Process Intensification for the Green Solvent Ethyl Lactate ...

Process Intensification for the Green Solvent Ethyl Lactate Production based on Simulated Moving Bed and

Pervaporation Membrane Reactors

A Dissertation Presented to the Faculdade de Engenharia da Universidade do Porto

for the degree of PhD in Chemical and Biological Engineering

by

Carla Sofia Marques Pereira

Supervised by Professor Alírio Egídio Rodrigues and

Dr. Viviana Manuela Tenedório Matos da Silva

Laboratory of Separation and Reaction Engineering, Associate Laboratory LSRE/LCM

Department of Chemical Engineering, Faculty of Engineering, University of Porto

October 2009

FEUP-LSRE/LCM - Universidade do Porto

© Carla Sofia Marques Pereira, 2009

All rights reserved

Acknowledgements

First of all, I want to thank my supervisors, Professor Alírio Rodrigues and Dr Viviana Silva.

Professor Alírio, thank you for all the friendship, constant support and for always challenging

me to reach higher goals within my work. Dr. Viviana, I want to thank you for all the

encouragement, motivation, constant support, for all the long discussions and great ideas that

make me go further and further within my work, and, also, for being a truly and special

friend.

I am very grateful to Professor Simão Pinho, for the friendship and all the support in the

framework of the project “POCI/EQU/61580/2004” and to Professor Madalena Dias for the

support whenever needed.

To all my LSRE colleagues, especially Israel Pedruzzi, Pedro Sá Gomes, Michael Zabka,

João Santos, Miguel Granato, João Pedro Lopes, Alexandre Ferreira and Nuno Lourenço for

the friendship, collaboration, and support whenever I needed.

To my primary school teacher, Professor João Aveiro, for always believing in me and keeping

me motivated along the years.

To Fundação para a Ciência e Tecnologia, for the financial support (Research Fellowship:

SFRH / BD / 23724 / 2005).

To Sofia Rodrigues, Marta Abrantes, Fátima Mota, Miguel Teixeira and Nuno Garrido for all

the great moments spent after work!!!

Last, but not least, I would deeply like to thank my family and friends, for all giving love,

support and trust, especially to my grandmother, Dulcelina, that will always stay in my heart.

To my parents and sisters

“Green chemistry represents the pillars that hold up our sustainable future. It is

imperative to teach the value of green chemistry to tomorrow’s chemists.”

Daryle Busch, President of the American Chemical Society (June 26, 2000)

Resumo

O principal objectivo deste trabalho foi o desenvolvimento de um novo processo eficiente para a

produção do solvente verde, lactato de etilo, através da reacção de esterificação entre etanol e

ácido láctico, utilizando tecnologias híbridas de reacção/separação baseadas em reactores de leito

móvel simulado e processos de membranas por pervaporação. De forma a atingir a meta proposta,

foram abordados os seguintes temas:

Aquisição de dados fundamentais: A resina de permuta iónica, Amberlyst 15-wet, foi avaliada

tanto como catalisador para a reacção de esterificação, como adsorvente selectivo para a água. Os

dados cinéticos e de equilíbrio da reacção foram medidos, na gama de temperaturas 50ºC-90ºC, e

usados para a determinação da constante de equilíbrio e da lei cinética da reacção como função da

temperatura, baseadas em actividades descritas pelo modelo UNIQUAC. Os dados de adsorção

foram também medidos, a 20ºC e a 50ºC, e ajustados a uma isotérmica de Langmuir

multicomponente, tendo-se assumido uma capacidade volumétrica da monocamada igual para

todas as espécies, reduzindo o número de parâmetros de ajuste de 8 para 5, para cada temperatura.

Membranas comerciais hidrófilas da Pervatech foram avaliadas para a desidratação do etanol,

ácido láctico e lactato de etilo, por pervaporação. As permeâncias de todas as espécies foram

determinadas em função da composição e temperatura na gama 48ºC-72ºC.

Intensificação de processo: Modelos matemáticos, considerando resistências internas e externas à

transferência de massa e velocidade variável devido à mudança das propriedades da mistura

multicomponente, foram desenvolvidos para reactores cromatográficos — reactores de leito fixo e

de leito móvel simulado, SMBR — e validados pelos dados experimentais. O modelo matemático

do reactor de membranas por pervaporação, PVMR, considera, adicionalmente, a permeação

através da membrana, os efeitos de polarização por concentração e temperatura e operação não

isotérmica. A avaliação teórica do comportamento da unidade SMBR foi realizada para analisar o

efeito da configuração, da composição da alimentação e tempo de comutação nas regiões de

separação/reacção e/ou no desempenho do processo nos pontos operacionais óptimos. O

desempenho do PVMR foi avaliado para operação isotérmica e não isotérmica, e foram

determinadas condições apropriadas para a maximização quer da conversão do ácido láctico quer

da pureza do lactato de etilo. Finalmente, uma nova tecnologia foi desenvolvida e submetida a

registo de patente, o reactor de membranas de leito móvel simulado, PermSMBR, o qual integra

membranas selectivas dentro das colunas do SMBR.

Abstract

The main objective of this work was the development of a new efficient process to produce the

green solvent ethyl lactate from the esterification reaction between ethanol and lactic acid by

using hybrid reaction/separation technologies based on simulated moving bed reactors and

pervaporation membrane processes. To accomplish this target, the following topics were

addressed:

Basic data acquisition: The acidic ion exchange resin Amberlyst 15-wet was evaluated as both

catalyst for esterification and selective adsorbent for water. Equilibrium and kinetic data were

measured in the temperature range 50-90ºC, and used to obtain the equilibrium constant and

kinetic law as function of temperature, which are based on liquid activities described by the

UNIQUAC model. Adsorption data was also obtained and fitted to a multi-component Langmuir

isotherm assuming a constant monolayer capacity in terms of volume for all species, reducing the

adjustable parameters from 8 to 5, for each temperature. Pervatech hydrophilic commercial

membranes were evaluated for the dehydration of ethanol, lactic acid and ethyl lactate, by

pervaporation. The permeances of all species were determined as function of composition and

temperature in the range 48-72ºC.

Process intensification: Mathematical models, considering external and internal mass-transfer

resistances and velocity variations due to the change of multi-component mixture properties, were

developed for chromatographic reactors — fixed bed and simulated moving bed reactor, SMBR

— and validated by experimental data. The pervaporation membrane reactor, PVMR, model also

takes into account, the permeation through the membrane, concentration and temperature

polarization effects, and non-isothermal operation. The theoretical assessment of the SMBR unit

behaviour was performed to analyse the effect of SMBR configuration, feed composition and

switching time on the reactive/separation regions and/or on the process performance at the

optimal operating points. The performance of the PVMR was evaluated for isothermal and non-

isothermal operation, and suitable conditions for maximization of both lactic acid conversions and

ethyl lactate purity were examined. Finally, a new technology was developed and submitted to

patent registration, the simulated moving bed membrane reactor, PermSMBR, which integrates

perm-selective membranes inside the SMBR columns.

Zusammenfassung

Ziel dieser Arbeit war es einen völlig neuartigen und effizienten Prozess für die Produktion von

“Grünem Lösungsmittel” und Ethyl-Lakton, mittels der Verästerung aus Ethanol und Milchsäure über

die Hybrid-Technologie, zu entwickeln. Die Reaktion/Trennung basiert auf simulierten

Fliessbettreaktoren und Membranprozessen mittels Pervaporization. Um diese Zielsetzung zu

erreichen, wurden folgende Themen diskutiert:

Folgende wichtige Daten wurden gesammelt: Für den Ionen-Austausch wurde das Harz, Amberlyst

15-wet, genutzt. Es wurde sowohl als Katalysator für die Reaktivverästherung, wie auch als selektiver

Adsorbent für Wasser ausgewertet. Die Kinetischen- und chemischen Gleichgewichtsdaten wurden

zwischen 50ºC und 90ºC gemessen. Als Vorlage für die Bestimmung der chemischen

Gleichgewichtskonstante wie auch der Kinetik als Funktion der Temperatur, wurde das UNIQUAC

Model genutzt. Die Adsorptionsdaten wurden zwischen 20ºC und 50ºC gemessen, und entsprechend

einer Langmuir-Isotherme für Multikomponenten angeglichen, wobei eine gleich grosse

Volumenkapazität der Monoschicht für alle Spezies angenommen wurde. Die Anzahl der zu

justierenden Parameter wurde hierbei für jede Temperatur von 8 auf 5 reduziert. Kommerzielle

hydrophile Membranen der Firma Pervatech wurden für die Dehydratisierung von Ethanol,

Milchsäure und Ethyllakton mittels Permeation ausgewertet. Die Permeation aller Spezies wurde als

Funktion der Zusammensetzung und Temperatur, zwischen 48ºC und 72ºC ermittelt.

Intensivierung des Prozesses: Es wurden mathematische Modelle für chromatographische Reaktoren

entwickelt (Modelierte Festbett- und Fliessbettreaktoren SMBR) und experimentell ausgewärtet, unter

Berücksichtigung des internen und externen Massenaustausches, sowie der variablen

Geschwindigkeiten. Es herrschen unterschiedlichen Vermischungseigenschaften der

Multikomponenten. Das mathematische Modell der Pervaporationsmembrane (PVMR) berücksichtigt

gleichfalls die folgenden Effekte: Polarization aufgrund von Temperatur- und

Konzentrationsgradienten sowie nicht-isotherme Reaktionsführung. Das theoretische Verhalten der

SMBR Einheit wurde unter Berücksichting der Konfiguration, der Mischungszusammensetzung beim

Eintritt und der Komutationszeit in den Trennungs- und Reaktionszonen und/oder für die

Prozessleistung an den operationallen Optima ausgewertet. Der Wirkungsgrad der PVMR wurde für

den isothermen und nicht-isothermen Betrieb ermittelt, und es wurden entsprechende Bedingungen für

den maximale Umsatz der Milchsäure und der maximalen Reinheit für Ethyl-Lakton bestimmt. Der

lezte Schritt war die Entwicklung einer neuen Technolgie, ein Membranreaktor mit simuliertem

Fliessbett (PermSMBR), welche als Patent angemeldet wurde. Dieser Reaktor integriert selektive

Membranen innerhalb der SMBR Säule.

Table of contents Pag. 1. Introduction .......................................................................................................................1

1.1 Relevance and Motivation..........................................................................................1 1.2 Objectives and Outline ...............................................................................................3

2. State of the art on Green Solvent Ethyl Lactate.............................................................5

2.1 Green Chemistry ........................................................................................................5 2.2 Ethyl lactate applications ...........................................................................................7 2.2.1 Solvent Market Analysis ................................................................................................................. 8 2.3 Synthesis of ethyl lactate............................................................................................9 2.3.1 Renewable Resources.................................................................................................................... 11

2.3.1.1 Ethanol Platform................................................................................................................................. 12 2.3.1.2 Lactic acid Platform............................................................................................................................ 13

2.3.2 Patented Processes Overview........................................................................................................ 13 2.3.3 Reactive Separations ..................................................................................................................... 16

2.3.3.1 Reactive Distillation (RD) .................................................................................................................. 17 2.3.3.2 Simulated Moving Bed Reactor (SMBR) ........................................................................................... 19 2.3.3.3 Pervaporation Membrane Reactor (PVMR)........................................................................................ 21

2.4 References ................................................................................................................26 3. Batch Reactor: Thermodynamic Equilibrium and Reaction Kinetics.......................35

3.1 Introduction ..............................................................................................................36 3.2 Experimental Section ...............................................................................................40 3.2.1 Chemicals and Catalyst ................................................................................................................. 40 3.2.2 Experimental set-up....................................................................................................................... 41 3.2.3 Analytical method ......................................................................................................................... 42 3.3 Thermodynamic Equilibrium Results ......................................................................42 3.3.1 Thermodynamic equilibrium constant........................................................................................... 42

3.3.1.1 Activity coefficients estimation .......................................................................................................... 44 3.3.2 Equilibrium constant and reaction enthalpy for the synthesis of Ethyl Lactate............................. 45 3.3.3 Application of this methodology to other works ........................................................................... 47 3.4 Kinetic Studies .........................................................................................................49 3.4.1 Preliminary Studies ....................................................................................................................... 50

3.4.1.1 Evaluation of external mass transfer limitations (effect of stirring speed).......................................... 50 3.4.1.2 Evaluation of internal mass transfer limitations (effect of particle size) ............................................. 50 3.4.1.3 Evaluation of catalyst deactivation (effect of catalyst reusability)...................................................... 51

3.4.2 Kinetic Model................................................................................................................................ 52 3.4.2.1 Parameter estimation from experimental data..................................................................................... 54

3.4.3 Modelling and discussion of results .............................................................................................. 55 3.4.3.1 Effect of catalyst loading .................................................................................................................... 56 3.4.3.2 Effect of initial molar ratio of reactants .............................................................................................. 57 3.4.3.3 Effect of reaction temperature............................................................................................................. 58

ii TABLE OF CONTENTS

3.4.3.4 Effect of Lactic acid and Ethyl Lactate oligomers...............................................................................58 3.4.3.5 Effect of polar species .........................................................................................................................62

3.5 Conclusions.............................................................................................................. 63 3.6 Notation.................................................................................................................... 64 3.7 References Cited ...................................................................................................... 66

4. Fixed Bed Adsorptive Reactor....................................................................................... 71

4.1 Introduction.............................................................................................................. 72 4.2 Experimental Section ............................................................................................... 73 4.2.1 Chemicals and Catalyst / Adsorbent ............................................................................................. 73 4.2.2 Experimental Apparatus................................................................................................................ 74

4.2.2.1 Bed Porosity and Peclet Number .........................................................................................................75 4.3 Modelling of Fixed Bed ........................................................................................... 76 4.3.1 Multi-component viscosity ........................................................................................................... 81 4.4 Results and Discussion ............................................................................................ 84 4.4.1 Adsorption Isotherm ..................................................................................................................... 84

4.4.1.1 Binary Adsorption experiments ...........................................................................................................85 4.4.2 Kinetic experiments ...................................................................................................................... 90

4.4.2.1 Fixed Bed Reactor ...............................................................................................................................90 4.5 Conclusions.............................................................................................................. 94 4.6 Notation.................................................................................................................... 94 4.7 References................................................................................................................ 97

5. Simulated Moving Bed Reactor................................................................................... 101

5.1 Introduction............................................................................................................ 102 5.2 Modelling Strategies .............................................................................................. 104 5.2.1 SMBR mathematical model ........................................................................................................ 104 5.2.2 SMBR performance parameters.................................................................................................. 108 5.2.3 Numerical Solution ..................................................................................................................... 108 5.3 Experimental Section ............................................................................................. 109 5.3.1 Chemicals and Catalyst / Adsorbent ........................................................................................... 109 5.3.2 The SMBR LICOSEP 12-26 Unit............................................................................................... 109 5.4 Results and Discussion .......................................................................................... 111 5.4.1 Experimental Results .................................................................................................................. 111 5.4.2 Simulated results......................................................................................................................... 115

5.4.2.1 Comparison of SMBR and TMBR models ........................................................................................115 5.4.2.2 Reactive/separation regions ...............................................................................................................116 5.4.2.3 Separation Region vs Reactive/Separation Region............................................................................117 5.4.2.4 Effect of the Feed Composition .........................................................................................................118 5.4.2.5 Effect of the SMBR columns arrangement ........................................................................................120 5.4.2.6 Effect of Switching Time...................................................................................................................121

5.5 Conclusions............................................................................................................ 123 5.6 Notation.................................................................................................................. 124 5.7 References Cited .................................................................................................... 127

PROCESS INTENSIFICATION FOR THE GREEN SOLVENT ETHYL LACTATE PRODUCTION iii

6. Pervaporation Membrane Reactor..............................................................................131 6.1 Introduction ............................................................................................................132 6.2 Experimental Section .............................................................................................135 6.2.1 Materials...................................................................................................................................... 135 6.2.2 Pervaporation Membrane Reactor Unit....................................................................................... 135 6.3 Pervaporation Studies.............................................................................................136 6.3.1 Pervaporation Transport .............................................................................................................. 137 6.3.2 Preliminary Studies ..................................................................................................................... 138

6.3.2.1 Evaluation of the membrane quality ................................................................................................. 138 6.3.2.2 Evaluation of mass transfer limitations in the boundary layer .......................................................... 139

6.3.3 Detailed Studies .......................................................................................................................... 140 6.3.3.1 Water/Ethanol System ...................................................................................................................... 140 6.3.3.2 Water/Ethyl lactate System............................................................................................................... 140 6.3.3.3 Water/Lactic acid System ................................................................................................................. 141 6.3.3.4 Membrane performance evaluation................................................................................................... 141

6.3.4 Parameters estimation ................................................................................................................. 143 6.3.4.1 Permeance temperature dependence ................................................................................................. 143 6.3.4.2 Permeance temperature and water content dependence .................................................................... 144 6.3.4.3 Estimation of the boundary layer mass transfer coefficient (kbl)....................................................... 148

6.4 Modelling ...............................................................................................................149 6.4.1 Batch Pervaporation Model......................................................................................................... 149 6.4.2 Pervaporation Membrane Reactor model .................................................................................... 152 6.5 Results and Discussion...........................................................................................155 6.5.1 Batch Pervaporation .................................................................................................................... 155 6.5.2 Pervaporation Membrane Reactor............................................................................................... 157 6.6 Conclusions ............................................................................................................161 6.7 Notation..................................................................................................................162 6.8 References ..............................................................................................................165

7. PermSMBR – A New Hybrid Technology ..................................................................171

7.1 Introduction ............................................................................................................172 7.2 Technical description of the PermSMBR technology............................................174 7.3 PermSMBR mathematical model...........................................................................178 7.4 PermSMBR geometrical specifications .................................................................181 7.5 Simulated Results...................................................................................................182 7.5.1 Reactive/Separation Region: PermSMBR vs SMBR .................................................................. 183 7.5.2 PermSMBR 3 zones .................................................................................................................... 184 7.5.3 Comparison between PermSMBR, SMBR and RD technologies ............................................... 187 7.6 Conclusions ............................................................................................................187 7.7 Notation..................................................................................................................188 7.8 References ..............................................................................................................191

8. Conclusions and Suggestions for Future Work..........................................................193 APPENDIX A. Safety Data ..................................................................................................A1

iv TABLE OF CONTENTS

APPENDIX B. Thermodynamic Properties ...................................................................... B1 APPENDIX C. Calibration ................................................................................................. C1 APPENDIX D. Binary adsorption experiments at 293.15 K ........................................... D1

1. Introduction

1.1 Relevance and Motivation

Petroleum (“black gold”) is at the heart of today’s economics and politics problems. The

traditional petroleum reserves are in decline; moreover, the environmental regulation is every

day more severe, being, therefore, a great challenge to the design and implementation of

green products and processes.

Green solvents, which are produced from the processing of agricultural crops, were

developed as a more environmentally friendly alternative to petrochemical solvents. Lactate

esters solvents are 100% biodegradable, easy to recycle, non-corrosive, non-carcinogenic and

non-ozone depleting. Lactate esters have found industrial applications in specialty coatings,

inks, cleaners and straight cleaning use.

Ethyl lactate is a green solvent derived from nature-based feedstocks and it is so benign that

the U.S. Food and Drug Administration approved its use in food products. Ethyl lactate could

replace a range of environment-damaging halogenated and toxic solvents, including ozone-

depleting chlorofluorocarbons, carcinogenic methylene chloride, and toxic ethylene glycol

ethers and chloroform. In Figure 1.1 the ethyl lactate life-cycle is shown.

Ethyl lactate is produced from the esterification of lactic acid with ethanol through a

reversible reaction, having water as a by-product. Traditionally, ethyl lactate is synthesized in

a reactor followed by separation units in order to recover it, to remove the by-product (water)

and to recycle the unconverted reactants to the reactor; however, this represents high costs.

The objective of this work is to study equipments and techniques that are more compact,

2 CHAPTER 1. Introduction

energy efficient, and environment-friendly sustainable processes for the ethyl lactate

production.

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Figure 1.1- Ethyl Lactate life-cycle.

Process intensification, regarding the integration of reaction and separation processes into a

single device, provides the most feasible engineering solution to the sustainable synthesis of

ethyl lactate, since at least one of the products is being removed from the reaction medium to

lead to depletion of the limiting reactant. In this perspective, continuous chromatographic

reactor and membrane reactor will be considered for ethyl lactate production, namely the

Simulated Moving Bed Reactor (SMBR) and the Pervaporation Membrane Reactor (PVMR),

respectively. The SMBR is a competitive technology for systems involving equilibrium

controlled reactions catalysed by ion exchange resins, which are also selective adsorbent for

water (the by-product formed in the ethyl lactate synthesis), given that the products are

formed and simultaneously separated and removed from the reaction medium. The PVMR

technology is a clean and economic alternative to conventional processes, since equilibrium

could be shifted by continuously removing water through a selective membrane, allowing

costs reduction and higher product purity. Combining the advantages of both technologies,

finally, a new hybrid technology will be developed, the Simulated Moving Bed Membrane

Reactor (SMBMembR or PermSMBR) that combines a reactor with two different separation

techniques into a single device: continuous counter-current chromatography (SMB) with a

selective permeable membrane (Pervaporation or Permeation).

PROCESSES INTENSIFICATION FOR THE GREEN SOLVENT ETHYL LACTATE PRODUCTION 3

1.2 Objectives and Outline

The main goal of this work is the development of a process for simultaneous reaction and

separation in a single device for ethyl lactate production with a high purity, yield and

complete reactants conversion. Three technologies will be studied: the Simulated Moving

Bed Reactor (SMBR) using a catalyst that is also a selective adsorbent for water/ethyl lactate

separation, the Pervaporation Membrane Reactor (PVMR) using a hydrophilic water

permselective membrane for continuous removal of the by-product water and the Simulated

Moving Bed Membrane Reactor (PermSMBR), where the SMBR is integrated with the

PVMR by using selective permeable membranes inside the columns of the SMBR.

The thesis comprises 8 chapters dealing with different aspects of ethyl lactate production, in

addition to the present one.

In Chapter 2, the state of the art of production process aspects as patented processes for esters

production; the advantages of heterogeneous catalyst, such as ion-exchange resins; the

methods used to displace equilibrium towards ester formation are addressed. An overview in

some reactive separations as reactive distillation, chromatographic reactors and membrane

reactors applied to the production of oxygenates is reviewed in order to improve the overall

efficiency of the process of ethyl lactate synthesis.

Chapter 3 addresses the kinetic studies for ethyl lactate production by heterogeneous

catalysis; the influence of catalyst loading, temperature and initial molar ratio of reactants are

analysed. A methodology based in the UNIQUAC model for determination of the

thermodynamic equilibrium constant is developed.

Experimental and simulated results for the ethyl lactate production in a fixed bed adsorptive

reactor are shown in Chapter 4. Dynamic adsorption experiments of binary non-reactive

mixtures were performed in order to obtain multicomponent adsorption equilibrium isotherms

of Langmuir type. The reaction kinetics and adsorption data were used in the mathematical

model of the adsorptive reactor, which also included axial dispersion, velocity variations and

external and internal mass-transfer resistances.

In Chapter 5, the simulated moving bed reactor technology for the ethyl lactate production is

evaluated by experiments as well as by simulations. In order to describe the dynamic

behaviour of this unit, a mathematical model considering external and internal mass-transfer

resistances and variable velocities is developed. The influence of operational parameters, as

4 CHAPTER 1. Introduction

feed composition, SMBR configuration and switching time, on the SMBR performance is

presented.

Pervaporation processes using hydrophilic silica membranes are evaluated in Chapter 6 for

the ethyl lactate system. The effects of feed composition and operating temperature on the

membrane performance are analyzed. Mathematical models, considering concentration and

temperature polarization and non-isothermal effects, are developed and applied to analyze the

performance of batch pervaporation and continuous pervaporation membrane reactor, in both

isothermal and non-isothermal conditions.

In Chapter 7, a new technology, the simulated moving bed membrane reactor, is presented

and applied for the ethyl lactate synthesis. The potential of this new equipment is

demonstrated by comparing its performance to other reactive separation processes.

Finally, the general conclusions drawn from this work and the suggestions for future work

will be presented in Chapter 8.

2. State of the art on Green Solvent Ethyl Lactate

In this chapter, a review on the green chemistry principles is made and a literature survey on

applications of ethyl lactate and production processes (renewable resources, patents, reactive

separations) is presented.

2.1 Green Chemistry

Green chemistry is the best use of chemistry for pollution prevention. More specifically,

green chemistry is the design of chemical products and processes that reduce or eliminate the

use and generation of hazardous substances. It is a highly effective approach to pollution

prevention because it applies innovative scientific solutions to real-world environmental

situations.

Currently, a great challenge is the design and implementation of completely green products

and processes. There is not a systematic and reliable method for ensuring that the chemistry

being implemented is green, since the number of chemicals synthesis pathways is enormous.

Indeed, it is more correct to verify if a proposed manufacturing process is “greener” than

other alternatives. Anastas and Warner have developed the “Twelve Principles of Green

Chemistry” to aid one in assessing how green is a product or a process (Anastas and Warner,

1998), which are:

1. It is better to prevent waste than to treat or clean up waste after it is formed.

2. Synthetic methods should be designed to maximize the incorporation of all materials

used in the process into the final product.

6 CHAPTER 2. State of the art on Green Solvent Ethyl Lactate

3. Wherever practicable, synthetic methodologies should be designed to use and generate

substances that possess little or no toxicity to human health and the environment.

4. Chemical products should be designed to preserve efficacy of function while reducing

toxicity.

5. The use of auxiliary substances (e.g. solvents, separation agents, etc.) should be made

unnecessary whenever possible and, innocuous when used.

6. Energy requirements should be recognized for their environmental and economic

impacts and should be minimized. Synthetic methods should be conducted at ambient

temperature and pressure.

7. A raw material or feedstock should be renewable rather than depleting whenever

technically and economically practical.

8. Unnecessary derivatization (blocking group, protection/deprotection, and temporary

modification of physical/chemical processes) should be avoided whenever possible.

9. Catalytic reagents (as selective as possible) are superior to stoichiometric reagents.

10. Chemical products should be designed so that at the end of their function they do not

persist in the environment and break down into innocuous degradation products.

11. Analytical methodologies need to be further developed to allow for real-time in-

process monitoring and control prior to the formation of hazardous substances.

12. Substances and the form of a substance used in a chemical process should be chosen so

as to minimize the potential for chemical accidents, including releases, explosions, and

fires.

Based on these twelve principles, this thesis focuses:

1) The synthesis of the green solvent Ethyl lactate produced from renewable raw

material that is a more environmentally friendly alternative to petrochemical

solvent: 7th principle.

2) Ethyl lactate is 100% biodegradable, easy to recycle, non-corrosive, non-

carcinogenic and non-ozone depleting: 3rd, 4th and 10th principles.

3) The use of solid acid catalysts to improve the reaction kinetics without using

increase the stoichiometric of reactants, and it is more advantageous then

PROCESS INTENSIFICATION FOR THE GREEN SOLVENT ETHYL LACTATE PRODUCTION 7

homogenous catalysts, since these are more corrosive and require a further step

of neutralization: 1st and 9th principles.

4) The process intensification by using hybrid technologies where reaction and

separation of at least one product take place in a single unity (SMBR, PVMR

and PermSMBR) will reduce/eliminate the use of solvents and requires less

energy consumption: 5th and 6th principles.

2.2 Ethyl lactate applications

The reaction between an alcohol and a carboxylic acid to form an ester and water is of

considerable industrial interest (Dhanuka et al., 1977). Organic esters are a very important

class of chemicals having applications in a variety of areas in the chemical industries such as

perfumes, flavours, pharmaceuticals, plasticizers, solvents and intermediates (Weissermel and

Arpe, 1997).

