 500 75 57 0.35 0.29 4.4 1.2 70 7 TNO-6c 600 75 240 0.20 0.08 0.7 0.7 5 23

♣ Retentate water concentration is 3.3 wt. %


23

Figure 2.4. Water permeance and selectivity in the dehydration of DMF as a function of time, 75-98°C, 5 wt. % water, membrane TNO-4b.

Large variations in selectivity are observed in Table 2.2 when comparing different membranes and the behavior of the selectivity with time appears inconsistent. These observations are not surprising, given that the mobility of the large DMF molecules (>5 Å) will strongly depend on the pore size distribution. Hence, small differences in pore morphology of the different membranes will give rise to large differences in DMF permeance.

2.5.3.2. Effect of sintering temperature on the flux

The pore morphology and hydroxyl concentration of silica is directly related to the sintering temperature. A higher sintering temperature results in a denser material with a smaller pore size14. The experimental data in Table 2.1, for membranes TNO-6a, b and c, suggest that an increase in sintering temperature from 350 to 600 ºC results in a reduction of the water permeance from 3.7 to 0.7 kg/m2 h bar. This is likely due to densification of the silica material accompanied by a decrease in the number of hydroxyl groups. The data also suggests, albeit rather indistinct, an increase in selectivity. This is possibly due to a decrease in the interactions of DMF with the denser and less hydrophilic silica.

2.5.3.3. Dehydration of dimethylformamide with a tubular membrane

The evolution of the water permeance in time for the dehydration of DMF with a tubular membrane is presented in Figure 2.5. Similar to hollow fibre membranes, the water permeance decreases from 14 kg/m2 h bar to reach a constant value of about 11 kg/m2 h bar. The selectivity strongly decreases from an initial value of 200 to a stable value of

Chapter 2

24

around 7. As compared to hollow fibre membranes, the measured water permeance is higher but selectivity is in the range of the low values measured for the hollow fibres (Table 2.2).

Figure 2.5.Water permeance and selectivity in the dehydration of DMF as a function of time, 70 °C, 30 wt. % water, with a tubular silica membrane. Lines are to guide the eye.

2.5.3.4. Stability measurements with a tubular silica membrane

A series of experiments was performed on a single tubular membrane to observe the evolution of the water permeance with several mixtures. Ethanol and methanol with a high and similar dipole moment (1,7 D) and 1,4-dioxane with a low dipole moment (0 D) were chosen. Figure 2.6 shows that the water permeance for ethanol-water and subsequently for methanol-water mixtures considerably drops from 16 kg/m2 h bar to a stable value of 1 kg/m2 h bar. However, the water permeance rises to around 14 kg/m2 h bar when the membrane was placed in the 1,4-dioxane-water mixture. This value is close to the initial value obtained with the ethanol-water mixture. This recovery of the water permeance in the aged membrane suggests a partially reversible interaction between the alcohols and the silica layer.


25

Figure 2.6. Water permeance in the dehydration of several organic-water mixtures as a function of time using a tubular silica membrane. Between the experiments listed in the legend the membrane was kept in the liquid mixture at room temperature. Water concentrations are 30, 1 and 5 wt. % in ethanol, methanol and 1,4-dioxane, respectively.

2.6. Conclusions Thin microporous silica membranes have been prepared on the outer surface of

hollow fibre ceramic membrane substrates. The membranes were characterized using SEM, SNMS, single gas permeance and dehydration of DMF via pervaporation. The thickness of the silica on top of the intermediate γ-alumina layers is in the order of 20-60 nm. Gas permeance data suggests that the membranes are microporous, with a pore size close to the molecular dimensions of small molecules, and posses a small number of defects. Strong interactions between the DMF and the silica result in a low water flux and selectivity. As compared to hollow fibre membranes a higher water permeance, and a lower selectivity, was measured with a tubular silica membrane. Stability measurements using a single tubular membrane for the dehydration of ethanol and sequentially methanol in a tubular membrane show a decrease in water permeance on time, which afterwards increased when a 1,4-dioxane -water mixture was used. This recovery in water permeance in the aged membrane suggests a partially reversible interaction between the alcohols and the silica layer.

2.7. Notation iC = concentration, mol/m3

N = total flux, kg/m2 h

iN = flux of component i , kg/m2·h

p = pressure, Pa

Chapter 2

26

*ip = partial equilibrium vapor pressure, Pa oip = vapor pressure of pure component i , Pa

iΓ = permeance, (mol/m2 s Pa) or (kg / m2 h bar)

t = time, (h) or (s) T = temperature, K TNO-Yx = membrane x from membrane series Y

ix = molar fraction retentate

iy = molar fraction permeate

Greek letters

α = separation factor

iγ = activity coefficient

Subscripts

c = calcination i = component “ i ” o = initial p = permeate pr = process r = retentate t = final

2.8. Reference list (1) Van Gemert, R. W.; Petrus-Cuperus, F. Newly developed ceramic membranes for

dehydration and separation of organic mixtures by pervaporation. J. Membr. Sci. 1995, 105, 287 - 291

(2) Verkerk, A. W.; Male, P.; Vorstman, M. A. G.; Keurentjes, J. T. F. Properties of high flux ceramic pervaporation membranes for dehydration of alcohol/water mixtures. Sep. Purif. Technol. 2001, 22-23, 689 - 695

(3) Wynn, N. Dehydration with silica pervaporation membranes. Mem. Tech. 2001, 10 - 11

(4) Irving, J. P.; Butt, J. B. An experimental study of the effect of intraparticle temperature gradients on catalytic activity. Chem. Eng. Sci. 1967, 22, 1859 - 1873


27

(5) Petersen, E. E. A general criterion for diffusion influenced chemical reactions in porous solids. Chem. Eng. Sci. 1965, 20, 587 - 591

(6) Uhlhorn, R. J. R.; Keizer, A. J. Gas transport and separation with ceramic membranes. Part II. Synthesis and separation properties of microporous membranes. J. Membr. Sci. 1992, 66, 271 - 287

(7) De Lange, R. S. A.; Hekkink, J. H. A.; Keizer, A. J.; Burggraaf, A. J. Formation and characterization of supported microporous ceramic membranes prepared by sol-gel modification techniques. J. Membr. Sci. 1995, 99, 57 - 75

(8) Weisz, P. B.; Hicks, J. S. The behaviour of porous catalyst particles in view of internal mass and heat diffusion effects. Chem. Eng. Sci. 1962, 17, 265 - 275

(9) Barrer, R. M. Activated diffusion in membranes. Trans. Faraday. Soc. 1939, 35, 644 - 656

(10) Kaerger, J. Diffusion in zeolites and other microporous solids; Wiley: New York, 1992.

(11) Van den Broeke, L. J. P.; Bakker, W. J. W. Binary permeation through a silicalite-1 membrane. AIChE J. 1999, 45, 976 - 985

(12) Calderbank, P. H.; Moo Young, M. B. The continuous phase heat and mass transfer properties of dispersions. Chem. Eng. Sci. 1961, 16, 39 - 54

(13) Deans, H. A.; Lapidus, L. A computational model for predicting and correlating the behavior of fixed bed reactors: I. derivation of model for nonreactive systems. AIChE J. 1960, 6, 656 - 663

(14) De Vos, R. M.; Verweij, H. Improved performance of silica membranes for gas separation. J. Membr. Sci. 1998, 143, 37 - 51

(15) Brenner, H. The diffusion model of longitudinal mixing in beds of finite length. Numerical values. Chem. Eng. Sci. 1962, 17, 229 - 243

(16) Gmehling, J., Onken, U., and Wolfgang, A. DeChema Chemistry Data Series. 1977, Vapor liquid equilibrium data collection.

(17) De Vos, R.; Verweij, H. High selectivity, high flux silica membranes for gas separation. Science. 1998, 279, 1710 – 1711

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(18) Nair, B. N.; Keizer, A. J.; Suematsu, H.; Sun, Y. M.; Naneko, N.; Ono, S.; Okubo, T.; Nakao, S. I. Synthesis of gas and vapor molecular sieving silica membranes and analysis of pore size and connectivity. Langmuir. 2000, 16, 4558 - 4562

(19) Elshof, J. E.; Rubio, C.; Sekulic, J.; Chowdhury, S. R.; Blank, D. H. A. Transport mechanisms of water and organic solvents through microporous silica in the pervaporation of binary liquids. Microp. Mesopor. Mater. 2003, 65, 197 - 208

________________________________ This chapter has been submitted for publication: J. Fontalvo, E. Fourcade, P. C. Cuellar, J. G. Wijers, and J. T. F. Keurentjes. (2005).

3. Study of the hydrodynamics in a pervaporation module and

implications for the design of multi-tubular systems

Abstract CFD simulations have been carried out to describe pervaporation of organic-

water mixtures with a tubular membrane. The calculated density and temperature profiles have been compared with experiments. The profiles have been measured by ultrasound computer tomography (UCT) for pervaporation of pure water and an ethanol-water mixture. Good agreement between the measured and calculated density and temperature profiles has been found. Several flow patterns have been detected depending on the density ratio between the non-permeating and the permeating component in the mixture. For light organic-water mixtures an inversion point has been found at the membrane surface below which the flow moves downwards and above which the flow moves upwards. The inversion point is caused by the opposite effects of water concentration and temperature on the mixture density and consequently on the natural convection. For heavy organic-water mixtures the inversion point disappears and the flow on the membrane surface is only downward. The flux through the membrane tube increases as the density ratio increases. For the design of multi-tubular pervaporation modules, the results suggest that for separation of heavy organic-water mixtures the flux is maximized with a squared configuration. On the other hand, a triangular configuration is preferred for dewatering of light organic-water mixtures.

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3.1. Introduction Pervaporation is becoming an important technique for the separation of liquid

mixtures, especially because of the reduced energy consumption compared with conventional processes such as distillation1,2. For the design of pervaporation modules concentration and temperature polarization have to be considered and minimized. Typically, the polarization effects are reduced using high liquid velocities and turbulent conditions. However, in some applications low liquid flow velocities are used, e.g. in membrane reactors or modules where long residence times are required. In this case the concentration and temperature gradients, generated around the membrane, induce density differences. Consequently, natural convection or secondary flow will influence the performance of the module.

Computational fluid dynamics (CFD) has been successfully applied for the simulation

of several membrane processes such as microfiltration3, gas separation4,5 and fuel cells6. For pervaporation of a pure component, the hydrodynamics in a pervaporation module have been studied, comparing measured and calculated temperature profiles7, leading to similar profiles. The temperature profiles have been measured with ultrasound computer tomography (UCT)8. Owing to the pervaporation process, the liquid around the membrane surface is cooled producing a downwards movement of liquid. This buoyancy effect increases the flux compared to a situation where no buoyancy occurs.

Using CFD, it has been shown for a glycerol-water mixture that secondary flows

affect the performance of pervaporation membrane tubes9. The hydrodynamics in mixtures is more complex due to the composition and temperature gradients induced by the pervaporation process. Temperature and composition can have a similar or opposite effect on the mixture density affecting the natural convection and the flux.

The aim of this work is to study the hydrodynamics of the liquid during the

pervaporation of various mixtures, its influence on the flux and the resulting implications for the design of multi-tubular pervaporation modules. First CFD simulations are validated with experimental temperature and density profiles measured with UCT. Then, CFD is used to simulate the effect of the density ratio between the components on the hydrodynamics and flux. Finally, the implications for the design of multi-tubular pervaporation modules are discussed.

Study of the hydrodynamics in a pervaporation module and implications for the design of multi-tubular systems

31

3.2. Setup The pervaporation module used for the experiments is presented in Figure 3.1. The

insulated external glass tube was made of two sections, with a total length of 165 cm with an inner diameter of 4 cm. The inner tube was made of three sections: a nonporous glass pipe with a length of 35 cm, the ceramic membrane and an outlet pipe of 25 cm length. The inner tube could be shifted to several axial locations. The outlet pipe connected the internal side of the membrane tube to the vacuum system. The membrane (provided by ECN, The Netherlands) was 90 cm long, with an outer and inner diameter of 1.4 and 0.8 cm, respectively. The membrane consisted of a support tube, made of α-Al2O3 with an intermediate layer of γ-Al2O3, externally coated with an amorphous water selective silica layer with a thickness of 200 nm and an average pore diameter smaller than 0.5 nm. The system operated with a feed reservoir of 20 L, giving virtually constant feed conditions over several hours of operation. Pure water and an ethanol - water mixture containing 42 wt. % water were fed to the annulus of the horizontally placed pervaporation module. Feed temperatures of 28 °C for water and 50 °C for the ethanol-water mixture and liquid feed flows of 500 mL/min were used. A pressure of 2 mbar was applied at the permeate side. The permeate stream was condensed in a cold trap system with liquid nitrogen and the flux was determined gravimetrically. Feed compositions were obtained by standard Karl Fischer titration and permeate concentrations by gas chromatography.

Figure 3.1. Schematic representation of the pervaporation module used for the experiments

To measure temperature and density profiles with UCT the setup also included

auxiliary equipment. A set of transducers (microphones and speakers) mounted in a cylindrical aluminum block (Figure 3.2) were used to measure the time of flight of sound (TOF). Figure 3.2b shows a picture where the transducers can be seen. The block was

Chapter 3

32

placed between the two glass tube sections that made up the external tube as shown in Figure 3.1. The aluminum block could be rotated independently from the glass tube.

Figure 3.2. Location of the eight transducers relative to the membrane (a) and a photograph representing the location of 8 speakers and their 8 microphones (b). Distances are given in millimeters.

The location of the transducers in the block relative to the membrane is shown in

Figure 3.2a. In the Z direction, eight pairs of transducers with a diameter of 5 mm each, were evenly distributed over 7 mm. In the axial direction, denoted as Y, the transducers were evenly distributed over 28 mm.

A data acquisition system controlled and synchronized all actions. A pulse generator

Yokogawa FG120/2 MHz generated a sound pulse for the speaker with a frequency of 1.2 MHz. The signal, as collected by the microphone was transferred to a PC for post-processing with a digital filtering technique. The obtained TOF was corrected by calibration.

3.3. Ultrasound reconstructions and calibration The sound velocity measured by the time of flight (TOF) between speaker and

microphone is a function of temperature and composition of the liquid. Temperature profiles can be reconstructed when the composition is constant or known in the 2D region. For pervaporation of pure water, temperature profiles were obtained. Due to changes in composition and temperature, for mixtures it is more convenient to express the sound velocity as a function of density and compressibility. Consequently, density profiles can be reconstructed for relatively small changes in composition and temperature where the compressibility is approximately constant. Density profiles were reconstructed for the dewatering of ethanol-water mixtures. The speed of sound in ethanol-water mixtures is


33

independent of compressibility, within an error of about 2%, in the density range of 790 to 930 kg/m3 and temperatures between 25 and 50 °C.

Because the transducers were distributed over a length of 28 mm (Figure 3.2), for the

reconstruction of the 2D temperature and density profiles it was assumed that all transducers were present in one single plane. Measurements of TOF and 2D reconstructions were performed in a single transversal plane at three different axial positions from the membrane inlet: 46, 60.5 and 85.5 cm. Figure 3.3a shows the 8 lines between speakers and microphones that were used for a given position of the block of transducers. By rotating the block of transducers a complete screening of the transversal section was achieved as shown in Figure 3.3b. A set of 96 lines was measured by sequential rotation of the block with transducers for angles of 30° on a single plane perpendicular to the liquid flow direction. Note that close to the membrane surface (Figure 3.3b), in a range of about 2 mm, no measurements were taken. The set of TOF measured for the lines was used as the input for a computer program that was developed to reconstruct profiles based on an algebraic reconstruction technique (ART)8 and calibration data.

Figure 3.3. Ray distribution for: (a) one angular position and (b) after rotation of the block over 360° in 12 steps

The calibration data consisted of values of the length of each line between speaker

and microphone and the delay time of the signal. The calibration was achieved by measuring the apparent TOF under conditions where the sound velocities are known. Water, ethanol, butanol and isopropyl alcohol, as pure substances, and ethanol-water mixtures were used. The values of sound velocities for pure substances were taken from Sakurai et al.10 and for the ethanol-water mixtures from Onori11.

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34

3.4. Numerical simulations 3D simulations of the system were carried out on a half-module, cut along the axis,

with symmetry boundary conditions using ANSYS CFX 5.7. All simulations were performed with an unstructured mesh containing 428,000 nodes forming 1,400,000 first-order tetrahedral elements. An attempt to run the calculations with a structured mesh surprisingly gave non-stable solutions. Simulations consisted of solving the coupled Navier-Stokes, heat and mass transfer equations in the laminar regime using the full buoyancy model. The buoyancy model makes the convergence to steady state difficult. Therefore, transient simulations were performed until steady state was reached, i.e. when the heat flux at the membrane surface became constant. To assess the independence of the results from the mesh, one simulation was performed on a finer mesh yielding no noticeable changes in the solution. Simulations were performed for pure water and for several organic-water mixtures.

In the simulations with pure water it is sufficient to remove only heat at the

membrane surface, because the removal of water does not result in a change of composition. This was done by using a boundary source condition (Neumann boundary condition) with the following heat flux term:

h v vp Hφ Γ= ⋅ ⋅ (1)

where Γ is the intrinsic water membrane permeance, pv is the water vapor pressure

and Hv is the heat of vaporization of water. Vapor pressure and heat of vaporization depend on the temperature at the membrane surface. The mass flux was calculated by dividing the heat flux by the heat of vaporization. The heat effect due to expansion in the permeate side was neglected. The intrinsic membrane permeance Γ was used as a fitting parameter to match the calculated flux at the membrane with the experimental one. Heat losses to the surroundings at the external wall were not considered. Using a time step of 0.5 s, for these calculations convergence was achieved in 10 hours with six parallel processors (2.8GHz Intel XEON).

The organic-water mixtures were modeled as a multicomponent fluid. Neumann

boundary conditions for mass flux are not available in CFX 5.7. Therefore, water removal was implemented as a sink term in the continuity equation within a thin layer or subdomain around the surface of the membrane. The sink term was defined as:

subwsub tS φ= (2)


35

where φw stands for the mass flux through the membrane and tsub is the thickness of the subdomain around the membrane surface, for which a value of 50 microns was used. The value of tsub was optimized by comparison between the heat flux calculated with a Neumann condition and the one calculated using a subdomain for pervaporation of pure water.

The heat flux was implemented similar to the simulations with pure water. The mass

and heat fluxes read:

m v w wp x aφ Γ= ⋅ ⋅ ⋅ (3)

h m vHφ φ= ⋅ (4)

where xw stands for the molar fraction of water and aw is the activity coefficient of

water in the mixture. The activity coefficient depends on the composition and the temperature at the membrane surface12. Ethanol removal was not taken into account in the simulations due to the high selectivity for water as obtained from experiments, being around 300. With a time step of 0.1 s, for these calculations convergence was achieved in 4 days with six parallel processors (2.8GHz Intel XEON). Heat losses to the surroundings were considered to be 10% of the heat flux at the membrane wall.

3.5. Comparison of experimental and calculated data

3.5.1. Pervaporation of pure water The experimental water flux measured for pervaporation of pure water and used to

adjust the intrinsic water permeance for CFD simulations is 1.03 kg/m2 h at 28 °C. This leads to a value of 29.52 kg/m2 h atm, which is used for all subsequent calculations for organic – water mixtures. The temperature profiles calculated with CFD and measured with UCT for an axial position of 46 and 85 cm are shown in Figure 3.4. The pervaporation process reduces the surface temperature on the membrane generating a secondary flow of cold liquid that moves to the bottom of the system. This cold liquid is identified in the Figure by green and blue for CFD and UCT profiles. The size, evolution and shape of the areas in red and yellow in the measured and calculated profiles are similar.

In the axial direction, the area of the cold zone increases, as can be seen from the evolution of the green and blue zones from Figure 3.4a to Figure 3.4b. The CFD calculation in Figure 3.4b presents a cold zone in light blue immediately below the

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membrane tube that was also measured using UCT. The dark blue zone on the membrane surface in the CFD calculations could not be measured by UCT because it is located within a range of 2 mm from the membrane surface. In this range the values of the reconstructed temperature were calculated from extrapolation. The temperature drop that has been obtained, close to the external wall from UCT reconstructions, indicate some heat losses to the surroundings.

The transversal average temperature at several axial positions evaluated from CFD

calculations and from UCT reconstructions are presented in Table 3.1. The average temperature decreases in the axial direction and the differences between calculated and measured values are smaller than 0.2 ºC. Obviously, CFD is able to predict very well the transversal and axial temperature profiles in a pervaporation module using pure water.