Ethyl lactate is an important organic ester, which is biodegradable and can be used as food

additive, in perfumery, as flavour chemicals and solvent, which can dissolve acetic acid

cellulose and many resins (Tanaka et al., 2002). It is a particularly attractive solvent for the

coatings industry as a result of its high solvency power, high boiling point, low vapour

pressure and low surface tension. Ethyl lactate is a desirable coating for wood, polystyrene

and metals and also acts as a very effective paint stripper and graffiti remover. It has replaced

solvents including N-methyl Pyrrolidone (NMP) (Reisch, 2008), toluene, acetone and xylene,

which has resulted in the workplace being made a great deal safer. In Table 2.1 solvating

properties of ethyl lactate and NMP are presented.

Table 2.1 Solvating properties of ethyl lactate and N-methyl Pyrrolidone.

Ethyl Lactate N-methyl pyrrolidone

Kauri Butanol(KB) Value >1000 350

Hildebrand 21.3 23.1

Disperse 7.8 8.8

Polar 3.7 6.0

Hydrogen 6.1 3.5

Solubility Miscible in Water and Hydrocarbons

Miscible in Water and Hydrocarbons


Other applications of ethyl lactate include being an excellent cleaner for the polyurethane

industry and for metal surfaces, efficiently removing greases, oils, adhesives and solid fuels.

Beyond all these applications, ethyl lactate can also be used in the pharmaceutical industry as

a dissolving/dispersing excipient for various biologically active compounds without

destroying the pharmacological activity of the active ingredient. It proves to be a very

effective agent for solubilising biologically active compounds that are difficult to solubilise in

usual excipients (Muse and Colvin, 2005).

Ethyl lactate can also be applied as a more environment friend alternative route to produce

1,2-propanediol, which is normally produced by the hydration of propylene oxide derived

from petrochemical resource (Huang et al., 2008). In Table 2.2 the major benefits of the ethyl

lactate are presented.

Table 2.2 Ethyl lactate major benefits.

Ethyl Lactate Benefits 100% Biodegradable Renewable - made from corn and other carbohydrates

FDA approved as a flavour additive EPA approved SNAP solvent

Non carcinogenic Non corrosive

Great penetration characteristics Stable in solvent formulations until exposed to water

Rinses easily with water High solvency power for resins, polymers and dyes

High boiling point Easy and inexpensive to recycle

Low VOC Not a Ozone Depleting Chemical

Low Vapor Pressure Not a Hazardous Air Pollutant

2.2.1 Solvent Market Analysis

Almost all manufacturing and processing industries depend on the use of solvents (see Figure

2.1). The world solvent market is estimated at 30 million pounds per year at prices from $0.90

to $1.70 per pound. The ethyl lactate green solvent has the potential to displace 80 % of these

solvents (Energetics Incorporated, 2003).


Figure 2.1 Solvent demand (AAE Chemie, 2009).

Selling prices for ethyl lactate have ranged from $1.50 to $2.00 per pound, but processing

advances could drive the price as low as $1.00 to $0.85 per pound (Argonne, 2006), enabling

ethyl lactate to compete directly with the petroleum-derived toxic solvents currently used.

Moreover, the crude prices have risen sharply, making of ethyl lactate green solvent more

commercially attractive. Among this, due to an environmental consciousness, some

consumers are willing to pay more for products that are less detrimental to the environment.

2.3 Synthesis of ethyl lactate

The conventional way to produce ethyl lactate is the esterification of lactic acid with ethanol

catalyzed by an acid catalyst, according to the reaction:

)()()()( WWaterELLactateEthylLaAcidLacticEthEthanol H +⎯⎯→←++

The use of these reactants (ethanol and lactic acid) has the advantage of both being produced

from renewable resources (by glucose or sugar fermentation processes).

Esterifications are self-catalyzed reactions, since the H+ cation released from the partial

dissociation of the carboxylic acid used as reactant catalyses the reaction. However, the use of

catalyst is favourable for the reaction rate as the kinetics of the self-catalyzed reaction is

extremely slow, since its rate depends on the autoprotolysis of the carboxylic acid. For

Agricultural chemicals 2% Dry Cleaning 1%

Others 8%

Paints 46%

Pharmaceuticals 9%

Adhesives 6%

Printing Inks 6%

Personal Care 6%

House/Car 6%

Metal/Industrial Cleaning 4%

Rubber/Polymer Manufacture 4%

Oil Seed Extract 2%


example, the lactic acid acidity constant is pKa=3.86 @ 25 ºC, and therefore an aqueous

solution with 85 % of lactic acid (about 10.8 M) has a pH=1.4. Typically, the catalytic

production of lactates is performed with homogeneous catalysts using acids, such as sulphuric

acid, phosphoric acid and anhydrous hydrogen chloride. However, the use of heterogeneous

catalyst (as for example, zeolites, ion-echange resins like Amberlyst 15-wet, Nafion NR50,

among others) has clear advantages:

- easy to separate from the reaction medium;

- long life time;

- higher purity of products (side reactions can be eliminated or are less significant);

- elimination of the corrosive environment caused by the discharge of acid

containing waste.

As previously mentioned, the esterification is a reversible reaction and, in order to obtain

acceptable ester yields, the equilibrium must be displaced towards the ester production, which

might be accomplished by different methods, such as:

1. to use a large excess of one of the reactants, in general the alcohol; however, this

results in a relatively inefficient use of reactor space and in very diluted products,

which will require an efficient separation afterwards;

2. to eliminate the water by azeotropic distillation between a solvent and water – the

solvent and water must be partially miscible and the boiling points of the different

components in the reaction medium must be compatible with that azeotrope;

3. to use reactive separations (as reactive distillation, simulated moving bed reactor,

pervaporation reactor, etc.) in order to remove the products from the reaction medium.

In reactions limited by chemical equilibrium where more than one product is formed

conversion can be enhanced in multifunctional reactor where the products are separated as

they are formed. Novel reactor configurations and choice of operating conditions can be used

to maximise the conversion of reactants and improve selectivity of desired product, thereby

reducing the costs associated with the separation step. Recently, reactive distillation,


chromatographic reactors and membrane reactors have been intensively applied to

esterifications processes, as it will be discussed latter within this chapter.

2.3.1 Renewable Resources

In recent years, an increasing demand on using biorenewable materials instead of petroleum

based feedstocks for producing chemicals, driven by environmental concerns and by the

concept of sustainability, has been noticed. Biobased products are one of the main pillars of a

sustainable economy. Nature produces 170 billion tons of biomass per year by

photosynthesis, 75 % of which belong to the class of carbohydrates; however, just 3-4 % of

these compounds are used by humans for food and non-food purposes (Röper, 2002).

Carbohydrates are very abundant renewable resources and they are currently considered as an

important feedstock for the Green Chemistry of the future (Lichtenthaler, 1998; Lichtenthaler,

2002; Lichtenthaler and Peters, 2004). Industrial plants, named as biorefineries, have been

created where biomass is converted economically and ecologically, in chemicals, materials,

fuels and energy (see Figure 2.2). The biorefineries could be the basis of the new bioindustry

and its concept is similar to the petroleum refinery; the difference is that the biorefinery is

based on conversion of biomass feedstocks instead of crude oil.

Figure 2.2 Schematic diagram of a biorefinery for precursor-contained

biomass. (Kamm and Kamm, 2004a; Kamm and Kamm, 2004b)


2.3.1.1 Ethanol Platform

Ethanol is an important raw material in the chemical industry and can also be used as

transportation fuel. It can be produced from a variety of biomass crops, including sugar crops

(e.g., sugarcane and sugar beet), starch crops (e.g., corn and cassava), or cellulosic feedstocks

(e.g., wood, grasses and agricultural residues).The production of ethanol from starch crops

involves as main steps: liquefaction and saccharification (conversion to sugar), milling,

pressing, fermentation and distillation. The production from cellulosic feedstocks is similar,

however it is significantly more difficult and costly to convert cellulose and hemicellulose

into their component sugars (glucose and xylose, respectively) than is the case for starches

(Sagar and Kartha, 2007). Currently, more than 37 billion litters of ethanol are produced

worldwide per year from starch and sugar crops (Rass-Hansen et al., 2007; Tilman et al.,

2006). In 2008, cellulosic ethanol industry developed some new commercial-scale plants. In

the United States, plants with 12 million liters capacity per year were operational, and an

additional 80 million liters per year of capacity (26 new plants) was under construction. In

Canada, capacity of 6 million liters per year was operational. In Europe, several plants were

operational in Germany, Spain, and Sweden, and capacity of 10 million liters per year was

under construction (REN21, 2009). Ethanol derived from cellulosic crops is appealing since it

broadens the scope of potential feedstocks beyond starch and sugar-based food crops.

Moreover, cellulosic ethanol can be more effective and promising as an alternative renewable

biofuel than corn ethanol because its use reduces even more the net greenhouse gas (GHG)

emissions when compared with the petroleum fuel (Wang et al., 2008) (see Figure 2.3).

Figure 2.3 Percent change in greenhouse gas emissions (adapted from Wang et al, 2008).


2.3.1.2 Lactic acid Platform

Lactic acid (2-hydroxypropionic acid) is an important platform chemical for the biorenewable

economy. It is an α-hydroxy acid containing a hydroxyl group adjacent to the carboxylic acid

functional group; a review on the lactic acid chemistry can be found in literature (Holten,

1971). Lactic acid can be produced through chemical synthesis or through the fermentation of

different carbohydrates, such as, glucose (from starch), maltose (produced by specific

enzymatic starch conversion), sucrose (from syrups, juices, and molasses), or lactose (from

whey) (Corma Canos et al., 2007). Nowadays, it is commercially produced by fermentation

of glucose. One of the most important steps in the lactic acid production is the recovery from

fermentation broths. The separation and purification stages represent about 50 % of the total

production cost. However, current advances in membrane-based separation and purification

technologies, particularly in microfiltration, ultrafiltration and electrodialysis, have originated

new processes which should reduce the lactic acid cost production (Wasewar et al., 2004).

The lactic acid production is around 350,000 tons per year and it is defended by some

observers that the worldwide growth per year is of 12-15 % (Wasewar et al., 2004). A lot of

products are derived from lactic acid; some of them are new chemical products and others,

which represent biobased routes to chemicals, currently produced from petroleum. The most

important ones are shown in Figure 2.4.

Figure 2.4 Some potential derivatives of lactic acid (Corma Canos et al., 2007).

2.3.2 Patented Processes Overview

There are a great number of patents related to esters production. A summary of those patented

processes is presented in Table 2.3.

Table 2.3 Patented processes for esters production. Commercial Solvents

Corporation (Bannister, 1936)

Lactic acid is dehydrated and mixed with ethanol and concentrated sulphuric acid (catalyst). This mixture is refluxed for one hour and the esterification takes place to a determined extent. Afterwards, distillation is started to separate the components of the reaction mixture.Temperature range of the process 135ºC-145ºC.

USA-Secretary of Agriculture

(Filachione and Fisher, 1951)

Process to produce esters from the reaction between the basic nitrogen salt of the carboxylic acid with an alcohol. It is a batch process and distillation is applied to separate the final mixture components. The process achieved from 61 to 92% ammonia removal and from 49 to 67% conversion to butyl lactate. Temperature range 89ºC-195ºC.

The American Oil Company (Jennings and Binning, 1960)

Process, where the esterification extent is enhanced, using pervaporation membranes in order to selective remove water from the reaction zone. It is a continuous method, which integrates reaction and separation in the same unit and uses as catalyst acid ion exchange resins. A temperature of 100ºC to 200ºC is maintained in the feed zone and the pressure is kept in order to maintain the water in liquid phase.

BASF Aktiengesellschaft (Bott et al., 1986)

Process to produce optically pure alkyl D- or L-lactates by reaction of calcium lactate with an alcohol in the presence of a strong acid; the water present in the reaction mixture or formed during the esterification is separated off by azeotropic distillation with the aid of an entraining agent.

Battelle Memorial Institute (Walkup et al., 1991; Walkup

et al., 1993)

Batch process for the preparation of esters of lactic acid directly from ammonium lactate and an alcohol. In this method the use of CO2 as acatalyst is required and the preferred range for the reaction mixture temperature is from 100ºC to 200ºC. A yield of lactate of about 75% is reported.

E. I. Du Pont de Nemours and Company

(Cockrem and Johnson, 1993)

Recovery of high purity lactate ester from fermentation broth containing ammonium lactate or other basic salt of lactic acid; acidifying in the presence of an alcohol using continues addition of sulphuric acid or other strong acid and crystallizing to precipitate out some or all of the basic salt of the strong acid; simultaneously or sequentially removing water while also esterifying the lactic acid with the alcohol to form impure lactate ester; removing the crystals formed; distilling the lactate ester to remove impurities.

Musashino Chemical Laboratory Ltd.

(Akira et al., 1994)

Method for producing a lactic ester by microorganic fermentation of lactic acid with a simple apparatus. A pressure in the range of 100 to 760 mmHg and a temperature of about 130ºC are recommended for this process.

BASF Aktiengesellschaft (Sterzel et al., 1995)

Process for the synthesis of lactates by fermentation of sugars mixtures, conversion of the lactic acid obtained during the fermentation to its salts, followed by esterification.

DAICEL CHEM IND LTD (Yukio, 1996)

In this process lactic acid is esterified with ethanol in the presence of a catalyst such as p-toluenesulfonic acid. The catalyst, water and unreacted ethanol are removed to give a solution (A). The solution A is neutralized with (B) a solution of an alkali metal salt in an alcohol and distilled to give the ethyl lactate.

(Feng et al., 1996) Rectification process to produce ethyl lactate from lactic acid and ethanol. This patented process includes technological steps, such as, determination of lactic acid, ethanol, sulphuric acid and benzene amounts, catalytic reaction, rectification for dewatering, neutralization and reduced distillation. Yield rate up to over 90% is reported.

Argonne National Laboratory

(Rathin and Shih-Perng, 1998)

Process for the synthesis of high purity ethyl lactate and other lactate esters from carbohydrate feedstock. This process consists in a reactor coupled with a pervaporation membrane unit for water removal and followed by separation of the reaction mixture in two consecutive distillation columns. Alternatively, the reactor is followed by a plurality of pervaporation steps. It is reported a conversion greater than 99 %, for an initial ethanol/lactic acid molar ratio of 2:1, a reaction mixture temperature of 95ºC, a permeate-side vacuum pressure less than 0.5 mbar and as catalyst an ion-exchange resin, Amberlyst XN-1010, at 10% of lactic acid weight.

Mitsubishi Gas Chemical Company, Inc.

(Abe et al., 1998)

Process to produce lactates from acetaldehyde and formate; The method described is characterized by the fact that there is no formation of ammonium salts as by-products as in the case of the conventional techniques.


A.E. Staley Manufacturing Co.

(Cockrem, 2001)

Process for the simultaneously production of an organic acid and of an ester of the organic acid. A mixture of ammonium salt of an organic acid with alcohol is rapidly heating in order to produce a liquid stream containing acid, ester, and unreacted ammonium salt.

Eastman Chemical Company (Arumugam et al., 2003)

Process where a solution of carboxylic acid in a solvent and an alcohol are fed to a simulated moving bed reactor (SMBR), that contains a solid(s) (adsorber/catalyst) to produce two streams, one comprising a solution of the ester of the carboxylic acid and the alcohol and other comprising the solvent. The process is particularly valuable for the preparation of an alkanol solution of an alkyl 2-keto-L-gulonate ester (AKLG). The simulated moving bed reactor is maintained at a temperature of 30 to 60ºC and a pressure of 3.5 to 20 bars.

Cargill, Incorporated (Eyal et al., 2003; Eyal et al.,

2006)

Techniques for processing lactic acid/lactate salt mixtures obtained from fermentation broths. These techniques generally concern the provision of separated lactic acid and lactate streams from the mixtures. In this patent preferred methods of separation and processing of each of the streams are provided.

A.E. Staley Manufacturing Co.

(Cockrem, 2003)

Process to produce an ester that comprises the following steps: (a) feeding to a first vessel a mixture of organic acid, alcohol, and water, where the organic acid and alcohol react to form monomeric ester and water (temperature of 150 to 220ºC) (b) feeding the mixture obtained in to a second vessel (temperature range 30 to 100ºC), where are produced a vapour stream, that comprises alcohol, ester and water, and a liquid one, that can be recycled to the first vessel.

La Chemical SpA (Ruggieri et al., 2003)

Process to produce esters in a chromatographic reactor in which the heterogeneous solid phase acts both as catalyst and as a means exhibiting preferential adsorption towards one of the reaction products (typically water). This process is particularly improved compared with the conventional technology since for regenerating the catalyst, it is used a desorbent mixed with a second compound, normally the anhydride of the acid used in the esterification reaction, which, by chemical reaction, completes the removal of the water adsorbed.

(Xueming and Jing, 2003) Disclose a technique that uses ammonium lactate as raw material to make ethyl lactate by rectifying. The adopted equipment includes rectifying tower, condenser on the top of tower and oil-water separator. It uses metal halide as catalyst, an initial molar ratio of lactic acid and ethanol being 1:1 and benzene being 30%-50% of lactic acid ammonium weight.

Arkema (FR) (Tretjak et al., 2006; Tretjak

and Teissier, 2004)

Process that relates to a continuous method to produce ethyl lactate from the esterification between lactic acid and ethanol in the presence of a catalyst (H2SO4 98%) ; this method consists in continuously extracting a mixture comprising ethyl lactate, ethanol, water and different heavy products from the reaction medium at partial lactic acid conversion rate and, then, fed the mixture to a reduced-pressure flash separation, producing an overhead stream containing a mixture of ethyl lactate, ethanol and water, that is subjected to a fractional distillation column. A purity higher than 94.6 % of ethyl lactate is reported for an initial ethanol/lactic acid molar ratio equal to 2.5; esterification carried out at 80ºC; flash separation at 85ºC and 50 mbar, and fractional distillation at a column bottom temperature of 155°C and top temperature of 77.2ºC.

Arkema (FR) (Martino-Gauchi and Teissier,

2004; Martino-Gauchi and Teissier, 2007)

Continuous method for preparing ethyl lactate which consists in reacting lactic acid with ethanol (ethanol/lactic acid molar ratio higher than 2.5) in the presence of a catalyst (H2SO4 98%) at a reflux of the reaction medium of about 100ºC under pressure ranging between 1.5 to 3 bars. This method is characterized by the continuous extraction of a near-azeotropic water/ethanol gas mixture from the esterification reaction medium, followed by dehydration of this mixture using molecular sieves and recuperation from the dehydration mixture an ethanol gas stream capable of being recycled to the esterification reaction medium and a flow consisting of water and ethanol which is fed to a distillation column.

Board of Trustees of Michigan State

(Miller et al., 2006)

Lactic acid esterification by continuous countercurrent reactive distillation with alcohols, especially ethanol (to produce ethyl lactate). In this invention recycle of dimmers and trimmers and other oligomers of lactic acid are provided in order to improve yields. For absolute ethanol fed near to the bottom of the column at 82ºC and lactic acid solution (85 wt % in water) fed near to the top of column at 25ºC (molar ratio ofethanol to lactic acid of 3.3) it is reported a lactic acid conversion of 83% and a ethyl lactate yield of 82%.

Roquette Freres (Fuertes et al., 2008)

Method for preparing a lactic acid ester composition based on a lactic acid composition involving two steps: (a) transforming of the composition into a lactic acid oligomeric composition; (b) mixing and reacting the oligomeric composition with an alcohol, in the presence of a transesterification catalyst, to esterify all or part of the lactic acid contained in the oligomeric composition. This invention also discloses the use of ethyl lactate as solvent for preparing gelified compositions.

The esterification reaction is usually catalyzed by a strong acid, being most common

sulphuric acid. However, the use of solid acid catalysts, as ion exchange resins, is also

mentioned. In order to overcome equilibrium limitations excess of one reactant is commonly

applied, normally the alcohol. Another technique is the use of a solvent, as benzene,

substantially immiscible with water in order to extract the ester. Reaction and separation are,

in almost all patented processes, separated steps, being distillation the most used separation

technology.

2.3.3 Reactive Separations

In the last years, chemicals, petrochemicals and pharmaceuticals industries have been gone

through a permanently increasing interest in the development of hybrid processes combining

reaction and separation mechanisms into a single, integrated operation known as ‘reactive

separation’. The combination of the two stages into a single unit brings important advantages,

such as energy and capital cost reductions, increased yield and removal of some

thermodynamic restrictions, e. g. azeotropes. A variety of separation principles and concepts

can be incorporated into a reactor, see Figure 2.5.

Figure 2.5 Separation functions integrated into a reactor.

Important examples of reactive separations are reactive distillation, reactive absorption,

reactive extraction or reactive membrane separation. Until now, such processes have had

industrial application, mainly in areas like the homogeneously catalysed synthesis of acetates

and the heterogeneously catalysed production of fuel additives. The potential is much wider;

however, optimal functioning depends on careful process design, with appropriately selected


column internals, feed locations and catalyst placement. Greater understanding of the general

and particular features of the process behaviour is equally essential.

2.3.3.1 Reactive Distillation (RD)

Reactive distillation (RD) is an unit operation that combines chemical reaction and distillation

within a single vessel, thereby reducing equipment and recycle costs. A typical RD column is

shown in Figure 2.6. Other advantages offered by reactive distillation include high selectivity,

reduced energy uses, and reduction or elimination of solvents (Malone and Doherty, 2000). It

is an effective method that has considerable potential for carrying out equilibrium-limited

reactions, such as esterification and ester hydrolysis reactions; conversion can be increased

far beyond chemical equilibrium conversion due to the continuous removal of reaction

products from the reactive zone.

B

D

Reactive Section

A

C

Figure 2.6 Typical Reactive Distillation Column ( b,C b,B b,A b,DT < T < T < T ).

Reactive distillation has received much attention in the last years (Sundmacher and Kienle,

2002; Taylor and Krishna, 2000; Tsai et al., 2008). It has been used for the esterification of

fatty acids (Dimian et al., 2008; Steinigeweg and Gmehling, 2003), as well as being devised

as a new method to clean industrial water from acetic acid (Bianchi et al., 2003). The RD

technology was, also, applied on the ethyl lactate synthesis, first by Asthana and collaborators

(Asthana et al., 2005), where it is reported higher lactic acid conversions (>95 %) and good

ethyl lactate yields (>85 %), and, more recently, by Gao and co-workers (Gao et al., 2007).


However, applications of this technology in industry are still limited to a few reactive

systems, mainly etherification (e.g. MTBE), esterification (e.g. methyl acetate), and

alkylation (e.g. ethylbenzene or cumene) (Tuchlenski et al., 2001).

The production of methyl acetate is a classic example of successful RD (Agreda and Partin,

1984; Agreda et al., 1990). Conventional processes use one or more liquid-phase reactors

with large excess of one reactant in order to achieve high conversions of the other. A typical

flow sheet of a conventional process for the methyl acetate production is shown in Figure 2.7

in which the reaction section is followed by eight distillation columns, one liquid-liquid

extractor and a decanter. This process requires a large capital investment, high energy costs

and a large inventory of solvents. In the reactive distillation process for methyl acetate, the

entire process is carried out in a single unit (see Figure 2.7), which represents one-fifth of the

capital investment of the conventional process and consumes only one-fifth of the energy

(Krishna, 2002).

Figure 2.7 Task-integrated methyl acetate column is much simpler than conventional plant (Stankiewicz and Moulijn, 2000).

In spite of all advantages of the RD technology, there are still some constraints and

difficulties in its implementation, mainly due to volatility limitations. In order to maintain

high concentrations of reactants and low concentrations of products in the reaction zone, the

reactants and products must have suitable volatility. Also, it is necessary that both products

have different boiling points to ensure the separation. The major disadvantage of RD


technology, for exothermic reactions, is that chemical reaction has to show significant

conversion at distillation temperature (Pöpken et al., 2000).

2.3.3.2 Simulated Moving Bed Reactor (SMBR)

Simulated Moving Bed (SMB) systems are used in industry for separations that are either

impossible or difficult using traditional techniques. This technology uses differences in the

adsorptivity of the different components involved rather than differences in their volatility,

being an interesting alternative to distillation when the species involved exhibit small

volatility differences, are non-volatile or are sensitive to temperature, as in the case of many

fine chemical and pharmaceutical applications.

The combination of SMB and chemical reaction has been, in the last years, a subject of

considerable attention in the scientific research, being this integrated reaction-separation

technology called Simulated Moving Bed Reactor (SMBR). A schematic diagram of a SMBR

unit is presented in Figure 2.8 where a reaction of type A+B↔C+D is considered, for the case

of D being more adsorbed than C. The SMBR consists of a set of columns connected in series

that are packed with a solid, which acts as both adsorbent and catalyst. Typically, there are

two inlets (feed and desorbent) and two outlets (extract and raffinate). The component A is

used as reactant and desorbent, therefore it is introduced in the system in the feed and

desorbent streams. The other reactant B is used as feed. The products D and C are collected in

the extract and the raffinate, respectively, since D is more adsorbed than C. At regular time

intervals, called switching time period, all streams are switched for one bed distance in

direction of the fluid flow. A cycle is completed when the number of switches is equal to a

multiple of the columns number. In this way, the countercurrent motion of the solid is

simulated with a velocity equal to the length of a column divided by the switching time.

According to the position of the inlet and outlet stream the unit can be divided in four

sections. In section I, positioned between the desorbent and extract nodes, the adsorbent is

regenerated by desorption of the more strongly adsorbed product (D) from the solid. In

section II (between the extract and feed node) and section III (between the feed and raffinate

node) the reaction is taking place and products (C and D) are formed. The more strongly

adsorbed product D is adsorbed and transported with the solid phase to the extract port. The

less strongly adsorbed product C is desorbed and transported with the liquid in direction of

the raffinate port. In section IV, positioned between the raffinate and desorbent node, the

desorbent is regenerated before being recycled to section I.


Desorbent (A) Raffinate (A+C)

Extract(A+D)

Direction of fluid flow and port switching

1 2

3 4

5 6

12 11

10 9

8 7

Feed(A+B)

A + B C + D

Figure 2.8 Scheme of a Simulated Moving Bed Reactor (SMBR).

The simultaneous reaction and separation in a SMBR has shown that considerable

improvements in some processes performance can be achieved, as for example, on the

synthesis of acetals (Pereira et al., 2008; Rodrigues and Silva, 2005; Silva and Rodrigues,

2005), fructose (Azevedo and Rodrigues, 2001; Da Silva et al., 2005; Zhang et al., 2004),

lactosucrose (Kawase et al., 2001; Pilgrim et al., 2006), methylacetate (Lode et al., 2003) and

MTBE (Zhang et al., 2001). In the last years, cation exchange resins are being widely used as

catalyst of esterifications and acetalizations. Moreover, those resins adsorb selectively water,

a by-product of those kind of reactions. Therefore, combining these two properties of acidic

resins, and knowing that esterifications are reversible reactions, chromatographic reactors

appear as promising technologies; in particular the SMBR, since the products are

continuously separated and removed from the reaction medium, leading to complete

conversion. This has motivated several studies on esterification reactions by means of the

SMBR technology; examples are the esterification of acetic acid with methanol (Lode et al.,

2003; Yu et al., 2003), ethanol (Mazzotti et al., 1996) and β-phenethyl alcohol (Kawase et

al., 1996), and the esterification of acrylic acid with methanol to form methyl acrylate

(Ströhlein et al., 2006).

Although the SMBR technology allows 100 % of conversion with 100 % of recovery of the

desired product, a further step is necessary to separate the product from the raffinate mixture,

and to recover the reactant A, used as desorbent, from both extract and raffinate streams, in

order to recycle it to the SMBR unit.


2.3.3.3 Pervaporation Membrane Reactor (PVMR)

Pervaporation is one of the membrane processes that can be employed for the separation of

liquid mixtures that are difficult or not possible to separate by conventional methods, such as

distillation. One example is the application of pervaporation to break the azeotrope of the

mixture ethanol-water, where it is much more economical to use pervaporation or vapour

permeation than other conventional methods (see Table 2.4).

Table 2.4 Dehydration costs of ethanol from 99.4 to 99.9 vol % by different methods (Drioli and Romano, 2001).

Utilities Vapor

Permeation ($/ton)

Pervaporation ($/ton)

Entrainer Distillation

($/ton)

Molecular Sieve

Adsorption ($/ton)

vapor - 12.8 120.0 80.0

electricity 40.0 17.6 8.0 5.2

cooling water 4.0 4.0 15.0 10.0

Entrainer - - 9.6 -

Replacement of membranes and molecular sieves

19.0 30.6 - 50.0

total costs 63.0 65.0 152.6 145.2

The pervaporation process has significant separation potential for various types of solutions,

being specially suited for organic-water and organic-organic separations (Feng and Huang,

1996; Fleming and Slater, 1992; Huang and Rhim, 1991; Neel, 1991; Neel, 1995). It is used

to separate a liquid mixture by partly vaporizing it through a nonporous permselective

membrane, as shown in Figure 2.9. The feed liquid mixture is allowed to flow along one side

of the membrane, and a fraction of it, the “permeate”, is recovered in the vapour state on the

other side of the membrane, by means of vacuum or sweep gas. The mass transport through

the membrane is induced by maintaining a low vapour pressure on the permeate side,

eliminating thereby the effect of osmotic pressure. The permeate stream, enriched in the most


permeating component, might be then condensed in order to recover it. The remaining feed

that does not permeate through the membrane, called the “retentate”, is depleted in the

permeating component (Neel, 1995).

Feed Retentate

Permeate

Pervaporation membrane

Liquid

Vapor

Figure 2.9 Schematic representation of the pervaporation process.

There are a number of reviews on pervaporation processes (Dutta et al., 1996; Song et al.,

2004; Van Hoof et al., 2004) and also on pervaporation membrane reactors (PVMRs) (Lim et

al., 2002; Lipnizki et al., 1999; Waldburger and Widmer, 1996), since its application to

equilibrium-limited reactions improves conversion by selectively removing one reaction

product e.g. (Benedict et al., 2006; Castanheiro et al., 2006; David et al., 1991; Domingues et

al., 1999; Lauterbach and Kreis, 2006; Peters et al., 2005a; Peters et al., 2005b; Sanz and

Gmehling, 2006; Tanaka et al., 2002). PVMRs are, therefore, a type of membrane reactors

that combines chemical reaction and separation by pervaporation, which is usually

implemented by two different semi-batch processes: (i) the pervaporation unit (PV) is

coupled to the reactor, i.e., the PV unit is an external process unit (see Figure 2.10a); and (ii)

the reactor and the membrane are integrated in the same unit (see Figure 2.10b).

Although there is a recent interest in PVMRs, its discovery goes back to 1960, according to

the first patent for a continuous process that integrates reaction and pervaporation in the same

unit applied to an esterification reaction catalyzed by an acid ion exchange resin and using

water selective membrane in order to remove it from the reaction zone and therefore

enhancing the conversion (Jennings and Binning, 1960). Also, in 1986, another process was

patented for the acetic acid esterification reaction with ethanol (Pearce, 1986), which reports

complete conversion of the acetic acid by using a pervaporation membrane reactor consisting

of two half-cells with a flat membrane disk (commercial PVA or Nafion) placed in the

middle, as shown in Figure 2.11.


Figure 2.10 Layout of a Semi-batch Pervaporation Membrane Reactor (SBPVMR):

(a) external pervaporation unit;

(b) membrane and reactor in the same unit.

Figure 2.11 Experimental set-up for the pervaporation membrane reactor (Pearce, 1986).

Even though PVMRs are finding broad uses, esterifications appear to be a key application

(Marcano and Tsotsis, 2002). The esterification reactions are a typical example of

equilibrium-limited reaction that produces by-product water. Considering a catalytic

esterification reaction scheme of the type:

HA B C D+

⎯⎯→+ +←⎯⎯

where C is the desired ester product and D is the by-product (water). Due to the

thermodynamic equilibrium limitation of the esterification reaction, a conventional reactor

Membrane

(a) (b)


Feed

(A+B) Retentate

(C)

Permeate (D)

Membrane

will operate at low conversion; however, if a membrane is integrated in the reactor, as shown

in Figure 2.12, wherein the water is removed through the permselective membrane from the

reaction zone to the other side of the membrane, the reaction will proceed in the forward

direction and therefore high conversion is expected to be attained in a reasonably short time.

Figure 2.12 Schematic representation of a membrane reactor for by-product withdrawal in a reversible reaction.

Zhu and collaborators studied the esterification reaction of acetic acid with ethanol both by

experiment and simulation in a continuous-flow PVMR using a polymeric/ceramic composite

membrane (Zhu et al., 1996). The same reaction was studied in a continuous tube membrane

reactor (Waldburger and Widmer, 1996). This reaction was also studied in a PVMR, housing

the reactor and membrane in the same unit, applying a zeolite T-membrane since it is stable

under acidic conditions (Tanaka et al., 2001). Nafion tubular membranes, which also act as

catalyst, were applied for the esterification of acetic acid with methanol and n-butanol, where

the equilibrium conversions of 73 % and 70 % were increased to 77 % and 95 %, respectively

(Bagnell et al., 1993). The esterification of acetic acid with butanol was more significantly

improved than with methanol, due to the higher membrane selectivity towards water in

butanol/water system. This particular esterification was also studied using Zr(SO4)·4H2O as

catalyst and using cross-linked polyvinyl alcohol (PVA) membranes (Liu et al., 2001; Liu

and Chen, 2002). Experiments and simulations were conducted to investigate the effects of

several operating parameters, such as reaction temperature, initial molar ratio of acetic acid to

n-butanol, ratio of the membrane area to the reacting mixture volume and catalyst

concentration.

Regarding the operating modes of PVMR, semi-batch esterification process coupled by

pervaporation is the most used (Xuehui and Lefu, 2001), being applied for the synthesis of

ethyl tert-butyl ether (ETBE) from tert-butyl alcohol (TBA) and ethanol (Kiatkittipong et al.,

2002); and applied for the esterification of acetic acid with isopropanol leading to conversions

higher than 90 % (Sanz and Gmehling, 2006). The esterification of lactic acid and succinic


acid with ethanol were studied in semi-batch well-mixed reactors coupled by pervaporation

(Benedict et al., 2006), using two different pervaporation membranes, a GFT-1005 membrane

(Deutsche Carbone AG) and a T-1b (TexSep 1) membrane (Texaco Research), both with

organic acid compatibility and very high water permeation selectivity. The conversion of

lactic acid obtained in the PVMR was 71 %, being the usual equilibrium value of 55 %.

Based on the synthesis of methyl acetate from methanol and acetic acid catalysed by

Amberlyst-15 and using a polyvinyl alcohol (PVA) membrane, three different configurations

of PVMRs were compared (Assabumrungrat et al., 2003): (i) semi-batch (SBPVMR), (ii)

plug-flow (PFPVMR) and (iii) continuous stirred tank (CSPVMR): Simulations, carried out

using the experimental determined kinetic and permeation parameters, conclude that the

PFPVMR is the most favourable mode, although there are some conditions (at low values of

Damkohler number (0.5 and 1)) where CSPVMR is superior to PFPVMR. A new concept of a

hybrid PVMR system, which integrates the pervaporation step through a membrane with

adsorption in the permeate side, proved to enhance in 5 % the conversion reported in the

PVMR in the absence of the adsorbent (Park and Tsotsis, 2004). Based on the concept of

catalytic membranes (Bagnell et al., 1993), the performance of a composite catalytic

membrane was examined for the esterification of acetic acid and butanol aiming to develop a

continuous composite catalytic pervaporation membrane reactor (Peters et al., 2005a). It was

proved by simulations that the outlet conversion for the catalytic pervaporation-assisted

esterification reaction exceeds the conversion of a conventional pervaporation membrane

reactor, with the same loading of catalyst dispersed in the liquid bulk. At the same conditions,

a conversion of 85 % for the catalytic pervaporation-assisted esterification reactor is reported

against 79 % for a conventional pervaporation membrane reactor.

The scientific literature on PVMRs is abundant, since it is an area where significant activity is

under way and many advances are expected in the future, due to the many advantages of

PVMRs: the simultaneous removal of a product from the reactor enhances the conversion;

undesired side reactions can be suppressed; high conversion is possible at almost

stoichiometric feed flow rates and the heat of reaction can be used for separation. Therefore,

lower energy consumption and higher product yields make of the pervaporation membrane

reactor an interesting alternative to conventional processes. However, no large-scale

industrial applications have been reported yet. The main reason for this should be related to

factors as small fluxes of the desired species and mechanical and thermal stability of the


membranes. Further developments in the field of materials engineering will certainly change

this picture.

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3. Batch Reactor: Thermodynamic Equilibrium and

Reaction Kinetics

Abstract. The heterogeneous catalysis of lactic acid (88 wt. %) esterification with ethanol

in presence of Amberlyst 15-wet, was studied for catalyst loading of 1.2 wt. % to 3.9 wt. %,

initial molar ratio of reactants of 1.1 to 2.8 and temperature from 50ºC to 90ºC.

In this work a methodology based in the UNIQUAC model was developed to determine the

thermodynamic equilibrium constant since in literature there is inconsistency concerning the

temperature dependence of the thermodynamic equilibrium constant. A simplified Langmuir-

Hinshelwood kinetic model was used to describe the experimental data. The proposed rate

law is ( ) 2)1(/ WWEthEthWELLaEthc aKaKKaaaakr ++−= ; the kinetic parameters are the

pre-exponential factor, 1170, min..1070.2 −−×= gmolkc and the activation energy,

molkJEa /98.49= . The equilibrium reaction constant is ( ))(/13.515exp35.19 KTK −= with

reaction enthalpy 4.28 kJ/mol. The model reasonably predicts the kinetic experimental data

and it will be very useful to apply for the design and optimization of industrial hybrid reactive

separation processes.

Adapted from: Pereira C. S. M., S. P. Pinho, V. M. T. M. Silva and A. E. Rodrigues, "Thermodynamic

Equilibrium and Reaction Kinetics for the Esterification of Lactic Acid with Ethanol Catalyzed by acid ion

exchange resin", Ind Eng Chem Res. 47: 1453-1463, 2008.

36 CHAPTER 3. Thermodynamic Equilibrium and Reaction Kinetics

3.1 Introduction

The ethyl lactate synthesis comprises a liquid-phase reversible reaction between ethanol and

lactic acid, wherein water is a sub-product:

)()()()( WWaterELLactateEthylLaAcidLacticEthEthanol H +⎯⎯→←++

The lactic acid contains a hydroxyl group adjacent to the carboxylic acid and because of its

bifunctional nature undergoes intermolecular esterification in aqueous solutions above

20 wt. % to form linear dimer and higher oligomer acids (Montgomery, 1952; Vu et al.,

2005):

WLaLa +⇔ 212 (lactic acid dimer formation)

WLaLaLa +⇔+ 321 (lactic acid trimer formation)

…

WLaLaLa nn +⇔+ −11 (lactic acid oligomer formation)

with 2≥n

where:

OH

O

OH

Lactic acid (La1) Lactic acid oligomers (Lan+1)

OH

O

O

O

OHn

The use of lactic acid as reactant is then complicated, since the extent of the self-esterification

increases with the increase of the acid concentration. A method to purify the lactic acid is

through the esterification with lower alcohols, such as methanol, ethanol or butanol; and then

the produced ester is separated and hydrolyzed back into pure lactic acid and the alcohol is

recovered and reused (Troupe and DiMilla, 1957).


For the process of ethyl lactate production it will be desirable to use a high lactic acid

concentration, with lower water content, in order to increase the yield of the reaction and the

concentration of the ethyl lactate. However, this will imply the presence of lactic acid

oligomers during the esterification, which will be converted into the corresponding esters:

WELEthLa +⇔+ 11 (ethyl lactate formation)

WELEthLa +⇔+ 22 (ethyl lactate dimer formation)

WELEthLa +⇔+ 33 (ethyl lactate trimer formation)

…

WELEthLa nn +⇔+ (ethyl lactate oligomer formation)

where:

OH

O

OC2H5

OH

O

O

O

OC2H5n

Ethyl lactate (EL1) Ethyl lactate oligomers (ELn+1)

For an aqueous solution with 88 wt. % of lactic acid, the molar percentage of the monomer

(La1), dimer (La2) and trimer (La3) are of about 43.5 mol %, 9.2 mol % and 1.8 mol %,

respectively; and about 45 mol % of water. While a 20 wt. % aqueous solution of lactic acid

is constituted only by monomer and water, being the monomer molar percentage of about

5.6 mol % (Asthana et al., 2006). However, the percentage of lactic acid and ethyl lactate

oligomers is less then 5% at equilibrium (Asthana et al., 2006; Tanaka et al., 2002) and the

use of an aqueous solution with a high lactic acid concentration is desirable to produce ethyl

lactate at industrial scale by means of a continuous hybrid process as reactive pervaporation,

reactive chromatographic processes, among others.

The kinetics of ethyl lactate production has been studied since 1957 (Troupe and DiMilla,

1957), but lately has deserved more attention since it is a green solvent and an alternative to

the traditionally petroleum derived solvents. Troupe and Dimilla (1957) studied the

esterification reaction between lactic acid and ethanol using sulphuric acid as catalyst.


However, this kind of homogeneous catalysts may be the origin of a lot of problems, because

of their miscibility with the reaction medium, which causes separation problems; in addition,

strong acid catalysts lead to corrosion of the equipment. The replacement of homogeneous

catalysts by heterogeneous catalysts is gaining importance due to their ecofriendly nature.

Besides being non-corrosive and easy to separate from the reaction mixture, the

heterogeneous catalyst can be used repeatedly over a prolonged period without any difficulty

in handling and storage. Many solid-acid catalysts have been used, such as acid treated clays

(Yadav and Krishnan, 1998), heteropolyacids (Dupont et al., 1995; Lacaze-Dufaure and

Mouloungui, 2000; Schwegler et al., 1991; Yadav and Krishnan, 1998), iodine (Ramalinga et

al., 2002), MCM-41 (Koster et al., 2001), zeolite-T membrane (Tanaka et al., 2002), smopex-

101 (Lilja et al., 2002; Mäki-Arvela et al., 1999), HY zeolite (Chen et al., 1989; Kirumakki et

al., 2003; Ma et al., 1996), zeolite beta (Kirumakki et al., 2003) and ZSM-5 (Kirumakki et

al., 2003; Ma et al., 1996; Wu and Chen, 2004; Yadav and Krishnan, 1998). However, ion-

exchange resins are the most commonly used solid catalysts and they have been proved to be

effective in liquid phase esterification (Benedict et al., 2003; Lee et al., 2002; Lee et al.,

2000; Liu and Tan, 2001; Yadav and Mujeebur Rahuman, 2002; Zhang et al., 2004). Since

the heterogeneous catalysis is clearly advantageous, some studies have been already

performed for the esterification of lactic acid with ethanol. Zhang and co-workers studied the

kinetic of esterification of lactic acid (20 wt. %) with ethanol catalyzed by five different

cation-exchange resins (Zhang et al., 2004). They proposed a simplified mechanism based on

Langmuir-Hinshelwood model to describe the kinetic behaviour. Delgado et al. (2007b) also

investigated the esterification of lactic acid with ethanol and the hydrolysis of the ethyl lactate

in the presence of a commercial cation-exchange resin; an aqueous lactic acid solution of

20 wt. % and a mechanism based on the Langmuir-Hinshelwood model to describe the

kinetics was used. This esterification with/without a solid catalyst (Amberlyst XN-1010) was

also investigated by Benedict and their collaborators. A kinetic model based in concentrations

to describe the behaviour of the reaction between an 88 wt. % of lactic acid solution with

ethanol was used (Benedict et al., 2003). The presence of oligomers was not mentioned in

their work. In Tanaka and co-workers studies about this esterification, the oligomers presence

was considered and the reactions were described by simple nth-order reversible rate

expressions based on the species concentration (Tanaka et al., 2002). Three different

solutions of lactic acid (20 wt. %, 50 wt. % and 88 wt. %) were used in the esterification

reaction with ethanol in Asthana’s and collaborators work (2006). The oligomers presence


was also taken into account and a similar model based on the species concentration was used

to describe the reaction kinetics.

In spite of the number of kinetic studies available in the literature, the thermodynamic

equilibrium of the reaction in the liquid phase has not been clearly studied, and some of the

authors do not report the values of the equilibrium constant (Zhang et al., 2004). In some

works (Delgado et al., 2007b; Troupe and DiMilla, 1957) different equilibrium constants

based in concentration (Kx) at the same temperature are reported. The authors of those studies

have concluded that the equilibrium constant Kx vary significantly with the initial molar ratio

of the reactants, being less sensitive to the temperature. However, the thermodynamic

equilibrium constant defined as function of the species liquid activities, which is only

temperature dependent, is not presented in their works. In order to overcome the lack of

thermodynamic data, Delgado and co-authors have studied the vapor-liquid reactive equilibria

for the ethyl lactate synthesis, and they have proposed the following expression to describe

the reaction equilibrium constant (Delgado et al., 2007a):

)(2.2431893.7)(ln

KTK −= (3.1)

Nevertheless, they have found some difficulties in measuring the vapor phase composition,

that affect significantly the experimental values of the experimental equilibrium constant,

leading to high deviations between experimental values and those predicted by Equation 3.1,

as shown in Figure 3.1.

0.00

0.50

1.00

1.50

2.00

2.50

0.00265 0.0027 0.00275 0.0028 0.002851 / T (K)

experimental [26]

ln (K

)

ln K = 7.893 - 2431.2 / T (K)

experimental

Figure 3.1 Representation of experimental values of ln K as function of 1/T. (Data

collected from Delgado et al. (2007b)).


This work was undertaken to obtain the reaction equilibrium and kinetic data for the synthesis

of the ethyl lactate in the liquid phase, avoiding the vaporization of all species by working at

6 bar (helium pressurization). The pressure influence in the value of equilibrium constant is

negligible for this system for the temperature and pressure operating range, as it can be

estimated by the correction factor PK (Smith and Van Ness, 1987). Therefore the

thermodynamic equilibrium constant was estimated by the UNIQUAC method and the

esterification reaction catalyzed by the ion-exchange resin Amberlyst 15-wet was described

by a simple activity based kinetic model, which will be applied in the modelling of some

reactive separation processes, such as membrane reactors. In this work a high lactic acid

concentration was used with the objective of maximizing the ethyl lactate productivity for an

industrial process. The presence of oligomers was neglected, since at equilibrium the total

amount of lactic acid and ethyl lactate oligomers represents less than 5 % according to

Tanaka et al. (2002) and Asthana et al. (2006) studies. However, in the final section of this

paper the presence of oligomers will be addressed.

3.2 Experimental Section

3.2.1 Chemicals and Catalyst

The chemicals used were ethanol (>99.9% in water), lactic acid (>85% in water) and ethyl

lactate (>98% in water) from Sigma-Aldrich (U.K.). A commercial strong-acid ion-exchange

resin named Amberlyst 15-wet (Rohm & Haas) was used as catalyst and adsorbent. This resin

is a bead-form macroreticular polymer of styrene and divinylbenzene, with particle diameter

varying between 0.3 to 1.2 mm, an ion exchange capacity of 4.7 meq H+/g of dry resin and

inner surface area of 53 m2/g. According to Ihm et al. (Ihm et al., 1988) only 4 % of the

active sites are located at the macropores (surface of the microspheres) and the others 96 %

are inside gel polymer microspheres. Since the water adsorbed on the catalyst surface

decreases the reaction kinetics, because it is one of the reaction products, it was necessary to

guarantee anhydrous resin. For that, the resin was washed several times with deionised water

and dried at 90ºC until the mass remains constant. This method is of value for the reuse

experiments, since the water is abundantly present due to the lactic acid solution.


3.2.2 Experimental set-up

The experiments were carried out in a glass-jacketed 1 dm3 autoclave (Büchi, Switzerland),

operating in a batch mode, mechanically stirred at 600 rpm, equipped with pressure and

temperature sensors and with a blow-off valve (Figure 3.2). The temperature was controlled

by thermostated ethylene glycol/water solution (Lauda, Germany) that flows through the

jacket of the reactor and feed vessel. To maintain the reacting mixture in liquid phase over the

whole temperature range, the pressure was set at 0.6MPa with helium. The lactic acid solution

is charged into the reactor and heated to the desired reaction temperature. The dry catalyst is

placed in a basket at the top of the stirrer shaft. Ethanol is heated up to the desired

temperature into the feed vessel and then charged to the reactor opening the on/off valve. The

agitation is immediately turned on and the basket of catalyst falls down in the reactant

solution. This time is considered to be the starting time of the esterification reaction. One of

the outlets of the reactor was connected directly to a liquid sampling valve (Valco, USA),

which injects 0.2 µl of pressurised liquid to a gas chromatograph.

TB

GCHe

He

BR

BV

TT

PTM

NVPM

NV

V2

V1vacuun

vent

vent

FV

Figure 3.2 Experimental set-up for kinetic studies. BR-batch reactor; FV-feed vessel; M motor; TT-temperature sensor; PT-pressure sensor; PM-manometer; BV-blow-off valve; V1-sampling valve; V2-injection valve; NV-needle valve; GC-gas chromatograph; TB-thermostatic bath.


3.2.3 Analytical method

All the samples were analysed in a gas chromatograph (Chrompack 9100, Netherlands) using

a fused silica capillary column (Chrompack CP-Wax 57 CB, 25 m x 0.53 mm ID, df = 2.0

μm) to separate the compounds and a thermal conductivity detector (TCD 903 A) to quantify

them. The column temperature was programmed with a 1.5 min initial hold at 110ºC,

followed by a 50ºC/min ramp up to 190ºC and held for 8.5 min. The injector and detector

temperature were maintained at 280ºC and 300ºC, respectively. Helium N50 was used as the

carrier gas with flowrate 10.50 ml/min. In order to analyze the lactic acid and ethyl lactate

oligomers at equilibrium a HPLC system from Gilson (France) using an ICSep ION-300

column held at 20ºC was used. A 0.0085 N H2SO4 solution was used as mobile phase (0.4 ml

/ min) and species were quantified by a refractive index detector.

3.3 Thermodynamic Equilibrium Results

The experiments to measure the equilibrium constant were done in a temperature range of

323-363 K. At each temperature different experiments were performed using different initial

molar ratios ( 0.1/ =LaEthR to 8.2/ =LaEthR ) and different mass of catalyst (2.3 wt. % to

6.0 wt. %) (see Table 3.1). All the experiments lasted long enough to ensure that the

equilibrium was reached.

3.3.1 Thermodynamic equilibrium constant

In its most general form the chemical equilibrium constant ( K ) for a reaction is given by:

=⎟⎟⎠

⎞⎜⎜⎝

⎛ Δ−=

RTGK exp ∏

ii

iaν (3.2)

where GΔ is the reaction standard free Gibbs energy, R the ideal gas constant, T the

absolute temperature, ia is the activity of species i, and iν its stoichiometric coefficient in the

reaction.

For the esterification reaction, occurring in the liquid phase at low pressure:

γγγγγ

KKxxxx

aaaa

K xLaEth

WEL

LaEth

WEL

LaEth

WEL === (3.3)


being x and γ the mole fraction and the activity coefficient of each species, respectively.

The thermodynamic chemical equilibrium constant is only temperature dependent. From the

experimental point of view, however, changing isothermically the initial mass of each

reactant will give, most probably, different values for the equilibrium constant. Naturally, that

is a consequence of experimental errors as well as deficiencies in the thermodynamic models

used to calculate the activity coefficients. For example, as can be observed in Table 3.1 three

different runs were carried out at 323.15 K. Estimating the activities coefficients by the

UNIFAC method (Fredenslund et al., 1977), using the relative molecular volume, surface

area and the interaction parameters presented in literature (Reid et al., 1987); the resulting

equilibrium constants are quite different; 4.126, 3.813, and 4.637, respectively, from top to

bottom.

Table 3.1 Conditions of the experiments performed to measure the thermodynamic equilibrium constant.

Initial number of moles Equilibrium number of moles

T (K)

Catalyst loading (wt. %) Ethanol

Lactic acid Water Ethanol

Lactic acid

Ethyl Lactate Water

323.15 5.67 0.8154 0.4051 0.2968 0.5459 0.1356 0.2695 0.5662

323.15 5.95 0.7329 0.4051 0.2966 0.4786 0.1508 0.2543 0.5509

323.15 5.63 0.4033 0.4051 0.2957 0.1983 0.2000 0.2051 0.5008

333.15 5.67 0.8154 0.4051 0.2968 0.5403 0.1300 0.2751 0.5719

333.15 5.95 0.7329 0.4051 0.2966 0.4744 0.1465 0.2586 0.5551

333.15 5.63 0.4033 0.4051 0.2957 0.1963 0.1981 0.2070 0.5027

343.15 5.67 0.8154 0.4051 0.2968 0.5351 0.1249 0.2802 0.5770

343.15 5.95 0.7329 0.4051 0.2966 0.4730 0.1452 0.2599 0.5565

343.15 5.63 0.4033 0.4051 0.2957 0.1969 0.1986 0.2065 0.5022

343.41 2.82 5.6988 2.0197 1.4837 4.1506 0.4716 1.5481 3.0319

343.41 2.82 5.6988 2.0197 1.4837 4.1361 0.4570 1.5627 3.0464

353.15 2.37 3.9033 3.4444 2.5154 1.9555 1.4966 1.9477 4.4632

362.87 2.82 5.6988 2.0197 1.4837 4.1173 0.4382 1.5815 3.0652

362.87 2.82 5.6988 2.0197 1.4837 4.1029 0.4239 1.5958 3.0796

One possibility to get a single equilibrium constant value for each temperature would be to fit

all the equilibrium constants using an equation of the type:


TbaK +=ln (3.4)

This, however, presents a big deficiency. In fact, using a unique value for the equilibrium

constant at a given temperature, the equilibrium composition must be recalculated for each

specific initial condition, which immediately will introduce changes in the magnitude of the

activity coefficients.

Assuming a given value for the equilibrium constant at 323.15 K, solving Equation 3.3 in

order to the equilibrium composition involves an iterative procedure. First, all the activity

coefficients must be assumed equal to one, obtaining the equilibrium composition of an ideal

solution. After the activity coefficients can be determined using a model such as UNIQUAC

or UNIFAC, and a new equilibrium composition can now be calculated. The procedure is

repeated until convergence. It must be stressed that this final equilibrium composition will

certainly not be the same as the one presented in Table 3.1, which was used to calculate the

equilibrium constant and to regress the coefficients in Equation 3.4. So this inconsistency

must be avoided.

To overcome this problem, in this work it is suggested to obtain the coefficients in Equation

3.4 that allow the calculation of the equilibrium composition as closer as possible to that

observed experimentally. Therefore, the coefficients were estimated minimizing the following

objective function (Fob):

∑ ⎟⎟⎠

⎞⎜⎜⎝

⎛ −=

k k

calckk

XXX

Fob2

exp

exp

(3.5)

where expkX and calc

kX are the experimental and the calculated equilibrium conversion for

experiment k, respectively.

3.3.1.1 Activity coefficients estimation

In this work, instead of using the UNIFAC method it was preferred to apply the UNIQUAC

model. Indeed, there are some available experimental vapor-liquid equilibrium data involving

mixtures of species involved in the reaction under study, which makes preferable to use a

correlation model instead of a pure predictive method.