Table 3.1. Average cross section temperature for several axial positions from CFD calculations and UCT reconstructions.

Axial position, cm Temperature, °C CFD UCT

46.0 26.9 26.7 60.5 26.8 26.6 85.5 26.5 26.4


37

Figure 3.4. Temperature distribution for pervaporation of water at 27.6 °C for two axial positions measured from the inlet, Feed flow 500 ml/min. Left from CFD calculations and right from UCT measurements. Axial positions of (a) 46.0 cm and (b) 85.5 cm.

Figure 3.5. Density distributions for pervaporation of water-ethanol. Feed flow of 500 ml/min, 42 wt. % water and 50.0 °C. Left from CFD calculations and right from UCT reconstruction. Axial position of 46.0 cm.

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38

3.5.2. Pervaporation of ethanol-water mixtures The density profile calculated using CFD and reconstructed from UCT measurements

for the dehydration of ethanol at 50 °C at an axial position of 46 cm is presented in Figure 3.5. The CFD calculation shows a horizontal stratification while in the reconstruction profile the stratification is annular. The average density and density range however, are similar in both profiles. The calculated and reconstructed profiles show a zone of low density in green at the top of the system that covers most of the transversal area. Both profiles show a band of high density at the bottom in red. A light liquid zone is also present in both profiles in blue at the top of the membrane surface. This zone is less extended in the CFD calculations as can be seen in the blow-up picture in Figure 3.5. The presence of this blue zone means that despite the cooling effect on the membrane surface the pervaporation process generates a secondary flow of light liquid that moves upwards in the system. This behavior is different than for pure water where a cool liquid flows downwards around the membrane tube.

A difference of only 0.1% has been found between the calculated and the

measured water flux through the membrane tube of 1.64 kg/m2 h. For the calculation of the flux, the same intrinsic water permeance as for the simulation of pervaporation of pure water has been used. If no concentration and temperature polarization occurs the expected flux would be around 4 kg /m2 h.

3.6. Hydrodynamics during pervaporation of several organic-water mixtures

CFD is used to evaluate the hydrodynamics of organic-water mixtures in which the organic compound has the same transport properties as ethanol but has a different density. Density ratios between the organic compound and water from 0.7 to 1.5 have been included in the simulations. The influence of the density ratio on the calculated water flux through the membrane is presented in Figure 3.6, from which it can be concluded that the flux through the membrane increases as the density of the organic compound increases.


39

Figure 3.6. Relation between the water flux through the membrane and the density ratio between the organic component and water calculated from CFD.

This relation between density ratio and flux results from the combined effects of

temperature and water concentration on the density of the mixture. At the membrane surface the cooling effect induces an increase in the density of the mixture. However, the water removal induces an increase or a decrease in the density depending on the organic compound being the heavier or the lighter compound in the mixture, respectively.

The effect of the density ratio on the velocity field is presented in Figure 3.7. For high

density ratios the lower temperature and water concentration on the membrane surface produces a downward movement of liquid as shown in the Figure for a density ratio of 0.93. The intensity of the downward velocity field controls the rate at which liquid from the bulk is transported to the membrane surface. This intensity increases with the density ratio resulting in higher fluxes. For low density ratios there is an inversion point at the membrane surface below which the flow goes down and above which the flow goes up as shown for a density ratio of 0.68. The inversion point moves to higher positions as the density ratio increases and finally disappears when the density ratio reaches a value of approximately 0.9. At higher density ratios the flow only moves downwards. For low density ratios, owing to the opposite effect of temperature and water concentration on the density of the mixture, the intensity of the velocity field around the membrane is damped reducing the flux as compared to high density ratios.

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40

Figure 3.7. Velocity profiles calculated with CFD with density ratios of 0.68 and 0.93 at an axial position of 46 cm. The thick lines, next to the numbers, indicate the direction of the flow close to the membrane surface.

3.7. Implications for the design of multi-tubular pervaporation modules

The observed secondary flow depends on the density ratio, the intrinsic water permeance and the trans-membrane partial vapor pressure difference. The secondary flow produced around each membrane tube will generate an overall secondary flow pattern in a multi-tubular pervaporation module. For either light or heavy organic-water mixtures, the pervaporation module should be designed to promote an optimal secondary flow.

It has been shown that the secondary flow pattern for pervaporation of light organic-

water mixtures is upwards at the top of the membrane tube and downwards at the bottom. This pattern suggests that in a multi-tubular system with a square configuration the liquid flow that is descending from a given row of tubes is damped by the upward flow from the tubes of the row below. Consequently, the rate at which liquid is transported from the bulk to the membrane surface is reduced. This effect results in a lower flux as compared to the values presented in Figure 3.6 for a single tube. For dewatering of a light organic-water mixture a triangular configuration is more effective as compared to a


41

square configuration. In a triangular configuration the distance between neighboring tubes along a vertical line increases, reducing the interaction of the flows from adjacent rows of tubes. When using a square configuration, the distance between adjacent rows of tubes should be increased and consequently the packing density will be lower. Figure 3.8a shows the flow stream-lines and the direction of the flow around a triangular configuration. The streamlines were calculated with FEMLAB® using temperature gradients to generate the secondary flows. As compared to Figure 3.8b the distance between tubes on a vertical line increases. The liquid that drops from one tube interacts with the liquid that flows upwards from the tube below generating several vortexes around each tube. Figure 3.8. Representation of stream-lines inside of a multi-tubular pervaporation module for two configurations a) triangular and a low density ratio mixture, b) squared and a high density ratio mixture. Stream-lines calculated with FEMLAB® using an analogy with heat transfer.

For pervaporation of heavy organic-water mixtures either a square or triangular

configuration can be used. The expected overall performance of the pervaporation module in any configuration will be superior to the one presented in Figure 3.6. The square configuration is favored because the intensity of the downwards velocity field of a given row of tubes is increased by the flow coming from the row of tubes immediately above. This effect can be seen in Figure 3.8a where some vortexes are present.

a) b)

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3.8. Conclusions Transversal temperature and density profiles at several axial positions have been

reconstructed from ultrasound computer tomography (UCT) measurements. The transversal temperature profiles calculated by CFD for pervaporation of pure water are similar to the ones reconstructed, with a difference in the transversal averaged temperatures smaller than 0.2 °C. For pervaporation of an ethanol-water mixture the predicted water flux through the membrane is only marginally higher than the measured value (< 0.1%). The density profile calculated with CFD is similar to the profile reconstructed from UCT. For dewatering of a light organic-water mixture CFD predicts an inversion point at the membrane surface below which the flow moves downwards and above which the flow moves upwards. This inversion point has also been measured for dewatering of ethanol-water mixtures where the density profile reconstructions show a zone of low density at the top of the membrane tube. The inversion point is caused by the opposite effect of water concentration and temperature on the mixture density and consequently on the natural convection. The simulations performed for pervaporation of organic-water mixtures show that the natural convection, and as a consequence the flux, in the pervaporation module decreases as the density ratio between the non-permeating and the permeating components decreases. The experimental and theoretical results suggest that for the design of multi-tubular pervaporation modules for dewatering of light organic-water at low liquid flow rates a triangular configuration is more convenient. For pervaporation of heavy organic-water mixtures a square configuration is preferred where higher fluxes than for a single tube can be obtained.

3.9. Reference List (1) Jonquieres, A.; Clement, R.; Lochon, P.; Neel, J.; Chretien, B.; Dresch, M. Industrial

state-of-the-art of pervaporation and vapor permeation in the western countries. J. Membr. Sci. 2002, 206, 87 - 117


(3) Rahimi, M.; Madaeni, S. S.; Abbasi, K. CFD modeling of permeate flux in cross-flow microfiltration membrane. J. Membr. Sci. 2005, 255, 23 - 31

(4) Takaba, H.; Nakao, S. I. Computational fluid dynamics study on concentration polarization in H2/CO separation membranes. J. Membr. Sci. 2005, 249, 83 - 88


43

(5) Takaba, H.; Nishimura, T.; Fanatsu, K.; Nakao, S. I. Development of a desing tool for gas separation membrane modules using computational fluid dynamics simulation. Transactions of the Materials Research Society of Japan. 2004, 29, 3279 - 3282

(6) Sivertsen, B. R.; Djilali, N. CFD based modelling of proton exchange membrane fuel cells. Journal of Power Sources. 2005, 141, 65 - 78

(7) Van der Gulik, G. J. S.; Wijers, J. G.; Keurentjes, J. T. F. Measurement of 2D temperature distributions in a pervaporation membrane module using ultrasonic computer tomography and comparison with computational fluid dynamics calculations. J. Membr. Sci. 2002, 204, 111 - 124

(8) Williams, R. A.; Beck, M. S. Process tomography : principles, techniques and applications; Butterworth-Heinemann: London, 1995.

(9) Van der Gulik, G. J. S.; Janssen, R. E. G.; Wijers, J. G.; Keurentjes, J. T. F. Hydrodynamics in a ceramic pervaporation membrane reactor for resin production. Chem. Eng. Sci. 2001, 56, 371 - 379

(10) Sakurai, M.; Nakamura, K.; Takenaka, N. Apparent molar volumes and apparent molar adiabatic compressions of water in some alcohols. Bulleting of the Chemical Society of Japan. 1994, 67, 352 - 359

(11) Onori, G. Adiabatic compressibility and structure of aqueous solutions of ethyl alcohol. J. Chem. Phys. 1988, 89, 4325 - 4332

(12) Gmehling, J., Onken, U., and Wolfgang, A. DeChema Chemistry Data Series. 1977, Vapor liquid equilibrium data collection.

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________________________________ This chapter has been submitted for publication: J. Fontalvo, M. A. G. Vorstman, J. G. Wijers, and J. T. F. Keurentjes. (2005).

4. Heat supply and reduction of polarization effects in pervaporation

by two-phase feed

Abstract Gas-liquid and vapor-liquid two phase pervaporation have been

experimentally studied using a lab-scale and a bench scale pervaporation setup. Pervaporation experiments with low liquid flow rates in internally coated tubes have been carried out for dewatering of 1,4-dioxane and isopropyl alcohol. Relatively small amounts of gas or vapor are sufficient to reduce concentration and temperature polarization and to increase the total flux and selectivity more than twofold compared with single phase experiments. The beneficial effects of a two-phase feed are also demonstrated by calculations. Additionally, by condensation the vapor effectively supplies the heat required for the selective evaporation through the membrane. As a consequence, this eliminates the liquid temperature drop that reduces the performance of pervaporation modules. The values of the total flux achieved using vapor-liquid two-phase pervaporation at laminar flow conditions are close to those for turbulent conditions.

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46

4.1. Introduction Pervaporation has emerged as a promising separation technique for aqueous and

organic solutions removing preferentially water or one of the organic compounds1. Heat is required for the evaporation of the compounds that are removed, which is generally supplied from the sensible heat carried by the liquid feed. The resulting temperature drop2 reduces the flux through the membrane. As a consequence, the membrane area required for a given separation increases as compared to the hypothetical isothermal case.

Polarization effects also reduce the performance of pervaporation systems and have

been studied for several types of membrane processes. Bhattacharya and Hwang3 have shown a generalized equation relating a modified Peclet number to the concentration polarization occurring in the boundary layer. The resulting equation is applicable to the majority of membrane separation processes like gas separation, reverse osmosis, ultrafiltration, pervaporation, and dissolved gas permeation in liquid. The negative effect of concentration polarization on the performance of pervaporation membranes (flux and selectivity), especially for the removal of traces of organic compounds has been studied by Baker et al.4, Mi and Hwang5, Michaels 6 and Psaume et al.7. Favre8 has studied the temperature polarization towards the membrane surface due to the resistance for heat transport. The temperature drop that occurs in the liquid due to the pervaporation processes has been studied by Rautenbach9 and more recently by Ito and Feng2.

Several methods have been explored to diminish the temperature drop in

pervaporation systems, e.g. a series of alternating pervaporation modules and heat exchangers9. When heat is supplied directly into the pervaporation module in order to maintain a constant retentate temperature, the membrane area could be reduced by 20% to 40 %10, and no inter-stage heat exchangers will be needed.

Various designs have been presented in the literature to supply energy directly into

the pervaporation module, for instance using electrical resistances11 or external heating with steam or another fluid12. Because of safety reasons electrical resistances cannot be applied with inflammable media. External heating in multi-tubular modules requires a system of concentric tubes where the heating fluid is fed to the shell side and the retentate is fed to the annulus of each set of concentric tubes. In general these pervaporation modules have a complex construction increasing the maintenance and total cost.

Heat supply and reduction of polarization effects in pervaporation by two-phase feed

47

Vapor-liquid two-phase operation might offer an effective way to control the retentate temperature drop in pervaporation units. Two-phase operation is created by vaporizing a fraction of the feed or retentate. The vapor flow is combined with the liquid feed stream to the pervaporation unit. The vapor phase has a multifunctional role, especially in low liquid flow rate applications: it reduces temperature and concentration polarization near the membrane surface by promoting turbulence and it supplies latent heat to the liquid by condensation. Two-phase operation in multi-tubular systems can be applied for flow inside tubes or for transversal flow externally to tubes in co-current or counter-current. Pervaporation units with counter-current vapor-liquid operation could be placed directly inside a distillation column replacing a section of packing or trays.

Slug flow using an inert gas has been applied in several membrane processes such

as ultrafiltration13-17, microfiltration18 and nanofiltration13. It has been found that gas, as secondary phase, is efficient in reducing concentration polarization and fouling. Watanabe and Fuchigami19 patented a system where the liquid feed is supplied as a thin film on flat sheet pervaporation membranes in contact with a vapor phase.

This chapter shows the benefits of vapor-liquid and gas-liquid two-phase

pervaporation to increase the performance of multi-tubular pervaporation modules. The focus is on two-phase operation inside pervaporation tubes at low liquid flow rates with liquid and vapor or gas in co-current. Low liquid flow rates are often relevant, for instance in membrane reactors20 where long residence times are necessary or in pervaporation systems where low pressure drops or short membrane lengths are required.

Three kinds of experiments have been performed: using air in a lab scale setup and using vapor and air in a bench scale pervaporation unit. The influence of the second phase on temperature and concentration polarization is demonstrated by measuring flux and selectivity in experiments with air. For the lab scale experimental setup, the extent of concentration and temperature polarization has also been calculated, based on measurements of bubble rise velocity and bubble and liquid slug sizes. Using vapor, the additional effect of heat supply by condensation is shown in the bench scale pervaporation unit. Using silica membranes, dehydration of 1,4-dioxane has been performed in the lab scale setup and dehydration of isopropyl alcohol (IPA) has been carried out in the bench scale setup.

Chapter 4

48

4.2. Theory This section describes some parameters of pervaporation membranes, some

characteristics of two-phase flow in the slug regime and the calculation of the mass and heat transfer coefficients for experiments in the slug regime. These coefficients are used to evaluate concentration and temperature polarization.

Pervaporation Pervaporation processes are usually characterized by the permeance and selectivity

of the membrane. For dehydrations, the water permeance wΓ is defined as the ratio of

the water flux ( wJ ) through the membrane and the water partial vapor pressure

difference over the membrane.

pw

vwww

ww pypx

J−

=γ

Γ 4.1

where vwp is the vapor pressure of pure water, pp is the total pressure on the

permeate side and wγ is the water activity coefficient in the retentate.

The selectivity is defined according to equation 4.2. x and y are the molar fractions

in the liquid and permeate side, respectively, for water w and the other compound “ i ”.

i

i

w

w

xy

xy

=α

4.2

For the intrinsic selectivity, the local retentate molar fractions on the membrane surface are used. The feed liquid molar fraction is used to calculate the process selectivity.

The concentration polarization index (CPI) is defined as the ratio of the local water

concentration on the membrane surface and the local water concentration in the liquid bulk. Temperature polarization is defined as the local difference between the bulk temperature in the retentate and the membrane surface temperature.

Two phase flow Several regimes can be found for two-phase flow inside tubes21. However, this

section will focus on the slug flow regime which has been used in the lab scale setup. A


49

schematic representation of a slug unit is shown in Figure 4.1. The slug unit consists of a Taylor bubble zone with a falling liquid film, and a liquid slug zone22. The liquid slug consists of a wake zone23 and a remaining liquid slug. Mass and heat transfer to the membrane surface in the slug flow regime are determined by the liquid and gas velocities and the length of each zone.

Figure 4.1. Description of a slug unit. Adapted from Fernandes et al.22

The falling film is formed when liquid from the liquid slug zone is pushed to the tube

wall by the Taylor bubble forming a thin film that travels downwards around the bubble. The velocity field in the liquid slug ahead of the Taylor bubble then changes. Polonsky et al.24 found that the effect of the rising bubble on the liquid velocity field ahead of it is restricted to distances of about 1D from the bubble tip, where D is the diameter of the tube. The falling film enters into the liquid slug zone generating a zone of high turbulence called wake zone. The size of the wake zone has been measured by Pinto et al.25 to be between 5D and 6D. The high velocities in the wake zone decrease in a sufficiently long liquid slug (>6D-7D) to low values causing a relatively calm liquid behind the wake zone.

Theory and experiments suggest26 that the Taylor bubble rise velocity is about 2.3

times the total superficial velocity if there is no interaction between the bubbles27. The total superficial velocity is defined as the sum of the superficial velocity of the gas and the liquid phase. Taitel et al.21 have calculated the minimum distance between bubbles to avoid interaction as 16D.

Chapter 4

50

Several mathematical descriptions of gas and liquid velocities and lengths for each zone of a slug unit are found in the literature22,28-30. In this study the model of Fernandes at al.22 is used to calculate the liquid velocity in the falling film. For the falling film thickness δL around the Taylor bubble the equations of Wallis31 have been used. The average axial velocity in the wake zone has been calculated according to Ghosh and Cui15 which is based on the axial velocity of a two-dimensional jet that enters a stagnant pool of liquid. The Taylor bubble rise velocity, the Taylor bubble length and the liquid slug length have been taken from measurements described in the “experimental” section.

Mass and heat transfer coefficients in slug flow The wake zone14,21 is believed to be the zone with the highest mass transfer

coefficient15 especially at low gas flow rates32. The mass transfer coefficient in the falling film is similar or slightly lower than in the wake zone. Mass transfer in the falling film is relatively constant and does not increase with longer bubbles because the terminal velocity in the falling film is reached at relatively short Taylor bubbles22,29. The remaining liquid slug has the lowest mass transfer15 due to the low turbulence within this zone resulting in a reduced average mass transfer coefficient for very long liquid slugs.

Mass transfer coefficients in the falling film and in the wake zones have been obtained according to Flaschel et al.33, requiring information of the liquid velocities in the falling film and wake zones. Assuming that the establishment of concentration and temperature profiles is independent of those profiles in other slug units, the axial length for laminar conditions has been taken as the Taylor bubble length (lTB) or the liquid slug length (lLS) to obtain the mass transfer in the falling film or the wake zone, respectively.

An average mass transfer coefficient from the liquid to the membrane surface can

be calculated from the mass transfer coefficients ( mk ) and the relative length of the Taylor

bubble (lTB) and the wake zone (lW):

( ) Wm

LTBm

avem kkk ββ −+= 1 4.3

where β is the length ratio between the Taylor bubble and the slug unit. The mass

transfer in the remaining liquid slug zone is omitted because in this study experimental conditions have been adjusted to create short liquid slugs with a size around the size of the wake zone. Similar to equation 4.3 an average heat transfer coefficient has been evaluated using the Chilton-Colburn analogy for each zone. Mass and heat transfer


51

coefficients in single phase laminar flow are calculated from a correlation for flow inside tubes34.

4.3. Experimental The influence of the addition of a gas phase on concentration and temperature

polarization in a pervaporation process was determined experimentally in a single membrane tube based on the dehydration of 1,4-dioxane and in a bench scale pervaporation unit for dehydration of IPA. It was planned to also use 1,4-dioxane in the bench scale setup. However, due to the size of the equipment and safety regulations this was not possible. The additional effect of heat supply by a condensing vapor was measured for dehydration of IPA in the bench scale pervaporation unit, where vapor was obtained by partially vaporizing the IPA-water feed mixture.