Initially the parameters between ethanol and ethyl lactate were estimated based on the data

published by Peña-Tejedor et al. ( 2005) and Vu et al. (2006). Following, the parameters


between water and ethyl lactate were obtained based on the data by Vu et al. ( 2006). Finally

the parameters between lactic acid and all other species were estimated using data from the

quaternary system measured by Delgado et al. (2007a). The parameters were estimated

minimizing the following objective function (Fob):

∑∑ ⎟⎟⎠

⎞⎜⎜⎝

⎛ −=

j i ji

calciiFob

2

exp

exp

γ

γγ (3.6)

where i is the species and j the experimental data point. It should be mentioned that the

parameters between water and ethanol were found in the DECHEMA books (Gmehling et al.,

1981). The interaction parameters are given in Table 3.2.

Table 3.2 UNIQUAC interaction parameters (K).

Ethanol Ethyl lactate Water Lactic acid

Ethanol 0.0000 -152.319 -17.554 + 0.2797 T (K)* -35.008

Ethyl Lactate 264.990 0.0000 207.789 -20.986

Water -21.987 + 0.2276 T (K)* -13.093 0.0000 -99.183

Lactic Acid 33.741 219.89 213.19 0.0000

*DECHEMA (Gmehling et al., 1981)

The average relative deviation found for the activity coefficients were 12.9 %, but special

difficulties were found when describing the behaviour of diluted solutions, which was never

the case when the chemical equilibrium experimental studies were carried out in this work.

3.3.2 Equilibrium constant and reaction enthalpy for the synthesis of Ethyl Lactate

Using the data found experimentally for the chemical equilibrium compositions in 14

different runs, at 4 different temperatures in the range between 323.15 and 362.87 K, it was

found the following relation for the equilibrium constant:

515.13ln( ) 2.9625( )

KT K

= − (3.7)

Using this result and the iterative procedure described before, the deviations found between

the experimental and calculated equilibrium compositions for all experiments are given in

Table 3.3. It can be seen that the higher deviation found between the experimental and

calculated equilibrium composition is of about 5.6 %.


Table 3.3 Experimental and calculated equilibrium compositions for all the experiments performed and the correspondent deviation percent.

Equilibrium composition (mole fraction)

Experimental Calculated T (K)

Ethanol Ethyl lactate Water Lactic

acid Ethanol Ethyl lactate Water Lactic

acid

average deviation

(%)

323.15 0.3598 0.1776 0.3732 0.0894 0.3561 0.1813 0.3769 0.0857 2.06

323.15 0.3336 0.1773 0.384 0.1051 0.3264 0.1845 0.3913 0.0978 3.77

323.15 0.1812 0.1854 0.4527 0.1808 0.1931 0.1734 0.4407 0.1928 5.58

333.15 0.3561 0.1813 0.3769 0.0857 0.3528 0.1846 0.3802 0.0824 1.87

333.15 0.3307 0.1802 0.3869 0.1021 0.3228 0.1881 0.3948 0.0943 4.11

333.15 0.1794 0.1871 0.4544 0.179 0.1881 0.1784 0.4457 0.1878 4.08

343.15 0.3527 0.1847 0.3803 0.0823 0.3498 0.1876 0.3832 0.0794 1.67

343.15 0.3297 0.1812 0.3879 0.1012 0.3194 0.1915 0.3982 0.0909 5.41

343.15 0.1799 0.1866 0.4539 0.1796 0.1835 0.1831 0.4504 0.183 1.64

343.41 0.451 0.1682 0.3295 0.0512 0.451 0.1683 0.3295 0.0512 0.01

343.41 0.4495 0.1698 0.3311 0.0497 0.451 0.1683 0.3295 0.0512 1.18

353.40 0.1983 0.1975 0.4525 0.1517 0.202 0.1937 0.4488 0.1555 1.78

362.87 0.4474 0.1719 0.3331 0.0476 0.4474 0.1718 0.3331 0.0477 0.07

362.87 0.4459 0.1734 0.3347 0.0461 0.4474 0.1718 0.3331 0.0477 1.30

In Table 3.4 the activity coefficients for the equilibrium composition and the thermodynamic

equilibrium constant are presented. In terms of equilibrium conversion the average relative

deviation found was 2.60%, being possible to find an average reaction enthalpy of

4.28 kJ·mol-1 in that temperature interval. Some authors (Zhang et al., 2004) have also

indicated the reaction is slightly endothermic, but uncertainty is high. In fact using the values

found for the standard state enthalpy of formation in the DIPPR 801 database (DIPPR, 1998)

(see Table 3.5) it is possible to calculate -20.97 ± 186.7 kJ·mol-1 for the reaction enthalpy at

298.15 K. This considerable high error (± 186.7 kJ·mol-1) is mainly due to the high

uncertainty on the ethyl lactate standard state enthalpy of formation. Nevertheless the value

found is in good agreement to values given in the literature (Benedict et al., 2003; Zhang et

al., 2004).


Table 3.4 Activity coefficients for the equilibrium composition and the thermodynamic equilibrium constant.

Activity coefficients T (K)

Ethanol Lactic acid

Ethyl lactate Water

K

323.15 1.0563 1.2391 1.3557 1.6928 3.9291

323.15 1.0563 1.2632 1.3941 1.6629 3.9291

323.15 1.0263 1.3238 1.7011 1.5288 3.9291

333.15 1.0652 1.2397 1.3386 1.6847 4.1217

333.15 1.0660 1.2643 1.3749 1.6547 4.1217

333.15 1.0398 1.3285 1.6644 1.5200 4.1217

343.15 1.0737 1.2396 1.3222 1.6771 4.3116

343.15 1.0755 1.2649 1.3564 1.6467 4.3116

343.15 1.0531 1.3325 1.6300 1.5117 4.3116

343.41 1.0607 1.1543 1.2338 1.7832 4.3165

343.41 1.0607 1.1543 1.2338 1.7832 4.3165

353.40 1.0749 1.3350 1.5295 1.5267 4.5035

362.87 1.0707 1.1511 1.2125 1.7715 4.6781

362.87 1.0707 1.1511 1.2125 1.7715 4.6781

Table 3.5 Standard state enthalpy of formation of the different species (Gmehling et al., 1981).

Ethanol Lactic acid Ethyl lactate Water

fH 0Δ (kJ·mol-1) -234.950 -682.960 -695.084 -241.814

Error < 1 % < 10 % < 25 % < 0.2 %

3.3.3 Application of this methodology to other works

The equilibrium compositions and equilibrium conversions of other studies (Benedict et al.,

2003; Delgado et al., 2007b; Troupe and DiMilla, 1957) were predicted by the UNIQUAC

model and the equilibrium constant proposed in this work. It should be noticed that those

equilibrium compositions were not used in the estimation of the UNIQUAC parameters nor in

the estimation of chemical equilibrium constant presented by Equation 3.7. The comparison


between the experimental and calculated equilibrium compositions is shown in Figure 3.3.a

for ethanol and water; and Figure 3.3.b for lactic acid and ethyl lactate.

0.0

0.2

0.4

0.6

0.8

1.0

0.0 0.2 0.4 0.6 0.8 1.0xexp

xcal

c

Ethanol

Water

Molar Fraction

Figure 3.3.a Experimental and calculated equilibrium molar fraction

of ethanol and water species.

0.00

0.05

0.10

0.15

0.20

0.00 0.05 0.10 0.15 0.20

xexp

xcal

c

Ethyl LactateLactic Acid

Molar Fraction

Figure 3.3.b Experimental and calculated equilibrium molar fraction

of ethyl lactate and lactic acid species.

The equilibrium compositions are very well predicted for ethanol and water; while for the

ethyl lactate and lactic acid deviations are higher, but even so satisfactory. The larger

deviations can be due to the fact that the interaction parameters used in UNIQUAC model for

those particular species are not based in a large and consistent set of experimental vapour


liquid equilibrium (VLE) data as is the case for water and ethanol species. Furthermore the

composition range is now much restricted (maximum mole fraction around 0.20), for which

the model presents more difficulties to calculate accurately the activity coefficients and the

analytical method for these species is not as precise as it is for water and ethanol compounds.

In terms of the deviation of experimental and calculated equilibrium conversion (Figure 3.4),

it can be seen that when 85 wt. % lactic acid feed is used, the model prediction is quite good.

In the remaining cases, for 44 wt. % lactic acid and mainly for 20 wt. % lactic acid solution

the deviations are more significant. This is maybe due to the fact that diluted solutions in

lactic acid and, obviously, in ethyl lactate are used, and like mentioned before the UNIQUAC

model does not describe with the desired accuracy the behaviour of that kind of solutions. If

possible it would be preferable to use activity coefficients at infinite dilution.

0.0

0.2

0.4

0.6

0.8

1.0

0.0 0.2 0.4 0.6 0.8 1.0Xexp

Xcal

c

La 20 wt.% (Delgado et al., 2007a) La 44 wt.% (Troupe and Dimilla, 1957) La 85 wt.% (Troupe and Dimilla, 2003) La 85 wt.% (Benedict et al., 2003) La 85 wt.% (This work)

Equilibrium conversion

Figure 3.4 Experimental and calculated equilibrium conversion. (Data collected

from Troupe and Dimilla (1957), Benedict et al. (2003), Delgado et al. (2007) and this work).

3.4 Kinetic Studies

The experimental results of the reaction kinetics of the esterification of lactic acid and ethanol

catalyzed by the Amberlyst 15-wet resin are presented in this section. The effect of various

conditions, such as, catalyst loading, initial molar ratio between ethanol and lactic acid and

reaction temperature on lactic acid conversion as function of time is studied. This study was


performed varying the condition under evaluation and keeping constant the remaining

conditions, in absence of mass transfer limitations and catalyst deactivation as shown by the

preliminary studies performed.

3.4.1 Preliminary Studies

3.4.1.1 Evaluation of external mass transfer limitations (effect of stirring speed)

To quantify the influence of external mass transfer resistance preliminary experiments at

different stirring speed were run using a molar ratio of ethanol to lactic acid of 1.82 at

324.18 K and 2.4 wt. % of Amberlyst 15-wet with particle diameter 0.5 < dp < 0.6 mm.

Figure 3.5 shows the conversion of lactic acid as a function of time at different stirring speed.

With a stirring speed of 600 rpm, there is no limitation due to external resistance, so all

further experiments were done at 600 rpm.

0

0.1

0.2

0.3

0.4

0.5

0.6

0 200 400 600 800

Conv

ersi

on

Time (min)

600 rpm

800 rpm

Figure 3.5 Experimental data obtained at different stirrer speeds for a

molar ratio of ethanol to lactic acid of 1.82 at 324.18 K using 2.4 wt. % of Amberlyst 15-wet as catalyst with particle diameter 0.5 < dp < 0.6 mm.

3.4.1.2 Evaluation of internal mass transfer limitations (effect of particle size)

The Amberlyst 15wet was separated by particle size and three classes with different diameters

were obtained: 0.425mm<dp<0.5mm; 0.5mm<dp<0.6mm and 0.6mm<dp<0.85mm. Figure

3.6 shows the effect of the different particle diameters, including the unsieved resin, which


has an average diameter of 0.685 mm, on the conversion of lactic acid; it can be seen that

there are no significant internal diffusion limitations for the experiments performed.

Therefore, the unsieved resin was used for the following kinetic experiments performed in

this work. This is in agreement with several works performed with this kind of resin and type

of reaction (Delgado et al., 2007b; Liu and Tan, 2001; Pöpken et al., 2000; Zhang et al.,

2004).

0

0.1

0.2

0.3

0.4

0.5

0.6

0 200 400 600 800

Conv

ersi

on

Time (min)

unsieved resin

0.6<dp<0.85

0.5<dp<0.6

0.425<dp<0.5

Figure 3.6 Effect of catalyst particle size on the conversion of lactic acid history

for a molar ratio of ethanol to lactic acid of 1.82 at 324.18 K using 2.4 wt. % of catalyst and stirrer speed of 600 rpm.

3.4.1.3 Evaluation of catalyst deactivation (effect of catalyst reusability)

The Amberlyst 15-wet catalyst reusability was studied at 344.05 K. The resin was reused up

to three times. First, a reaction was carried out using fresh catalyst. Then the catalyst used

was separated from the reaction mixture by filtration, washed several times with deionised

water and dried at 90ºC until the mass remained constant and charged again to the reactor. A

new esterification reaction was performed using the same conditions that the first one and so

on. The lactic acid conversion as a function of time was analyzed and the three conversion

histories are presented in Figure 3.7. As it can be seen no significant changes were observed.

The resin activity was kept the same in the three runs. However, each experiment was

performed using fresh Amberlyst 15-wet catalyst. Some authors (Dixit and Yadav, 1996)

studied the reusability of the Amberlyst 15-wet in the alkylation reaction of o-xylene with

styrene and they observed a drastic reduction in the conversion of styrene due to the direct

deposition of the by-products on the active sites and the loss of accessibility of the active sites


due to pore blockage. So, it can be concluded that the catalytic activity of the resins depends

on its interaction with the reaction medium.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0 200 400 600 800 1000 1200

Con

vers

ion

Time (min)

fresh catalyst2nd use3rd use

Figure 3.7 Effect of the Amberlyst 15-wet reusability on the conversion of lactic acid

history at 344.05 K for a molar ratio of ethanol to lactic acid of 1.82 using 2.4 wt. % of Amberlyst 15-wet catalyst with an average particle diameter of 0.685 mm and stirrer speed of 600 rpm.

3.4.2 Kinetic Model

The esterification reactions have been described using different models, such as, pseudo-

homogeneous, Eley-Rideal and Langmuir-Hinshelwood model (L-H model) (Lee et al., 2002;

Lilja et al., 2005; Pöpken et al., 2000; Sanz et al., 2002; Teo and Saha, 2004). However, the

L-H model has been considered the most appropriate model for the reaction between lactic

acid and ethanol (Delgado et al., 2007b; Zhang et al., 2004). Therefore, the model developed

is based on an L-H mechanism and it considers the following steps:

Ethanol and Lactic acid adsorption:

SEthSEth EthsK .,⎯⎯ →←+

SLaSLa LasK .,⎯⎯ →←+

Surface reaction between the adsorbed species of ethanol and lactic acid:

SWSELSLaSEthK

....1

++ ⇔

desorption of ethyl lactate and water:

SELSEL ELsK +⎯⎯ →←*,.


SWSW WsK +⎯⎯ →←*,.

Surface reaction is assumed to be the controlling step, with the other steps remaining in

equilibrium. Multi-component Langmuir adsorption isotherms, written in terms of activities,

are assumed to describe the adsorption behaviour of the compounds of the reaction mixture in

the surface of the resin. Taking into account the above considerations the following rate

expression, written in terms of activities of components, due to the non-ideality of the

reaction mixture, is obtained:

2

,1 ⎟⎠

⎞⎜⎝

⎛+

−=

∑=

W

Ethiiis

WELLaEth

c

aK

Kaaaa

kr (3.8)

where kc is the kinetic constant, isK , is the adsorption constant for species i and K is the

equilibrium reaction constant.

In order to reduce the number of optimization parameters it was taken into account only those

components that had the strongest adsorption. It was considered that the most polar

molecules, water and ethanol, have the strongest adsorption strength on the Amberlyst 15-wet

surface, being therefore neglected the adsorption of lactic acid and ethyl lactate. This

consideration is corroborated by several works of adsorption in similar resins (Mazzotti et al.,

1997; Sanz et al., 2002; Silva and Rodrigues, 2002; Zhang et al., 2004). Thus, the simplified

rate expression used to describe the experimental data is:

2,, )1( WWsEthEths

WELLaEth

c aKaKKaa

aakr

++

−= (3.9)

In this kinetic model (Equation 3.9) there are three parameters to be estimated, at each

temperature, the kinetic constant (kc) and the two adsorption parameters (Ks,Eth and Ks,W),

instead of five if the rate Equation 3.8 was used to describe the experimental data.

The temperature dependence of kinetic constant was fitted with the Arrhenius equation:

⎟⎠

⎞⎜⎝

⎛−=RTE

kk aCC exp,0 (3.10)

where aE is the reaction activation energy, ck ,0 is the pre-exponential constant, R is the gas


constant and T is the temperature.

The temperature dependence of the adsorption equilibrium constants were fitted with:

⎟⎠

⎞⎜⎝

⎛ Δ−=

RTH

KK sss exp,0 (3.11)

where sK ,0 is the constant for Equation 3.11 and SHΔ is the adsorption enthalpy.

3.4.2.1 Parameter estimation from experimental data

The mass balance in the batch reactor for a component i, in liquid phase, at constant

temperature is given by:

rwvtdnd

catii = (3.12)

where in is the number of moles of component i, t is the time, catw is the mass of catalyst

and r is the reaction rate expressed in moles i / (mass of catalyst .min).

The number of moles of component i ( in ) as a function of the conversion X of the limiting

reactant ( l ) is:

⎟⎟⎠

⎞⎜⎜⎝

⎛+=

lilili

XRnnν

ν/0, (3.13)

where 0,ln and lν are, respectively, the initial moles number and the stoichiometric

coefficient of the limiting reactant and /i lR is given by:

0,

0,/

l

ili n

nR = (3.14)

where 0,in is the initial number of moles of component i .

Introducing Equation 3.13 into Equation 3.12 we get:

0,l

catl

nrw

tdXd ν= (3.15)

with the initial condition: 00; Xt == .


The differential Equation 3.15 combined with the suggested rate expression, Equation 3.9,

was solved numerically with the DASOLV integrator implemented in gPROMS-general

PROcess Modelling System version: 3.0.3, which is a commercial package from Process

Systems Enterprise. For all simulations a tolerance equal to 10-5 was fixed. The estimation of

the unknown parameters ( aE , ck ,0 , EthK ,0 , WK ,0 , EthHΔ and WHΔ ) was carried out using the

“Parameter estimation in gProms” that attempts to determine the values for the parameters, in

order to maximize the probability that the mathematical model will predict the values

obtained from experiments. Assuming independent, normally distributed measurements

errors, ikε , with zero means and standard deviations, ikσ , this maximum likelihood goal can

be captured through the following objective function:

⎪⎭

⎪⎬⎫

⎪⎩

⎪⎨⎧

⎥⎦

⎤⎢⎣

⎡ −++=Φ ∑∑

= =

NE

i

NM

k ik

ikikik

i XXN1 1

2

'2 )ln(

21)2ln(

2 min σσπ

θ (3.16)

Where, N is the total number of measurements taken during the experiments ( 290=N ), θ is

the set of model parameters to be estimated ( 6=θ ), NE is the number of experiments

performed ( 15=NE ), iNM is the number of measurements of the conversion in the ith

experiment, 2ikσ is the variance of the kth measurement of the conversion in experiment i

( 321083.1 −×=σ ) , '

ikX is the kth measured value of conversion in experiment i and,

finally, ikX is the kth (model-) predicted value of conversion in experiment i. The quality of

the model fit was tested through the mean relative deviation (MRD) between the calculated

conversion values (Xcalc) and the experimental ones (Xexp) (see Equation 3.17).

%100*1

exp

exp

⎟⎟

⎠

⎞

⎜⎜

⎝

⎛ −= ∑

N

calc

XXX

NMRD (3.17)

3.4.3 Modelling and discussion of results

The simplified L-H model (Equation 3.9) was used to describe the kinetic behaviour of the

esterification of lactic acid with ethanol using Amberlyst 15-wet as catalyst. In order to obtain

the unknown parameters at least two different experiments (different initial molar ratio) were

performed for each temperature, which varied from 323.15 K to 363.15 K. The values of the


optimized parameters along with the MRD value are presented in Table 3.6. A value of MRD

between experimental and calculated conversion of 6.8 % was obtained.

As mentioned before in section 3.4.2 the Arrhenius equation was used to fit the kinetic (kc)

constant (Figure 3.8) and the Equation 3.11 was used to fit the adsorption constants (Ks,Eth and

Ks,W) which are given by ( ))(/01.12exp19.15 KTKW = and ( ))(/63.359exp22.1 KTK Eth = . From

the temperature dependence of the kinetic constant the apparent activation energy was

calculated being the correspondent value 49.98 kJ/mol.

Table 3.6 Parameters of the kinetic model expressed in terms of activities (Equation 3.9) and the mean relative deviation, MRD.

k0,c (mol.g-1.min-1)

Ea (kJ.mol-1) K0, W WHΔ

(J.mol-1) K0, Eth EthHΔ

(J.mol-1) MRD (%)

2.70×107 49.98 15.19 - 99.85 1.22 - 29.95×102 6.84

-2

-1.5

-1

-0.5

0

0.5

1

2.7 2.8 2.9 3 3.1 3.2

ln kc

1000/T (K-1) Figure 3.8 Representation of experimental values of ckln as

function of T/1 and linear fitting.

3.4.3.1 Effect of catalyst loading

The catalyst loading was varied from 1.2 wt. % to 3.9 wt. % (weight of catalyst/total weight

of reaction mixture) keeping the rest of the experimental conditions similar. The effect of the

catalyst loading on the esterification reaction between lactic acid and water is shown in

Figure 3.9. As may be observed the reaction rate increases with the percentage of Amberlyst

15-wet. This was expected since the increase of catalyst implies an increase in the number of

active sites available for the reaction.


0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0 50 100 150 200 250 300

Conv

ersi

on

Time (min)

1.2 wt.%2.4 wt.%3.9 wt.%L-H model

Figure 3.9 Effect of catalyst loading on the conversion of lactic acid history for a

molar ratio of ethanol to lactic acid of 1.82 at 353.49 K and using an average particle diameter of 0.685 mm and stirrer speed of 600 rpm.

3.4.3.2 Effect of initial molar ratio of reactants

To study the effect of the initial molar ratio of reactants (REth/La) on the conversion of lactic

acid, this condition was varied from 1.1 to 2.8, as presented in Figure 3.10. It can be seen that

the equilibrium conversion increases with the increase of the initial molar ratio of ethanol to

lactic acid and that the equilibrium is achieved faster for larger initial molar ratio values.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

0 500 1000 1500 2000 2500

Conv

ersi

on

Time (min)

REth/La = 1.1

REth/La = 1.8

REth/La = 2.8

L-H model

Figure 3.10 Effect of initial molar ratio of ethanol to lactic acid on the conversion

of lactic acid history at 353.40 K and using 2.4 wt. % of Amberlyst 15-wet catalyst with an average particle diameter of 0.685 mm and stirrer speed of 600 rpm.


It may be observed from Figure 3.10 that the experiments are initially well predicted by the

model (until about 120 minutes); and due to the methodology developed in this work, the

equilibrium is also well described. However, there is a transient state (between 120 minutes

till the equilibrium), where the model fails to describe the experiments, being more significant

for higher values of initial molar ratio between ethanol and lactic acid.

3.4.3.3 Effect of reaction temperature

The effect of reaction temperature is shown in Figure 3.11. The reaction rate increases with

the reaction temperature. The same effect is noticed on the equilibrium conversion. Once

again, it can be noticed that in the transient state, the model predicts higher conversion of

lactic acid values than those obtained experimentally.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0 500 1000 1500 2000 2500 3000

Conv

ersi

on

Time (min)

T = 324.14 KT = 333.11 KT = 344.05 KT = 353.40 KT = 363.50 KL-H model

Figure 3.11 Effect of the reaction temperature on the conversion of lactic acid

history for a molar ratio of ethanol to lactic acid of 1.82 and using 2.4 wt. % of Amberlyst 15-wet catalyst.

3.4.3.4 Effect of Lactic acid and Ethyl Lactate oligomers

The proposed model shows a kinetic behaviour with two limiting situations:

a) first slope at the beginning of the experiment, corresponding to the initial reaction

rate;

b) final plateau, corresponding to the reaction equilibrium.


However, experimentally, it seems that there are three limiting situations, where the first and

third are the same that the ones described by the proposed model, but the transient state is

represented by a line with very small slope. This could be due to the presence of the lactic

acid oligomers, and consequently ethyl lactate oligomers that where formed. To confirm this

assumption an evaluative modelling study was made considering the oligomers presence

using the kinetic model and the correspondent’s parameters reported in literature (Tanaka et

al., 2002).

In Figure 3.12a the molar fractions of the ethyl lactate, lactic acid, ethanol and water as

function of time are presented, being the correspondent’s oligomers presented in Figure

3.12b. Analysing the simulated kinetic curve for the ethyl lactate monomer, here denominated

just as ethyl lactate, one can conclude a similar behaviour with the experimental results

shown in this work, where there are three limiting steps:

a) initial conversion of lactic acid oligomers in ethyl lactate oligomers, but the kinetic

rate is much higher for the monomers than for the dimers and trimers. Therefore the

experiments are initially well predicted;

b) transient state, where the ethyl lactate monomer concentration increases slowly, since

several reversible reactions are occurring and the equilibrium is shifted towards lactic

acid monomer formation and consequently ethyl lactate monomer. In this transient

step, firstly, the dimers concentrations of lactic acid and ethyl lactate decrease due to

their hydrolysis, and then the same happens to the trimers that are converted into

dimers and finally again into monomers, till the equilibrium is reached;

c) reaction equilibrium, when all the reversible reactions are in equilibrium, the total

amount of lactic acid and ethyl lactate oligomers are less then 0.4%. Therefore, the

experiments are well predicted at the equilibrium since oligomers presence is

negligible.


0 10 20 30 40 50 480 490 5000.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

Mol

ar fr

actio

n

Time (min)

Ethyl lactate Ethanol Lactic acid Water

Figure 3.12a Molar fraction histories of the compounds: ethyl lactate monomer, ethanol,

lactic acid monomer and water. KT 15.363= and 3/ =LaEthR . The kinetic model and the parameters used were taken from Tanaka et al. (2002).

0 10 20 30 40 50 480 490 5000.00

0.01

0.02

0.03

0.04

0.05

0.06

Mol

ar fr

actio

n

Time (min)

Ethyl lactate dimer Ethyl lactate trimer Lactic acid dimer Lactic acid trimer

Figure 3.12b Molar fraction histories of the oligomers: ethyl lactate dimer, ethyl lactate trimer, lactic acid dimer and lactic acid trimer.

KT 15.363= and 3/ =LaEthR . The kinetic model and the parameters used were taken from Tanaka et al. (2002).


The transient behaviour, where the hydrolysis of dimers and trimers of lactic acid and ethyl

lactate are dominant to produce the ethyl lactate monomer, is even more noticeable for higher

values of initial molar ratio of ethanol/lactic acid, as it can be shown in Figure 3.13a. This is

due to the fact that the excess of ethanol benefits more the ethyl lactate dimer formation as

shown in Figure 3.13b, that will be further hydrolysed and the lactic acid dimer formed will

be converted into lactic acid monomer that will be finally converted into ethyl lactate

monomer. Nevertheless, in the equilibrium composition the excess of ethanol leads to smaller

amounts of oligomers (2.4 molar % in the case of 1/ =laEthR and 0.4 molar % in case

of 3/ =laEthR ).

0 10 20 30 40 50 480 490 5000.00

0.05

0.10

0.15

0.20

0.25

Mol

ar fr

actio

n

Time (min)

Ethyl lactate, REth/La= 1 Lactic acid, REth/La= 1 Ethyl lactate, REth/La= 3 Lactic acid, REth/La= 3

Figure 3.13a Molar fraction histories of ethyl lactate and lactic acid for different

initial molar ratios. KT 15.353= . The kinetic model and the parameters used were taken from Tanaka et al. (2002).


0 10 20 30 40 50 480 490 5000.000

0.002

0.004

0.006

0.008

0.010

0.012

0.014

0.016

0.018M

olar

frac

tion

Time (min)

Ethyl lactate dimer, REth/La= 1 Ethyl lactate dimer, REth/La= 3

Figure 3.13b Molar fraction histories of ethyl lactate dimer for different initial

molar ratios. KT 15.353= . The kinetic model and the parameters used were taken from Tanaka et al. (2002).

3.4.3.5 Effect of polar species

The activity of the resin varies with the polarity of the reaction medium, since it influences

the number of available sulfonic groups and their acidity (Fite et al., 1998). The polarity of

the medium, mainly due to the water and alcohol concentrations, can affect the reaction rate

in two ways:

i) the more adsorbed species (water and alcohol) inhibit the others to adsorb onto the

active sites.

ii) water can ionize, solvate and dissociate the acidic protons of the sulfonic groups,

depending on their concentration; when the sulfonic sites are completely dissociated

the reaction occurs in the liquid phase as in the case of homogeneous catalysis

(Gomez et al., 2004).