Lab scale setup; air as secondary phase The setup for the air-liquid two-phase experiments is shown in Figure 4.2 The ceramic

pervaporation membrane supplied by Pervatech (The Netherlands) consisted of an alumina support tube with 7 mm internal diameter and 50 cm length, internally coated with a silica separation layer. Heating oil was supplied to the jacket of the storage vessel from an external heating bath to control the feed temperature to the pervaporation membrane at 70 °C. The feed liquid, a solution of 1,4-dioxane containing 9.0 wt. % water, was pumped to a chamber at the bottom of the vertically placed membrane. The liquid superficial velocity in the tube was 0.048 m/s (110 mL/min), for both laminar flow and two phase experiments. For two-phase experiments air was injected in the same chamber using a solenoid valve. Liquid and air were recycled to the storage vessel after passing the pervaporation membrane and air was released in the storage vessel.

Chapter 4

52

Figure 4.2. Lab scale experimental setup used for dewatering of 1,4-dioxane by gas-liquid two phase pervaporation.

Prior to the pervaporation experiments the pervaporation tube was replaced by an

identical transparent glass tube. A digital camera was positioned at a height of 40 cm from the tube bottom to measure bubble size, liquid slug size and bubble rise velocity at air flow rates between 40 and 250 mL/min and several solenoid valve frequencies. Under the experimental conditions used in this study slug flow was created. The experimental results in the glass tube were used to obtain a desired size of the Taylor bubbles and liquid slugs in the membrane tube by adjusting the gas flow rate, and the opening and closure frequency of the solenoid valve.

In order to characterize the membrane tube, additional single phase pervaporation

experiments were performed in the turbulent regime using high liquid superficial velocities of above 3 m/s. At these superficial velocities the concentration and temperature drop in the liquid and the concentration and temperature polarization on the membrane are negligible.

The permeate side of the experimental setup consisted of a system of cold traps and

a vacuum pump, shown in Figure 4.2 A metering needle valve was used to adjust the vacuum pressure at 10 mbar at the permeate side outside the membrane tube. The vapor at the permeate side was condensed using liquid nitrogen and was weighted as a function of time to determine the total flux. The water concentration in the retentate was measured by Karl-Fischer titration. The water concentration in the permeate stream was measured by gas chromatography using a TCD detector. For the lab and the bench scale


53

setups, the experimental error increases as the selectivity rises. For the highest selectivities the experimental error has been estimated to be around 7%. The feed water concentration was virtually constant during one experiment and was readjusted when necessary.

Bench - scale setup; vapor as secondary phase The bench scale unit, built by GTI (The Netherlands) is sketched in Figure 4.3. Two

pervaporation modules identified as Perv1 and Perv2, supplied by Pervatech (The Netherlands), were placed vertically on top of the reboiler. Each membrane module contained 7 pervaporation membrane tubes, which were the same tubes as used in the lab scale setup, in a shell and tube configuration. At the shell side the vacuum was adjusted to 85 mbar. At this pressure cooling water can be used to condense the permeate stream that is collected in independent cold trap systems for each pervaporation module. In the setup single phase experiments and two-phase experiments using air and vapor were performed.

Figure 4.3. Bench scale experimental setup used for dewatering of IPA.

An IPA-water solution of 17 wt. % water was fed at 7 L/hr to the reboiler and the

membrane module with a feed pump, as shown Figure 4.3. The heating coil duty in the reboiler was used to control either the liquid feed temperature to the pervaporation modules for single-phase experiments and experiments with air or the vapor feed flow rate for vapor-liquid two-phase experiments using a PID controller. For gas-liquid two-phase pervaporation experiments, air was used with flow rates between 60 and 700 mL/s. The

Chapter 4

54

hydrodynamic regimes achieved were different from the lab scale experiments. Calculations according to Taitel et al.21 indicated that slug flow was obtained at low vapor and air flow rates while for higher flow rates churn flow and annular flow were reached. The inlet and outlet temperatures, the permeate pressure in each module and the liquid levels in the reboiler and storage vessel were monitored using the software InTouch®. The liquid-vapor mixture remaining at the top of the pervaporation modules was condensed using cooling water and was recycled to the storage tank. The retentate and permeate water concentrations were measured with Karl-Fisher titration and gas chromatography, respectively. The water content in the liquid feed was monitored and was kept constant by recycling the permeate stream to the storage tank.

4.4. Results

Dioxane - water system; gas - liquid pervaporation Experimental values of the total flux through the membrane as a function of the air

flow rate are presented in Figure 4.4 for two different liquid slug sizes at a liquid flow rate of 110 mL/min. The average total flux for laminar flow at zero gas flow is shown for comparison. The total flux through the membrane strongly increases when a small amount of gas is injected in the feed stream. Low air flow rates, as low as 0.7 mL/s, are sufficient to increase the total flux twofold as compared to single-phase pervaporation in laminar flow. The retentate temperature drop in the slug flow experiments is around 9 K. Experimental values of the total flux are generally higher for a liquid slug of 5.7 cm than for 4.6 cm. The total flux decreases slightly at increasing gas flow rate, in particular for a liquid slug length of 4.6 cm because of the higher ratio of the Taylor bubble length to the slug unit length (Equation 4.3). For liquid slugs of 5.7 cm this reduction is less important since the mass transfer coefficient in the wake zone is higher for a liquid slug of 5.7 cm than for 4.6 cm slugs. The velocities in the wake zone of a liquid slug shorter than 7D, i.e. 4.6 cm, are reduced by the next ascending Taylor bubble.


55

Figure 4.4. Experimental values of the total flux as a function of the air flow rate for two-phase pervaporation of 1,4-dioxane containing 9 wt. % water at 343 K in slug flow.

The process selectivity as a function of air flow rate is presented in Figure 4.5. It shows a low process selectivity for single phase in laminar flow. However, the selectivity for water increases when air is injected and increases slightly at increasing gas flow rates.

Figure 4.5. Experimental process selectivity as a function of the gas flow rate using air-liquid two phase pervaporation for dewatering of 1,4-dioxane containing 9 wt. % water at 343 K. The lines are a guide to the eye.

Using a small amount of air, as low as 0.7 mL/s, increases the flux at least twofold and the selectivity by 50 % compared with laminar single phase experiments.

Chapter 4

56

Experimental values on slug flow and theoretical prediction of polarization.

As mentioned in the section 4.2, the mass and heat transfer are calculated based on experimental data of Taylor bubble rise velocity and Taylor bubble and liquid slug sizes. The mass and heat transfer coefficients are used to calculate the concentration and temperature polarization for the dehydration of 1,4-dioxane at the experimental conditions used. The mass balance equation for the liquid in the membrane tube has been integrated using finite differences. Flux and polarization effects have been calculated with a Maxwell-Stefan approach 35. The approximate method of Krishna 36 was applied by considering the Maxwell-Stefan diffusivity and the thermodynamic factors constant along the diffusion path in the mass transfer boundary layer. The thermodynamic factors 37,38 and activity coefficients were calculated using UNIFAC. Fluxes through the membrane were calculated as the product of permeability and local partial pressure difference over the membrane. Permeabilities are considered constant for each component.

The bubble sizes measured for two different liquid slug sizes are presented in Figure

4.6 as a function of the gas superficial velocity. The bubble size increases with the gas superficial velocity. For the two liquid slug lengths the bubbles sizes are similar at higher superficial velocities.

Figure 4.6. Experimental values of bubble size as a function of the air superficial velocity for 1,4-dioxane containing 9 wt. % water at 343 K with two liquid slug lengths. The lines are a guide to the eye.

Figure 4.7 shows the bubble rise velocity as a function of the total superficial velocity.


57

The bubble rise velocity increases at a rate of 2.9 times the total superficial velocity as compared to the theoretical value of 2.3. This higher rate suggests bubble-bubble interactions due to the short liquid slug length used in the experiments39.

Figure 4.7. Experimental values of the bubble rise velocity as a function of the total superficial velocity for 1,4-dioxane containing 9 wt. % water at 343 K.

Experimental values of the total flux and selectivity measured in single phase at

turbulent and laminar conditions in the lab-scale setup are presented in Table 4.1. The experimental values in turbulent conditions have been used as the intrinsic parameters of the membrane. The intrinsic parameters of the membrane coupled with the mass and heat transfer coefficients in laminar flow have been used to calculate the total flux and selectivity. These calculated results are also included in the Table. It can be concluded that with the intrinsic parameters of the membrane a good prediction of the total flux and process selectivity is obtained. The same intrinsic parameters have been included for the calculation of the flux in the slug flow regime.

Table 4.1. Experimental and calculated total flux and selectivity for dewatering of 1,4-dioxane containing 9.0 wt. % water at 70 °C in single phase.

Total Flux kg / m2 h

Selectivity α

Turbulent flow, experimental 5.1 2600 Laminar flow, experimental♣ 1.6 779

Laminar flow, calculated♣ 1.7 718

♣ Feed flow rate is 110 ml/min.

Calculated values of the total flux are presented in Figure 4.8 for slug flow together with experimental values of the total flux for single phase in the turbulent and laminar flow

Chapter 4

58

regime and the total fluxes in slug flow regime. Calculated values of total flux for slug flow with isothermal conditions are also included. All calculated and experimental values for slug flow are for a liquid slug length of 5.7 cm. Good predictions of the total flux for slug flow have been obtained with a maximum difference of 6%. For liquid slug lengths of 4.6 cm similar good predictions have been obtained, however, the predicted process selectivities are 40% higher than the experimental values.

Figure 4.8. Comparing experimental and predicted values of the total flux for air-liquid two-phase pervaporation in slug flow for dewatering of 1,4-Dioxane containing 9 wt. % water at 343 K with a liquid slug length of 5.7 cm.

The calculated concentration polarization index and temperature polarization are

0.39 and 5.8 K, respectively, for laminar flow experiments. Under slug flow conditions, the concentration polarization index increases to at least 0.83 and the temperature polarization reduces to 1.8 K at most. The addition of gas has reduced concentration polarization by more than 50% and temperature polarization by more than 70% for the dehydration of 1,4-dioxane. The total flux in the slug regime using air, with a liquid temperature drop of 8.5 K, is 20% lower compared with experimental values in the turbulent regime. However, Figure 4.8 shows that an isothermal operation, e.g. using vapor instead air, in slug flow increases the total flux resulting in a performance that is only 10% lower than under turbulent conditions.

IPA - water system; vapor-liquid pervaporation For the IPA-water system, the total flux as a function of the vapor feed flow rate is

presented in Figure 4.9 for both pervaporation modules connected in series in the bench scale pervaporation setup with a liquid flow rate of 117 mL/min. In single phase


59

experiments total fluxes of 0.85 and 0.60 kg/m2 h have been measured for Perv1 and Perv2, respectively. In the absence of vapor, the 30% lower performance of Perv2 is mainly due to the liquid temperature drop in the axial direction: the liquid feed temperature to Perv2 is around 9 K lower than to Perv1 reducing the driving force for mass transfer.

Figure 4.9. Total flux as a function of the vapor feed flow rate for dewatering of IPA containing 15 wt. % water at 353 K.

Figure 4.9 shows that when vapor is added to the liquid feed higher fluxes are

achieved and the differences in total flux between Perv1 and Perv2 are eliminated. The total flux through the membrane increases at least twofold using relatively low vapor flow rates, as low as 12 mL/s, as compared to single phase laminar flow experiments. The water transport through silica membranes for 1,4-dioxane-water mixtures40 is higher than for IPA-water mixtures41 as shown in Figure 4.1 and Figure 4.9, respectively. The total flux through the membranes increases with increasing vapor flow rate and in general the total flux in Perv2 is similar to the flux in Perv1. Some experimental points of the total flux in Perv2 are slightly higher than for Perv1, probably as a consequence of the experimental error. The condensation of the vapor supplies the heat required for the pervaporation process in Perv1 and Perv2. As a consequence, the performance of the two pervaporation modules is similar. Also, the improved performance at high vapor flow rates is the result of the higher turbulence in the liquid phase created by the vapor phase.

The process selectivity is shown in Figure 4.10 as a function of vapor flow rate. Vapor

flow rates higher than 86 mL/s increase the selectivity twofold as compared to single-phase laminar flow experiments. The process selectivity increases only slightly at higher vapor flow rates.

Chapter 4

60

Figure 4.10. Process selectivity as a function of the vapor feed flow rate for dewatering of IPA containing 15 wt. % water at 353 K.

Water – IPA system; experimental comparison of two-phase pervaporation using air and vapor.

Experimental results of the total flux for the dewatering of IPA using air-liquid two-phase pervaporation in the bench scale setup are shown in Figure 4.11. Also experimental results of the total flux using vapor have been included. When air is used, Perv2 has a lower performance than Perv1 due to the retentate temperature drop.

Figure 4.11. Comparison of the measured total flux as a function of the air and vapor flow rates for dewatering of IPA containing 15 wt. % water at 353.5 K. The Lines are a guide to the eye


61

The total flux increases by around 20% for Perv1 and 40% for Perv2 when a condensing vapor instead air is used (Figure 4.11). Also, using vapor, the differences in flux between the pervaporation modules are eliminated. The maximum total flux through the pervaporation membrane in Perv1 with air is 2.2 kg/ m2 h, however, this flux increases to 3.1 kg/ m2 h when vapor is used.

4.5. Conclusions Experiments performed on the dehydration of 1,4-dioxane and isopropyl alcohol

show that small amounts of air or vapor are sufficient to increase the total flux and selectivity in pervaporation modules at least twofold as compared to the performance in single phase laminar flow. For slug flow experiments, the total flux, selectivity, concentration and temperature polarization have been calculated based on experimental data of bubble size, liquid slug size and bubble rise velocity. Good predictions of the total flux and selectivity have been found for single phase and two-phase flow experiments. Compared with laminar flow experiments for the dehydration of 1,4-dioxane, calculations show that concentration polarization and temperature polarization are reduced by 50% and 70%, respectively.

Water flux increases by 20% and 40% when vapor instead of air is used for dehydration of isopropyl alcohol in two pervaporation modules connected in series due to the reduction of the liquid temperature drop. Vapor-liquid two-phase pervaporation proves to increase fluxes through internally coated membranes by reducing the liquid temperature drop in pervaporation systems and for relatively long residence times by enhancing mass and heat transfer towards the membrane surface.

4.6. Notation D = internal diameter of the membrane tube [m] de= equivalent diameter defined in equation A5 g = gravity [m/s2] J = membrane flux [kg/m2 h]

mk = mass transfer coefficient [m/s]

k = constant in equation A1 l = length [m] L = Length for laminar flow equivalent to lTB or lLS for the falling film and the wake

zone, respectively. m = constant in equation A1

Chapter 4

62

vp = vapor pressure [bar]

Perv1 = first pervaporation module in Figure 3 Perv2 = second pervaporation module in Figure 3 Re = Reynolds number for the liquid phase s = constant in equation A7, 1x10-7 [m] Sc = Schmidt number Sh = Sherwood number U = velocity [m/s] x = molar fraction

Greek letters α = membrane selectivity or process selectivity (equation 2) β = ratio between Taylor bubble length and slug unit length

Lδ = falling film thickness [m]

µ = viscosity [kg / m s]

ρ = density [kg/m3]

wΓ = membrane water permeance [kg/m2 h bar]

γ = activity coefficient

Subscripts ave = average i= organic compound G= gas phase GTB = gas in the Taylor bubble GLS = gas in the liquid slug L = liquid phase LLS = liquid in the liquid slug LS = liquid slug LTB = falling film liquid around the Taylor bubble N = bubble rising r = radial direction [m] SU = slug unit TB= Taylor bubble w = water W = wake zone z = axial direction [m]


63

Superscripts LTB = falling film liquid zone W = wake zone

4.7. Appendix

Hydrodynamic parameters Several mathematical descriptions of gas and liquid velocities and lengths for each

zone of a slug unit are found in the literature22,28-30. In this study the model of Fernandes at al.22 is used to calculate the gas and liquid velocities, the gas hold up and the sizes of the different zones in the slug unit. The Taylor bubble rise velocity UN, The Taylor bubble length and the liquid slug length have been taken from experiments described in the results chapter in the section “Experimental values on slug flow and theoretical prediction of polarization”. The falling film thickness δL around the Taylor bubble was calculated from31:

( )

m

L

LLTBL

LGL

LL UgD

kD

−

=µ

δρρρρ

µδ 43/1

3

2

(A1a)

where,

100Re3.1;909.0 <===L

LLTBLLTB

Uifmkµ

δρ 31 (A1b)

4300Re1003/2;0682.0 <<== LTBifmk 22

Mass and heat transfer coefficients An average mass transfer coefficient from the liquid to the pervaporation membrane

surface can be calculated from the mass transfer coefficient and the relative length of the Taylor bubble and the wake zones:

Wm

LTBm

avem kkk )1( ββ −+= (A2)

The mass transfer in the remaining liquid slug zone is omitted because in this study

experimental conditions have been adjusted to create short liquid slugs with a size around the size of the wake zone. Similar to equation A2, an average heat transfer coefficient is calculated using the Chilton-Colburn analogy for each zone. Mass transfer coefficients have been calculated using equations A3 and A433,42:

Chapter 4

64

2100ReRe632.13/1

<

= if

Ld

ScSh e (A3)

35000Re2100Re023.0 3/183.0 <<= ifScSh (A4)

For laminar flow the length (L) in equation A3 has been chosen as the Taylor bubble

length (lTB) or the liquid slug length (lLS) assuming that the building up of concentration and temperature profiles in the falling film and wake zones are independent of those profiles in other slug units. This simplified approach has proved to give good results15.

For the falling film zone the Reynolds number is defined by:

TBL

LeL

LLTBe lLandD

dwithUd

=

−==

δδµ

ρ14Re

(A5)

For the wake zone the Reynolds number is defined by:

LSL

LW lLwithDU

==µ

ρRe

(A6)

The average axial velocity UW in the wake zone is calculated, according to Ghosh

and Cui15, with equation A7 which is based on the axial velocity of a two-dimensional jet that enters a stagnant pool of liquid43.

( )∫ ∫

∫ ∫=LS

LS

l

s

D

l

s

D

LTBWdzdr

dzdrzr

UU 2/

0

2/

0

2 67.7sech

(A7)

Where s is a very small number used because the function inside the integral is not

defined for z=0. The mass transfer coefficients in each zone of a slug unit are combined to obtain an

average mass transfer coefficient using equation A2. This mass transfer coefficient is used to predict an average concentration polarization and fluxes through the membrane in slug flow.


65

4.8. Reference List

(1) Jonquieres, A.; Clement, R.; Lochon, P.; Neel, J.; Chretien, B.; Dresch, M. Industrial state-of-the-art of pervaporation and vapor permeation in the western countries. J. Membr. Sci. 2002, 206, 87 - 117

(2) Ito, A.; Feng, Y.; Sasaki, H. Temperature drop of feed liquid during pervaporation. J. Membr. Sci. 1997, 133, 95 - 102

(3) Bhattacharya, S.; Hwang, S. T. Concentration polarization, separation factor, and peclet number in membrane processes. J. Membr. Sci. 1997, 132, 73 - 90

(4) Baker, R. W.; Wijmans, J. G.; Athayde, A. L.; Daniels, R.; Ly, J. H.; Le, M. The effect of concentration polarization on the separation of volatile organic compounds from water by pervaporation. J. Membr. Sci. 12-24-1997, 137, 159 - 172

(5) Mi, L.; Hwang, S. T. Correlation of concentration polarization and hydrodynamic parameters in hollow fiber modules. J. Membr. Sci. 1999, 159, 143 - 165

(6) Michaels, A. S. effects of feed side solute polarization on pervaporative strippinf of volatile organic solutes from diluted aqueous solution: a generalized analytical treatment. J. Membr. Sci. 1995, 101, 117 - 126

(7) Psaume, R.; Aptel, Ph.; Aurelle, Y.; Mora, J. C.; Bersillon, J. L. Pervaporation: importance of concentration polarization in the extraction of trace organics from water. J. Membr. Sci. 1988, 36, 373 - 384

(8) Favre, E. Temperature polarization in Pervaporation. Desalination. 2003, 154, 129 - 138

(9) Rautenbach, R.; Albrecht, R. The separation Potential of pervaporation. Part 2. Process design and economics. J. Membr. Sci. 1985, 25, 25 - 54



(12) Schleger, M.; Sommer, S.; Melin, T. Module arrangement for solvent dehydration

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66

with silica membranes. Desalination. 2004, 163, 281 - 286

(13) Cui, Z. F.; Chang, S.; Fane, A. G. The use of gas bubbling to enhance membrane processes. J. Membr. Sci. 2003, 221, 1 - 35

(14) Taha, T.; Cui, Z. F. Hydrodynamic analysis of upward slug flow in tubular membranes. Desalination. 2002, 145, 179 - 182

(15) Ghosh, R.; Cui, Z. F. Mass transfer in gas sparged ultrafiltration: upward slug flow in tubular membranes. J. Membr. Sci. 1999, 162, 91 - 102

(16) Cui, Z. F.; Wright, K. I. T. Gas liquid two phase cross flow ultrafiltration of BSA and dextran solutions. J. Membr. Sci. 1994, 90, 183 - 189

(17) Cheng, T. W.; Yeh, H. M.; Gau, C. T. Enhancement of permeate flux by gas slugs for crossflow ultrafiltration in tubular membrane module. Sep. Sci. Tech. 1998, 33, 2295 - 2309

(18) Vera, L.; Delgado, S.; Elmaleh, S. Dimensionless numbers for the steady state flux of cross flow microfiltration and ultrafiltration with gas sparging. Chem. Eng. Sci. 2000, 55, 3419 - 3428

(19) Watanabe, K. and Fuchigami, Y. Separation of liquid mixtures. EP0294827, 1988.

(20) Peters, T. A.; Fontalvo, J.; Vorstman, M. A. G.; Keurentjes, J. T. F. Design directions for composite catalytic hollow fibre membranes for condensation reactions. Chem. Eng. Res. Des. 2004, 82, 220 - 228

(21) Taitel, Y.; Bornea, D.; Duckler, A. E. Modelling flow pattern transitions for steady Upward gas-liquid flow in vertical tubes. AIChE Journal. 1980, 26, 345 - 354

(22) Fernandes, R. C.; Semiat, R.; Duckler, A. E. Hydrodynamic model for gas-liquid slug flow in vertical tubes. AIChE Journal. 1983, 29, 981 - 989

(23) Taha, T.; Cui, Z. F. CFD modelling of gas sparged ultrafiltration in tubular membranes. J. Membr. Sci. 2002, 210, 13 - 27

(24) Polonsky, S.; Shemer, L.; Barnea, D. The relation between the Taylor bubble motion and the velocity field ahead of it. International Journal of Multiphase Flow. 1999, 25, 957 - 975

(25) Pinto, A. M. F. R.; Coelho, M. N.; Campos, J. B. L. On the interaction of Taylor bubbles rising in two phase co current slug flow in vertical columns: turbulent wakes.


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(26) Fabre, J.; Liné, A. Modeling of two-phase slug flow. Ann. Rev. Fluid Mech. 1992, 24, 21 - 46

(27) Collins, R.; Moraes, F. F.; Davidson, J. F.; Harrison, D. The motion of a large gas bubble rising through liquid flowing in a tube. Journal of Fluid Mechanics. 1978, 89, 497 - 514

(28) Abdul-Majeed, G. H.; Al-Mashat, A. M. A mechanistic model for vertical and inclined two phase slug flow. Journal of Petroleum Science and Engineering. 2000, 27, 59 - 67

(29) Barnea, D. Effect of bubble shape on pressure drop calculations in vertical slug flow. International Journal of Multiphase Flow. 1990, 16, 79 - 89

(30) Brauner, N.; Ullmann, A. Modelling of gas entrainment from Taylor bubbles. Part A. Slug flow. International Journal of Multiphase Flow. 2004, 30, 239 - 272

(31) Wallis, G. B. One-dimensional two phase flow; McGraw-Hill: New York, 1969.

(32) Yun, J.; Shen, Z.; Ming, P. Wall liquid Mass transfer for Taylor bubbles rising through liquid in a vertical tube. Chin. J. Chem. Eng. 2002, 10, 404 - 410

(33) Flaschel, E.; Wandrey, C.; Kula, M. R. Ultrafiltration for the separation of Biocatalysts. Fiechter, A. Springer: Berlin, 1983.

(34) Skelland, A. H. P. Diffusional mass transfer; John Wiley & Sons Inc :1974.

(35) Taylor, S.; Krishna, R. Multicomponent mass transfer; John Wiley & sons: New York, 1993.

(36) Krishna, R. A generalized film model for mass transfer in non-ideal fluid mixtures. Chem. Eng. Sci. 1977, 32, 659 - 667

(37) Taylor, R.; Kooijman H. Composition derivatives of activity coefficient models for estimation of thermodynamic factors in diffusion. Chem. Eng. Comm. 1991, 102, 87 - 106

(38) Mori, H.; Oda, A.; Ito, C.; Aragaki, T.; Liu, F. Z. Thermodynamic factors derived from group contribution activity coefficient models. J. Chem. Eng. of Japan. 1996 , 29, 396 - 398

Chapter 4

68

(39) Li, Q. Y.; Cui, Z. F.; Pepper, D. S. Effect of bubble size and frequency on the permeate flux of gas sparged ultrafiltration with tubular membranes. Chem. Eng. J. 1997, 67, 71 - 75


(41) Verkerk, A. W.; Male, P.; Vorstman, M. A. G.; Keurentjes, J. T. F. Description of dehydration performance of amorphous silica pervaporation membranes. J. Membr. Sci. 2001, 193, 227 - 238

(42) Perry, R. H.; Green, D. W. Chemical Engineers' Handbook; McGraw-Hill:1999.

(43) Bird, R. B.; Stewart, W. E.; Lightfoot, E. N. Transport Phenomena; John Wiley & Sons: New York, 2002.

________________________________ This chapter has been accepted for publication in Ind. Eng. Chem. Res.: J. Fontalvo, M. A. G. Vorstman, J. G. Wijers, and J. T. F. Keurentjes. (2005).

5. Separation of organic-water mixtures by co-current vapor-liquid

pervaporation with transversal hollow fibre membranes

Abstract The advantages of a liquid-vapor feed in a pervaporation unit are presented

in comparison to a single phase feed pervaporation unit that uses inter-stage heating. The comparison is based on calculations for the dehydration of isopropyl alcohol (IPA) from 13 wt. % to 1 wt. % in a transversal system with hollow fibre membranes. Due to the presence of the vapor phase high mass and heat transfer coefficients in combination with low pressure drops are achieved. Vapor also supplies the heat for the evaporation and expansion of the permeating components through the membrane. Mass transfer coefficients from the liquid to the membrane surface have been measured using an electrochemical method. At low liquid superficial velocities mass transfer increases up to 4 fold as compared with single-phase flow. Both the enhancement in mass and heat transfer to the membrane and the energy supply to the liquid results in a reduction of the required membrane area of 45%. This reduction in membrane area combined with the avoidance of inter-stage heat exchangers results in a strong reduction of capital cost, thus significantly expanding the application window of pervaporation.

Chapter 5

70

5.1. Introduction Pervaporation is becoming an important alternative for separation of liquid mixtures

because it can strongly reduce energy consumption as compared to conventional separation processes such as distillation1-3. However, drawbacks of pervaporation include high-pressure drops, concentration and temperature polarization and retentate temperature drop. When these drawbacks are reduced or eliminated the capital cost associated with pervaporation will be reduced and the application window can be expanded. The use of a liquid-vapor feed flow in combination with a transversal pervaporation unit with hollow fibre membranes is suggested to overcome the disadvantages of pervaporation processes with inter-stage heating.

Hollow fibre membranes are attractive because they combine relatively fast and

inexpensive production of large membrane surface areas, a low support resistance to the mass transfer and packing densities higher than 500 m2/m3. Externally coated hollow fibre membranes operating in transversal flow offer lower pressure drops than multi-tubular or flat sheet pervaporation membranes. For single phase pervaporation with the same membrane area, Figure 5.1 shows calculated pressure drops and mass transfer coefficients in a transversal membrane module4 with liquid flowing externally around the hollow fibres and in a multi-tubular pervaporation module5 with liquid flowing inside each tube. From the figure it will be clear that pressure drops in multi-tubular pervaporation units are at least one order of magnitude higher than in transversal pervaporation units.

Figure 5.1. Pressure drop and mass transfer calculated for transversal flow externally to hollow fibres and inside of a pervaporation tube for dewatering of IPA at several liquid flow rates. Pervaporation tube with 7 mm internal diameter. Hollow fibres of 3 mm external diameter in parallel in line-configuration. Calculated from Futselaar4 and Oliveira et al.5

Separation of organic-water mixtures by co-current vapor-liquid pervaporation with transversal hollow fibre membranes

71

In a pervaporation unit the heat carried by the liquid supplies the required energy for

vaporization and expansion of the permeating components. The resulting temperature drop in the liquid reduces the flux, thus increasing the membrane area that is required for a specific separation duty, as well as the associated operating and capital cost. The retentate temperature drop in pervaporation modules has been reduced by electrical resistances6, by external heating with another fluid7 or by using a series of alternating pervaporation modules and heat exchangers8. Electrical resistances are restricted to non-flammable media and external heating in multi-tubular modules requires a system of concentric tubes where the heating fluid is fed to one of the lumens. A series of heat exchangers increases the capital cost, piping and membrane area. The required membrane area with inter-stage heating is around 40% higher1,9 than in ideal isothermal operation. In general these pervaporation modules have a complex construction increasing the total cost of the unit.

A liquid-vapor feed operation offers an effective and simple way to diminish the

retentate temperature drop in pervaporation units while it simultaneously reduces polarization effects. For this, a fraction of the liquid product or feed is vaporized and introduced simultaneously with the liquid feed to the pervaporation unit. The vapor phase not only reduces the liquid temperature drop by supplying heat by condensation but also reduces concentration and temperature polarization, as we have shown experimentally before10.

This chapter is focused on demonstrating the advantages of a two-phase transversal

pervaporation system compared to a transversal pervaporation system with inter-stage heating. To calculate the concentrations and temperatures in the liquid and vapor phase and the required membrane area for a specific dewatering duty, a computer program has been built using a rate-based model. The effect of two-phase flow on the mass transfer between the liquid and the membrane surface has been measured using an electrochemical method. The resulting correlation is included in the computer program. The mass transfer coefficients between the liquid and the vapor phase have been obtained from literature.

Chapter 5

72

5.2. Description of the pervaporation module and the simulations

A transversal module using ceramic hollow fibres with a cross-in-line configuration has been chosen for the simulations of the liquid-vapor two-phase flow pervaporator and the single-phase flow pervaporator with inter-stage heating. Two crossed rows of hollow fibres of the pervaporation unit are shown in Figure 5.2a. A packed unit, shown in Figure 5.2b, consists of rows of hollow fibres placed above each other. The fibres, with an outside diameter of 3 mm and a length of 40 cm, are externally coated with a silica layer11 and have an arrangement with longitudinal and transversal pitches of 5.0 and 4.5 mm, respectively.

Figure 5.2. Schematic representation of two layers of pervaporation hollow fibres in cross configuration (a) and packed section with hollow fibres in cross in line configuration (b).

For simulation of single phase pervaporation modules, inter-stage heat exchangers

are located between identical pervaporation units with liquid flowing perpendicular at the outside of the hollow fibres. In each heat exchanger the retentate is reheated to the feed temperature. In the two-phase pervaporation unit liquid and vapor flow in co-current upwards, external and perpendicular to the hollow fibres. Vapor is fed together with the liquid feed and additional vapor is injected at several axial positions. An additional injection takes place when the vapor flow rate, due to condensation, reaches a preselected minimum value. The total amount of vapor to be used depends on the removal duty of the pervaporation unit. The total mass flow, considering vapor and liquid, fed to the pervaporation systems is the same for both with inter-stage heating and in two-phase flow. The vapor superficial velocity can be modified by variation of the amount of


73

vapor in each injection and by the number of injections. Relatively low vapor superficial velocities ( 170G <Re ) have been considered in the calculations to assure complete

wetting of the hollow fibres. In the computations it is assumed that the fibres are connected to a shell where vacuum is applied to the internal side of the tubes and that the obtained permeate stream is removed and condensed.

Concentrations of every component and temperatures in the liquid and vapor

phase are calculated with a rate-based model12 for two-phase pervaporation and single phase simulations. Figure 5.3 shows a schematic representation of a differential element of the transversal unit for two-phase flow. Differential elements of 1.3 mm length were used. The outlet temperatures and flows from each differential element are calculated from the mass and energy balances for each phase. Only the mass transfer between the liquid and the membrane surface and between the membrane surface and the permeate side were taken into account for the simulation of the single phase pervaporator.

Figure 5.3. Differential element of a transversal unit for a co-current liquid-vapor two phase pervaporation.

The Maxwell-Stefan relations for non-ideal mixtures13 are used to describe the mass

transport at the vapor-liquid interface and at the liquid-membrane interface. The approximate method of Krishna14 is used by considering the Maxwell-Stefan diffusivity and the thermodynamic factors constant along the diffusion path in the mass transfer

Chapter 5

74

boundary layer. The mass transfer coefficients in the gas and liquid phases on the vapor-liquid interface are calculated according to Fukushima and Kusaka15 using the churn regime for the calculation of the specific interface area. The mass transfer coefficient between the liquid and the membrane surface has been calculated using the correlation obtained from experimental data in this study that is presented in the “results” section. The thermodynamic factors16,17, activity and fugacity coefficients are calculated by UNIFAC and the Peng-Robinson EOS. Fluxes through the membrane have been calculated as the product of permeability and local partial pressure difference over the membrane.

Simulations of liquid-vapor two phase pervaporation units and single phase

pervaporation units with inter-stage heating were performed for the dewatering of a liquid feed of 1.5 ton/h18 of IPA from a water concentration close to the azeotropic point of 13.0 wt. % to a water concentration of 1.0 wt. % at a feed temperature of 353 K. The vapor, obtained from partial vaporization of the liquid feed, is distributed between the liquid feed and the additional injections. A vapor flow of 94 kg/h, corresponding to a gas phase Reynolds number ( GRe ) of 102, is added to the liquid feed. An additional injection takes

place when the amount of vapor in the system drops to 10 kg/h, corresponding to a GRe of

9. In total 6 additional injections are used, in which the vapor flow for each injection is 90% of the flow of the previous injection. The retentate pressure is 1.013 bar and the permeate pressure has been chosen as 80 mbar in order to use cooling water for condensing the permeate stream. A water and IPA permeability of 2.16 x 10-4 and 1.38 x 10-7 kmol/m2 s bar, respectively, have been used19.

5.3. Solid-liquid mass transfer measurements and gas void fractions

An illustration of the equipment for measuring mass transfer coefficients between the liquid and the membrane surface in gas-liquid two phase flow is depicted in Figure 5.4. Air and liquid were distributed uniformly over the column section area in the inlet section. The air distributor consisted of a square plate with 23 holes of 2 mm diameter in triangular configuration. The configuration of the packed section is the same as described in the simulations but no connection with vacuum was used. The packed section contained 23 rows of ceramic hollow fibres supplied by TNO (The Netherlands) with a length of 3.6 cm each and an external diameter of 3 mm.

The average gas void fraction in the column was measured on a mass basis, and

was determined by the weighing the column. The mass transfer coefficient was measured using a conventional electrochemical method20. The electrochemical reaction was the


75

cathodic reduction of ferricyanide ions. In order to carry out the electrochemical reaction the complete 4th row from above, containing 8 ceramic tubes, was replaced by nickel tubes acting as the anode and two middle tubes in the row above were replaced by nickel tubes acting as the working electrode. The reference electrode was placed on top of the outlet section with the tip at approximately 1 cm from the working electrode.

Figure 5.4. Experimental setup used for measurement of liquid-solid mass transfer coefficients in two phase flow by reduction of ferricyanide ions.

Aqueous solutions with different concentrations of Emkarox HV45 (C.H.Erbsloh

Benelux B.V, The Netherlands) were used to measure the influence of the Schmidt number, by modifying the solution viscosity, on the mass transfer coefficient. Water-Emkarox solutions are characterized by a Newtonian behavior21,22. Table 5.1 shows the Emkarox concentrations prepared, the corresponding Schmidt numbers and the ferricyanide ion diffusion coefficients at the experimental temperature of 21 °C. Schmidt numbers were calculated based on data reported by Vachon et al.21. Diffusivities were calculated using the Stokes-Einstein relation20 given in equation 5.1.

Ksmkg10312 2

15

⋅⋅×= −.

TDµ

5.1

Chapter 5

76

Table 5.1. Physical properties of the several working solutions prepared for mass transfer coefficient measurements at 21 °C.

Weight fraction of

EMKAROX HV45

D *1010,

m2/s

ν *106, m2/s Sc

0 7.08 0.96 1362

0.03 4.96 1.41 2773

0.06 3.97 1.82 4317

The electrolyte solution consisted of 0.01 mol/kg K3Fe(CN)6 and, in order to prevent

anodic polarization, 0.02 mol/kg K4Fe(CN)6 in combination with 0.5 mol/kg of NaOH. Air, saturated with water, and the liquid solution were fed at the bottom of the system in upward flow. Liquid superficial velocities between 0.46 and 3 cm/s and air superficial velocities between 0 and 0.26 m/s were used for gas void fraction and mass transfer coefficient measurements. The current, measured online using a computer interface, fluctuated due to velocity oscillations in the liquid phase produced by the gas stream. The current was averaged to determine the limiting-current curves for the various liquid and gas velocities, from which the limiting current and the mass transfer coefficients were calculated.

5.4. Results

5.4.1. Gas void fractions Experimental data of gas void fractions as a function of the Martinelli parameter

( ttX , equation 5.2) for two Froude numbers ( Fr , equation 5.3) are shown in Figure 5.1. Gas

void fractions increase at a decreasing Martinelli parameter and increasing Froude number. Equation 5.4 is a fit of the experimental data as a function of the Martinelli parameter and the Froude number. The equivalent diameter ( ed ) is calculated with

equation 5.5. 50

L

G

11

G

L

90

G

Ltt

...

ReReX

=

ρρ

µµ

5.2

eL

2

gdGFr

ρ= 5.3

15591tt

43640

G

G 12126 .. XFr. −=−ΦεΦ

5.4


77

te 14 dabd

−=

π

5.5

Figure 5.5. Experimental values of gas void fraction as a function of the Martinelli parameter for two Froude numbers. Lines have been calculated with equation 4.

Xu et al.23 suggest that a correlation in the form of equation 5.4 is independent of the

flow pattern. For the range of air and liquid flow rates used in the experiments in this study only the churn regime has been observed. The prediction of the gas void fraction using equation 4 has a maximum error of ±10% for 98% of the experimental data.

5.4.2. Mass transfer coefficients Experimental values of mass transfer coefficients expressed as Sherwood numbers

are presented in Figure 5.6. The experimentally determined liquid-solid mass transfer coefficients in single phase ( oSh ) have been compared in Figure 5.6a with the values

calculated using the correlation of Futselaar4 given by equation 5.7. Good agreement has been found between the experimental data and equation 6. The ratio of the Sherwood number in two-phase flow ( Sh ) to the Sherwood number in single phase flow ( oSh ) is

presented in equation 5.7 from a fit to the experimental data as a function of gas void fraction and Reynolds number for the gas phase24,25.

100020 with 7720 L342131

L5180

Lo ≤≤= RebScRe.Sh ./. 5.6

Chapter 5

78

170with 1031 GG

1585

G

G8

o

≤

−

×+= ReReShSh

.

ΦεΦ

5.7

Figure 5.6. Comparison between experimental and predicted values for Sherwood numbers in a transversal pervaporation module using hollow fibres: a) single phase. b) air-liquid two phase flow.

A comparison between experimental and predicted values using equation 5.7 is

shown in Figure 5.6b. Mass transfer coefficients between the liquid and the hollow fibres in the two phase regime are increased up to 4 fold as compared with single phase flow, due to the promoted turbulence in the liquid by the gas phase. Equation 5.7 has been used in the simulations in the next section and has a deviation in the prediction of mass transfer coefficients for two phase flow of about 15% for 98% of the experimental data.


79

5.5. Comparing single phase and two-phase flow pervaporation units

5.5.1. Single phase pervaporation units with inter-stage heating The required membrane area in a single phase transversal pervaporation unit to

dewater 1.5 ton/h of IPA, from 13.0 wt. % to 1.0 wt. % at a feed temperature of 353 K, is shown in Figure 5.7a as a function of the number of inter-stage heat exchangers. The required membrane area increases as the number of inter-stage heat exchangers decreases due to the increasing temperature drop in the retentate. The required removal duty cannot be reached when no heat exchangers are used.

Figure 5.7. a) Total required membrane area for dewatering of 1.5 ton/h of IPA from 13 wt. % to 1 wt. % as function of the number of inter-stage heat exchangers used in a conventional pervaporation unit. b) Axial temperature profile in the liquid and axial water flux profile with 6 inter-stage heat exchangers.