In this case, the reaction medium has a high water concentration; besides the water initially

present in the reaction due to the lactic acid solution, more water is being formed during the

course of the esterification of lactic acid with ethanol. Although the kinetic model proposed in

this work has taken into account the inhibitory effect caused by adsorption, it doesn’t consider

the remaining effects that could be due to water or ethanol presence. Therefore, this could


also be a reason for the deviations between the experimental results and the predicted by the

kinetic model in the transient stage. Françoisse and Thyrion studied the ETBE synthesis

catalyzed by Amberlyst 15 (Françoisse and Thyrion, 1991). In that study the influence of

ethanol in the reaction rate was taken into account. A kinetic model to describe the behaviour

of the reaction for low and high ethanol concentrations was developed. However, this case is

different from the one presented in this work, since the most adsorbed component is ethanol

and there is no water as reaction product. They use a non-polar solvent (n-pentane) and the

other reactant (isobutene) is also non-polar. In the case under study, besides ethanol there is

also lactic acid concentrated solution that has high water content. Therefore, at the beginning

of the reaction there are three polar species competing to the acid sites; it was observed that

the initial kinetic rate increases with the ethanol concentration till a plateau, which was not

the behaviour observed in the ETBE kinetics. Most of the works about the esterification

reaction of ethanol and lactic acid considers that water and ethanol are the most adsorbed

species that supports our kinetic model assumptions (Delgado et al., 2007b; Zhang et al.,

2004). Moreover, in a later work for the esterification of lactic acid with butanol catalyzed by

Amberlyst 15, the authors didn’t observe different mechanisms for high and low alcohol

concentration, and similarly with our model, they neglect the oligomers presence and only

considered the water adsorption (Dassy et al., 1994).

The proposed model describes quite well the experimental data up to 80 % of the equilibrium

conversion as well as the equilibrium stage, which are the most important to apply to a hybrid

reactive separation technology.

3.5 Conclusions

The equilibrium composition for the liquid phase reaction of ethyl lactate synthesis catalyzed

by the acid ion exchange resin Amberlyst 15-wet was measured in the temperature range of

50-90ºC, at 6 bars. The thermodynamic equilibrium constant estimated by the UNIQUAC

method was ( )KTK 13.5159625.2ln −= and the average reaction enthalpy was of

4.28 kJ·mol-1 in that temperature interval. This relation was also successfully applied to

describe the equilibrium compositions of other published studies; better prediction was found

for systems where high concentrations of lactic acid was used.


Because of the strong non-ideality of the liquid reaction mixture, the reaction rate model was

formulated in terms of activities. The rate-controlling step for the esterification reaction

between lactic acid and ethanol, heterogeneously catalyzed by the Amberlyst 15-wet, was the

surface reaction, since external and internal mass resistances were insignificant for the

temperature range of 50-90ºC. The catalyst reusability was also studied and it was not

verified catalyst deactivation till 3 usages of the same resin sample.

A three-parameter model based on a Langmuir-Hinshelwood rate expression was proposed to

describe the experimental kinetic results:

( ) 2)1(/ WWEthEthWELLaEthc aKaKKaaaakr ++−= ; and the model parameters are

( )7 1 12.70 10 exp 6011.55 / ( ) ( . .min )ck T K mol g − −= × − , ( ))(/01.12exp19.15 KTKW =

and ( ))(/63.359exp22.1 KTKEth = . The agreement between experimental and simulated

results was good for the following operating conditions: catalyst loading from 1.2 wt. % to

3.9 wt. %, initial molar ratio of reactants from 1.1 to 2.8 and temperature from 50ºC to 90ºC.

3.6 Notation

a liquid phase activity

pd pellet diameter (m)

df film thickness (µm)

aE apparent reaction activation energy (J mol-1)

K equilibrium reaction constant

xK equilibrium constant based on molar fractions

γK equilibrium constant based on activity coefficients

sK equilibrium adsorption constant

ck kinetic constant (mol g-1 min-1)

ck ,0 pre-exponential factor (mol g-1 min-1)

sK ,0 constant for Equation 3.11


MRD mean residual deviation

n number of moles (mol)

N total number of measurements taken during the experiments

NE number of experiments performed

iNM number of measurements of the conversion in experiment i

sHΔ enthalpy of adsorption (J mol-1)

fH 0Δ standard enthalpy of formation (J mol-1)

ºGΔ standard Gibs energy of reaction (J mol-1)

R gas constant (J mol-1 K-1)

r reaction rate (mol g-1 min-1)

LaEthR / initial molar ratio of ethanol to lactic acid

t time coordinate (min)

T temperature (K)

X conversion of the limiting reactant

x molar fraction

catw mass of dry catalyst (g)

ikX kth (model-) predicted value of conversion in experiment i

'ikX kth measured value of conversion in experiment i

Greek letters

ν stoichiometric coefficient

γ activity coefficient

θ set of model parameters to be estimated

2ikσ variance of the kth measurement of conversion in experiment i

2σ average of variance


Subscripts

0 initial value

Eth ethanol

La lactic acid

EL ethyl lactate

W water

exp experimental

calc calculated

i relative to component i

l relative to limiting reactant

3.7 References Cited

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Pöpken T., L. Götze and J. Gmehling, "Reaction kinetics and chemical equilibrium of homogeneously and heterogeneously catalyzed acetic acid esterification with methanol and methyl acetate hydrolysis", Ind. Eng. Chem. Res. 39(7): 2601-2611, 2000.

Ramalinga K., P. Vijayalakshmi and T. N. B. Kaimal, "A mild and efficient method for esterification and transesterification catalyzed by iodine", Tetrahedron Lett. 43(5): 879-882, 2002.

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Vu D. T., C. T. Lira, N. S. Asthana, A. K. Kolah and D. J. Miller, "Vapor - Liquid equilibria in the systems ethyl lactate + ethanol and ethyl lactate + water", J. Chem. Eng. Data 51(4): 1220-1225, 2006.

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4. Fixed Bed Adsorptive Reactor

Abstract. Multi-component adsorption equilibrium data were measured through binary

adsorption experiments performed in a fixed bed column packed with Amberlyst 15-wet for

the ethyl lactate system, at 20ºC and 50ºC. A novel approach based on multi-component

Langmuir isotherm was used assuming a constant monolayer capacity in terms of volume for

all species, reducing the adjustable parameters from 8 to 5, for each temperature. Reactive

adsorption experiments were performed and used to validate a mathematical model developed

for both fixed bed and simulated moving bed reactors, which involves velocity variations due

to the change of multi-component mixture properties.

Adapted from: Pereira C. S. M., V. M. T. M. Silva and A. E. Rodrigues, "Fixed Bed Adsorptive Reactor for

Ethyl Lactate Synthesis: Experiments, Modelling, and Simulation", Sep. Sci. Technol. 44(12): 2721 - 2749,

2009.

72 CHAPTER 4. Fixed Bed Adsorptive Reactor: Experiments, Modelling and Simulation

4.1 Introduction

The conventional way to produce ethyl lactate is by the esterification of lactic acid with

ethanol in the presence of an acid catalyst. The conversion of this kind of reactions is limited

by the chemical equilibrium; therefore, the technology for the industrial preparation of esters

involves two consecutive steps. The first is the reaction itself, which stops when the

equilibrium is reached. The second is the separation of the products from the equilibrium

mixture containing products and unconverted reactants. The disadvantage of this technology

is in its economics, because of the high energy costs and investment in several reaction and

separation units. Multifunctional reactors, where reaction and separation take place into a

single unit, allow, in addition to obvious savings in equipment costs, significant

improvements in process performance for reactions limited by chemical equilibrium. By

removing one of the products from the reaction zone, the equilibrium limitation can be

overcome and the conversion can be driven to completion.

In the esterification reaction between lactic acid and ethanol the catalyst used is usually

concentrated sulphuric acid (Zhang et al., 2004). However, its application has several

drawbacks (as separation problems and corrosion of equipment) and, thus, the use of

heterogeneous catalysts is preferable. Among this type of catalysts, strongly acid resins, like

Amberlyst 15-wet (A15), are of great interest in the case of reversible reactions, as

esterifications and acetalizations, since they can act as catalyst and also as selective adsorbent

(Funk et al., 1995; Gandi et al., 2006; Kawase et al., 1996; Mazzotti et al., 1996; Ruggieri et

al., 2003; Silva and Rodrigues, 2002). Accordingly, the synthesis of ethyl lactate in a

chromatographic reactor using the A15 resin is very attractive. The simulated moving bed

reactor (SMBR) is one of the most interesting chromatographic reactors and has been applied

to several reversible reactions catalyzed by acidic resins (Borges da Silva et al., 2006;

Kawase et al., 1996; Lode et al., 2003; Mazzotti et al., 1996; Minceva et al., 2008; Pereira et

al., 2008; Silva and Rodrigues, 2005; Yu et al., 2003) improving the reaction conversion,

given that the products are formed and simultaneously separated and removed from the

reaction medium in two different streams (extract and raffinate). In order to provide a better

understanding of the performance of the SMBR and to validate kinetic and adsorption data it

is appropriate to evaluate first the dynamic behaviour of the fixed bed adsorptive reactor

(Gandi et al., 2006; Kawase et al., 1999; Lode et al., 2001; Mazzotti et al., 1997; Silva and

Rodrigues, 2002). The present work addresses the detailed experimental study of the ethyl

lactate synthesis in a fixed bed adsorptive reactor in order to validate the mathematical model


to be later used to define operating conditions of the SMBR. Binary adsorption experiments

were carried out on a fixed bed packed with A15 resin, in absence of reaction, at 293.15 and

323.15 K to determine the multi-component adsorption parameters.


4.2.1 Chemicals and Catalyst / Adsorbent

The chemicals used in the experiments were ethanol (>99.9% in water), lactic acid (>85% in

water) and ethyl lactate (>98% in water) from Sigma-Aldrich (U.K.)

The column was packed with Amberlyst 15-wet (A15), which is a highly cross-linked

polystyrene-divinylbenzene ion exchange resin functionalized with sulfonic groups (SO3H),

that acts as catalyst and adsorbent in this system. The properties of the A15 resin are

presented in Table 4.1. A15 is a macroreticular-type resin, but according to Ihm et al. (Ihm et

al., 1988) only 4 % of the active sites are located at the macropores (surface of the

microspheres) and the others 96 % are inside gel polymer microspheres.

Table 4.1 Physical and chemical properties of resin A15.

Properties A15

Manufacturer Rohm and Haas

Concentration of acid sites (eq H+/kg) 4.7

Surface area (m2/g) 53

Average pore diameter (nm) 30

Particle diameter (mm) 0.3-1.2

This kind of resins swell selectively in the presence of a liquid phase constituted of a

multi-component mixture. The swelling is due to the sorption of the different components of

the mixture, depending on their relative affinities to the resin. The polymeric resins that

contain sulfonic acid functional groups, like the A15, exhibit a strong selectivity for polar

species. Some authors (Mazzotti et al., 1997) studied the sorption equilibrium of acetic acid,


ThermostaticBath

Reservoir

Resin

Refrigeratingwater

Three-wayvalve

Sampling line

Feed Stream

ThermostaticBath

ReservoirReservoir

Resin

Refrigeratingwater

Three-wayvalve

Sampling line

Feed Stream

ethanol, water and ethyl acetate on A15. They verified that the components can be listed by

the following decreasing order of affinity to the resin: water, ethanol, acetic acid and ethyl

acetate and that the swelling ratio is, respectively, 1.52, 1.48, 1.30 and 1.22. This conclusion

agrees with the polarity of the components. So, it is expected for the system hereby presented

that the order of decreasing affinity of the components to the resin will be: water, ethanol,

lactic acid and ethyl lactate. The swelling effect might change the length and the bulk

porosity of the fixed bed reactor; however, in the system under study, just a insignificant

variation in the bed length was noticed; and, therefore, a constant bed length was considered.

4.2.2 Experimental Apparatus

The experiments were performed in a laboratory-scale jacketed glass column that was

maintained at constant temperature through a thermostatic bath (293.15 or 323.15 K), at

atmospheric pressure (see Figure 4.1). The experimental breakthrough curves were obtained

by analysing with a gas chromatograph, small samples withdrawn at different times from the

column exit.

Figure 4.1 Experimental set-up (configuration: top-down flow direction).


4.2.2.1 Bed Porosity and Peclet Number

The bed porosity and the Peclet number were determined by pulse experiments of tracer using

a blue dextran solution (5 kg/m3), which is a polymer whose molecule is large enough

(M.W. = 2,000,000) to diffuse only in the bulk fluid phase between resin particles. Samples

of the blue dextran solution (0.2 cm3) were injected under different flow rates and the column

response was monitored using a UV-VIS detector at 300 nm. The bed porosity was calculated

from the stoichiometric time of the obtained experimental curves. An average Peclet number

was obtained for the range of flow rates to be used in the fixed bed experiments by

calculating the second moment of the experimental curves. Figure 4.2 shows the different

tracer experiments performed and the respective values of bed porosity and Peclet number are

presented in Table 4.2. The characteristics of the fixed bed column are summarized in Table

4.3.

0.00

0.20

0.40

0.60

0.80

1.00

0 4 8 12 16 20

C/C 0

Time (min)

run 1run 2run 3

Figure 4.2 Tracer experiments using a blue dextran solution. Points

are experimental values and lines are simulated curves.

Table 4.2 Results obtained from tracer experiments.

Q (mL/min) τ (min) ε σ 2 Pe

run 1 1.985 11.23 0.347 3.528 71.48

run 2 4.981 4.69 0.364 0.556 79.24

run 3 8.026 2.92 0.365 0.229 74.54


Table 4.3 Characteristics of the fixed bed column.

Solid weight (A15) 25 g

Length of the bed (L) 12 cm

Internal diameter (Di) 2.6 cm

Average radius of resin beads (rp) 372.5 μm

Bulk density (ρ b) 390 kg/m3

Bed porosity (ε) 0.36

Resin particle porosity (εp) 0.36 (Lode et al., 2001)

Peclet number 75

All the samples removed from the column exit were analysed in a gas chromatograph; the

analytical method is described in Chapter 3.

4.3 Modelling of Fixed Bed

Several mathematical models have been developed to explain the kinetic behaviour of the

fixed adsorptive reactor and to predict the breakthrough curves. In order to interpret correctly

the behaviour of a fixed bed adsorptive reactor, it is necessary to characterize on one hand the

partitioning equilibrium on the solid sorbent and on the other hand the reaction kinetics on the

solid catalyst. A rigorous modelling of the sorption in the swelling polymer should include an

appropriate model to predict the polymer-phase activities. Mazzoti and co-workers used the

extended Flory–Huggins model to determine the chemical activities of the species in the

polymer phase for the esterification of acetic acid to ethyl acetate in the presence of A15 ion

exchange resin catalyst (Mazzotti et al., 1996). They assumed, in their kinetic model, that the

activities of the species in the bulk liquid phase (estimated by UNIFAC method) were in

equilibrium with the ones in the resin polymer phase (estimated by the extended Flory–

Huggins model). However, in latter papers (Lode et al., 2001; Pöpken et al., 2000), it is

mentioned that the extended Flory–Huggins model is not well suited for high cross-linked

resins bearing highly polar groups on almost all monomer and should be regarded as an

empirical tool to correlate the equilibrium data and to predict the behaviour of


multi-component mixtures. Yu and collaborators state that using the extended Flory–Huggins

model based in adsorption equilibrium data is not viable for most adsorption systems, given

that non-reactive binary mixtures are scarce in the literature and the model involves

complexity and inconvenience in computation (Yu et al., 2004). Thus, an alternative

approach based on the multi-component Langmuir model was considered in this work, since

it is able to represent satisfactorily the experimental adsorption data and is simpler than the

Flory-Huggins model. Nevertheless, it has to be mentioned that one of the assumptions of the

Langmuir model considers that the monolayer adsorption has energetically equal binding

sites, which do not describe the actual physical adsorption phenomena. Therefore, some

authors (Pöpken et al., 2000) proposed a multi-component Langmuir adsorption isotherm

which assumes a constant monolayer capacity in terms of mass for all species, after having

measured mono-component adsorption data expressed in terms of volumes, masses and moles

of different species per gram of A15. In this work we will consider a constant volumetric

monolayer capacity for all species, which will reduce the adjustable adsorption parameters

from 8 (one molar monolayer capacity and one equilibrium constant for each component) to 5

(one volumetric monolayer capacity for all components and one equilibrium constant for each

component).

A detailed model was used to describe the dynamic behaviour of the fixed bed adsorptive

reactor taken into account the following assumptions:

- Isothermal operation;

- The flow pattern is described by axial dispersed plug flow model;

- External and internal mass transfer for absorbable species is combined in a global

resistance;

- Constant column length and packing porosity;

- Extended Langmuir Isotherm model for multi-component adsorption;

- Velocity variations due to changes in bulk composition.

The model equations are constituted by the following system of four second order partial

differential equations in the bulk concentration of the ith component ( iC ), four ordinary

differential equations in the average concentration of the ith component into the particle pores


( ,p iC ) and four algebraic equations in the adsorbed concentration in equilibrium with

,p iC ( iq ):

Bulk fluid mass balance to component i:

( ) ( ), ,(1 ) 3ii i

L i i p i ax Tp

u CC xK C C D Ct z r z z

εε

∂ ⎛ ⎞∂ ∂− ∂+ + − = ⎜ ⎟∂ ∂ ∂ ∂⎝ ⎠

(4.1)

where iLK , is the global mass transfer coefficient of the component i, ε is the bed porosity, t

is the time variable, z is the axial coordinate, axD and u are the axial dispersion coefficient

and the interstitial velocity respectively, pr is the particle radius and xi is the component

molar fraction in liquid phase.

The axial dispersion coefficient axD was calculated from the Peclet number:

ax

u LPeD

= (4.2)

The interstitial fluid velocity variation is calculated using the total mass balance:

( ) ( ),, ,1

1 3 NC

p iL i mol i iip

du K V C Cdz r

εε =

−= − −∑ (4.3)

where Vmol,i is the molar volume of component i .

Pellet mass balance to component i:

( ) ( )ipbii

pip

pipiiLp

Crt

qt

CCCK

r ,,

,, 1)1(3

ερν

εε−

−∂

∂−+

∂

∂=− (4.4)

Where iν is the stoichiometric coefficient of component i , bρ is the bulk density, pε is the

particle porosity, iq is the average adsorbed phase concentration of species i in equilibrium

with ipC , , and r is the kinetic rate of the chemical reaction relative to the average particle

concentrations in the fluid phase given by (see Chapter 3):

2

,1

1

Eth La EL W eqc NC

s i ii

a a a a Kr k

K a=

−=

⎛ ⎞+⎜ ⎟

⎝ ⎠∑

(4.5)


where kc is the kinetic constant, isK , is the adsorption constant for species i and Keq is the

equilibrium reaction constant, a is the species activity (calculated by UNIQUAC model) and

the subscripts WELLaEth and , , refer to ethanol, lactic acid, ethyl lactate and water,

respectively. Kinetic and thermodynamic parameters, essential for the Equation 4.5, were

taken from Chapter 3 and are given in Table 4.4.

Table 4.4 Kinetic and thermodynamic parameters.

Langmuir Adsorption equilibrium isotherm to component i:

,

,1

1

p ii ii NC

p jjj

Q K CqK C

=

=+∑

(4.6)

where ,/i V mol iQ Q V= , QV is the volumetric monolayer capacity, ,mol iV is the molar volume of

species i and iK is the equilibrium constant for component i.

Initial and Danckwerts boundary conditions:

0,,0 iipi CCCt === (4.7)

,0

0 ii ax T i F

z

xz u C D C u Cz

=

∂= − =

∂ (4.8a)

0u u= (4.8b)

0=∂

∂=

=Lz

i

zC

Lz (4.9)

where subscripts F and 0 refer to the feed and initial states, respectively.

The proposed model considers a global mass transfer coefficient ( LK ) defined, for each

component, as:

ipeL kkK ε111

+= (4.10)

Temperature (K)

kc (mol/(g.min)) Ks,w Ks,Eth Keq

293.15 0.035 15.82 4.16 3.34

323.15 0.225 15.77 3.71 3.93


wherein ek and ik are, respectively, the external and internal mass transfer coefficients.

Santacesaria and co-workers (Santacesaria et al., 1982) showed that the internal mass

transfer coefficient varies in time, and the calculation of the rigorous values requires the

solution of the complete model equations inside particles. As an approximation, the mean

value estimated by the Equation 4.11 (Glueckauf, 1955) was used:

5 /mi

p

Dkr

τ= (4.11)

The external mass transfer coefficient was estimated by the Wilson and Geankoplis

correlation (Ruthven, 1984):

( ) 550015.009.1 33.0 <<= ppp ReScReShε

(4.12)

where pSh and pRe are, respectively, the Sherwood and Reynolds numbers, relative to

particle:

m

pep D

dkSh = (4.13)

pp

d uRe

ρη

= (4.14)

and Sc is the Schmidt number:

m

ScD

ηρ

= (4.15)

The infinite dilution diffusivities were estimated by the Scheibel correlation (Scheibel, 1954)

which modified the Wilke-Chang equation in order to eliminate the association factor:

2/38

,0 2, 1/3

, ,

38.2 10( / ) 1 mol BA B

B mol A mol A

VTD cm sV Vη

− ⎡ ⎤⎛ ⎞× ⎢ ⎥= + ⎜ ⎟⎜ ⎟⎢ ⎥⎝ ⎠⎣ ⎦ (4.16)

where 0,A BD is the diffusion coefficient for a dilute solute A into a solvent B, T is the

temperature, ,mol iV is the molar volume of the component i, Bη is the viscosity of solvent B.

Table 4.5 shows the diffusion coefficients for all binary systems:


Table 4.5 Diffusion coefficients for the binary mixtures estimated using the Scheibel correlation (equation 4.16).

Di,jº (cm2/s) Ethanol Lactic acid Ethyl lactate Water

Ethanol 0 1.6796E-06 2.583E-05 2.3655E-05

Lactic acid 2.5056E-05 0 2.1087E-05 2.0236E-05

Ethyl lactate 1.8034E-05 9.8526E-07 0 1.5444E-05

Water 8.2195E-05 4.6733E-06 7.4567E-05 0

For binary systems, Vignes equation (Vignes, 1966) was used to predict ,A BD in

concentrated solutions from the infinite dilute coefficients as a simple function of

composition:

2 10 02,1 1,2 1,2 2,1( ) ( )x xD D D D= = (4.17)

For concentrated multi-component systems there are several mixing rules (Umesi and

Danner, 1981) to predict the molecular diffusivity coefficient of a solute in a mixture. In this

work it was used the Perkins and Geankoplis method (Perkins and Geankoplis, 1969):

0.8 0 0.8, ,

1#

n

A m m i A i iii A

D x Dη η=

= ∑ (4.18)

where iη is the viscosity of pure component i and mη is the viscosity of the mixture. The

viscosities of the mixtures play an important role in the estimation of the diffusion coefficient.

The methods used to predict binary and multi-component mixtures viscosities are detailed in

the next sub-section.

In this model, the mixture viscosity and the components diffusivities in the liquid mixture

were calculated at each time and at every axial position; therefore, the mass transfer

parameters also depend on the liquid composition (estimated at each time and axial position

from the equations presented).

4.3.1 Multi-component viscosity

One of the most used and recommended liquid mixture viscosity correlation is the one by

Grunberg–Nissan (Grunberg and Nissan, 1949), which for binary mixtures is:


1 1 2 2 1 2 1,2ln( ) ln( ) ln( )m x x x x Gη η η= + + (4.19)

where ix is the mole fraction of component i and 1,2G is an empirical interaction parameter

adjusted by experimental data.

For aqueous lactic acid solutions, the interaction parameters were found to be

, (20º ) 5.240LA WG C = and , (50º ) 4.369LA WG C = using viscosities determined experimentally

presented in literature (Troupe et al., 1951). As can be seen in Figure 4.3, the Equation 4.19

describes well the viscosity in a large range, from diluted solutions till concentrated lactic

acid (85% mass fraction).

0

5

10

15

20

25

30

35

40

0.0 0.2 0.4 0.6

Visc

osity

(cP)

Mole fraction of lactic acid

Exp. 20 ºC (Troupe et al, 1951)Exp. 50 ºC (Troupe et al,1951)Theoretical (Eq. 4.19)

Figure 4.3 Viscosities of aqueous lactic acid solutions.

For the system water ethanol, experimental data at 20 ºC (Gonzalez et al., 2007) and 50 ºC

(Motin et al., 2005) are not well fitted by Equation 4.19 ( , (20º ) 2.568Eth WG C = and

, (50º ) 1.799Eth WG C = ), as it can be seen in Figure 4.4. It was stated by several authors, that

the Grunberg–Nissan correlation is not suitable for several aqueous system (Li and Carr,

1997; Macias-Salinas et al., 2003). Normally, the correlation used to describe the viscosity of

those systems involves at least 3 adjustable parameters (Gonzalez et al., 2007; Motin et al.,

2005). We propose a new correlation with just one adjustable parameter similar to that of

Grunberg–Nissan, but the excess viscosity function is averaged by the volume fractions

instead of molar ones:

1 1 2 2 1 2 1,2ln( ) ln( ) ln( )m x x Gη η η φ φ= + + (4.20)


where ix is the mole fraction of component i , iφ is the volume fraction of component i and

1,2G is an empirical interaction parameter adjusted by experimental data. Using the new

interaction parameters in Equation 4.20 ( , (20º ) 3.833Eth WG C = and , (50º ) 2.698Eth WG C = ), the

ethanol/water system viscosities are very well predicted, as it can be seen in Figure 4.5.

0.0

0.5

1.0

1.5

2.0

2.5

3.0

3.5

0.0 0.2 0.4 0.6 0.8 1.0

Visc

osity

(cP

)

Mole fraction of ethanol

Exp. 20 ºC (Gonzalez et al, 2007)Exp. 50 ºC (Motin et al, 2005)Theoretical (Eq. 4.19)

Figure 4.4 Viscosities of ethanol/water mixtures. Solid line calculated by Equation 4.19.

0.0

0.5

1.0

1.5

2.0

2.5

3.0

3.5

0.0 0.2 0.4 0.6 0.8 1.0

Visc

osity

(cP)

Mole fraction of ethanol

Exp. 20 ºC (Gonzalez et al, 2007)

Exp. 50 ºC (Motin et al, 2005)

Theoretical (Eq. 4.20)

Figure 4.5 Viscosities of ethanol/water mixtures. Solid line calculated by Equation 4.20.

Since there are no experimental data available for the systems ethanol/ethyl lactate and

ethanol/lactic acid, these systems are considered as ideal, i.e., , 0Eth ELG = and , 0Eth LAG = .


The liquid viscosity of the multi-component system ethanol, lactic acid, ethyl lactate and

water is calculated by the following model:

, ,ln( ) ln( ) ln( ) ln( ) ln( )m Eth Eth LA LA EL EL W W LA W LA W Eth W Eth Wx x x x x x G Gη η η η η φ φ= + + + + + (4.21)

The viscosities of pure components at 20 and 50ºC are presented in Table 4.6 (see appendix B

for details in calculation).

Table 4.6 Viscosities (cP) of ethanol, lactic acid, ethyl lactate and water at 20ºC and 50ºC.