Chapter 5

80

The economical optimal number of inter-stage heat exchangers1,9 is around 6 with

an optimum required membrane area of 520 m2. The axial liquid temperature and water flux profiles are shown in Figure 5.7b. The temperature drop is 35 K for the first module and decreases to 3 K for the last module. Low temperature drops are found at the end of the pervaporation unit where only a small amount of water is removed. The concentration polarization, defined as the fractional decrease of the water concentration between the bulk and the membrane surface, is 15% and 2% at the inlet and outlet position of the pervaporation module, respectively. The maximum temperature difference between the bulk and the membrane surface is 2.2 K at the inlet of the pervaporation unit. The flux decreases in the axial direction and as a consequence the polarization effects are reduced.

5.5.2. Liquid-vapor two-phase pervaporation unit Vapor and liquid flows are presented in Figure 5.8 as a function of the axial position in

a two-phase transversal pervaporation unit. Between injections vapor condenses until the preselected minimum amount of vapor of 10 kg/h is reached. This amount of vapor keeps the mass transfer coefficient to a level of around twice the mass transfer coefficient in single phase pervaporation. The vapor condensation rate decreases in the axial direction due to the reduction of the flux through the membrane.

Figure 5.8. Vapor and liquid flow rates as function of the axial position for dewatering of IPA in a liquid-vapor two-phase flow pervaporation module. Conditions as in Figure 5.7.


81

The axial liquid temperature profile and the axial profile of the water flux through the membrane in the pervaporation unit are shown in Figure 5.9. The axial liquid temperature closely follows the corresponding bubble temperatures at the different axial concentrations. At these temperatures and concentrations the maximum driving force for water transport through the membrane is achieved. The axial water fluxes are higher than in a single pervaporation unit with inter-stage heating as shown in Figure 5.7b. The maximum concentration and temperature polarization calculated are 10% and 1 K, respectively. These values are lower than for single phase pervaporation with inter-stage heating, for which values of 15% and 2 K have been determined, respectively.

Figure 5.9. Axial liquid, vapor and membrane surface temperature profiles and water flux profile in a two phase pervaporation module using 6 vapor injections. Total amount of liquid and vapor used is 1.5 ton/h. Inlet temperatures and dewatering duty as in Figure 5.7.

In an ideal pervaporation unit no vapor will remain at the top of the pervaporation

unit. As a consequence, no additional energy is required compared to a single phase pervaporation unit with inter-stage heat exchangers. The remaining amount of vapor can be calculated from Figure 5.8 as 4% of the total outlet mass flow. When the amount of supplied vapor in the last injection is reduced by 80% no vapor remains at the top of the pervaporation unit and no additional energy consumption takes place. The required membrane area for a two-phase pervaporation unit for dewatering of 1.5 ton/h of an IPA-water mixture from a water concentration of 13 wt. % to 1 wt. % is around 300 m2. This is about 45% lower than the required membrane area of 520 m2 for a pervaporation unit with inter-stage heating.

Chapter 5

82

5.6. Conclusions Gas void fractions and mass transfer coefficients have been measured in a

transversal pervaporation module as a function of the liquid and gas flow rates at several Schmidt numbers. Mass transport to the membrane surface increases up to 4 fold due to the turbulence induced in the liquid by the vapor phase compared to single phase flow at low liquid flow rates. A correlation for mass transfer coefficients, based on the gas void fraction, was obtained from the experimental data. This correlation has been implemented in a computer program that uses a rate-based model, to simulate liquid-vapor two-phase transversal pervaporation units and single phase transversal pervaporation units with inter-stage heating. With the computer program the axial temperature profile in the liquid and vapor phases, the axial profile of the water flux through the membrane and the required membrane area have been calculated for dewatering of 1.5 ton/h of an IPA-water mixture from a water content of 13 wt. % to 1 wt. %.

High mass transfer coefficients are obtained in the liquid-vapor two-phase transversal

pervaporation module due to the turbulence induced by the vapor flow that also supplies energy to the liquid by condensation. The calculated maximum concentration and temperature polarization is 10% and 1 K, respectively. These polarization effects are lower than for single phase pervaporation modules with inter-stage heating for which values of 15% and 2 K have been determined, respectively.

The reduction of both the polarization effects and the liquid temperature drop results

in a reduction of about 45% in the required membrane area compared to a single phase transversal pervaporation module that uses an economical optimal number of heat exchangers. The lower required membrane area and the avoidance of inter-stage heat exchangers results in a strong reduction of capital cost that can expand the application window of pervaporation.

5.7. Notation

t

td

sa = = ratio between transversal pitch and tube diameter

t

td

sb = = ratio between longitudinal pitch and tube diameter

td = tube diameter, m

te 14 dabd

−=

π= equivalent diameter, m


83

D = diffusion coefficient in the liquid phase, m2/s e = energy flux through the interface, kJ/m2 s f = molar feed flow kmol/s

Fr = Froude number, define in eq. 3. g = gravity acceleration, 9.81 m/s2

G = total mass flow, kg/m2 s H = enthalpy, kJ/s k = mass transfer coefficient for two-phase flow, m/s

ok = mass transfer coefficient for single phase flow, m/s

L = liquid flow rate, kmol/s N = molar flux through the interface, kmol/m2 s Q = Heat losses, kJ/h m2

µρeeudRe = = Reynolds number

DdkSh eo

o = = Sherwood number for single phase flow

DdkSh e= = Sherwood number for two-phase flow

Sc = Schmidt number T = temperature, K

−

=

ab

uu

41

se π

= equivalent velocity, m/s

su = superficial velocity, m/s

V = vapor flow rate, kmol/s x = molar fraction in the liquid phase

ttX = Martinelli parameter, defined in eq. 2.

y = molar fraction in the vapor phase

Greek symbols ε = packing void fraction, m3/m3

GΦ = gas void fraction, m3/m3

iΓ = permeance of component i, kmol/m2 s bar

µ = viscosity, kg/m s

ν = kinematic viscosity, m2/s ρ = density, kg/m3

Chapter 5

84

Subscripts L = liquid G = gas phase e = equivalent s = superficial i = component “i” j = differential element “j”

Superscripts L = liquid phase V = vapor phase VF = feed in vapor phase LF = feed in liquid phase P = liquid-membrane interface

5.8. Reference list (1) Fontalvo, J.; Cuellar, P.; Timmer, J. M. K.; Vorstman, M. A. G.; Wijers, J. G.; Keurentjes,

J. T. F. Comparing pervaporation and vapor permeation hybrid distillation processes. Ind. Eng. Chem. Res. 2005, 44, 5259 - 5266



(4) Futselaar, H., Ph.D. Thesis/Dissertation, Universiteit Twente,1993.

(5) Oliveira, T. A. C.; Cocchini, U.; Scarpello, J. T.; Livingston, A. G. Pervaporation mass transfer with liquid flow in the transition regime. J. Membr. Sci. 2001, 183, 119 - 133


(7) Schleger, M.; Sommer, S.; Melin, T. Module arrangement for solvent dehydration with silica membranes. Desalination. 2004, 163, 281 - 286

(8) Rautenbach, R.; Albrecht, R. The separation Potential of pervaporation. Part 2. Process design and economics. J. Membr. Sci. 1985, 25, 25 - 54


85


(10) Fontalvo, J., Vorstman, M. A. G., Wijers, J. G., and Keurentjes, J. T. F. Heat supply and reduction of polarization effects in pervaporation by two-phase feed. Submitted. 2005,

(11) Peters, T. A.; Fontalvo, J.; Vorstman, M. A. G.; Benes, N. E.; Van Dam, R. A.; Vroon, Z. A. E. P.; Van Soest-Vercammen, E. L. J.; Keurentjes, J. T. F. Hollow fibre microporous silica membranes for gas separation and pervaporation. Synthesis, performance and stability. J. Membr. Sci. 2005, 248, 73 - 80

(12) Krishnamurthy, R.; Taylor, R. A nonequilibrium model of multicomponent separation processes. Part I: Model description and method of solution. AIChE J. 1985, 31, 451 - 456



(15) Fukushima, S.; Kusaka, K. Gas liquid mass transfer and hydrodynamic flow region in packed columns with cocurrent upward flow. J. Chem. Eng. of Japan. 1979, 12, 296 - 301



(18) Sommer, S.; Melin, T. Design and optimization of hybrid separation processes for the dehydration of 2-propanol and other organics. Ind. Eng. Chem. Res. 2004, 43, 5248 - 5259

Chapter 5

86


(20) Selman, J. R.; Tobias, C. W. Mass transfer measurements by the limiting-current technique. Advances in Chemical Engineering. 1978, 10, 211 - 318

(21) Vachon, G.; Legentilhomme, P.; Lamer, T. Experimental determination of mass transfer at the active surface of a membrane pervaporation cell. Chem. Eng. Tech. 2000, 23, 243 - 249

(22) Legrand, J.; Dumont, E.; Comiti, J.; Fayolle, F. Diffusion coefficients of ferricyanide ions in polymeric solutions- comparison of different experimental methods. Electrochimica Acta. 2000, 45, 1791 - 1803

(23) Xu, G. P.; Tso, C. P.; Tou, K. W. Hydrodynamics of two- phase flow in vertical up and down flow across a horizontal tube bundle. International Journal of Multiphase Flow. 1998, 24, 1317 - 1342

(24) Mochizuki, S.; Matsui, T. Liquid-solid mass transfer rate in liquid - gas upward cocurrent flow in packed beds. Chem. Eng. Sci. 1974, 29, 1328 - 1330

(25) Specchia, V.; Baldi, G.; Gianetto, A. Solid liquid mass transfer in concurrent two-phase flow through packed beds. Ind. Eng. Chem. Process Des. Dev. 1978, 17, 362 - 367

________________________________ This chapter has been published: J. Fontalvo, P. Cuellar, J. M. K. Timmer, M. A. G. vorstman, J. G. Wijers, and J. T. F. Keurentjes. Ind. Eng. Chem. Res. 44 (2005) 5259-5266.

6. Comparing pervaporation and vapor permeation hybrid distillation

processes

Abstract Previous studies have shown that hybrid distillation processes using either

pervaporation or vapor permeation can be very attractive for the separation of mixtures. In this chapter, a comparison between these two hybrid processes has been made. A tool has been presented that can assist designers and engineers to decide which process is more convenient for a specific application. Water removal from acetonitrile (ACN) has been used as an example. A hybrid process with vapor permeation is preferred when the membrane is used either for water removal at high water concentration or just for overcoming the azeotropic composition. When the membrane removes water at water concentrations lower than the azeotropic point, pervaporation is more effective. Recycling part of the product as permeate (product sweep) and applying different pressures in the distillation columns and the membrane unit strongly reduce the required membrane area and the total cost of the process. Relatively low membrane selectivities are required for an economically optimal hybrid membrane - distillation process.

Chapter 6

88

6.1. Introduction The chemical industry is looking for economical ways to reduce energy demand,

and to avoid auxiliary agents for the separation of mixtures. Pervaporation (PV) and vapor permeation (VP) are highly suitable to separate close boiling mixtures1,2 and in breaking azeotropic mixtures where distillation is energy intensive or requires the use of an auxiliary agent. Distillation is a well-known operation with lower capital cost than membrane operations. Hybrid membrane-distillation processes exploit the advantages of distillation and membrane operations, while overcoming the disadvantages of both3.

Previous studies have shown the advantages of distillation processes in combination

with either pervaporation4-7 or vapor permeation8,9 for various mixtures. Strategies for the optimization of hybrid processes have been developed in the past1,5,10. However, the comparison between the various types of hybrid processes on the same calculation basis can help designers and engineers to decide which hybrid process is more convenient for a specific application.

This work presents an economical and technological comparison between

distillation - vapor permeation and distillation – pervaporation hybrid processes using silica membranes. Silica membranes are of interest for the dehydration of several solvents because they are resistant to harsh conditions11,12, high temperatures and high pressures. As a model case the dewatering of acetonitrile (ACN) is studied.

The model assumptions for the simulations are described followed by results for

conventional distillation based processes and hybrid processes using pervaporation or vapor permeation. Product sweep, high membrane operating pressures combined with low distillation pressures and relatively low selectivities are included to reduce the total cost of a hybrid process. This study shows the influence of these parameters on the required membrane area, total energy consumption and total cost of the hybrid process. Guidelines for the selection of PV or VP are suggested followed by an economical evaluation.

6.2. Process modeling Conventional and hybrid processes were simulated for the separation of 16.000

ton/year of 50 wt. % ACN-water mixture at 1 atm and 333 K. The purity of the two product streams is set to 99.9 wt. %. ASPEN PLUS 11.1 (AspenTech) was used for the simulations of the distillation tower and other basic equipment in connection with a rigorous subroutine

Comparing pervaporation and vapor permeation hybrid distillation processes

89

(written by us in MATLAB ® - The Mathworks Inc.) describing the behavior of the membrane unit including concentration and temperature polarization. The size of the columns, the reboiler and condenser duties, and the water composition in the column and in the outlet streams were calculated using the subroutine RADFRAC in ASPEN PLUS 11.1 (AspenTech). For the membrane unit the outlet temperatures and final water concentrations, retentate total pressure drop and total membrane area were calculated using the MATLAB subroutine. The economic evaluation was performed according to Appendix A presented at the end of this chapter.

The pervaporation and vapor permeation units consist of a set of membrane

modules (and heat exchangers in the case of pervaporation) connected in series. Each membrane module has several ceramic tubes mounted in a shell like a shell & tube heat exchanger. The silica separation layer is located on the inner side of the 7 mm internal diameter tube. Liquid in pervaporation or vapor in vapor permeation is fed to the inside of the tubes and reduced pressure is applied at the shell side.

The developed MATLAB program allows calculating a series of pervaporation units

where the retentate temperature drop or the membrane area per module is fixed. Co-current or counter-current operation can be simulated but only results for co-current operation between permeate and retentate flows will be presented. The software accounts for local concentration and temperature polarization in the retentate by linking the mass transfer in the retentate (equation 6.1) with the mass and heat flux through the membrane (equations 6.2, 6.3). The mass transfer in the retentate (equation 6.1) was calculated using a Maxwell-Stefan equation expressed in difference form13. Conventional correlations for flow inside pipes were used for the evaluation of the local mass and heat transfer coefficients. Fluxes through the membrane were computed using equation 6.1 where the flux for water11 and ACN are proportional to its vapor pressure difference between retentate and permeate. Activity coefficients ( iγ ) were obtained by UNIFAC

and the Poynting factor was taken into account to calculate the vapor pressures of each compound in the retentate.

∑≠

−=

∆−

jiji

jiij

i

ii

Ck

NxNxx _

,γγ

6.1

( )GiviIiIiIii PyPxN −Γ= γ

6.2

( ) ∑∑ =+−=i

iGii

iLiILL HNHNTThq

6.3

Chapter 6

90

( )( )IACNIw

IACNIw

xxyy

//

=α

6.4

For VP the concentration and temperature polarization were neglected because of

the low mass and heat transfer resistance in the vapor phase. Analogous to pervaporation the flux was calculated proportional to the difference in vapor pressures. The partial vapor pressures in the retentate were calculated by means of fugacity coefficients using the Peng Robinson EOS. The permeate pressure was set at 0.11 atm for both pervaporation and vapor permeation processes. At this pressure the permeate stream can be condensed using cooling water.

Values for the water permeance (Γwater) and the membrane separation factor (α )

defined in equation 6.4 were, as a first assumption, taken as constant in the simulations. These values were calculated from typical data supplied by PERVATECH BV, The Netherlands (Table 6.1)

Table 6.1. Total flux and separation factor for the system ACN-water using silica membranes.

Water wt. % 9 Temperature, K 343 Permeate pressure, mbar 17 Total flux, kg/ m2 h 3.9 Separation factor, α 331 Limiting operating temperature, K 448 Limiting pressure (module), bar 50

In order to reduce the total cost of the PV process the feed flow rate per membrane

tube was economically optimized. Concentration and temperature polarization and as a consequence capital cost increase as the feed flow rates decrease, whereas increasing the flow rate will increase the pumping cost.

For vapor permeation no optimization of the feed flow was performed because

concentration and temperature polarization are much less significant than for pervaporation processes. The feed flow rate per tube was set to obtain a retentate pressure drop lower than 10 mbar.

The feed temperature of the PV process was taken as the bubble temperature of the

mixture at the feed pressure and composition while for VP it corresponds to the dew


91

temperature. For vapor permeation an overheating of 5 °C is used in the feed stream. The MATLAB subroutine evaluates if a change of phase occurs in the retentate inside the membrane modules; if so, the temperature is modified accordingly.

The retentate temperature drops along a PV unit as a consequence of the

vaporization and expansion of the different components at the permeate side. The heat supply necessary to perform an ideal isothermal operation was calculated. The total cost for an actual non-isothermal PV operation, using a series of inter-stage heat exchangers, is presented in the economic evaluation for comparison with the ideal operation. The retentate temperature drop in VP was neglected because no latent heat is involved in the transport of the components through the membrane.

6.3. Conventional distillation based processes Distillation-based processes for the separation of ACN-water mixtures have to be

designed to overcome the azeotropic composition. This mixture presents a pressure dependent low boiling azeotrope, which is close to 30 molar % water at atmospheric pressure. Two conventional processes are described here, i.e. pressure swing distillation and azeotropic distillation.

The pressure swing distillation process is presented in Figure 6.1. This process consists of

two distillation columns operating at different pressures. The azeotropic composition is shifted to higher water concentrations as the total pressure increases. As a consequence, the two columns together can operate in the whole range of concentrations between the water rich stream and the ACN rich stream. Figure 6.1 shows that both the feed flow to the second distillation column and the recycle stream flow rate are high. This is due to the high water concentration in the azeotrope at the top of the first and second distillation column.

Chapter 6

92

Figure 6.1. Pressure swing process for dehydration of ACN.

Russell14 patented an azeotropic distillation process using benzene as an entrainer for

the separation of ACN – water mixtures, see Figure 6.2. Water and ACN form a ternary azeotrope with benzene. Upon condensation of the ternary vapor, phase separation occurs into a benzene-free phase and an ACN-rich phase. The ACN-rich phase is fed to the second distillation column. The first distillation column produces a water-rich stream and the second distillation column produces an ACN – rich stream. The amount of benzene involved in the process is 350 kg per ton of feed. This amount can be calculated by minimizing the reboiler duties as a function of the benzene flow rate.

Figure 6.2. Azeotropic distillation process for dehydration of ACN using benzene as entrainer.

In the economical evaluation section these processes will be compared with the

hybrid processes.


93

6.4. Membrane – distillation hybrid processes Two different hybrid membrane-distillation processes are of interest for this kind of

azeotropic mixtures in which either distillation or the membrane unit is used as the final dewatering step15. The structural and parametric optimization of these hybrid processes is relatively straightforward because the top composition (and feed to the membrane) is limited by the thermodynamic conditions10. It has been suggested in the literature that the permeate flow should be minimized16. We have found that in these processes the membrane unit should be operated in such a way that the driving force is as high as possible in order to reduce the required membrane area and thus the capital cost. High driving forces are accomplished by using high feed temperatures or by using a product sweep. The required membrane area and energy consumption are analyzed as a function of the final retentate concentration, membrane feed temperature and membrane feed pressure when distillation is used as the final dewatering step. A similar analysis is presented for product sweep and the separation factor when the membrane is used as the final dewatering step. Because qualitatively similar results are expected for PV and VP only results for hybrid PV-distillation processes are presented in this section.

6.4.1. Distillation as the final dewatering step Simulations have been performed for processes where pervaporation is used for

splitting the azeotropic composition in the hybrid pervaporation-distillation process (Figure 6.3). The first distillation column produces a water rich stream and an azeotropic top stream. The liquid top stream is pressurized, heated and fed to a pervaporation membrane that splits the azeotropic composition. Permeate is condensed, pumped and recycled to the first distillation column and the retentate is treated further in a second distillation column. The second distillation column will produce an ACN rich stream and an azeotropic top stream that is recycled to the pervaporation unit.

Figure 6.3. Column-pervaporation-column (CPC) hybrid process for dewatering of ACN

Chapter 6

94

6.4.1.1. Final retentate concentration.

The final retentate concentration after the pervaporation module is a design variable. Figure 6.4 shows the effect of the final retentate water concentration on reboiler duty and required membrane area for a pressure of the feed to the pervaporation unit of 5 atm and a pressure in the columns of 1 atm. Low final water concentrations in the retentate will require large membrane areas and low reboiler duties for the second distillation column. The energy duty required for the first distillation column is approximately constant while the total energy consumption is reduced when the final retentate water concentration decreases resulting in an economically optimal final retentate concentration.