Temperature Ethanol Lactic acid Ethyl lactate Water

20 ºC 1.1617 53.6564 2.5986 1.0254

50 ºC 0.6885 13.9919 1.1233 0.5530

Numerical Solution

The model equations were solved numerically using the gPROMS-general PROcess

Modelling System version: 3.0.3 (www.psentreprise.com). gPROMS is a software package

for modelling and simulation of processes with both discrete and continuous as well as

lumped and distributed characteristics. The mathematical model involves a system of partial

and algebraic equations (PDAEs). Third order orthogonal collocation over twenty one finite

elements was used in the discretization of axial domain. The system of ordinary differential

and algebraic equation (ODAEs) was integrated over time using the DASOLV integrator

implementation in gPROMS. For all simulations was fixed a tolerance equal to 10-7.

4.4 Results and Discussion

4.4.1 Adsorption Isotherm

As the resin A15 acts simultaneously as adsorbent and as catalyst, in order to have

information on the adsorptive equilibrium alone it is necessary to perform experiments with

non reactive binary mixtures. So, the breakthrough curves of ethanol, lactic acid, ethyl lactate

and water were measured in the absence of reaction. The resin was saturated with a certain

component A and then the feed concentration of component B was changed stepwise. The


experimental data were used to calculate the number of moles adsorbed/desorbed of each

component for all the experiments. The adsorption parameters were optimized by minimizing

the difference between experimental and theoretical values, according to Equation 4.26:

[ ]∫∞

−=0exp )( dttCCQn outF

ads (4.22)

[ ]∫∞

−=0exp )( dtCtCQn Fout

des (4.23)

[ ]( ) [ ]( )VCqCqCCn FpFpadstheo )()()1()1()1( 00 −−−+−−+= εεεεε (4.24)

[ ]( ) [ ]( )VCqCqCCn FpFpdestheo )()()1()1()1( 00 −−−+−−+= εεεεε (4.25)

( ) ( )2 2

exp exp1

NEads ads des des

theo theok

fob n n n n=

⎡ ⎤= − + −⎢ ⎥⎣ ⎦∑ (4.26)

4.4.1.1 Binary Adsorption experiments

As mentioned above, for the determination of the adsorption parameters over the A15 resin it

was necessary to perform experiments with non reactive pairs of species. The possible binary

mixtures to run the breakthrough experiments in absence of reaction are ethanol / water, ethyl

lactate / ethanol and lactic acid / water. In all the experiments, the correct liquid flow

direction (bottom-up or top-down) was considered in order to obtain reproducibility in the

experimental results, since the hydrodynamic regime has an important effect on them and the

difference in densities of the species can lead to axial backmixing driven by natural

convection. The densities of ethanol, lactic acid, ethyl lactate and water at 293.15 K are

0.79 g/cm3, 1.21 g/cm3, 1.04 g/cm3 and 1 g/cm3, respectively. The concentration fronts moving

within the column are hydrodynamically stable if the component above the front is less dense

than the component below the front (Silva and Rodrigues, 2002). The multi-component

adsorption equilibrium was measured at two different temperatures, 293.15 K and 323.15 K.

In Table 4.7 the optimized adsorption parameters and the molar volumes at the two

temperatures are presented.


Table 4.7 Adsorption parameters over A15 resin.

Component QV (ml/lwet solid) 20 ºC / 50 ºC

Q (mol/lwet solid) 20 ºC / 50 ºC

K (l/mol) 20 ºC / 50 ºC

Vmol (ml/mol) 20 ºC / 50 ºC

Ethanol 6.70 / 6.30 5.443 / 3.068 58.17 / 60.87

Lactic acid 5.22 / 4.94 4.524 / 4.085 74.64 / 77.56

Ethyl lactate 3.42 / 3.23 1.117 / 1.815 113.99 / 118.44

Water

390.0/383.5

21.57 / 20.58 15.353 / 7.055 18.08 / 18.63

In order to compare the selectivity of the resin to the components, two different experiments

were performed. One, where ethanol is fed to the column initially saturated with water, and

other, where ethanol is fed to the column initially saturated with ethyl lactate (see Figure 4.6).

In the first case (Figure 4.6a) the concentration front of ethanol has a dispersive character and

in the second one (Figure 4.6b) the ethanol concentration front is self-sharpening. This is due

to the fact that water is the preferentially adsorbed component and ethyl lactate is the weakest

adsorbed component.

02468

1012141618

0 10 20 30 40

Out

let c

once

ntra

tion

(mol

L)

Time (min)

Ethanol

Theoretical

02468

1012141618

0 10 20 30 40

Out

let c

once

ntra

tion

(mol

/L)

Time (min)

Ethanol

Theoretical

Figure 4.6 Breakthrough experiments: outlet concentration of ethanol as a function of

time; Q = 5 mL/min; T = 293.15 K; top-down direction flow; ( a) ethanol displacing water; (b) ethanol displacing ethyl lactate.

The breakthrough curves of the four components at 323.15 K are presented next in this

section, while the remaining breakthrough curves at 293.15 K used for the determination of

the adsorption parameters are shown in appendix D.

(a) (b)


0

10

20

30

40

50

60

0 50 100 150 200 250 300 350

Out

let c

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(mol

/L)

Time (min)

WaterEthanolTheoretical

0

10

20

30

40

50

60

0 20 40 60 80 100 120 140 160

Out

let c

once

ntra

tion

(mol

/L)

Time (min)

WaterEthanolTheoretical

Figure 4.7 Breakthrough experiments: outlet concentration of ethanol and water as a

function of time; Q = 5 mL/min; T = 323.15 K; (a) water displacing ethanol; Bottom up flow direction; (b) ethanol displacing water; Top-down flow direction.

The experimental and theoretical concentration history for the binary mixture ethanol / water

is shown in Figure 4.7. Comparing the experimental amount of water adsorbed in the

experiment of Figure 4.7a with the experimental amount of water desorbed in the experiment

of Figure 4.7b, the error obtained is 0.3 %. For ethanol, the error between the amount

adsorbed, Figure 4.7b, and the amount eluted, Figure 4.7a, is 0.31 %. The deviations found

between the experimental amount adsorbed of water and the one predicted by the model is of

4 %, while in the case of ethanol the difference between experimental amount adsorbed and

predicted is of 0.6 %.

(a)

(b)


0

2

4

6

8

10

12

14

16

18

0 40 80 120 160 200

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(mol

/L)

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Ethyl lactateEthanolTheoretical

0

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4

6

8

10

12

14

16

18

0 20 40 60 80 100

Out

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(mol

/L)

Time (min)

Ethyl lactate

Ethanol

Theoretical

Figure 4.8 Breakthrough experiments: outlet concentration of ethanol and ethyl lactate

as a function of time; Q = 5 mL/min; T = 323.15 K; (a) ethyl lactate displacing ethanol; Bottom up flow direction (b) ethanol displacing ethyl lactate; Top-down flow direction.

The experimental and simulated results for the non reactive binary mixture ethanol / ethyl

lactate are shown in Figure 4.8. The difference obtained between the experimental amount of

ethyl lactate eluted in the experiment of Figure 4.8b, and the experimental amount of ethyl

lactate adsorbed in the experiment of Figure 4.8a, was 0.21 %. In the case of ethanol, the

error obtained between the experimental amount adsorbed, Figure 4.8b, and the experimental

amount eluted, Figure 4.8a, was 4.22 %.

For the non reactive pair lactic acid / water (Figure 4.9) the deviation found between the

experimental amount of lactic acid eluted (Figure 4.9b) and adsorbed (Figure 4.9a) was

0.45 % and between the experimental amount of water adsorbed (Figure 4.9b) and eluted

(Figure 4.9a) was -1.07 %.

(a)

(b)


0

10

20

30

40

50

60

0 40 80 120 160

Out

let c

once

ntra

tion

(mo/

L)

Time (min)

WaterLactic acidTheoretical

0

10

20

30

40

50

60

0 10 20 30 40 50 60

Out

let c

once

ntra

tion

(mol

/L)

Time (min)

Water

Lactic acid

Theoretical

Figure 4.9 Breakthrough experiments: outlet concentration of water and lactic acid as

a function of time; Q = 5 mL/min; T = 323.15 K; (a) lactic acid displacing water; Bottom up flow direction (b) water displacing lactic acid; Top-down flow direction.

All the binary adsorption experiments performed are very well described by the model,

except for the case of the experiment of Figure 4.9b, where pure water displaces an 86 wt %

lactic acid solution. This could be explained by the water/lactic acid high selectivity in A15

(about 7.2); however, this behaviour was not observed in the binary experiments

ethanol/water where the selectivity is very similar (about 7.5). Therefore, that behaviour

might be due to viscous fingering phenomenon, since the highest viscous fluid (85% lactic

acid solution: 9.40 cP at 323.15 K), is being displaced by the lowest viscous fluid (water:

0.55 cP at 323.15 K), in the top-down direction, inducing an unstable interface between them.

(b)

(a)


The difference in their viscosities is of about 9 cP and in this case fingering effects are

significant (Catchpoole et al., 2006). Viscous fingering is not accounted in the global mass

transfer coefficient (Equation 4.10). In order to verify this assumption, the effect of viscous

fingering can be described increasing the axial dispersion (Mallmann et al., 1998). Indeed, the

theoretical curve using an axial dispersion coefficient increased by a factor of 7 (thin line in

Figure 4.9b) fits the experimental data very well.

From the data reported it can be concluded that the most adsorbed component in A15 is water

followed by ethanol and lactic acid, being the less adsorbed the ethyl lactate, as expected due

to the polarity of the species.

4.4.2 Kinetic experiments

4.4.2.1 Fixed Bed Reactor

A reaction experiment was performed in the chromatographic reactor packed with A15

initially saturated with ethanol, at 20ºC. As shown in Figure 4.10, the lactic acid conversion is

very low due to both mass transfer and reaction kinetics limitations. Therefore, the fixed bed

reactor was operated at 50ºC in order to reduce mass transfer resistance and to increase

reaction kinetics.

0

2

4

6

8

10

12

14

16

18

0 10 20 30 40 50 60

Out

let c

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tion

(mol

/L)

Time (min)

EthanolWaterEthyl lactateLactic acidTheoretical

Figure 4.10 Concentration histories at the outlet of the fixed bed adsorptive reactor;

column initially saturated with ethanol and then fed with a mixture of ethanol and lactic acid solution: 4.6 / minQ mL= , 293.15T K= ,

, 6.76 /La FC mol L= , , 6.76 /Eth FC mol L= and , 5.45 /W FC mol L= .


A mixture of ethanol and lactic acid was fed to the reactor saturated with ethanol and the

concentration history of ethanol, lactic acid, ethyl lactate and water at the end of the column

is shown in Figure 4.11a.

0

2

4

6

8

10

12

14

16

18

0 20 40 60 80 100 120 140

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let c

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(mol

/L)

Time (min)


0

2

4

6

8

10

12

14

16

18

0 10 20 30 40 50 60 70

Out

let c

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tion

(mol

/L)

Time (min)


Figure 4.11 Concentration histories at the outlet of the fixed bed adsorptive reactor for

production steps. a) Experiment 1: Column initially saturated with ethanol and then fed with a mixture of ethanol and lactic acid solution.

1.0 / minQ mL= , 323.15T K= , , 5.88 /La FC mol L= , , 6.60 /Eth FC mol L= and , 5.87 /W FC mol L= . b) Experiment 2: Column initially saturated with ethanol and then fed with lactic acid solution (86 wt % in water). 1.3 / minQ mL= , 323.15T K= .

(a)

(b)


As the lactic acid solution enters the column it is adsorbed, reacts with ethanol to produce

ethyl lactate and water in the stoichiometric amount. Ethyl lactate is first eluted, since it has

less affinity with the resin than water. This process continues until the equilibrium is reached;

the resin is completely saturated with water and lactic acid. At this moment the selective

separation of ethyl lactate and water is no longer possible. The local compositions remain

constant and the steady state is achieved being the outlet stream constituted by a reactive

mixture at the equilibrium composition. Analysing the outlet concentration curves of ethyl

lactate and water, shown in Figure 4.11a, a difference between them is noticed although they

are formed at the same stoichiometric amount. This is related to the difference in the adsorbed

amount of these species in the resin and, mainly, due to the fact that in the feed there is

already some water content due to the lactic acid solution.

A second experiment was performed by only feeding the lactic acid solution (86 wt % in

water) to the fixed bed adsorptive reactor (see Figure 4.11b), in order to validate the model

under very different conditions. Similarly to the previous experiment, as the lactic acid and

water enter the column, they are adsorbed and start displacing the initially adsorbed ethanol.

Simultaneously, the lactic acid reacts with ethanol in the resin phase until complete depletion

of adsorbed ethanol and the formed ethyl lactate is adsorbed, forming a dispersive front with

ethanol once this is more adsorbed than ethyl lactate. The adsorption of lactic acid solution

continues till the resin saturation, displacing the ethyl lactate. Finally, lactic acid and water

exit the column, being water the last eluted component due to high affinity of A15 towards

this specie.

After the steady state is achieved and to perform a new reaction experiment it is necessary to

first regenerate the column. This step requires the use of a solvent to displace the adsorbed

species. The solvent used was ethanol, since lactic acid solution could not be used, because it

contains water and then the column wouldn’t be completely regenerated. In Figure 4.12, the

regeneration steps for each experiment are presented. As expected, ethyl lactate and lactic

acid are rapidly desorbed; however, a large amount of ethanol is required in order to

completely desorb all water.

Al the experimental results obtained for both production and regeneration steps, at 293.15 and

323.15 K, are well described by the proposed mathematical model.


0

2

4

6

8

10

12

14

16

18

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tion

(mol

/L)

Time (min)


0

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4

6

8

10

12

14

16

18

0 10 20 30 40

Out

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tion

(mol

/L)

Time (min)


Figure 4.12 Concentration histories at the outlet of the fixed bed adsorptive reactor

for regeneration steps with ethanol ( 4.3 / minQ mL= , 323.15T K= ). a) Experiment 1. b) Experiment 2.

(a)

(b)


4.5 Conclusions

A detailed study of the ethyl lactate synthesis in a fixed bed adsorptive reactor was carried

out. Adsorption experiments in absence of reaction at 20ºC and 50ºC were performed. For

each temperature, the adsorption parameters were obtained: a single volumetric monolayer

capacity for all components and one equilibrium constant for each component. A

mathematical model for the fixed bed adsorptive reactor was developed which includes: axial

dispersion, external and internal mass transfer resistances, multi-component Langmuir

adsorption isotherm, the reaction rate measured in a previous work and velocity variations

due to mixture molar volume variations. This model proved to be efficient in the prediction of

the reaction and regeneration steps performed in the fixed bed reactor and it will be very

useful for the study of the SMBR for the synthesis of ethyl lactate.

4.6 Notation

a liquid phase activity

C liquid phase concentration (mol/m3)

TC total liquid phase concentration (mol/m3)

pC average liquid phase concentration inside the particle (mol/m3)

df film thickness (µm)

pd particle diameter (m)

0,A BD diffusion coefficient for a dilute solute A into a solvent B (m2/s)

,A BD mutual diffusion coefficient for binary concentrated solutions (m2/s)

axD axial dispersion (m2/s)

mD molecular diffusivity (m2/s)

1,2G interaction parameter in Equation 4.19

K Langmuir equilibrium parameter (m3/mol)


ik internal mass transfer coefficient (m/s)

ek external mass transfer coefficient (m/s)

ck kinetic constant (mol/kgres.s)

eqK equilibrium constant

SK adsorption constant in Equation 4.5.

LK global mass transfer coefficient (m/s)

L bed length (m)

n number of moles (mol)

Pe Peclet number

q average solid phase concentration in equilibrium with pC (mol/m3res)

Q molar adsorption capacity, defined as ,/i V mol iQ Q V= (mol/m3res)

QV volumetric monolayer capacity (m3/m3res)

r reaction rate relative to fluid concentration inside the particle (mol/kgres.s)

pr particle radius (m)

pRe Reynolds number relative to particle

pSh Sherwood number relative to particle

Sc Schmidt number

T temperature (K)

t time (s)

t* switching time (min)

u interstitial velocity (m/s)

molV molar volume in the liquid phase (m3/mol)

V volume of the bulk (m3)

x liquid phase molar fraction


z axial coordinate (m)

Greek letters

ε bulk porosity

pε porosity of pellet

η fluid viscosity (kg/m.s)

mη mixture viscosity (kg/m.s)

φ volume fraction

ν stoichiometric coefficient

ρ liquid density (kg/m3)

bρ bulk density (kgres/m3)

τ tortuosity

Subscripts

0 relative to initial conditions

exp experimental

F relative to the feed

i relative to component i (i= Eth, La, EL, W)

out at the end of the fixed bed column

p relative to particle

theor theoretical

Eth relative to ethanol

La relative to lactic acid

EL relative to ethyl lactate

W relative to water


Superscripts

ads adsorbed

des desorbed

4.7 References

Borges da Silva E. A., A. A. Ulson de Souza, S. G. U. de Souza and A. E. Rodrigues, "Analysis of the high-fructose syrup production using reactive SMB technology", Chem Eng J. 118(3): 167-181, 2006.

Catchpoole H. J., R. Andrew Shalliker, G. R. Dennis and G. Guiochon, "Visualising the onset of viscous fingering in chromatography columns", J Chrom A 1117(2): 137-145, 2006.

Funk G. A., J. R. Lansbarkis and A. K. Chandhok, "Process for concurrent esterification and separation using a simulated moving bed", Patent 5,405,992 (1995).

Gandi G. K., V. M. T. M. Silva and A. E. Rodrigues, "Synthesis of 1,1-dimethoxyethane in a fixed bed adsorptive reactor", Ind Eng Chem Res. 45(6): 2032-2039, 2006.

Glueckauf E., "Theory of chromatography", Trans. Faraday Soc 51: 1540-1551 1955.

Gonzalez B., N. Calvar, E. Gomez and A. Dominguez, "Density, dynamic viscosity, and derived properties of binary mixtures of methanol or ethanol with water, ethyl acetate, and methyl acetate at T = (293.15, 298.15, and 303.15) K", J. Chem. Thermodyn. 39(12): 1578-1588, 2007.

Grunberg L. and A. H. Nissan, "Mixture law for viscosity", Nature 164(4175): 799-800, 1949.

Ihm S. K., M. J. Chung and K. Y. Park, "Activity difference between the internal and external sulfonic groups of macroreticular ion-exchange resin catalysts in isobutylene hydration", Ind Eng Chem Res. 27(1): 41-45, 1988.

Kawase M., Y. Inoue, T. Araki and K. Hashimoto, "The simulated moving-bed reactor for production of bisphenol A", Catal Today 48(1-4): 199-209, 1999.

Kawase M., T. B. Suzuki, K. Inoue, K. Yoshimoto and K. Hashimoto, "Increased esterification conversion by application of the simulated moving-bed reactor", Chem Eng Sci. 51(11): 2971-2976, 1996.

Li J. and P. W. Carr, "Accuracy of Empirical Correlations for Estimating Diffusion Coefficients in Aqueous Organic Mixtures", Anal. Chem. 69(13): 2530-2536, 1997.

Lode F., G. Francesconi, M. Mazzotti and M. Morbidelli, "Synthesis of methylacetate in a simulated moving-bed reactor: Experiments and modeling", AIChE J. 49(6): 1516-1524, 2003.


Lode F., M. Houmard, C. Migliorini, M. Mazzotti and M. Morbidelli, "Continuous reactive chromatography", Chem Eng Sci. 56(2): 269-291, 2001.

Macias-Salinas R., F. Garcia-Sanchez and G. Eliosa-Jimenez, "An equation-of-state-based viscosity model for non-ideal liquid mixtures", Fluid Phase Equilib. 210(2): 319-334, 2003.

Mallmann T., B. D. Burris, Z. Ma and N. H. L. Wang, "Standing wave design of nonlinear SMB systems for fructose purification", AlChE J. 44(12): 2628-2646, 1998.

Mazzotti M., A. Kruglov, B. Neri, D. Gelosa and M. Morbidelli, "A continuous chromatographic reactor: SMBR", Chem Eng Sci. 51(10): 1827-1836, 1996.

Mazzotti M., B. Neri, D. Gelosa and M. Morbidelli, "Dynamics of a Chromatographic Reactor: Esterification Catalyzed by Acidic Resins", Ind Eng Chem Res. 36(8): 3163-3172, 1997.

Minceva M., P. S. Gomes, V. Meshko and A. E. Rodrigues, "Simulated moving bed reactor for isomerization and separation of p-xylene", Chem Eng J. 140(1-3): 305-323, 2008.

Motin M. A., M. H. Kabir and M. E. Huque, "Viscosities and excess viscosities of methanol, ethanol and n-propanol in pure water and in water + surf excel solutions at different temperatures", Phys. Chem. Liq. 43(2): 123-137, 2005.

Pereira C. S. M., P. S. Gomes, G. K. Gandi, V. M. T. M. Silva and A. E. Rodrigues, "Multifunctional Reactor for the Synthesis of Dimethylacetal", Ind Eng Chem Res. 47(10): 3515-3524, 2008.

Perkins L. R. and C. J. Geankoplis, "Molecular diffusion in a ternary liquid system with the diffusing component dilute", Chem. Eng. Sci. 24(7): 1035-1042, 1969.

Pöpken T., L. Götze and J. Gmehling, "Reaction kinetics and chemical equilibrium of homogeneously and heterogeneously catalyzed acetic acid esterification with methanol and methyl acetate hydrolysis", Ind Eng Chem Res. 39(7): 2601-2611, 2000.

Ruggieri R., G. Ranghino, G. Carvoli, A. Tricella, D. Gelosa and M. Morbidelli, "Process for esterification in a chromatographic reactor", Patent 6,586,609 (2003).

Ruthven D. M., "Principles of Adsorption and Adsorption Processes", Wiley & Sons, New York (1984).

Santacesaria E., M. Morbidelli, A. Servida, G. Storti and S. Carra, "Separation of xylenes on Y zeolites. 2. Breakthrough curves and their interpretation", Ind Eng Chem Process Des Dev 21(3): 446-451, 1982.

Scheibel E. G., "Correspondence. Liquid Diffusivities. Viscosity of Gases", Ind Eng Chem 46(9): 2007-2008, 1954.


Silva V. M. T. M. and A. E. Rodrigues, "Novel process for diethylacetal synthesis", AIChE J. 51(10): 2752-2768, 2005.


Troupe R. A., W. L. Aspy and P. R. Schrodt, "Viscosity and Density of Aqueous Lactic Acid Solutions", Ind Eng Chem 43(5): 1143-1146, 1951.

Umesi N. O. and R. P. Danner, "Predicting diffusion coefficients in nonpolar solvents", Ind Eng Chem Proc DD 20(4): 662-665, 1981.

Vignes A., "Diffusion in binary solutions: Variation of diffusion coefficient with composition", Ind. Eng. Chem. Fundam. 5(2): 189-199, 1966.

Yu W., K. Hidajat and A. K. Ray, "Modeling, Simulation, and Experimental Study of a Simulated Moving Bed Reactor for the Synthesis of Methyl Acetate Ester", Ind Eng Chem Res. 42(26): 6743-6754, 2003.

Yu W., K. Hidajat and A. K. Ray, "Determination of adsorption and kinetic parameters for methyl acetate esterification and hydrolysis reaction catalyzed by Amberlyst 15", Appl Catal A: Gen. 260(2): 191-205, 2004.

Zhang Y., L. Ma and J. Yang, "Kinetics of esterification of lactic acid with ethanol catalyzed by cation-exchange resins", React Funct Polym 61(1): 101-114, 2004.

5. Simulated Moving Bed Reactor

Abstract. A novel approach for the synthesis of ethyl lactate using a simulated moving bed

reactor was evaluated by experiments as well as by simulations. A mathematical model

considering external and internal mass-transfer resistances and variable velocity due to

change of liquid composition was developed to describe the dynamic behaviour of the SMBR

and it was validated by the experiments performed; it was observed that the experimental

results were well predicted by the model. The effect of operating parameters, as the feed

composition, SMBR configuration and switching time on the SMBR performance parameters

at the optimal operating points and/or reactive/separation regions was studied. It was shown

that the SMBR is a very attractive technology for the production of ethyl lactate, since under

appropriate conditions the lactic acid conversion can be driven to completion and productivity

as high as 32 kgEL/(Lads.day) and purity of 95 % can be obtained.

Adapted from: Pereira C. S. M., M. Zabka, V. M. T. M. Silva and A. E. Rodrigues, "A novel process for the

ethyl lactate synthesis in a simulated moving bed reactor (SMBR)", Chem. Eng. Sci. 64(14): 3301-3310, 2009.

102 CHAPTER 5. Simulated Moving Bed Reactor

5.1 Introduction

Esterification and transesterification (Ferreira and Loureiro, 2004; Ma et al., 2008; Pereira et

al., 2008b; Schmid et al., 2008; Yadav and Devi, 2004), acetalization (Chopade and Sharma,

1997; Gandi et al., 2005; Silva and Rodrigues, 2001; Yadav and Pujari, 1999), and

etherification (Cruz et al., 2005; Yadav and Lande, 2005) reactions, that are limited by the

thermodynamic equilibrium should be performed using hybrid technologies where reaction

and separation take place in a single unit, like chromatographic reactors or membrane

reactors, given that the reaction yield will be increased by removing one or more products.

Moreover, the use of this kind of technologies allows cost reduction and higher product

purity. In order to improve the process the use of solid acid catalysts is preferable. Since they

are less corrosive and, in opposition to homogeneous ones, they not require a further step of

separation and neutralization. Among the solid acid catalysts, ion exchange resins, like

Amberlyst 15-wet (A15), offer to the ethyl lactate production the advantage of having a

double role, they act as catalyst and as selective adsorber between the two reaction products,

water and ethyl lactate. Therefore, the use of the Amberlyst 15-wet resin in a

chromatographic reactor to produce ethyl lactate seems to be interest. The principle of the

chromatography batch reactor has been developed by Dinwiddie and Morgan in the beginning

of the 1960s (Dinwiddie and Morgan, 1961). The reactant is injected as a sharp pulse into a

fixed-bed column; during its propagation along the column it reacts and leads to different

products. The different affinity of the components with the solid phase leads to different

propagation velocities and, because of that, it is possible to separate the products from the

reaction mixture. For equilibrium limited reactions where complete separation of the products

is achieved the conversion can overcome the thermodynamic equilibrium value. However, the

batch chromatographic reactor is not the best technology to overcome that limitation, since it

has the usual drawbacks of a batch operation, where high quantity of eluent is needed to

perform the separation, resulting in highly diluted products. These led to the development of

continuous chromatographic reactors in order to decrease the solvent consumption and to

enhance the productivity, where the continuous mode of operation can be achieved by

moving the adsorbent or simulating its motion. However, the flow of solid particles in the

columns leads to some difficulties such as particle attrition, fluid velocity limited by

fluidization phenomena, considerable pressure drops, lack of efficiency, that are overcome by

the simulated moving bed (SMB) technology (Broughton and Gerhold, 1961), in which the

counter-current movement of the solid is simulated by using a series of packed beds where


the inlet/outlet ports are synchronously shifted in direction of the liquid flow. The extension

of the SMB technology to integrate reaction and separation steps into a single apparatus led to

the Simulated Moving Bed Reactor (SMBR). This chromatographic reactor gained more and

more attention in the last years because of its potential to improve process efficiency, as in the

case of esterifications (Dunnebier et al., 2000; Kawase et al., 1996; Lode et al., 2003; Lode et

al., 2001; Mazzotti et al., 1996; Meissner and Carta, 2002; Migliorini et al., 1999; Yu et al.,

2003), etherifications (Zhang et al., 2001), sugars (Azevedo and Rodrigues, 2001; Da Silva

et al., 2005; Da Silva et al., 2006; Dunnebier et al., 2000; Pilgrim et al., 2006; Zhang et al.,

2004), bisphenol A (Kawase et al., 1999), acetals (Dubois, 2008a; Pereira et al., 2008a;

Rodrigues and Silva, 2005; Silva and Rodrigues, 2005) and polyacetals (Dubois, 2008b) and

will be applied in this work on the synthesis of the green solvent ethyl lactate. A schematic

diagram of a SMBR unit is presented in Figure 5.1 where a reaction of type A+B↔C+D is

considered. Similarly to the SMB, the SMBR consists of a set of columns connected in series

packed with a solid, which could be a mixture of catalyst and selective adsorbent or a solid

that acts both as catalyst and as adsorbent. Typically, there are two inlets, feed and desorbent,

and two outlets, extract and raffinate. In this case, the component A is used as reactant and

desorbent, therefore it is introduced in the system in the feed and desorbent streams. The

other reactant B is used as feed. At regular time intervals, called the switching time, all

streams are switched for one column distance in direction of the fluid flow. In this way, the

countercurrent motion of the solid is simulated and its velocity is equal to the length of a

column divided by the switching time. A cycle is completed when the number of switches is

equal to a multiple of the total number of columns. According to the position of the inlet and

outlet stream the unit can be divided in four sections. In section 1, positioned between the

desorbent and extract nodes, the adsorbent is regenerated by desorption of the more strongly

adsorbed product (D) from the solid. In section 2 (between the extract and feed node) and

section 3 (between the feed and raffinate node) the reaction is taking place and the products C

and D are formed. The more strongly adsorbed product, D, is adsorbed in sections 2 and 3

and transported with the solid phase to the extract port. The less strongly adsorbed product, C,

is desorbed in sections 2 and 3 and transported with the liquid in direction of the raffinate

port. In section 4, positioned between the raffinate and desorbent node, before being recycled

to section 1, the desorbent is regenerated by adsorption of the less adsorbed product (C).


Desorbent (A) Raffinate (A+C)

Extract(A+D)

Direction of fluid flow and port switching

1 2 3 4

5 6

12 11

10 9

8 7

Feed(A+B)

A + B C + D

Figure 5.1 Schematic diagram of a SMBR with three columns per section.

In this work an alternative technology, the SMBR, is proposed for the ethyl lactate

production. The ethyl lactate synthesis is experimentally studied using a SMBR pilot unit and

theoretically using the SMBR model proposed. This model results are verified by the

experiments performed and the effect of various operating parameters, as feed composition,

SMBR configuration and switching time, into the SMBR performance at the optimal

operating points and/or reactive/separation regions is evaluated. The SMBR model is

compared with the equivalent True Moving Bed Reactor (TMBR) model. All this study will

allow to future compare the SMBR and the membrane reactor processes in economic and

efficiency terms.

5.2 Modelling Strategies

5.2.1 SMBR mathematical model

The SMBR modelling strategy allows the visualization of the axial movement of

concentration profiles and the variations in extract and raffinate concentrations within a

period. The dynamic calculation of concentration profiles until the unit achieves the cyclic

steady state, accounting for the inlet and outlet discontinuous shift, was made considering

axial dispersion flow for the bulk fluid phase, linear driving force (LDF) approximation for

the inter and intra-particle mass transfer rates, multi-component adsorption equilibrium and


velocity variations due to adsorption/desorption rates leading to more accurate results. The

porosity of the packed bed and its length (packed bed length) were assumed to be constant.

The SMBR model equations are:

Bulk fluid mass balance to component i in column k:

( ) ( ),, , ,

1 3ik ik k ikp ikL ik ik ax k T k

p

C C u xK C C D Ct z r z z

εε−∂ ∂ ∂∂ ⎛ ⎞+ + − = ⎜ ⎟∂ ∂ ∂ ∂⎝ ⎠

(5.1)

where ikC and ,p ikC are the bulk and average particle concentrations in the fluid phase of

species i in column k respectively, ikLK , is the global mass transfer coefficient of the

component i, ε is the bulk porosity, t is the time variable, z is the axial coordinate, kaxD , , and

ku are the axial dispersion coefficient and the interstitial velocity in column k, respectively,

and pr is the particle radius.

The global mass transfer coefficient ( LK ) is defined as:

ipeL kkK ε111

+= (5.2)

wherein ek and ik are, respectively, the external and internal mass transfer coefficients. The

values of the external, internal and global mass transfer coefficients are shown in Table 5.1.

The methods used to obtain these values and the determination of axial dispersion

coefficient, axD , are presented in detail in Chapter 4.

Table 5.1 Mass transfer coefficients estimated at 50 ºC at the inlet of the SMBR for a simulation where the feed is lactic acid solution (86 % in water) at a flowrate of 7 mL/min.

Component ke (cm/min) ki (cm/min) KL (cm/min)

Ethanol 0.387 0.138 0.045

Lactic Acid 0.351 0.119 0.039

Ethyl Lactate 0.296 0.092 0.031

Water 0.623 0.280 0.090


Interstitial fluid velocity variation is calculated using the total mass balance:

( ) ( ),, ,1

1 3 nk

p ikL ik mol i ikip

du K V C Cdz r

εε =

−= − −∑ (5.3)

where Vmol,,i is the molar volume of component i that is 60.87 mL/mol, 77.56 mL/mol,

118.44 mL/mol and 18.63 mL/mol at 50 ºC for ethanol, lactic acid, ethyl lactate and water,

respectively.

Pellet mass balance to component i, in column k:

( ) ( ),, ,,

31 ( )p ik ikp ik p ikp p L ik ik i p

p

qC K C C r Ct t r

ε ε υ ρ∂∂+ − = − +

∂ ∂ (5.4)

where ikq is the average adsorbed phase concentration of species i in column k in equilibrium

with ,p ikC , pε the particle porosity, iυ the stoichiometric coefficient of component i, pρ the

particle density and r is the chemical reaction rate relative to the average particle

concentrations in the fluid phase and is given by (Chapter 3):

2

,1 ⎟⎠

⎞⎜⎝

⎛+

−=

∑=

W

Ethiiis

WELLaEth

c

aK

Kaaaa

kr (5.5)

where kc is the kinetic constant, isK , is the adsorption constant for species i and Keq is the

equilibrium reaction constant, a is the species activity (calculated by UNIQUAC model) and

the subscripts WELLaEth and , , refer to ethanol, lactic acid, ethyl lactate and water,

respectively.

Multi-component adsorption equilibrium isotherm:

,,

,1

1

p ikads i iik n

p ikll

Q K Cq

K C=

=⎛ ⎞

+⎜ ⎟⎝ ⎠

∑ (5.6)

where, iadsQ , and iK represent the total molar capacity per unit volume of resin and the

equilibrium constant for component i, respectively, and n the total number of components.

The adsorption parameters were determined and presented in Chapter 4.


Initial and Danckwerts boundary conditions:

0=t : , ,0p ikik ikC C C= = and ,0ik ikq q= (5.7)

0=z : , ,0

ikk ik ax k k ik F

z

Cu C D u Cz =

∂− =

∂ (5.8a)

,0k ku u= (5.8b)

cLz = : 0=∂

∂

=Lcz

ik

zC (5.8c)

where F and 0 refer to the feed and initial states, respectively.

Mass balances at the nodes of the inlet and outlet lines of the SMBR:

Desorbent node: 1( 4, ) ( 1, 0)

4 4

DDi j z Lc i j z i

u uC C Cu u= = = == − (5.9a)

Extract(j=2) and Raffinate (j=4) nodes: )0,(),1( ==− = zjiLczji CC (5.9b)

Feed node: Fi

FziLczi C

uuC

uuC

2)0,3(

2

3),2( −= == (5.9c)

where,

Dsuuu += 41 Desorbent (Ds) node ; (5.10a)

Xuuu −= 12 Extract (X) node ; (5.10b)

Fuuu += 23 Feed (F) node ; (5.10c)

Ruuu −= 34 Raffinate (R) node ; (5.10d)

The ratio between the fluid interstitial velocity, uj, and the simulated solid velocity, Us,

(defined by the column length and switching time relation,*t

LUs = ) could be defined for

each section giving a new parameter:

Usu j

j =γ (5.10e)


5.2.2 SMBR performance parameters

The SMBR process performance is calculated over a complete cycle according to the next

equations:

Raffinate Purity:

( )

*

*(%) 100

c

c

t N tR

ELt

t N tR R R

La EL Wt

C dtPUR

C C C dt

+

+=

+ +

∫

∫ (5.11a)

Extract Purity:

( )

*

*(%) 100

c

c

t N tX

Wt

t N tX X X

La EL Wt

C dtPUX

C C C dt

+

+=

+ +

∫

∫ (5.11b)

Lactic acid Conversion:

⎟⎟⎟⎟⎟

⎠

⎞

⎜⎜⎜⎜⎜

⎝

⎛+

−=∫ ∫

+ +

*1100(%)

* *

tNCQ

dtCQdtCQX

cFlaF

tNt

t

tNt

t

RLaR

XLaX

c c

(5.12)

Raffinate Productivitysin

EL

re

kgdayL

⎛ ⎞⎜ ⎟⎝ ⎠

: ( ) *1

*

tNV

dtCQPR

cunit

tNt

t

RELR

c

ε−=

∫+

(5.13)

Desorbent Consumption (LEth/kgEL):

∫+

−+= *

,,, )(* tNt

t

RELR

FLaFEthFDEthDc c

dtCQ

XCCQCQtNDC (5.14)

where Nc represents the total number of columns (the complete cycle). The productivity is

defined considering the ethyl lactate (desired product) produced and withdrew from the

raffinate stream. The ethanol consumed in the reaction is not taken into account to calculate

the desorbent consumption, being only considered the amount of ethanol used as desorbent.

5.2.3 Numerical Solution

The above model equations were solved numerically by using the gPROMS-general PROcess

Modelling System version: 3.0.3. The mathematical model involves a system of partial and

algebraic equations (PDAEs). The axial domain was discretized using third order orthogonal

collocation in finite elements method (OCFEM). Ten finite elements per column with two


collocation points in each element were used. The system of ordinary differential and

algebraic equations (ODAEs) was integrated over time using the DASOLV integrator

implementation in gPROMS. For all simulations was fixed a tolerance equal to 10-5. It is

assumed that an SMB simulation has reached the cyclic steady state when the columns

profiles in two consecutive cycles have less than 1.0 % of relative deviation and the global

mass balance is verified with less than 1.0 % of relative error.


5.3.1 Chemicals and Catalyst / Adsorbent

The chemicals used in the experiments were ethanol (>99.5 % in water), lactic acid (>85 % in

water) and ethyl lactate (>98 % in water) from Sigma-Aldrich (U.K.).

The columns were packed with Amberlyst 15-wet (Rohm & Haas, France), which is a highly

cross-linked polystyrene-divinylbenzene ion exchange resin functionalized with sulfonic

groups (SO3H) that acts as catalyst and adsorbent in this system.

5.3.2 The SMBR LICOSEP 12-26 Unit

All SMBR experiments were performed in a pilot unit LICOSEP 12-26 by Novasep

(Vandoeuvre-dès-Nancy, France), where 12 columns Superformance SP 230 x 26 (length x

internal diameter, mm), by Götec Labortechnik (Mühltal, Germany), packed with the acid

resin Amberlyst 15-wet were connected. The characteristics of the SMBR columns are

presented in Table 5.2 and a side view of the pilot unit is shown in Figure 5.2. The bed

porosity and the Peclet number were determined as in previous Chapter (Chapter 4). These

columns can withstand up to 60ºC of temperature and 60 bar of pressure. The operating

temperature was 50ºC; this temperature was ensured in the jacketed columns using a

thermostatic bath (Lauda, Germany). Between every two columns exists a four-port valve

(Top-Industrie, France) actuated by the control system. When required, according to the

operating conditions set into the system, the valves allow either pumping of the feed or

desorbent into the system or withdrawal of extract or raffinate from the system. HPLC pumps

are used to pump each of the inlet streams (feed or desorbent) and outlet streams (extract or

raffinate). The maximum flow rate in the desorbent and extract pumps is 30 mL/min, while in


the feed and raffinate pumps is 10 mL/min. The recycling pump, which is a positive

displacement three-headed membrane pump (Dosapro Milton Roy, France), can deliver flow

rates as low as 20 mL/min up to 120 mL/min and it can hold up to 100 bar pressure. A six-

port valve, located between the twelfth and the first columns, was used to collect the samples

for internal concentration measurements (Novasep). All the samples were analysed in a gas

chromatograph using the analytical method described in Chapter 3.

Figure 5.2 Licosep 12-26 unit with the 12 packed columns.

Table 5.2 Characteristics of the SMBR columns.

Solid weight (A15) 47.6 g

Length of the bed (L) 23 cm

Internal diameter (Di) 2.6 cm

Average radius of resin beads (rp) 342.5 μm

Bulk density (ρ b) 390 kg/m3

Bed porosity (ε) 0.4

Resin particle porosity (εp) 0.36 (Lode et al., 2001)

Peclet number 300



5.4.1 Experimental Results

Three different experiments of ethyl lactate synthesis were performed at 50ºC in the SMBR

unit. In all the experiments the feed composition was a lactic acid solution (85 % in water).

The flow rates of feed and raffinate streams were QF = 1.80 mL/min and QR = 8.80 mL/min,

while for desorbent and recycle the flow rates were QD = 27.40 mL/min and

QRe c = 24.70 mL/min. Two different switching times (2.9 and 3.0 min) and two different

configurations (3-3-3-3 and 3-3-4-2) where tested, as shown in Table 5.3. The SMBR

performance parameters obtained experimentally are presented in Table 5.3, as well as the

corresponding values obtained by simulation. In Figure 5.3, the cyclic steady-state (CSS)

concentration profiles obtained experimentally at the middle of the switching time after 13

cycles are shown and compared with the profiles simulated by the SMBR model.

Table 5.3 Operating conditions for SMBR experiments and performance parameters obtained experimentally (bold) and by simulation (inside brackets).

Run 1 Run 2 Run 3

t* (min) 2.9 2.9 3

Configuration 3-3-3-3 3-3-4-2 3-3-4-2

PUR (%) 72.75 (79.24) 73.54 (79.80) 75.17 (81.32)

PUX (%) 95.54 (99.89) 95.22 (99.83) 97.76 (99.85)

X (%) 99.10 (97.46) 99.91 (97.87) 99.96 (98.00)

PR (kgEL.Lresin-1.day-1) 3.20 (3.66) 3.31 (3.68) 3.36 (3.68)

DC (LEth/kgEL) 13.41 (11.74) 12.99 (11.70) 12.77 (11.68)

For the first experiment the lactic acid conversion is about 99 %, being for the second and

third experiments around 100 %. However the raffinate purities are low, this is due to the fact

that the SMBR experiments realized were performed under conditions of incomplete

adsorbent regeneration in section 1 (see Figure 5.3) caused by equipment limitations

(maximum allowable desorbent flow rate of 30 mL/min).


b)

a)

c)

Figure 5.3 Experimental concentration profiles in SMBR unit at the middle of switching time at cyclic steady state (13th cycle) and simulated curves.

(a) Run 1; (b) Run 2; (c) Run3.


The feed is a lactic acid solution (85 % in water), so water is already introduced into the

system and as the reaction proceeds more water is being formed. Since water is the most

adsorbed component and there is a significant amount of this component in the SMBR, the

desorbent flow rate used was not high enough to completely regenerate (removing the water

adsorbed by the resin) the first section; because of that the adsorbed water was carried by the

resin entering in section 4 and transported to section 3, hydrolysing the ethyl lactate

producing ethanol and lactic acid (the opposite reaction). As consequence the ethyl lactate

productivity and its concentration in the raffinate stream is low. The use of a higher desorbent

flow rate will permit the complete regeneration of section 1 leading to higher purities and

productivity. From Figure 5.3, it can be perceived that the reaction takes place in sections 2

and 3; since the esterification reaction between lactic acid and ethanol is very slow, the

change of configuration from 3-3-3-3 to 3-3-4-2 increases the lactic acid conversion, due to

the increase of the number of columns available in the reactive sections 2 and 3. Increasing

the switching time (Figure 5.3c), increases the lactic acid conversion once the contact time

between liquid and solid streams is higher. Moreover, the ethyl lactate purity increases since

the value of γ1 is higher and therefore the resin is better regenerated. The CSS concentration

profiles obtained experimentally at 25 %, 50 % and 75 % of the switching time, at the 13th

cycle are compared with the simulated ones, in Figure 5.4. It is noticed that all the species

concentration fronts propagate along the column without changing its shape presenting,

therefore, a constant behaviour along the time. For the third experiment, the experimental

composition of extract and raffinate streams collected for a whole cycle at the 3rd, 5th, 7th,

9th and 11th cycles, and the extract and raffinate concentration histories calculated from the

SMBR model are shown in Figure 5.5.

As it can be observed, the mathematical model used predicts with reasonable accuracy the

experimental concentration profiles (Figure 5.3 and Figure 5.4), the raffinate and extract

histories (Figure 5.5) and the performance parameters (Table 5.3).


Figure 5.4 Experimental concentration profiles in SMBR unit at the 25 %, 50% and 75 % of the switching time at cyclic steady-state (13th cycle) and simulated curves for Run 3.

Figure 5.5 Experimental and theoretical average concentration of all species (ethanol, lactic acid, ethyl lactate and water) for the conditions of the third experience in the extract (a) and raffinate (b) streams.

a) b)


5.4.2 Simulated results

5.4.2.1 Comparison of SMBR and TMBR models

The equivalent TMBR model was also studied since it is a simpler model that implies less

computational time than that required by the SMBR model. However, comparing the

concentration profiles at the cyclic steady state calculated from the SMBR model with those

calculated from the equivalent TMBR model, as shown in Figure 5.6, some difference is

noticed and, because of that, the SMBR model was used for all the simulations presented in

this work. The operating conditions used in the simulations are presented in Table 5.4 and the

feed flow rate and raffinate flow rate used were 5 mL/min and 17 mL/min, respectively.

Figure 5.6 Comparison of the concentration profiles at the middle of the switching time at cyclic steady state determined with the SMBR model (solid lines) and steady state concentration profiles calculated with the TMBR model (dashed lines), for the operating conditions presented in Table 5.


Table 5.4 Operating conditions.

Operating conditions

Configuration (analyzed in section 5.4.2.5) 3-3-4-2

Feed Concentration (mol/L) (analyzed in section 5.4.2.4)

CEth,F =0.0; CEL,F = 0.0 CLA,F =10.75; CW,F = 9.48

Desorbent flow rate (mL/min) 58.0

Recycle flow rate (mL/min) 27.0

Switching time (min) (analyzed in section 5.4.2.6) 2.7

5.4.2.2 Reactive/separation regions

The proper design of a SMBR unit involves the correct choice of the operating conditions as

flow rates in each section of the unit, switching time, columns arrangement (SMBR

configuration) and feed concentration, among others. In the next sub-sections, the

reactive/separation regions and the optimal operating points for different operating conditions

are presented in order to understand the behaviour of the SMBR process for ethyl lactate

synthesis. The operating conditions used in the simulations indicating which parameter is

being changed in each sub-section are presented in Table 5.4. The reactive/separation regions

setting a criteria of 95 % for extract and raffinate purity and, also, for lactic acid conversion

were determined from the average concentrations over a cycle obtained by the SMBR model

at cyclic steady state. The cyclic steady state SMBR model was successively solved for

several values of γ2 and γ3, keeping the values of γ1 (4.698) and γ4 (1.492). The

reactive/separation region is located within the region between the diagonal γ2=γ3, the

horizontal line γ3 =1.492 and γ3 axis. The γ3 value must be higher than γ2, since the diagonal

γ2=γ3 corresponds to zero feed flow rate. The algorithm used in the construction of the

reactive/separation region begins by setting a feed flow rate of 0.01 ml/min and the value of

γ2 equal to 1.492. Then, the feed flow rate was kept constant and the γ2 values were gradually

increased. The value of γ3 was calculated from the mass balance in the feed node for each

value of γ2. For each set of γ2 and γ3, the conversion and the purities of extract and raffinate

were estimated and the values that satisfy the criteria of 95 % were selected to build the

reaction/separation region. After this set of simulations, the feed flow rate was increased and


the same procedure was repeated. The simulations procedure ends when the maximum value

of feed flow rate that gives required product purities and conversion is achieved. Above that

feed flow rate value the requirements can not be fulfilled for any pair of values of γ2 and γ3.

5.4.2.3 Separation Region vs Reactive/Separation Region

One can use a SMB unit to perform the separation of ethyl lactate and water, in absence of

reaction, considering the purity criteria of 95 %, to process a feed with composition of

CW = 12.17 mol/L and CEL = 6.47 mol/L that corresponds to complete conversion of the

lactic acid solution (CAL = 10.75 mol/L and CW = 9.48 mol/L) into ethyl lactate and water,

taking into consideration the expression ( ),i i n m kk

C x x V= ∑ , where ix is the molar fraction of

species i and nmV , is the molar volume of species n.

In Figure 5.7, the comparison between the separation region of the SMB and the

reactive/separation region of the SMBR indicates that the process is limited by the reaction.

As mentioned by Fricke and Schmidt-Traub, the reactive separation region is reduced by

decreasing either the equilibrium constant or the reaction kinetics (Fricke and Schmidt-Traub,

2003). The lactic acid esterification reaction is slow and, because of that, it is necessary to

keep the reactant as long as possible in sections 2 and 3 (between the extract and raffinate

ports). If the lactic acid is not fully consumed it will preferentially contaminate the raffinate

stream, since the resin selectivity between lactic acid and ethyl lactate is smaller than the one

between water and lactic acid. Thus, in the SMBR it is necessary to ensure the separation of

the products (ethyl lactate and water), but is crucial to guarantee the complete conversion of

the limiting reactant (lactic acid). One way of improving reaction kinetics is by increasing the

system temperature that will also benefit the equilibrium constant since the ethyl lactate

synthesis is an endothermic reaction. However, this leads to higher energy requirement and it

will affect the products separation: (i) if water adsorption is more affected than ethyl lactate

adsorption, less desorbent consumption is required, but maximum productivity decreases; (ii)

in the opposite case, productivity is enhanced while the consumption of desorbent increases.

The trade-off between those parameters should be evaluated through an economical

assessment. Another way to get the full conversion is increasing the reaction rate by using a

mixture of two stationary phases that increases the catalytic properties without losing


efficiency on the separation of the products (Silva and Rodrigues, 2008; Ströhlein et al.,

2004).

1.0

1.5

2.0

2.5

3.0

3.5

4.0

1.0 1.5 2.0 2.5 3.0 3.5 4.0

γ 3

γ2

Feed: lactic acid solution (SMBR)

Feed: ethyl lactate and water (SMB)

Figure 5.7 Reactive/separation region for the case where lactic acid solution is fed to

the SMBR (CLA,F = 10.75 mol/L; CW,F = 9.48 mol/L) and separation region for ethyl lactate and water fed to the SMB (CEL,F = 6.47 mol/L; CW,F = 12.17 mol/L). (γ1 = 4.698; γ4 =1.492; 95 % purity).

5.4.2.4 Effect of the Feed Composition

In order to study the influence of the feed composition, this was varied by adding pure

ethanol to the lactic acid solution remaining constant the other operation conditions (see

Table 5.4). As it can be seen in Figure 5.8, the size of the reactive/separation region decreases

with the increase of the lactic acid concentration in the feed. This could be justified by a lack

of ethanol in the sections 2 and 3, where reaction occurs, limiting the lactic acid conversion to

the thermodynamic equilibrium and so higher lactic acid concentrations will lead to higher

quantities of unreacted lactic acid and therefore the raffinate stream will be contaminated.

This could be confirmed by analyzing the influence of the lactic acid feed molar fraction in

the ethyl lactate productivity and the desorbent consumption for the optimal operating points

(see Figure 5.9). It can be seen that by increasing the lactic acid feed molar fraction until

72 %, where the lactic acid is the limiting reactant, the productivity increases with consequent

decrease in desorbent consumption; however, further increase in the lactic acid feed molar

fraction has a negative impact in the SMBR performance parameters since, as it was

mentioned, ethanol is the limiting reactant in sections 2 and 3. Thus, for the values of γ1 and


γ4 and SMBR configuration defined in Table 5.4, the feed molar composition of 72 % of

lactic acid leads to the optimal values of productivity (17.33 kgEL/(Lads.day)) and desorbent

consumption (5.36 LEth/kgEL). So, it can be concluded that the lactic acid concentration in the

SMBR feed should be high enough to increase the productivity without leading to a lack of

ethanol in the reaction zone (section 2 and 3).

1.5

2.0

2.5

3.0

3.5

1.5 2.0 2.5 3.0 3.5

γ 3

γ2

100 % La solution

88 % La solution

38 % La solution

Figure 5.8 Reactive/separation regions for different feed concentrations

of lactic acid solution. (γ1 = 4.698; γ4 =1.492; 95 % purity).

5.0

5.8

6.6

7.4

8.2

9.0

12.0

13.0

14.0

15.0

16.0

17.0

18.0

35 45 55 65 75 85 95

DC

(LEt

h/K

g EL)

PR (K

g EL/

(Lad

s.day

))

Lactic acid solution (85 % in water) feed composition (%)

PR

DC

Figure 5.9 SMBR performance for the optimal operating points as a

function of the lactic acid solution molar fraction in feed.


5.4.2.5 Effect of the SMBR columns arrangement

The influence of the columns arrangement on the SMBR system performance was studied by

simulating different SMBR configurations (3-3-3-3, 3-3-4-2 and 3-3-5-1), keeping the

remaining operating conditions as mentioned in Table 5.4. The performance parameters

corresponding to the optimal operating point (vertex of the reactive/separation region) for

each case are presented in Table 5.5, and the reactive/separation regions are shown in Figure

5.10. It can be perceived that, the introduction of more columns into the section 3 by taking of

the section 4 and keeping the total number of columns constant, does not improve

significantly the process performance. This is due to the fact that after about 3 mL/min of

lactic acid solution feed flow rate the ethanol is the limiting reactant and even when the

reaction zone is increased is not possible to completely convert the lactic acid, which will

cause contamination of the raffinate stream. Although the increase of the reaction zone in

these conditions does not improve much the ethyl lactate production on the SMBR unit it was

noticed that only one column in section 4 is enough to allow the complete regeneration of the

desorbent (see Figure 5.11).

The kinetics of the ethyl lactate synthesis is very slow and to allow full conversion of the

lactic acid is necessary not only to increase the reaction zone, but also to feed to the SMBR a

lactic acid concentration that does not lead to a lack of ethanol in the reaction zone.

1.5

2.0

2.5

3.0

3.5

1.5 2.0 2.5 3.0 3.5

γ 3

γ2

conf 3-3-3-3conf 3-3-4-2conf 3-3-5-1

Figure 5.10 Reactive/separation regions for different SMBR configurations.

(γ1 = 4.698; γ4 =1.492; 95 % purity).


Table 5.5 Performance parameters of the SMBR unit at cyclic steady state for the optimal operation points for the different SMBR configurations.

Columns arrangement 3-3-3-3 3-3-4-2 3-3-5-1

Raffinate productivity (kgEL/LA15/day) 14.97 15.33 15.36

Desorbent consumption (LEth/kgEL) 5.83 5.68 5.67

Figure 5.11 Concentration profiles at the middle of the switching time at cyclic

steady state for the optimal operating point of the SMBR configuration 3-3-5-1.

5.4.2.6 Effect of Switching Time

The switching time effect on the SMBR performance for the optimal operating point was

studied by two different ways: one varying the switching time and keeping constant the

remaining operating parameters of Table 5.4 (Figure 5.12); the other varying the switching

time and the desorbent and recycle flow rate in order to keep constant the γ1 (4.698) and

γ4 (1.492) values (Figure 5.13). For the first case (Figure 5.12), as the switching time

decreases until 2.1 minutes, the solid flow rate increases leading to a better SMBR

performance. However, decreasing the switching time below 2.1 min, the restriction of γ1 is

being violated and, therefore, is no longer possible to achieve the minimum purity of 95 % for

the raffinate stream, since the solid is not being completely regenerated in section 1, and,

because of that, the water appears in the raffinate stream, leading to a low purity. The SMBR


operation using a switching time of 2.1 min instead of 2.8 min enhances the productivity in

22.0 % (maximum of 18.06 kgEL/(Lads.day)) and decreases the desorbent consumption in

19.6 % (minimum of 4.75 LEth/kgEL), and will reduce downstream costs since the products are

less diluted.