Figure 6.4. Influence of the final retentate water concentration from the pervaporation unit on the membrane area and reboiler duty. Operating pressure of the PV and distillations columns are 5 atm and 1 atm respectively. No product sweep applied. Total energy duty considers the energy supplied to the pervaporation unit and reboilers.

6.4.1.2. Liquid feed temperature and pressure.

To reduce the required membrane area, the driving force for the water transport through the membrane can be increased using high retentate temperatures. To avoid vaporization of liquid feed high pressures are required. Under these conditions the required membrane area decreases linearly with pressure when the feed is at the bubble temperature. The main effect of high temperatures is reducing the required membrane area although a small increase in total energy consumption is involved. High feed temperatures to the membrane unit, but lower than the maximum allowed operation temperature as given in Table 6.1 (448 K), were used in the calculations of all the hybrid membrane-distillation processes. For vapor permeation the maximum operation temperature is obtained at lower pressures than for pervaporation.


95

6.4.2. Membrane as a final dewatering step (column – membrane CM) In this type of system the membrane unit is used for splitting the azeotropic

composition and as a polishing step. Permeate is recycled to the distillation column and retentate is the ACN rich product stream in the configuration given in Figure 6.5a. High temperatures (430 and 448 K) and operating pressures (10 and 15 atm) are necessary in the membrane unit in order to reduce the required membrane area. Low operating pressures in the distillation column, however, are advantageous for reducing the water concentration at the column top stream and so the amount of water to be removed by the membrane unit. Table 6.2 shows temperatures, pressures and flow rates for the calculations performed for the most effective flow diagram (Figure 6.5b).

Figure 6.5. Column-pervaporation (CP) hybrid process for dewatering of ACN. a) Schematic representation. b) Flow diagram for process P2b in Figure 6.10. See Table 6.2 for details.

Chapter 6

96

Table 6.2. Stream properties for the flow diagram of Figure 6.5b, process P2b in Figure 6.10.

Stream Temperature Pressure Vapor fraction Mass flow Mass fractions °C atm Kg/h

1 43.6 0.3 0.0 2000 0.5 0.5 2 43.8 0.3 0.0 1208 0.134 0.866 3 46.5 0.3 0.0 208 0.773 0.227 4 69.2 0.3 0.0 1000 0.999 0.001 5 46.0 15.0 0.0 1208 0.134 0.866 6 23.5 0.3 0.0 208 0.773 0.227 7 23.5 0.1 0.0 208 0.773 0.227 8 175.7 15.0 0.0 1208 0.134 0.866 9 163.3 0.1 1.0 208 0.773 0.227 10 23.1 0.1 0.0 20 0.001 0.999 11 175.6 15.0 0.0 1020 0.001 0.999 12 23.1 0.1 0.5 20 0.001 0.999 13 175.6 15.0 0.0 20 0.001 0.999 14 175.6 15.0 0.0 1000 0.001 0.999

The influence of product sweep and membrane separation factor on the energy consumption and required membrane area are analyzed below.

6.4.2.1. Product sweep.

When using product sweep (PS) a fraction of the retentate stream is vaporized and recycled at reduced pressure to the permeate side of the pervaporation module (Figure 6.5a). Product sweep decreases the water concentration at the permeate side of the pervaporation membrane. This reduction in water content increases the driving force for the water transport, while the driving force for the ACN transport is approximately constant due to the high ACN concentration in the retentate side. A PS of around 20% of the retentate flow rate results in an important reduction of membrane area while the total energy consumption, including reboiler and pervaporation duties, stays low (Figure 6.6). At low PS a strong reduction in required membrane area is obtained but further membrane area reductions are almost negligible for high PS (>20%) where the reboiler duty strongly increases.


97

Figure 6.6. Effect of product sweep on the membrane area required and reboiler duty. Column pressure 0.3 atm. Retentate pressure 10 atm. Total energy duty considers the energy supplied to the pervaporation unit and reboiler.

Product sweep has only little influence when the membrane is used just to overcome

the azeotropic composition due to the high water driving force available along the membrane.

6.4.2.2. Membrane separation factor

The presented results have been calculated with a constant membrane separation factor of 331. However, the separation factor changes with temperature and concentration. For instance, it has been shown17,18 that higher selectivities occur at higher temperatures for different water-organic mixtures. Simulations using higher and lower separation factors than 331 are considered here. A difference to product sweep the effect of selectivity on hybrid processes using either pervaporation or vapor permeation has been shown in the literature8,15 for several mixtures.

A low membrane separation factor results in a low water content at the permeate

side. This “self-sweeping” increases the permeate stream and the reboiler duty in the distillation column but reduces the required membrane area, but very low separation factors require high membrane areas due to the high recycle flow rate to the distillation column and to the membrane (Figure 6.7). On the other hand, high separation factors produce low driving forces, increasing the membrane area but slightly reducing the total energy consumption. At separation factors of around 100, membrane area and total supplied energy (reboiler and pervaporation duties) are low. Fahmy et al.8 also found that

Chapter 6

98

membranes with very high selectivities are not desirable when high purity of the retentate is required using VP. For an economically optimal operation, a higher membrane separation factor requires more product sweep. On the other hand, if the permeate stream is used as the product15, high selectivities are more convenient.

Figure 6.7. Effect of membrane separation factor on the required membrane area and reboiler duty. Column pressure 0.3 atm. Retentate pressure 15 atm. Product sweep 2%.

Summarizing: Product sweep and relatively low separation factors have a positive

effect in a hybrid membrane – distillation process reducing the required membrane area in combination with low energy consumption. High operating pressures in the membrane units and low operating pressures in the distillation columns reduce the required membrane area if the membrane is used as the polishing step.

6.5. Guidelines for selecting hybrid pervaporation-distillation or hybrid vapor permeation-distillation processes

The feed composition to the membrane unit in the process presented in Figure 6.3 and Figure 6.5 corresponds to the top stream composition from the distillation columns. This composition is the same for hybrid processes using either pervaporation or vapor permeation. Figure 6.8 presents the driving force available for pervaporation and vapor permeation at 1 and 10 atm at different feed water concentrations. VP has a higher driving force for water removal than PV in the range of water concentrations higher than the azeotropic point. In the range of water concentrations lower than the azeotropic point the opposite occurs and the difference in driving force increases in favor of PV as the pressure increases. At the azeotropic point the driving forces are equivalent for PV and VP.


99

Figure 6.8. Water driving force for pervaporation and vapor permeation as function of the water concentration at 1 and 10 atm. Driving force calculated as the difference in water partial pressure between permeate and retentate at bubble point (PV) and dew point (VP).

The difference in water driving force for these membrane processes (Figure 6.8) is

used as a tool for selecting either (or both in some hybrid systems) pervaporation or vapor permeation for a hybrid process. The qualitative concepts introduced in this section are discussed quantitatively in the economic evaluation. Three general cases are discussed below. The membrane unit is used for water removal at high water concentrations, at low water concentrations or for overcoming the azeotropic composition.

Vapor permeation is more efficient for water removal at concentrations higher than

the azeotropic point because it requires less membrane area than pervaporation and the energy consumption is also lower due to the availability of vapor from the distillation column9,19.

For water removal at low water concentrations, pervaporation is preferable,

especially at high pressures where it requires a smaller membrane and lower energy consumption than VP. In a hybrid VP process three design alternatives can be identified which are all them economically less favorable: A partial condenser could be used followed for a compressor (in this case the compression cost is high), the column top stream is condensed, pumped and vaporized (energetically more expensive than PV because the whole column top stream has to be vaporized) and a high operating pressure could be applied in the distillation column in combination with a partial condenser (under these conditions the water concentration at the top of the distillation column is shifted at high values increasing the amount of water to be removed). In

Chapter 6

100

conclusion, when pervaporation is used as the final dewatering step at low water concentrations and high pressures the saving on required membrane area and energy consumption is higher compared to VP. However, a hybrid PV process needs only a total condenser after the distillation tower and a liquid pump to increase the pressure of the column top stream (Figure 6.5). As a consequence, total and capital costs of the hybrid PV-distillation process are lower than for a hybrid VP-distillation process. The costs of hybrid processes, where VP is used as a final dewatering step, are not included in the economic evaluation.

When the membrane is used just for splitting the azeotropic composition, vapor

permeation is preferable. In this situation pervaporation and vapor permeation require similar membrane areas where high pressures are applied at the distillation towers and the membranes units. VP is superior to PV because vapor is available at the top of the distillation columns resulting in a lower energy consumption with similar capital costs (see economical evaluation).

The membrane driving force at several pressures has been presented as a tool for

selecting pervaporation or vapor permeation in a hybrid process. This tool has been used for dehydration of ACN but it can be applied to systems with or without azeotrope depending on which component is preferentially transported through the membrane. For instance, if water is the component with the lower boiling point in a non-azeotropic mixture, vapor permeation using a water selective membrane is the best option for a hybrid process.

6.6. Economical evaluation Required membrane area and energy consumption have been presented in Figure

6.7 as a function of the separation factor. Figure 6.9 shows that the total cost of the hybrid process reduces when the separation factor increases reaching a constant value for a separation factor of 100. In general an optimal separation factor exists with a minimum total cost.


101

Figure 6.9. Total cost as function of membrane separation factor for the system T (0.3 atm) P (15 atm) and 2% product sweep.

Capital, operating and total costs are presented in Figure 6.10 for the hybrid

processes described in section 6.4. Capital and total costs are reduced when the operating pressure is increased as shown for hybrid processes at 2, 5 and 10 atm using pervaporation (P1a, P1b and P1c) and for hybrid processes at 5 and 10 atm using vapor permeation (V1a, V1b) where distillation is used as the final dewatering step. In this type of configuration the total and energy costs for VP are slightly lower than for PV at the same pressure, 5 and 10 atm (P1b vs. V1a, P1c vs. V1b).

The cost for a hybrid pervaporation – distillation process using a membrane unit with

inter-stage heating, where the membrane modules and heat exchangers are connected in series (P3), is presented in Figure 6.10. Similar operating costs are obtained compared to an ideal isothermal membrane operation (P2b). If the number of inter-stage heat exchangers is optimized, the capital cost of the process P3 is 10% higher than for an ideal operation (P2b). It is found that the membrane area using inter-stage heating is around 1.34 times the theoretically minimal membrane area (isothermal operation) which is similar to the value suggested by Bausa et al.4 between 1.2 and1.3.

Chapter 6

102

Figure 6.10. Comparison of total, capital and operating costs between conventional distillation based processes and hybrid membrane-distillation processes for dewatering of ACN.

Hybrid processes show a total cost reduction between 25 % and 60 % in comparison

with conventional processes (C1, C2). Capital costs of hybrid processes are higher than for conventional processes but the energy consumption and the operating costs are lower. Capital costs for hybrid PV processes are lower (P2a, P2b and P3) than for conventional processes (C1) when pervaporation is used as the final dewatering step (Figure 6.10). Processes where the membrane is used just to overcome the azeotropic composition are more economical using VP than PV (P1b vs. V1a, P1c vs. V1b). Hybrid processes using a pervaporation unit as a polishing step (P2b and P3) show the lowest total cost.

6.7. Conclusions The membrane driving force at several pressures has been used as a tool for the

economic selection of pervaporation or vapor permeation in hybrid processes. It has been applied for dewatering of ACN where for water removal at high water concentration and just for overcoming the azeotropic composition vapor permeation is more efficient. For water removal at water concentrations lower than the azeotropic point pervaporation is the best option. The combination of different operating pressures in the membrane unit and in the distillation column has proved to be cost efficient in the design of hybrid processes.


103

A low operating pressure in the distillation column reduces the azeotropic water concentration and a high pressure in the pervaporation unit reduces the required membrane area by increasing the driving force for the water transport through the membrane. Product sweep or the use of a membrane with relatively low separation factor also increases the driving force, reducing the water content in the permeate side of the membrane unit.

In general, hybrid processes using either pervaporation or vapor permeation are

economically more favorable than conventional process for the separation of ACN-water mixtures with an obtained reduction between 25% and 60% of the total cost of the conventional processes. The lower total cost is mainly due to the lower energy consumption of the hybrid processes.

6.8. Notation CONDA = Condenser area, (m2) BM = Bare module cost estimation (3.57)20 awh = Annual working hours, 8000 (h/year) C = Molar concentration (kmol/m3)

C = Average molar concentration between on the surface and in the liquid bulk (kmol/m3). pC = Thermal heat capacity (kJ/kmol K) CAPITALC = Capital cost, (€) CONDC = Condenser cost, (€) DC = Distillation column cost, including reboiler and condenser, (€)

FPC = Feed pump cost, (€) HEC = Inter-stage heat exchangers cost (€) MEMC = Membrane cost, (€) MODC = Module cost, (€) OPERATINGC = Operating cost, (€/year)

TOTALC = Treatment cost (€/year) UC = Utility cost, (€/year) VPC = Vacuum pump cost, (€)

Chapter 6

104

totalF = Volumetric liquid flow, (m3/s) Tg = Total flow in the permeate (kmol/h) Lh = heat transfer coefficient (kJ/h m2 K) GH = Permeate Enthalpy (kJ / h) iH = Partial molar enthalpy for component “I”. (kJ / kmol) LH = Retentate enthalpy (kJ / h)

GH∆ = Enthalpy change from outlet permeate temperature from module to 25 °C, (kJ/h) jLH∆ = Enthalpy change in the retentate in the heat exchanger for the stage “j”, (kJ/h) jik , = Mass transfer coefficient between component “I” and “j” in the mixture (m/h) kr = Heat capacity ratio, 1.33 LMTD = Logarithmic media temperature difference, (K) mr = Membrane replacement period in Table 3, (years) MSxx = Marshall and Swift index values for the year 19xx MW = Molecular weight, (kg/kmol) iN = Flux of component “I” through the pervaporation membrane (kmol/h m2) P = Pressure (bar) DISP = Discharge pressure, 1.01325x102 (kPa) v

iP = Vapor pressure of the component “I” oP = Pressure at the standard condition, 1.01325x102 (kPa) OPP = Permeate operating pressure, (kPa) P∆ = Pressure drop for the retentate, (Pa) pc = Power cost, 0.09 (€/kWh) q = Heat flux through the membrane (kJ/h m2) R = Gas law constant, 8.31 m3 (kPa / kmol K) rt = ratio Euro to Dollar as 1.137 T = Temperature (K) CONDT∆ = Temperature difference for cooling water in the condenser, 10 K

intakeT = Absolute temperature of vapor at intake conditions, 298 K To= Temperature at the standard condition, 273 K


105

U = Overall heat transfer coefficient in the condenser, 1800 kJ/h m2 K wc = Cooling water cost, 5x10-5 €/kg ws = Steam cost 18x10-3 €/kg CONDW = Condenser power consumption, (€/h) FPW = Feed pump power consumption, (kW)

VPW = Vacuum pump consumption, (kW) x = Molar fraction in the retentate x = Average molar fraction between on the surface and in the liquid bulk. y = Molar fraction in the permeate HEW = Heat exchangers power consumption, (€/h) RW = Reboilers power consumption, (€/h)

Greek variables

wλ = Heat of vaporization for water, (kJ/kg) FPε = Pump efficiency , 0.6

iΓ = Permeance of component i (kmol/h m2 bar)

= Separation factor for the system ACN-water

γ = Activity coefficient γ∆ = Activity coefficient difference between the surface membrane and the bulk liquid

Subscripts I = Interfacial on the membrane L = Liquid bulk phase or retentate G = Gas bulk phase or permeate i = Component “i” r = reference w = water

( )( )IACNIw

IACNIw

xxyy

//

=α

Chapter 6

106

6.9. Appendix The capital costs of the distillation columns, condensers and reboilers were

calculated using the program Aspen Icarus Process Evaluator 12. The total cost was calculated using the method employed by Oliveira et al.21 and Ji

et al.22, where the total cost (CTOTAL) is based on the following components: capital depreciation and taxes, annual maintenance and labor requirements, module replacement and energy consumption. Depreciation and taxes have been assumed to be 15% of capital costs (eq. A1). Annual maintenance and labor have been taken to be 10% of the capital cost (eq. A1).

OPERATINGCAPITALCAPITALTOTAL C C . C . C +⋅+⋅= 100150 (A1)

)DHECONDVPFPMODMEMBCAPITAL CCCCCC BM(CC ++++++= (A2)

)(1

MEMUOPERATING Cmr

CC +=

(A3)

The operating cost (eq. A3) takes into account the utility costs (CU) and membrane

replacement. The estimated replacement period is presented in Table A1 for several operating temperatures. No supervision, laboratory or insurance costs have been included.

Table A1. Membrane replacement periods used in the economic evaluation

Membrane replacement (mr) , years

Operating temperature, °C

6 150 4 160 2 175

A bare module cost estimation technique20 was used in which the capital cost is

estimated to be 3.57 times the cost of the major equipment (membranes, modules, feed pump, vacuum pump, condenser and inter-stage heat exchangers). This factor (BM) includes piping, valves, instrumentation and peripheral equipment as well as engineering site preparations and other installation costs.

The membrane cost was provided by Pervatech as 1400 €/m2 and the module cost

as 2000 €/m2. The feed pump cost was calculated according to Oliveira et al.21 as follows:


107

93.13.215

6803

52.0

modmod ⋅

∆⋅⋅⋅⋅⋅=series

totalseriesFP nPFn

MSMSrtC

(A4)

Where rt is the ratio Euro to Dollar and MS are the Marshall and Swift index values,

1115 for 2003 and 280 for 1968. The vacuum pump cost and the condenser cost were calculated according to Ji et al.22 and presented in eq. A5 and A6. A similar approach that was used for the condenser was used for evaluating the capital cost of the inter-stage heat exchangers but using steam at 16 bar .

55.0

0

0

360060

94034200

⋅

⋅⋅⋅⋅⋅=

PTRg

rtMSMSC T

VP

(A5)

( )CONDCOND AMSMSrtC ⋅+⋅⋅= 1.1287.1176

9403

(A6)

Where,

LMTDUH

A GCOND ⋅

∆=

(A7)

The utility cost (eq. A8) was calculated for the feed pump, the vacuum pump,

condensers, reboilers and the inter-stage heat exchangers. The cost for electrical power is 0.09 €/KWh, for steam at 16 bar is 18 €/Ton and for cooling water is 0.05 €/Ton with a maximum temperature raise of 10 K for the cooling water.

( )[ ]RHECONDVPFPU WWWWWpcawhC ++++⋅= (A8)

Where awh are the annual working hours and pc is the power cost. The power consumption for the feed pump (WFP) was calculated according to 21:

∆⋅=1000

1 PFW totalFP

FP ε

(A9)

The power consumption for the vacuum pump (WVP) and for the condenser

(WCOND) was calculated according to Ji et al.22 by means of equation A10. This equation involves the total permeate flow. Because the permeate stream is condensed before the

Chapter 6

108

vacuum pump the power consumption is over-predicted but the influence in the total cost is small.

−

−

⋅⋅⋅=

−

113600

%10

1

ntake

krkr

OP

DISi

TVP P

Pkr

krTRg

W

(A10)

Water-cooling consumption and steam consumption were calculated as follows:

pWCOND

wGCOND CT

MWHwcW

⋅∆⋅∆⋅

=

(A11)

exchangersheat stage-inter ofnumber theis where;

1

NSH

wsWNS

j w

jLHE ∑

=

∆=

λ

(A12)

6.10. Reference list (1) Stephan, W.; Noble, R. D.; Koval, C. A. Design methodology for a membrane

distillation column hybrid process. J. Membr. Sci. 1995, 99, 272



(4) Bausa, J.; Marquardt, W. shortcut design methods for hybrid membranes/distillation processes for the separation of nonideal multicomponent mixtures. Ind. Eng. Chem. Res. 2000, 39, 1658 - 1672

(5) Szitkai, Z.; Lelkes, Z.; Rev, E.; Fonyo, Z. Optimization of hybrid ethanol dehydration systems. Chem. Eng. Proc. 2002, 41, 631 - 646

(6) Sommer, S. Hybrid separation processes for dewatering of solvents with inorganic membranes. IVT-Information. 2001, Jhrg 31, 15 - 24

(7) Goldblatt, M. E.; Gooding, C. H. An engineering analysis of membrane aided distillation. AIChE Symp. Ser. 1986, 82, 51 - 69


109

(8) Fahmy, A.; Mewes, D.; Ebert, K. Design methodology for the optimization of membrane separation properties for hybrid vapor permeation-distillation processes. Sep. Sci. Tech. 2001, 36, 3287 - 3304

(9) Pettersen, T.; Lien, K. M. Design of hybrid distillation and vapor permeation processes. J. Membr. Sci. 1995, 102, 21 - 30



(12) Van Hoof V.; Van den Abeele, L.; Buekenhoudt, A.; Dotremont, C.; Leysen, R. Economic comparison between azeotropic distillation and different hybrid systems combining distillation with pervaporation for the dehydration of isopropanol. Sep. Purif. Technol. 2004, 37, 33 - 49

(13) Wesselingh, J. A.; Krishna, R. Mass Transfer in multicomponent mixtures; Delft University press: Delft, 2000.

(14) Russell, H. Purification of acetonitrile by and azeotropic distillation method. United States 3,451,899, 1969.


(16) Hommerich, U.; Rautenbach, R. Design and optimization of combined pervaporation / distillation processes for the production of MTBE. J. Membr. Sci. 1998, 146, 53 - 64


(18) van Veen, H. M.; van Delft, Y. C.; Engelen, C. W. R.; Pex, P. P. A. C. Dewatering of organics by pervaporation with silica membranes. Sep. Purif. Technol. 1-3-2001 , 22-23, 361 - 366

(19) Fakhri, Ph.D Thesis/Dissertation, Technische Universiteit Delft,1999.

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(20) Lipski, C.; Cote, P. The use of pervaporation for the removal of organic contaminants from water. Environ. Prog. 1990, 9, 254 - 261

(21) Oliveira, T. A. C.; Cocchini, U.; Scarpello, J. T.; Livingston, A. G. Pervaporation mass transfer with liquid flow in the transition regime. J. Membr. Sci. 2001, 183, 119 - 133

(22) Ji, W.; Hilaly, A.; Sikdar, S. K.; Hwang, S. T. Optimization of multicomponent pervaporation for removal of VOCs from water. J. Membr. Sci. 1994, 97, 109 - 125

________________________________ This chapter has been submitted for publication: J. Fontalvo, M. A. G. Vorstman, J. G. Wijers, and J. T. F. Keurentjes. (2005).

7. Separation of multicomponent mixtures using an integrated

distillation – pervaporation column

Abstract This chapter proposes a hybrid distillation – pervaporation process that

contains a pervaporation unit in a single column (DPSU). The vapor from the distillation column reduces concentration and temperature polarization from the liquid to the membrane surface while it supplies the energy required for the pervaporation process by condensation. Three different mixtures are considered: dewatering of ethylene diamine (EDA) and isopropyl alcohol (IPA) and the removal of methanol in the production of methyl-tert-butyl ether (MTBE). A rate-based model has been built to describe the distillation and the pervaporation processes in the column. The DPSU has been compared with hybrid distillation processes where the membrane is externally connected to the column (DPEC). For separation of binary mixtures a DPEC is more convenient; however, for the removal of a component from a multicomponent mixture the DPSU is more efficient because the required membrane area is lower. The difference in performance between a DPEC and a DPSU for the multicomponent system lies in the higher methanol concentration that occurs in the pervaporation section in a DPSU compared to an externally connected pervaporation unit.

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7.1. Introduction Hybrid processes of distillation and pervaporation have gained considerable

attention in the recent years1-3. Pervaporation is especially suitable for the separation of close boiling4 and azeotropic mixtures5 where distillation is energy intensive or requires the use of an auxiliary agent. However, compared with distillation the capital cost of pervaporation processes are higher. Hybrid membrane-distillation processes exploit the advantages of distillation and pervaporation operations, while overcoming the disadvantages of both1.

A conventional hybrid distillation - pervaporation process consists of a distillation

column and an externally connected pervaporation module. A conventional pervaporation unit consists of a series of pervaporation modules and inter-stage heat exchangers where the retentate flows in liquid phase. Several hybrid configurations have been studied by Stephan et al.5, Pettersen et al.4, Pressly et al.6 and more recently based on a formal mathematical methodology by Kookos7. In general, the pervaporation module can be used in a hybrid distillation system to remove a specific component from a side stream from the distillation column, to overcome the azeotropic composition or as a final treatment stage. Applications for binary mixtures3,8-11 and multicomponent mixtures1,12 can be found and the effect of the membrane properties and operating variables on the performance of the hybrid process have been discussed.

Recently, we have shown that a retentate in liquid-vapor flow reduces the required

membrane area compared with conventional pervaporation units13,14 and membrane savings higher than 40% can be achieved. Also an important reduction in capital cost is expected due to avoidance of inter-stage heat exchangers. Vapor induces turbulence in the liquid phase increasing the mass and heat transfer between the liquid and the membrane surface while the vapor supplies the energy required for the pervaporation process by condensation. Such a module can be externally connected to a distillation column and the hybrid system follows the same guidelines and behavior as the conventional ones6,15.

The liquid-vapor operation in pervaporation modules opens the possibility to have

pervaporation and distillation operations in one single column. A pervaporation module with hollow fibre membranes, which are coated on the outside, replaces a section of packing or trays in a distillation column. The conventional hybrid distillation - pervaporation process with a pervaporation unit externally connected (DPEC) and the hybrid process containing a pervaporation unit in a single column (DPSU) are depicted in Figure 7.1a and Figure 7.1b respectively. In a DPSU (Figure 7.1b) the reboiler supplies the vapor required for

Separation of multicomponent mixtures using an integrated distillation – pervaporation column

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the pervaporation process while liquid and vapor flow in counter-current in the pervaporation section. In a single column the azeotropic composition and distillation boundaries can be overcome and the separation of more than two components is feasible reducing the involved capital and operating cost. Unlike single-phase, in the two-phase pervaporation unit a higher driving force is expected because the liquid along the pervaporation section is at or close to saturated conditions.

Figure 7.1. Hybrid distillation-pervaporation processes with a pervaporation unit externally connected - DPEC (a) and a pervaporation unit in a single column-DPSU (b). The dotted line indicates that the permeate stream can be recycled to the distillation column. The vacuum pump, the condenser, the liquid pump and the pre-heater in the permeate line have been excluded from the sketch for clarity.

This chapter analyses the advantages and disadvantages of a hybrid process that

contains the pervaporation and the distillation units in a single column. Three examples are studied: dewatering of ethylene diamine (EDA) and isopropyl alcohol (IPA) and the removal of methanol in the production of methyl-tert-butyl ether (MTBE). A rate-based model has been built to describe the distillation and the pervaporation processes in the column.

7.2. Theory and simulations A pervaporation section consists of several layers of hollow fibre membranes as

shown Figure 7.2a where two crossed rows of fibres are presented. The rows of hollow fibres are placed above each other to constitute a packed pervaporation section as shown in Figure 7.2b. The fibres are externally coated with a silica layer16 and connected to a shell where vacuum is applied in order to obtain the permeate stream. The pervaporation and the distillation process can be combined in a single unit as shown in Figure 7.1b. If required, the permeate stream shown in Figure 7.1 can be recycled to the

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distillation column prior to condensation. If the permeate is recycled, it is fed to the point in the column where the local and the permeate stream concentrations are similar.

Figure 7.2. Schematic representation of two layers of pervaporation hollow fibres in cross configuration (a) and a packed section with hollow fibres in cross in line configuration (b).

Concentrations of every component and temperatures in the liquid and vapor

phases were calculated with a rate-based model17,18. Figure 7.3 shows a schematic representation of a differential element of the transversal unit for two-phase flow. Differential elements of 5 mm length in the pervaporation section were used for the calculations. Complete wetting was considered in the pervaporation section. The distillation sections contain packing and differential elements of 0.5 and 0.3 m were employed for the simulations with binary and multicomponent mixtures, respectively. A system of simultaneous equations18 is obtained for the material balances for the vapor and the liquid phases for each interface, the energy balance for each phase and each interface, the mass and energy transfer rate, the summation relations and the liquid-vapor interface equilibrium. Pressure was considered constant along the column.

The set of equations known as MERSHQ18 and more specifically in this chapter

MERSQ, because no pressure drop was considered, was solved using a Newton-Broyden method. However, in contrast with other authors19 the complete set of equation was not solved simultaneously. From initial values of flows per component and temperatures for the bulk of each phase the interfacial balance, summation and equilibrium equations were solved in an internal iteration. The external iteration solved only the flow and temperatures for the bulk of each phase. In this way the Newton-Broyden method is more stable,

a) b)


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although the computation time could increase, especially for a large number of components.

Figure 7.3. Differential element of a transversal unit for a counter-current liquid-vapor two-phase pervaporation.

The Maxwell-Stefan relations for non-ideal mixtures18 were used to describe the mass

transport at the vapor-liquid interface and at the liquid-membrane interface. The approximate method of Krishna20 was applied by considering the Maxwell-Stefan diffusivity and the thermodynamic factors constant along the diffusion path in the mass transfer boundary layer. The mass and heat transfer coefficients in the vapor and liquid phases on the vapor-liquid interface were calculated with the equations of Billets and Schultes21 with the parameters for packing "Montz B1-200" for the distillation sections and "Ralu flu, 2 mm" for the pervaporation sections. The load limits were calculated and taken into account for the results presented in this chapter. The pervaporation section consisted of hollow fibres of 3 mm external diameter with equal longitudinal and transversal pitch to fibre diameter ratio of 3.5. The values of the superficial area and void fraction of packing for "Ralu flu, 2 mm" are close to the ones calculated for the pervaporation section of 110 m2/m3 and 0.92, respectively, which was the criterion for selection of the parameters of this packing due to the unavailability of more accurate information. The permeate pressure was fixed at 80 mbar to allow the use of water for condensation of the permeate stream.

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Mass transfer coefficients between the liquid and the membrane surface in two-phase co-current flow have been measured with hollow fibre membranes14. The mass transfer in two-phase flow increases up to fourfold compared to single phase when no vapor is present keeping the same liquid superficial velocity. In this chapter the mass and the heat transfer coefficient between the liquid and the membrane surface were taken as three times the values for single phase and were calculated from correlations given by Futselaar22. The thermodynamic factors23,24, activity and fugacity coefficients were calculated by UNIFAC and the Peng-Robinson EOS. Fluxes through the membrane were calculated as the product of permeance and local partial pressure difference over the membrane. Permeabilities are considered constant for each component and the values used in this study are presented in Table 7.1 and are based on data from literature 12,25-27. The composition at the permeate side was calculated based on the flux for each component. As flux values are unknown, an iterative procedure was necessary.

Table 7.1. Values of permeance used in the simulations for three different mixtures.

Component Permeance System kmol/m2 h bar

Water 0.12 1 Ethylene diamine 5 x 10-4

Water 0.78 2 Isopropyl alcohol 4.9 x 10-4

Methanol 0.10 MTBE 5.0 x 10-5

3

1-butene 1.0 x 10-5

7.3. Results In this section two binary (EDA-water and IPA-water) and one ternary system

(Methanol – 1-butene – MTBE) are presented. For each mixture it is briefly explained how the separation is conventionally performed with distillation-based processes. Afterwards, the results for the hybrid distillation-pervaporation process in a single unit (DPSU) are shown, discussed and compared with hybrid processes where the pervaporation unit is externally connected to the distillation column (DPEC).


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7.3.1. Binary systems

7.3.1.1. Ethylene diamine (EDA)-water mixture

At atmospheric pressure, EDA-water mixtures can form a maximum boiling point azeotrope, which is pressure sensitive and breaks at 4.0 bar. At high pressure, the separation can be performed only by one single distillation column with a high number of trays due to the low relative volatility at low water concentrations. To obtain EDA with a purity of 99 wt. % from a feed stream of 1470 kg/h containing 73 wt. % of water requires a packed column of 84 m based on our calculations.

In a DPSU the bottom packed section of the distillation column is replaced by a

packed section of hollow fibre pervaporation membrane with a membrane area of 128 m2. EDA with 99 wt. % can then be obtained with a column of 35 m. Figure 7.4 shows in a McCabe-Thiele diagram the water molar fractions in the vapor and liquid streams. The permeate stream has been recycled to the column. In contrast to conventional distillation, the operating line crosses the 45° line due to the water removal from the liquid phase in the pervaporation unit. EDA, which is the heavier component, condenses and water vaporizes as soon as water is removed from the liquid by the membrane. In the DPSU the membrane shifts the water composition thus strongly increasing the driving force for the liquid-vapor mass transfer and consequently reducing the length of the column, in this case by about 60%.

Figure 7.4. Operating line for a hybrid distillation-pervaporation process in a single unit (DPSU) for dewatering of ethylene diamine (EDA) at 4.0 bar.

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A hybrid process with an externally connected pervaporation unit is nevertheless more convenient than the DPSU because in the last case more membrane area is required. In the DPSU the complete liquid stream that is flowing inside the column in the stripping zone is purified by the pervaporation section. However, only a fraction of this purified stream is withdrawn as a product while the remaining liquid is mixed, upwards in the column, with streams with higher water content. The fraction of liquid that returns to the column corresponds to the boilup ratio. In a DPEC process only the bottom stream from the distillation column is treated in the pervaporation module, thus reducing the required membrane area.

There are two purified water streams in this process: the permeate stream from the

pervaporation section and the distillate stream. In general, DPSU is not convenient when the product stream from the pervaporation section is also a product stream from the distillation column itself. DPSU should therefore be used in a process where the compound to be preferentially removed by the membrane is not expected from the distillate or bottom streams. In this way the amount of a specific compound that should be removed with either an externally connected or in a single unit pervaporation membrane is equal. However, comparing both hybrid processes, the driving force in the pervaporation unit is different and consequently the required membrane area.

7.3.1.2. Isopropyl alcohol (IPA)-water mixture

The purification of IPA with distillation-based processes occurs through azeotropic distillation using three distillation columns and benzene as entrainer9. Hybrid processes for the separation of this mixture using pervaporation can reduce the total cost by 50%1,3,9 compared to the distillation-based processes. A DPSU is described in this section for the separation of 1900 kg/h of an IPA-water mixture with a water content of 20 wt. %.

The operating line of the DPSU overcomes the azeotropic composition as shown in

Figure 7.5. The total height of the column considering the packing and the pervaporation zone is 28 m. The pervaporation section, that constitutes the longest section of the column with 20 m, has been positioned in the rectifying zone. The water concentration mainly increases in the direction of the liquid flow except close to the condenser. Close to the top of the column IPA instead of water condenses. Because the azeotropic composition has been overcome the temperature profile shows a minimum of around 355 K that is close to the azeotropic temperature (Figure 7.6).


119

Figure 7.5. Operating line for a hybrid distillation-pervaporation process in a single unit (DPSU) for dewatering of isopropyl alcohol (IPA) at 1.01 bar.

Figure 7.6. Temperature profile at the liquid-vapor interface in the DPSU column for dehydration of IPA. The positions are measured from the top of the column.

From the foregoing it can be concluded that DPSU is not convenient for the

separation of binary mixtures as compared to DPEC. DPSU is more efficient than a DPEC in a process where the compound to be preferentially removed by the membrane is not expected from the distillate or bottom streams. The next case shows that for the removal of methanol from a multicomponent mixture, a DPSU process is more efficient.

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7.3.2. Multicomponent system

7.3.2.1. Methyl tert butyl ether(MTBE) – methanol – n-butene (C4) mixture

In the process licensed by Huels AG, the separation of MTBE is performed as shown in Figure 7.7. Methanol in excess and i-butene are fed to the reactor to produce MTBE. The i-butene, which is available from the C4-raffinate of a steamcracking or catalytic cracking process, is completely converted into MTBE. The resulting stream mainly consists of methanol, MTBE and linear butanes (nC4). Three separation units are required due to the formation of azeotropes between the methanol with both the MTBE and the nC4. In the first distillation column, operating at 6 bar, the feed is separated in a distillate stream containing nC4 and methanol and the bottom stream with methanol and MTBE. The MTBE-methanol mixture is purified in a second distillation column operating at 12 bar. The distillate with azeotropic composition from the second distillation column is returned to the reactor. The distillate from the first distillation column is fed to a separation unit for production of nC4.

Figure 7.7. Huels process for the production of MTBE

A conventional hybrid distillation-pervaporation process is based on the treatment of

a side stream from the first distillation column to remove methanol28,29. We have used the approach of Hömmerich and Rautenbach12, which consists of purifying the MTBE stream (bottom stream) from the first distillation column eliminating the necessity of the second distillation column.

The methanol concentration and temperature profiles in column 1 are presented in

Figure 7.8. Table 7.2 shows the values of molar flux and concentrations of the feed, distillate and bottom. A maximum methanol concentration is shown close to the reboiler,


121

where the temperature is also high. At this point the methanol partial vapor pressure and the driving force for the pervaporation process are the highest. The methanol concentration profile in the vapor phase also presents a maximum at this point. The pervaporation section for the DPSU process should be placed at this maximum at a position of 10.5 m from the top of the column. Also in a DPEC process the side liquid stream should be taken from this position.

Table 7.2. Molar fractions and flows for the first distillation column in Figure 7.

Feed Distillate Bottom Flow, kmol/h 50 33.594 16.406 Molar fraction:

Methanol 0.06 0.0475 0.0851 MTBE 0.3 2.2e-5 0.9146

1-butene 0.64 0.9524 2.1e-4

Figure 7.8. Profiles of temperature at the liquid-vapor interface and molar fraction of methanol in the liquid in the first distillation column of Figure 7. The position is measured from the top.

The required membrane area to obtain a specified methanol concentration in the

bottom stream in the DPSU process is presented in Figure 7.9. In the Figure, the effect of two different positions have been plotted; 10.5 m (which is optimal) and 8 m. The required membrane area in the optimal location is more than 3 times lower than for the other position. The relation between membrane area and methanol composition in the bottom stream is almost linear as a consequence of the limited change of methanol concentration within the pervaporation section: the highest variation of the methanol

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molar fraction is 0.04 only. The reboiler duty of the DPSU column is about 5% higher than the first distillation column in a distillation-based process (Figure 7.7), while the condenser duty is about 2% lower. When the membrane fibres are not completed wetted, similar membrane areas are required and the same optimal position is obtained as compared to the presented case due to the equilibrium between the liquid and vapor phase. On dry spots vapor permeation takes place.

Figure 7.9. Required membrane area to obtain a given concentration in the bottom of a DPSU. The influence of the position of the pervaporation section on the required area is shown.

A hybrid distillation process that contains the pervaporation unit in a single column

(DPSU) is more convenient than a hybrid process with an externally connected pervaporation unit (DPEC) for the purification of MTBE by removal of methanol in a multicomponent mixture. The liquid flow rate that passes through the pervaporation membrane in a DPEC is lower than in a DPSU process because only a fraction of the liquid is withdrawn from the column. Because the amount of methanol to be removed is equal in both hybrid processes, the outlet methanol concentration and the methanol driving force for pervaporation of the externally connected unit is lower increasing the required membrane area. Additionally, in the pervaporation section of the DPSU the methanol driving force is at its maximum at the given pressure because the liquid follows the saturation conditions. The required membrane area in a DPEC decreases as the flow rate of the withdrawn side stream increases. Thus, the minimum required membrane area is obtained when all the liquid is withdrawn from the column, which corresponds to the required membrane area of a DPSU. However, in an externally connected pervaporation unit high side flow streams can not be withdrawn without disturbing the performance of the distillation column.


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The required membrane area of a pervaporation section in a hybrid single unit is

lower compared to an externally connected unit due to the differences in driving force for the methanol transport by pervaporation. If membranes are developed to remove specific components from a multicomponent mixture, the different membranes can be positioned at several optimal positions within the distillation column allowing the separation of the mixture in one single column.

7.4. Conclusions In this chapter a hybrid process has been proposed that contains the pervaporation

and the distillation units in a single column. The vapor from the distillation column reduces concentration and temperature polarization from the liquid to the membrane surface while it supplies the energy required for the pervaporation process by condensation. The performance of the hybrid process in a single unit has been compared with a conventional process where the pervaporation unit is externally connected. Three cases have been studied: the dehydration of ethylene diamine-water and isopropyl alcohol-water mixtures and the methanol removal from MTBE – 1-Butene – methanol mixtures. Simulations for the hybrid system in a single unit have been performed using a rate-based model. For purification of the binary mixtures, a conventional hybrid process with an externally connected pervaporation unit is more convenient than a hybrid process in a single unit due to the lower required membrane area. However, for the removal of methanol from the multicomponent mixture a hybrid process in a single unit requires a lower membrane area because the methanol concentrations within the pervaporation section are higher than in an externally connected unit. Also, the liquid in the pervaporation section follows saturation conditions. The pervaporation section has been positioned at a height in the column where the methanol partial pressure in the liquid is the highest.

A hybrid distillation-pervaporation process in a single unit is convenient for the

removal of one component from multicomponent mixtures when the component can not be obtained from the distillation column as pure distillate or bottom product. If membranes are developed to remove specific components from a multicomponent mixture, the different membranes can be positioned at various optimal positions within the distillation column allowing the separation of the mixture in one single column.

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7.5. Nomenclature e = energy flux through the interface, kJ/m2 h f = molar feed flow kmol/h

H = enthalpy, kJ/h L = liquid flow rate, kmol/h N = molar flux through the interface, kmol/m2 h Q = Heat losses, kJ/h m2

S = Sidestream, kmol/h T = temperature, K V = vapor flow rate, kmol/h x = molar fraction in the liquid phase y = molar fraction in the vapor phase

Greek symbols

iΓ = permeance of component i, kmol/m2 h bar

Subscripts i = component “i” j = differential element “j”

Superscripts L = liquid phase V = vapor phase VF = feed in vapor phase LF = feed in liquid phase P = liquid-membrane interface

7.6. Reference list (1) Lipnizki, F.; Field, R. W.; Ten, P. K. Pervaporation based hybrid process: a review of

process design, applications and economics. J. Membr. Sci. 1999, 153, 183 - 210


(3) Sommer, S.; Melin, T. Design and optimization of hybrid separation processes for the dehydration of 2-propanol and other organics. Ind. Eng. Chem. Res. 2004, 43, 5248 - 5259


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(5) Stephan, W.; Noble, R. D.; Koval, C. A. Design methodology for a membrane distillation column hybrid process. J. Membr. Sci. 1995, 99, 272



(8) Goldblatt, M. E.; Gooding, C. H. An engineering analysis of membrane aided distillation. AIChE Symp. Ser. 1986, 82, 51 - 69

(9) Van Hoof V.; Van den Abeele, L.; Buekenhoudt, A.; Dotremont, C.; Leysen, R. Economic comparison between azeotropic distillation and different hybrid systems combining distillation with pervaporation for the dehydration of isopropanol. Sep. Purif. Technol. 2004, 37, 33 - 49

(10) Szitkai, Z.; Lelkes, Z.; Rev, E.; Fonyo, Z. Optimization of hybrid ethanol dehydration systems. Chem. Eng. Proc. 2002, 41, 631 - 646

(11) Fontalvo, J.; Cuellar, P.; Timmer, J. M. K.; Vorstman, M. A. G.; Wijers, J. G.; Keurentjes, J. T. F. Comparing pervaporation and vapor permeation hybrid distillation processes. Ind. Eng. Chem. Res. 2005, 44, 5259 - 5266

(12) Hommerich, U.; Rautenbach, R. Design and optimization of combined pervaporation / distillation processes for the production of MTBE. J. Membr. Sci. 1998, 146, 53 - 64

(13) Fontalvo, J., Vorstman, M. A. G., Wijers, J. G., and Keurentjes, J. T. F. Heat supply and reduction of polarization effects in pervaporation by two-phase feed. Submitted. 2005

(14) Fontalvo, J., Vorstman, M. A. G., Wijers, J. G., and Keurentjes, J. T. F. Separation of organic-water mixtures by co-current vapor-liquid pervaporation with transversal hollow fibre membranes. Submitted. 2005.

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(16) Peters, T. A.; Fontalvo, J.; Vorstman, M. A. G.; Benes, N. E.; Van Dam, R. A.; Vroon, Z. A. E. P.; Van Soest-Vercammen, E. L. J.; Keurentjes, J. T. F. Hollow fibre microporous silica membranes for gas separation and pervaporation. Synthesis, performance and stability. J. Membr. Sci. 2005, 248, 73 - 80

(17) Krishnamurthy, R.; Taylor, R. A nonequilibrium model of multicomponent separation processes. Part I: Model description and method of solution. AIChE J. 1985, 31, 451 - 456


(19) Kooijman, H.; Taylor, R. The Chemsep Book; Libri books on demand: New York, 2000.


(21) Billet, R.; Schultes, M. Prediction of mass transfer columns with dumped and arranged packing. Updated summary of the calculation method of Billet and Schultes. Trans IchemE. 1999, 77, Part A, 498 - 504

(22) Futselaar, H., Ph.D. Thesis/Dissertation, Universiteit Twente,1993.



(25) Sommer, S. Hybrid separation processes for dewatering of solvents with inorganic membranes. IVT-Information. 2001, Jhrg 31, 15 - 24



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(27) Van Gemert, R. W.; Petrus-Cuperus, F. Newly developed ceramic membranes for dehydration and separation of organic mixtures by pervaporation. J. Membr. Sci. 1995, 105, 287 - 291

(28) Lu, Y.; Zhang, L.; Chen, H. L.; Qian, Z. H.; Gao, C. J. Hybrid process of distillation side-connected with pervaporation for separtion of methanol/MTBE/C4 mixture. Desalination. 2002, 149, 81 - 87

(29) Daviou, M. C.; Hoch, P. M.; Eliceche, A. M. Design of membrane modules used in hybrid distillation pervaporation systems. Ind. Eng. Chem. Res. 2004, 43, 3403 - 3412

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8. Conclusions and perspectives

Abstract Pervaporation modules operating in two-phase flow appear to be effective to

reduce polarization effects and to supply the energy required for the separation process. The membrane area, the amount of auxiliary equipment and the related capital and operating cost are reduced as compared with single phase flow expanding the application window of pervaporation. Industrial applications are already viable with the multi-tubular modules available in the market. For applications with hollow fibre membranes, it is necessary to obtain more information on wetting behavior as a function of hydrodynamics and mixture composition. For pervaporation-distillation operations in one single column information on dynamic behavior and multiple steady states are important for optimization and start up. Interesting developments are related with tuning of membrane properties with an external source, the use of catalytic membranes in distillation columns and the use of internal two-phase flow in capillary membranes.

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8.1. Introduction Pervaporation modules operating in two-phase flow appear to be effective to

reduce polarization effects and to eliminate the use of inter-stage heat exchangers. Industrial applications are already viable with the multi-tubular modules available in the market. For these applications pervaporation can operate as a stand alone system or in a hybrid process as an externally connected unit. Currently, pervaporation is mainly used in hybrid systems. In the future, this will probably not change due to the important capital cost of pervaporation modules and membranes.

The industrial application of two-phase flow pervaporation with hollow fibre

membranes will take some more time. This chapter describes the major findings presented in this thesis and the outlook for industrial applications, especially with hollow fibre membranes. The final sections of this chapter describe some general perspectives and a general conclusion.

8.2. Major findings The findings addressed in this thesis can, in general, be divided into three inter-

related subjects: silica membrane stability, the influence of hydrodynamics on mass and heat transfer in pervaporation modules, and the performance of pervaporation units in two-phase flow in stand alone and hybrid applications.

8.2.1. Membrane stability For the dehydration by pervaporation of dimethylformamide (DMF)-water, ethanol-

water and methanol-water mixtures with silica hollow fibre and tubular membranes, initially high fluxes and selectivities were observed. Subsequently, the pervaporation performance decreased with time, likely due to interactions of water and the organic compounds with the silica material. However, the water flux could be recovered when the aged membrane was used for dewatering of a 1,4-dioxane-water mixture. This recovery shows that the interaction between the alcohols or DMF with the silica layer is at least partly reversible.

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8.2.2. Hydrodynamics, mass and heat transfer Secondary flows in a pervaporation module have been calculated and measured

with CFD and ultrasound computer tomography, respectively. Several density ratios between the organic compound and water have been used. The flux through the membrane tube increases as the density ratio increases. The results show an inversion point on the tubular membrane surface above which the flow moves upwards and below which the flow moves downwards. For heavy organic-water mixtures the inversion point disappears and the flow on the membrane surface is only downward. For the design of multi-tubular pervaporation modules, the results suggest that for separation of heavy organic-water mixtures the flux is maximized with a squared configuration. On the other hand, a triangular configuration is preferred for dewatering of light organic-water mixtures.

This thesis proposes a two-phase operation in pervaporation modules for which the

retentate consists of vapor and liquid. Vapor is obtained from partial vaporization of the liquid feed stream. Experiments in two phase flow using vapor and gas have been performed in a lab-scale setup for dewatering of 1,4-dioxane water mixtures and in a bench scale setup for dewatering of isopropyl alcohol (IPA) water mixtures. For these experiments the retentate flows inside of silica tubular membranes. The second phase increases the turbulence in the liquid and thus the mass and heat transfer between the liquid and the membrane surface. The polarization effects are reduced and consequently flux and selectivity are increased by at least twofold as compared with laminar flow conditions. These findings have been supported by calculations of flux and selectivity in slug flow based on experimental data of bubble size, liquid slug size and bubble rise velocity. Experiments with gas have shown the effect of the second phase on concentration and temperature polarization, while experiments with vapor have shown the additional effect of heat supply by condensation. When using vapor, the heat required for the pervaporation process is supplied to the liquid by condensation, increasing the performance and avoiding the use of inter-stage heating. The performance of the tubular membrane at low liquid flow rates in liquid-vapor flow is close to the performance at turbulent conditions.

Moreover, two-phase flow has also been applied to pervaporation modules with

hollow fibre membranes, in which the liquid and the second phase flow around the fibres. Mass transfer coefficients and gas void fractions have been measured. Mass transfer coefficients in two-phase flow increase about four times as compared with single phase.

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8.2.3. Performance of pervaporation units in two-phase flow in stand alone and hybrid applications

Conventional hybrid processes which consist of a pervaporation unit that is externally connected to a distillation column have been studied (DPEC). Guidelines to decide whether pervaporation or vapor permeation is more convenient for a specific application are presented. The positive influence of relatively low selectivities on the total cost of the process is demonstrated for the dewatering of acetonitrile (ACN)-water mixtures. Also, recycling a fraction of the retentate into the permeate side leads to a strong reduction of the required membrane area and thus, total separation cost. For a pressure sensitive azeotropic mixture such as ACN-water, low pressures in the distillation column combined with high pressures in the pervaporation unit are economically convenient. As compared to distillation-based processes a reduction between 25% and 60% in the total separation cost can be achieved. A hybrid process with vapor permeation is preferred when the membrane is used either for water removal at high water concentration or just for overcoming the azeotropic composition. When the membrane removes water at water concentrations lower than the azeotropic point, pervaporation is more effective.

Based on experimental measurements of mass transfer coefficients and void

fractions a two-phase pervaporation module has been designed for dewatering of 1.5 ton/h of an IPA-water mixture. Liquid and vapor flow in co-current upwards in the pervaporation unit. The required membrane area is 45% lower as compared to a pervaporation module with inter-stage heat exchangers.

A hybrid distillation-pervaporation system integrated in one single (DPSU) column has

been proposed. A pervaporation section that consists of a bed of hollow fibre membranes, replaces a section of trays or packing in a distillation column. This section can be placed in an optimal position where the vapor pressure of the component to be removed is at its maximum. The hybrid system has been designed with a modified rate-based model for three different mixtures: Ethylene diamine (EDA)-water, IPA-water and methyl tert butyl ether (MTBE)-methanol-butene. The DPSU has been compared with a conventional hybrid process where the pervaporation unit is externally connected to the distillation column (DPEC). As compared with conventional hybrid systems, it has been found that the DPSU is more convenient when the component to be removed can not be obtained as pure bottom or top product. Thus, for the binary mixtures a DPEC is economically more convenient and, on the other hand, for the removal of methanol from the multicomponent mixture a DPSU is more efficient.


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8.3. Perspectives

8.3.1. Industrial application of two-phase flow Liquid-vapor flow in pervaporation systems in stand alone and in hybrid applications

appears to be promising due to the reduction in membrane area, auxiliary equipment and the related capital and operating cost as compared with conventional single phase systems. However, several steps have to be taken to arrive at an industrial application.

• It has been show that the water flux strongly decreases in silica membranes

before a steady value is achieved. The exact mechanism for this behavior is not well known but it is generally believed that it is due to interactions of the compounds of the mixture with silanol groups on the silica surface. However, there are several types of silanol groups and Si-O-Si groups are also present. The Si-O-Si groups are characterized by a high acidity and are important in the decomposition of alcohols. Chemisorption of organic molecules can take place and as a consequence the flux of the permeating compounds could be reduced as a result of blocking and/or chemical interaction.

• More insight in the hydrodynamics, both from an experimental and theoretical

point of view, will improve the design of pervaporation modules. Hydrodynamic studies can be focused on two-phase operation for tubular and hollow fibre membranes. The influence of the hydrodynamic conditions and mixture composition on the wetting behavior on the membrane surface should be taken into account. It is possible that the required membrane area will be unaffected when silica membranes are used, due to the equilibrium between the liquid and vapor phase. Nevertheless, the energy consumption will differ as a function of the wetting coverage. When a heterogeneous reaction is involved on the membrane surface, the properties of the liquid and vapor mixture will change, thereby modifying the hydrodynamic conditions in the pervaporation module and thus the performance.

• Several modules with hollow fibre membranes in transversal orientation have

already been constructed and have been described in the open literature. However, mechanical stability tests of the various module types are necessary.

• The behavior of a hybrid distillation-pervaporation unit in one column is

complex. It could be interesting to find out if several operating steady states are possible and which ones are the most convenient. Together with

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information on transient behavior this information is also valuable for the start up and operation of such hybrid processes.

8.3.2. Further developments in pervaporation processes The ideas presented below, related to the development of pervaporation processes,

are general and clearly some of them can be combined with the ideas presented above.

• Two-phase flow inside capillary pervaporation membranes could be interesting due to the high packing density, low mass and heat transfer resistance, and low pressure drop. The hydrodynamic behavior of the two-phase system will differ to the ones presented in this thesis due to capillary effects. An important amount of research has already been done in monoliths, knowledge that can be applied for this kind of pervaporation modules.

• Industrial companies that perform the separation of various mixtures with the

same equipment will require membranes which flux and selectivity properties can externally be modified. E.g. a magnetic field or ultrasound waves can be used for this purpose in combination with specially designed membranes.

• Membranes with different selective properties can be placed in one single

column at optimal locations for the purification of multicomponent mixtures. For this application, short-cut methods and rules of design and optimization are necessary.

• Hollow fibre membranes with a catalytic layer can be placed inside of

distillation removing one of the products of a reaction. This kind of operation will improve conversion and it can overcome azeotropes and distillation boundaries. Additionally, due to the removal of one of the products, the driving force for the liquid-vapor mass transfer is increased and consequently the performance of the distillation process improves.


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8.4. Conclusion A pervaporation unit operating in two-phase flow where the retentate consists of a

liquid and vapor phase has appeared to be effective to reduce polarization effects and supply the energy required for the pervaporation process. This kind of operation reduces the required membrane area, the amount of auxiliary equipment and the related capital and operating cost of pervaporation processes as compared with single phase flow. It can expand the application window of pervaporation processes in stand alone and hybrid applications. This allows to further reduce the energy consumption as compared to conventional pervaporation processes where the retentate is in single liquid phase.

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Acknowledgements

This book is the product of a continuous effort and the contribution in

several ways of a lot of people. I got a vision for this book and they helped me to

build it with their scientific or personal support. They helped me to learn a lot of

things in these four years which clearly was the idea from the very beginning.

I want to thank every one in the Process Development Group and especially

those who I worked with. They were always willing to help, very kind and

supportive. This book would not be a reality without the confidence that Prof. Jos

Keurentjes, Ir. Marius Vorstman and Ir. Johan Wijers put on me. They

continuously feed the project with nice ideas, suggestions and courage to the

future. Also, the way that this thesis is written has a strong influence of the

valuable comments of Ir. Johan Wijers.

Thanks to all the people who was involved in this project from CALDIC,

SHELL, NEDALCO, PERVATECH, CEPARATIONS, TNO and DELFT UNIVERSITY

OF TECHNOLOGY. Especial thanks to Dr. Michiel Nienoord, Dr. Zeger Vroon,

Frans Velterop, I. Braber, Prof.dr.ir. P. Jansens, Dr. Ž. Olujiċ and Paulo Perez.

Thanks to Dr. Otto Oudshoorn who wrote the patent application. The financial

support by the Programme Office on Economy, Ecology and Technology from the

Dutch Ministries of Economic affairs, Education and Environmental affairs under

contract EETK20046 is acknowledged.

My students Karin, Evert, Jelan, Michel, Alexander and Aart made an

essential contribution in this project with their work, ideas and enthusiasm. I

really appreciate your contribution, the discussion of ideas and the very nice

atmosphere you created. I wish for all of you the best of the futures, you deserve

it. I enjoyed so much working with you guys.

I want to thank Eric Fourcade, we no only worked together on the project on

ultrasound, but we have built a nice friendship. I also learned a lot of staff on

CFX from you. Thanks my “French friend”. Thijs, with his always spontaneous

personality, made me feel at home, he supported me with the laboratory staff and

we also had really interesting discussions about his and my findings. Frank

Gielens, thanks because from the very first moment I started with SPD your talks

were always comforting.

During this project I had to do some modifications in the experimental

setups and this was possible due to the help of Chris and Frank. Thanks for all

your help guys.

Thanks to all the people in the SPD group for making a pleasant

atmosphere: Maartje, Nieck, Peter, Carlijn, Joop, Marguerite, Martin, Zwannet,

Henny, Maikel, Leon, Ard, Ana, Vishal, Xaviera, Martijn, Arjan, Dick, Stefan and

Marcus.

This book would not be in your hands without the help of Mabel Caipa who

followed all the process. Mabel y Jos muchas gracias de todo corazón en primer

lugar por la amistad que nos brindaron y por su cariño. Este es un regalo que no

tiene precio. También les estoy muy agradecido por ayudarme con todos estos

últimos detalles que hacen realidad este sueño que se manifiesta en este

documento. Mabel gracias por las agradables charlas que tuvimos entorno a

nuestras familias, nuestro doctorado, nuestras alegrías y preocupaciones. Todas

ellas valieron la pena y me ayudaron a seguir adelante.

Miguel Angel, mi amigo, mi hermano gracias por todo el apoyo que

incondicionalmente he recibido de Ti. Una muestra más no solo de la grandeza de

la amistad sino de la confianza en el futuro. Esta realidad es el fruto de un sueño,

de hace ya varios años, que ambos hemos hecho realidad. Un esfuerzo grande

con una satisfacción personal inmensa.

Ana, Maria, Peter, Martin, Daniel e Isabel son mi familia en Holanda, los

llevo siempre en el corazón no solo por sus muestras de cariño

“Colombo-Holandés”, sino por el constante apoyo que me brindaron y por

compartir conmigo una parte de sus vidas. Tengo muchos recuerdos hermosos de

este tiempo y les agradezco de todo corazón porque estuvieron en todos y cada

uno de los pasos para conseguir esta meta.

Paula, mi amor, mi vida, la fuente de donde recojo fuerzas para seguir

adelante, muchas gracias. Solo Dios y tu sabes cuantas veces estuve con dudas y

tu me ayudaste en este proceso con esa, tu manera particular de la cual sigo

enamorado de ti. Me inspiraste con nuevas ideas y formas de mirar las cosas

para encontrar soluciones prácticas y cuando la fatiga me rodeaba me ayudaste a

ver los diferentes proyectos con nuevas perspectivas, esenciales en un proyecto

de investigación.

Mi familia, Cristonel, Dolly y Edison, siempre han estado ahí, conmigo,

dándome confianza, esperanza y certeza en el futuro. Su cariño, una

manifestación del amor de Dios, me ha soportado toda la vida hasta llegar a este

punto y continuar adelante. Esta tesis es la recopilación histórica de todos esos

cariños, preocupaciones, voces de aliento y apoyos.

Señor Dios gracias por la vida, por el regocijo de vivir, por el nuevo día, por

mi profesión y mis virtudes y por la finalización con éxito de este proyecto.

Curriculum Vitae

Javier Fontalvo Alzate was born in Bogotá, Colombia on June 13, 1971. He

earned his B.S. degree in Chemical Engineering from the National University of

Colombia (1995). He worked for Sucromiles in the production department of citric

acid and afterwards he started his M.Sc and received his degree from the National

University of Colombia in 1998. During his M.Sc. studies he was involved in the

development of pellets with non-uniform distribution of catalyst for the partial

and total oxidation of propylene under the supervision of Professor Luis Carballo.

We worked for the same university as a teacher and he was also involved in

several research projects in heterogeneous catalytic systems. He moved to

Netherlands in 2001 to start his Ph.D. with Professor Jos Keurentjes in the

Process Development Group. He was involved in preliminary studies on

membrane reactors and his core work was on the development of hybrid

distillation – pervaporation processes for dehydration of solvents.

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