4.4

4.8

5.2

5.6

6.0

14.0

15.0

16.0

17.0

18.0

19.0

2.1 2.2 2.3 2.4 2.5 2.6 2.7 2.8

DC (L

Eth/

KgEL

)

PR (K

g EL/

(Lad

s.da

y))

Switching time (min)

PRDC

Figure 5.12 SMBR performance for the optimal operating points as a function

of the switching time.

For the second case (Figure 5.13), where the γ1 and γ4 values are keeping the same, similarly

to the previous case, the productivity increases by decreasing the switching time; contrarily,

the desorbent consumption increases once the desorbent flow rate is increasing in order to

keep the value of γ1 constant. Changing the switching time from 2.8 to 2.1, increases the

productivity and desorbent consumption to 19.34 kgEL/(Lads.day) and 5.79 LEth/kgEL,

corresponding to a variation of +30.7 % and +2.0 %, respectively. The reduction of 2.8 min to

1.0 min, in switching time, leads to a significant increase in productivity in 114.4 %

(31.7 kgEL/(Lads.day)), while the desorbent consumption is increased in 33.3 %

(7.6 LEth/kgEL). The choice of the optimal switching time value requires an economical

assessment to the whole process, considering the downstream separation units, in order to

verify if the enhancement of productivity reimburses the increase of the cost associated to the

desorbent separation and recycle to the SMBR unit.


5.6

6.0

6.4

6.8

7.2

7.6

12.0

16.0

20.0

24.0

28.0

32.0

1.0 1.4 1.8 2.2 2.6

DC (L

Eth/

Kg E

L)

PR (K

g EL/

(Lad

s.day

))

Switching time (min)

PR

DC

Figure 5.13 SMBR performance for the optimal operating points as a function of

the switching time keeping γ1 = 4.698 and γ4 =1.492.

5.5 Conclusions

The ethyl lactate was produced in the Simulated Moving Bed Reactor pilot unit LICOSEP,

using Amberlyst 15-wet resin as catalyst and selective adsorbent. The experiments conducted

were not performed under optimal operating conditions since they were limited by restrictions

of the experimental set-up; therefore, the desorbent flow rate used was not enough to achieve

complete regeneration of the resin in section 1. The performance parameters and the

experimental profiles were predicted with good accuracy using the SMBR model. The

theoretical assessment of the SMBR unit behaviour was performed ensuring complete

regeneration of the resin (in section 1) and desorbent (in section 4), by using the mathematical

model to analyse the effect of SMBR configuration, feed composition and switching time into

the reactive/separation regions or/and into the process performance at the optimal operating

points. Since the kinetics of the ethyl lactate synthesis is very slow, even at 50ºC, its

production in the SMBR is determined by the reaction, and therefore, the complete

conversion of lactic acid is a decisive parameter in order to optimize the SMBR performance.

To get that target, several aspects should be regarded: to play with feed composition to avoid

a lack of ethanol in the reactive sections 2 and 3; to increase the number of columns in those

sections to increase the reaction zone; to choose the best switching time, that affects

significantly the SMBR process. The other key aspect that should be considered is the water


desorption in section 1 to ensure the complete regeneration of the resin; Amberlyst 15-wet is

very selective to water and therefore higher desorbent flow rates are required. The effects of

the various operating parameters, as the feed composition, the SMBR configuration and the

switching time, on the SMBR process showed that there is a complex interaction among all of

them in terms of their impacts on the performance of the process, as the extract and raffinate

purities, lactic acid conversion, ethyl lactate productivity and desorbent consumption.

Moreover, some set of parameters might act in conflicting ways, like the case of the switching

time for fixed values of γ1 and γ4, that increase significantly the productivity but also increases

the desorbent consumption. For a fixed switching time of 2.7 min the best feed molar

composition was 72 % of lactic acid leading to a productivity of 17.33 kgEL/(Lads.day) and

desorbent consumption of 5.36 LEth/kgEL. The higher value of productivity,

31.7 kgEL/(Lads.day), was obtained for a switching time of 1 min, 100% of lactic acid solution

in feed, γ1 = 4.698 and γ4 =1.492.

5.6 Notation

a liquid-phase activity

C liquid phase concentration (mol/L)

CT total liquid phase concentration (mol/L)

axD axial dispersion coefficient (m2/min)

DC desorbent consumption (L/mol)

ck kinetic constant (mol kg-1 min-1)

ek external mass transfer coefficient

eqK equilibrium reaction constant

ik internal mass transfer coefficient

iK Langmuir equilibrium constant of component i (L/mol)

LK global mass transfer coefficient


L column length (m)

n total number of components

Nc total number of columns

PR raffinate productivity (kgEL/(L resin.day))

PUR raffinate purity (%)

PUX extract purity (%)

q solid phase concentration in equilibrium with the fluid concentration inside the

particle (mol/L)

,ads iQ adsorption capacity of component i (mol/Lwet solid)

jQ volumetric flowrate in section j (L/min)

r rate of reaction (mol kg-1 min-1)


t time variable (min)

*t switching time (min)

Us solid velocity (m/min)

u interstitial velocity (m/min)

,mol iV molar volume of species i (L/mol)

unitV volume of adsorbent in SMBR (L)

x mole fraction

X lactic acid conversion


Greek letters

γ interstitial velocities ratio

ε bulk porosity


pε particle porosity

iυ stoichiometric coefficient of component i

pρ particle density

Subscripts


j relative to section in SMBR (j = 1, 2, 3, 4)

k relative to column in SMBR




W relative to water




R relative to raffinate

cRe relative to recycle

X relative to extract

Superscripts





5.7 References Cited

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Broughton D. B. and C. G. Gerhold, "Continuous sorption process employing fixed bed of sorbent and moving inlets and outlets.", Patent 2 985 589 (1961).

Chopade S. P. and M. M. Sharma, "Reaction of ethanol and formaldehyde: use of versatile cation-exchange resins as catalyst in batch reactors and reactive distillation columns", React. Funct. Polym. 32(1): 53-65, 1997.

Cruz V. J., J. F. Izquierdo, F. Cunill, J. Tejero, M. Iborra and C. Fité, "Acid ion-exchange resins catalysts for the liquid-phase dimerization/etherification of isoamylenes in methanol or ethanol presence", React. Funct. Polym. 65(1-2): 149-160, 2005.

Da Silva E. A. B., A. A. U. De Souza, S. G. U. De Souza and A. E. Rodrigues, "Simulated moving bed technology in the reactive process of glucose isomerization", Adsorption 11(1 SUPPL.): 847-851, 2005.

Da Silva E. A. B., A. A. U. De Souza, S. G. U. De Souza and A. E. Rodrigues, "Analysis of the high-fructose syrup production using reactive SMB technology", Chem. Eng. J. 118(3): 167-181, 2006.

Dinwiddie J. A. and W. A. Morgan, " Fixed bed type reactor", Patent 2 976 132 (1961).

Dubois J.-L., "Procede de synthese d' acetals par transacetalisation dans un reacteur a lit mobile simule", Patent FR 2909669 (A1) (2008a).

Dubois J.-L., "Process for the synthesizing polyacetals in a simulated mobile-bed reactor", Patent WO 2008/053108 (2008b).

Dunnebier G., J. Fricke and K. U. Klatt, "Optimal Design and Operation of Simulated Moving Bed Chromatographic Reactors", Ind. Eng. Chem. Res. 39(7): 2290-2304, 2000.

Ferreira M. V. and J. M. Loureiro, "Number of actives sites in TAME synthesis: Mechanism and kinetic modeling", Ind. Eng. Chem. Res. 43(17): 5156-5165, 2004.

Fricke J. and H. Schmidt-Traub, "A new method supporting the design of simulated moving bed chromatographic reactors", Chem. Eng. Process. 42(3): 237-248, 2003.

Gandi G. K., V. M. T. M. Silva and A. E. Rodrigues, "Process Development for Dimethylacetal Synthesis: Thermodynamics and Reaction Kinetics", Ind. Eng. Chem. Res. 44(19): 7287-7297, 2005.

Kawase M., Y. Inoue, T. Araki and K. Hashimoto, "The simulated moving-bed reactor for production of bisphenol A", Catal. Today 48(1-4): 199-209, 1999.



Lode F., G. Francesconi, M. Mazzotti and M. Morbidelli, "Synthesis of methylacetate in a simulated moving-bed reactor: Experiments and modeling", AlChE J. 49(6): 1516-1524, 2003.

Lode F., M. Houmard, C. Migliorini, M. Mazzotti and M. Morbidelli, "Continuous reactive chromatography", Chem. Eng. Sci. 56(2): 269-291, 2001.

Ma L., J. Hong, M. Gan, E. Yue and D. Pan, "Kinetics of esterification and transesterification for biodiesel production in two-step process", J. Chem. Ind. Eng. China 59(3): 708-712, 2008.

Mazzotti M., A. Kruglov, B. Neri, D. Gelosa and M. Morbidelli, "A continuous chromatographic reactor: SMBR", Chem. Eng. Sci. 51(10): 1827-1836, 1996.

Meissner J. P. and G. Carta, "Continuous regioselective enzymatic esterification in a simulated moving bed reactor", Ind. Eng. Chem. Res. 41(19): 4722-4732, 2002.

Migliorini C., M. Fillinger, M. Mazzotti and M. Morbidelli, "Analysis of simulated moving-bed reactors", Chem. Eng. Sci. 54(13-14): 2475-2480, 1999.

Novasep, "Licosep 12-26 Instructions manual, France",

Pereira C. S. M., P. S. Gomes, G. K. Gandi, V. M. T. M. Silva and A. E. Rodrigues, "Multifunctional Reactor for the Synthesis of Dimethylacetal", Ind. Eng. Chem. Res. 47(10): 3515-3524, 2008a.

Pereira C. S. M., S. P. Pinho, V. M. T. M. Silva and A. E. Rodrigues, "Thermodynamic equilibrium and reaction kinetics for the esterification of lactic acid with ethanol catalyzed by acid ion-exchange resin", Ind. Eng. Chem. Res. 47(5): 1453-1463, 2008b.

Pilgrim A., M. Kawase, F. Matsuda and K. Miura, "Modeling of the simulated moving-bed reactor for the enzyme-catalyzed production of lactosucrose", Chem. Eng. Sci. 61(2): 353-362, 2006.

Rodrigues A. E. and V. M. T. M. Silva, "Industrial process for acetals production in a simulated moving bed reactor", Patent PT 103123, 2004; WO 2005/113476A1 (2005).

Schmid B., M. Döker and J. Gmehling, "Esterification of ethylene glycol with acetic acid catalyzed by Amberlyst 36", Ind. Eng. Chem. Res. 47(3): 698-703, 2008.

Silva V. M. T. M. and A. E. Rodrigues, "Synthesis of diethylacetal: thermodynamic and kinetic studies", Chem. Eng. Sci. 56(4): 1255-1263, 2001.


Silva V. M. T. M. and A. E. Rodrigues, "Influence of Different Mixtures of Catalyst and Adsorbent on the Smbr Performance: Reactive-Separation and Regeneration Regions ", 2008.


Ströhlein G., F. Lode, M. Mazzotti and M. Morbidelli, "Design of stationary phase properties for optimal performance of reactive simulated-moving-bed chromatography", Chem. Eng. Sci. 59(22-23): 4951-4956, 2004.

Yadav G. D. and K. M. Devi, "Immobilized lipase-catalysed esterification and transesterification reactions in non-aqueous media for the synthesis of tetrahydrofurfuryl butyrate: comparison and kinetic modeling", Chem. Eng. Sci. 59(2): 373-383, 2004.

Yadav G. D. and S. V. Lande, "Rate intensive and selective etherification of vanillin with benzyl chloride under solid-liquid phase transfer catalysis by aqueous omega phase.", J. Mol. Catal. A: Chem. 244: 271, 2005.

Yadav G. D. and A. A. Pujari, "Kinetics of acetalization of perfumery aldehydes with alkanols over solid acid catalysts.", Can. J. Chem. Eng. 77, 1999.

Yu W., K. Hidajat and A. K. Ray, "Modeling, Simulation, and Experimental Study of a Simulated Moving Bed Reactor for the Synthesis of Methyl Acetate Ester", Ind. Eng. Chem. Res. 42(26): 6743-6754, 2003.

Zhang Y., K. Hidajat and A. K. Ray, "Optimal design and operation of SMB bioreactor: Production of high fructose syrup by isomerization of glucose", Biochem. Eng. J. 21(2): 111-121, 2004.

Zhang Z., K. Hidajat and A. K. Ray, "Application of Simulated Countercurrent Moving-Bed Chromatographic Reactor for MTBE Synthesis", Ind. Eng. Chem. Res. 40(23): 5305-5316, 2001.

6. Pervaporation Membrane Reactor

Abstract. Integrated membrane reactors are one of the most promising and sustainable

technologies to carry out equilibrium limited reactions. However, that integration leads to

lower process flexibility when trying to find suitable operating conditions for both reaction

and pervaporation process.

Pervaporation process using commercial microporous silica water selective membranes was

evaluated to contribute either for the separation of SMBR extract stream (water/ethanol

mixtures) or for the ethyl lactate process intensification by continuous pervaporation

membrane reactor (PVMR). Preliminary studies were performed in order to assess the

existence of membrane defects and mass transfer limitations, by studying the influence of

feed pressure and flowrate, respectively. After, in the absence of mass transfer limitations,

membrane performance was evaluated experimentally, at different composition and

temperature measuring the flux and selectivity of each species in binary mixtures

(water/ethanol, water/ethyl lactate and water/lactic acid). Thus, species permeances were

obtained for each experiment and correlated in order to account for the effect of temperature

and feed composition. Permeances of ethanol and ethyl lactate depend solely on the

temperature, following an Arrhenius equation; for water, its permeance follows a modified

Arrhenius equation taking into account also the dependence on the feed water content.

Mathematical models, considering concentration and temperature polarization, and non-

isothermal effects as well, were developed and applied to analyze the performance of batch

pervaporation and continuous pervaporation membrane reactor, in both isothermal and non-

isothermal conditions. The PVMR with 5 membranes in series, operating at 70ºC, leads to

98 % of lactic acid conversion and 96 % of ethyl lactate purity.

132 CHAPTER 6. Pervaporation Membrane Reactor

6.1 Introduction

In the past few years, the interest in the application of reactive separations to many chemical

processes has substantially increased. This is particularly true for equilibrium limited

reactions, where the removal of at least one of the reaction products shifts the equilibrium

towards the product formation. Among all the separation techniques, pervaporation is

becoming a promising technology, potentially useful in applications such as dehydration of

organic mixtures (Chapman et al., 2008; Slater et al., 2006; Van Hoof et al., 2006;

Yoshikawa et al., 2002), separation of organic mixtures (Lin et al., 2009; Smitha et al., 2004)

and removal of organic compounds from aqueous solutions (Khayet et al., 2008; Ohshima et

al., 2005). Aiming to produce esters, where water is a by-product, membranes suitable for

dehydration of organic mixtures can be applied for esters process intensification; being

polymeric (Hasanoglu et al., 2007; Krishna Rao et al., 2007; Namboodiri et al., 2006) and

ceramic (Asaeda et al., 2005; Casado et al., 2005; Li et al., 2009) membranes the most used.

Based on this, several works have been made regarding the use of pervaporation membrane

reactors for chemicals production (Feng and Huang, 1996; Lim et al., 2002; Park and Tsotsis,

2004; Peters et al., 2005; Zhu et al., 1996). In the literature, two main types of pervaporation

membrane reactors can be found; reactor and membrane housed in separate units (Figueiredo

et al., 2008; Korkmaz et al., 2009; Sanz and Gmehling, 2006) or membrane and reactor

incorporated into a single unit (de la Iglesia et al., 2007). Although the number of works

dealing with PVMR in the last years is significant, most of them use simplified mathematical

models to optimize and predict the behaviour of PVMR units. Usually, concentration and

temperature polarization effects on the pervaporation process are neglected, which can be

very significant under certain PVMR working conditions (Gómez et al., 2007).

In pervaporation processes, the transport of the components from the feed liquid mixture to

the vapor phase involves the following steps: (i) mass transfer from the feed bulk to the feed

membrane interface; (ii) partition of penetrants between feed and membrane; (iii) selective

transport (diffusion) through the membrane; and (iv) desorption into the vapor phase on the

permeate side. The partitioning and desorption steps are normally neglected and the

remaining steps are considered as the main contribution to the overall resistance to mass

transfer. The mass transfer resistance in membranes depends on its properties and also on the

chemical and physical properties of the feed components. In the past, the membrane materials

and dimensions were not optimized and, therefore, high membrane mass transfer resistance

was commonly observed. However, new developments in membrane materials and ultra-thin


composite membranes have led to much smaller membrane mass transfer resistances, which

are now mainly due to the diffusive transport at the boundary layer, like noticed in the

removal of volatile organic compounds from wastewaters (Wijmans et al., 1996) and in the

dehydration of cyclohexane (Ortiz et al., 2006). Diffusive transport depends on the

hydrodynamic conditions, physical properties of the fluid and of the solute, and the system

geometry. Several correlations based on Sherwood number for determination of the mass

transfer coefficient for transport in the boundary layer, blk , have been proposed over the years

on different membrane modules configurations (Bandini et al., 1997; Crowder and Cussler,

1998; Dotremont et al., 1994; Gekas and Hallstrom, 1987; Lipnizki and Field, 2001; Lipski

and Côté, 1990; Oliveira et al., 2001; Urtiaga et al., 1999).

Pervaporation is often considered as an isothermal process, but this process always induces a

temperature drop in the feed, since it involves vaporization. The heat transfer in pervaporation

involves: (i) heat transfer from the feed bulk to the feed membrane interface; (ii) heat transfer

through the membrane; and (iii) consumption of heat (vaporization followed by expansion) at

the permeate side of the membrane. In systems operating under laminar conditions, where the

membrane is thin and/or has high permeability, the temperature drop in the feed membrane

interface is considerable (Favre, 2003) and, therefore, the boundary layer heat transfer

resistance should be taken into account. The heat transfer coefficient in the boundary layer is

often determined by Nusselt correlations (Karlsson and Trägardh, 1996).

The ethyl lactate synthesis on pervaporation and vapour-permeation membrane reactors was

already studied by some authors. The two configurations adopted were: (i) batch reactor,

where the esterification reaction takes place, followed by a membrane for water removal, and

reflux of the retentate to the reactor (Benedict et al., 2003; Benedict et al., 2006; Rathin and

Shih-Perng, 1998; Wasewar et al., 2009) and (ii) membrane inside a batch reactor (Jafar et

al., 2002; Tanaka et al., 2002). In fact none of these studies considers a continuous integrated

membrane reactor for ethyl lactate production. Regarding the type of membranes tested,

polymeric membranes were used for pervaporation (Benedict et al., 2003; Benedict et al.,

2006; Rathin and Shih-Perng, 1998) and zeolites for vapour-permeation (Jafar et al., 2002;

Tanaka et al., 2002). In all of these studies, the pervaporation process is superficially

addressed; important membrane parameters, like selectivity and species permeances, at

different temperatures and feed concentrations were not determined. The most detailed

pervaporation study for the ethyl lactate system was performed by Delgado and co-authors for

binary and quaternary mixtures using a commercial polymeric membrane Pervap 2201 (from


Sulzer Chemtech) (2008; 2009, respectively). These studies, focused just on the separation,

still did not consider concentration and temperature polarization, and were not extended to the

reactive system.

Although most of studies involves the use of polymeric membranes that provide good

selectivity and flux, these membranes do not normally support reaction conditions (in terms

of concentrations, temperature, and pH, among others) and, therefore, are not appropriated for

applications under those conditions. Inorganic membranes, mainly those of silica and zeolites,

present better stability under acidic and high temperature conditions, and are the best

alternative to perform dehydration of a reaction medium. Since the commercial microporous

silica membrane (from Pervatech) revealed better selectivity and water flux than the

commercial membrane Pervap SMS (from Sulzer Chemtech) in the dehydration of aqueous

mixtures containing acetone and isopropanol (Casado et al., 2005), Pervatech membrane will

be considered in this work. Moreover, for the dehydration of ethanol, the Pervatech

membrane proved to have high flux and selectivity (Sommer and Melin, 2005).

The synthesis of ethyl lactate in a continuous integrated pervaporation membrane reactor,

using the commercial tubular membrane with higher flux and high selectivity (Pervatech), is

here, for the first time, implemented and evaluated. Furthermore, in order to better

characterize the pervaporation process and aiming to describe the PVMR unit, the effect of

feed pressure, flowrate, temperature and composition on the pervaporation performance is

studied by batch experiments (BP), testing the different binary mixtures involved in the

synthesis of ethyl lactate (ethanol/water, ethyl lactate/water and lactic acid/water).

Additionally, a new mathematical model accounting for the mass transport phenomena under

non-isothermal conditions is developed and applied to better understand and describe the

ethyl lactate production by means of the PVMR.


6.2.1 Materials

The chemicals used were ethanol (>99.9% in water), lactic acid (>85% in water) and ethyl

lactate (>98% in water) from Sigma-Aldrich (U.K.). A commercial strong-acid ion-exchange

resin named Amberlyst 15-wet (A15-wet) (Rohm & Haas) was used as catalyst and


adsorbent. Commercial hydrophilic membrane supplied by Pervatech BV (The Netherlands)

was used. It is resistant for any solvent at any concentration, but sensitive to acidic and

alkaline media (preferred pH range from 2 up to 9) and so, exposure to inorganic acids or

caustic should be avoided. This membrane has a modified silica (methyl silica) selective layer

coated onto gamma alumina. The separation layer is applied inside of an asymmetric ceramic

tube that has an outer diameter of 10 mm, an inner diameter of 7 mm and a length of 50 cm. It

has an effective membrane area per tube of about 110 cm2. Two equal membrane modules

were used to perform the experiments.

6.2.2 Pervaporation Membrane Reactor Unit

The experiments were carried out in a pervaporation membrane reactor unit, which can

operate in batch (pervaporation studies) or in continuous mode (reaction pervaporation

experiments). This unit is equipped with temperature (TI) (type K thermocouple, accuracy of

about +/- 2.2ºC) and pressure sensors (PI) in order to monitor and register these two

parameters. The absolute pressure is measured through 2 analogue dials (accuracy of about

+/- 0.5 bar), filled with glycerine (Nuova Firma), while the pressure in the permeate side is

measured by means of 1 digital dial ceraphant-T PTC31 (Endress+Hausser) (accuracy of

about +/- 1 mbar). A schematic representation of the pervaporation membrane reactor unit is

shown in Figure 6.1. The temperature was controlled by a thermostated bath (Lauda,

Germany) with ethylene glycol/water solution that flows through the jackets of feed vessels 1

and 2; pressure was set at 2 bars by applying an overpressure of helium to the system in order

to prevent vaporization of feed mixture over the whole temperature range.

Pervaporation experiments

The feed was charged into the feed vessel 1 (1 L capacity) and heated to the desired

temperature. In order to heat the whole system (tubing and membranes) to that same

temperature, the feed mixture is after re-circulated over the membrane modules using a

positive displacement diaphragm pump (Hydra Cell G-03, Wanner International), in the

absence of vacuum on the permeate side. When the steady state is reached, the pervaporation

experiment starts by applying vacuum to the permeate side by means of a vacuum pump (Boc

Edwards, U.K.). During the whole run, all vapour permeated was condensed on two parallel

glass cold traps filled with liquid nitrogen. Finally, the collected permeate was defrosted,

weighted and analyzed. The duration of the experiment is conditioned by the trade-off


between ensuring the nearly constant feed composition and enough amount of permeate. To

verify the assumption of constant feed composition samples were collected before and after

each experiment.

Figure 6.1 Set-up of the pervaporation membrane reactor unit.

6.3 Pervaporation Studies

The experimental results of the pervaporation studies for the different binary mixtures

(ethanol/water, ethyl lactate/water and lactic acid/water) involved in the esterification reaction

between lactic acid and ethanol are presented in this section. The effect of the absolute feed

pressure, feed flowrate, operation temperature and water feed molar fraction onto the

membrane performance is evaluated. The membrane permeabilities were determined in

absence of mass transfer limitations in the boundary layer like shown by the preliminary

studies performed.

6.3.1 Pervaporation Transport

The pervaporation performance of the membrane was evaluated in terms of pervaporation

flux and separation factor (membrane selectivity). The process separation factor (α ) is

defined as:

i j

i j

y xx y

α = (6.1)


where ix is the liquid mole fraction of component i on the feed side and iy is the mole fraction

of component i on the permeate side.

The partial flux of one component through the membrane is given by:

,i perm i totJ w J= (6.2)

where totJ it the total permeation flux expressed in kg/(m2.s) and ,perm iw is the mass fraction of

component i on the permeate.

The solution-diffusion model (Wijmans and Baker, 1995) was successfully applied in the

description of solvent dehydration using microporous silica membranes (Ma et al., 2009;

Sommer and Melin, 2005). This model provides the following transport equation for the

permeation molar flux of a component through the membrane:

0,Q ( )i memb i i i i i permJ x p y Pγ= − (6.3)

where ,Qmemb i is the permeance of component i through the membrane (mol/(m2.s.Pa)), which

is equal to /iP ( iP is the permeability coefficient of component i (mol/(m.s.Pa)), and is

the thickness of the selective layer of the membrane), iγ is the activity coefficient (calculated

by the UNIQUAC model using the parameters determined in chapter 3), 0ip is the saturation

pressure of component i, permP is the total pressure on the permeate side and iy is the molar

fraction of component i in the vapor phase. The saturation pressure of pure components was

estimated by the Antoine equation and it is presented in Appendix B.

Some works consider the temperature influence on the total flux to measure an apparent

activation energy (Delgado et al., 2009; Khayet et al., 2008; Slater et al., 2006). However, in

this work, the activation energy of permeation is determined by an Arrhenius-type equation

for the permeance temperature dependence (Feng and Huang, 1997):

,, ,0 exp perm i

memb i memb

EQ Q

RT−⎛ ⎞

= ⎜ ⎟⎝ ⎠

(6.4)


where ,0membQ is the pre-exponential factor, ,perm iE is the activation energy of permeation,

which is a combination of activation energy of diffusion and the heat of adsorption on the

membrane ( , ,perm i D i sE E H= + Δ ), T is the absolute temperature and R is the ideal gas constant.

6.3.2 Preliminary Studies

6.3.2.1 Evaluation of the membrane quality

The driving force for the transport through the membrane is based on the species chemical

potential difference over the membrane and, theoretically, the absolute feed pressure affects

the chemical potentials via the Poynting factor, which might affect either the flux or the

selectivity. However, this influence is often negligible since the Poynting factor is one at low

pressures and, therefore, a way to detect membrane imperfections is by performing

pervaporation experiments at different feed pressures. If the permeate composition and the

total flux remains constant over the studied pressure range the membrane quality can be, in

principle, guaranteed (Pera-Titus et al., 2008).

The influence of the feed pressure onto the pervaporation membrane performance, is shown

in Figure 6.2, where the total permeation flux and permeate composition are represented as a

function of feed pressure, keeping constant the remaining conditions (temperature, flowrate,

water feed concentration, permeate pressure). The variations observed either on the permeate

compositions or on the total flux are within the experimental error, and, therefore, it is

possible to conclude that pervaporation is not affected by the feed pressure. This is

corroborated by other work (De Bruijn et al., 2007) where the effects of feed pressure on a

silica membrane performance were evaluated. Within the pressure range studied, the Poynting

factor is almost one for all species, and therefore these results indicate that the silica

membranes used are free of significant defects.


0.94

0.96

0.98

1.00

0.80

0.82

0.84

0.86

0.88

0.90

0.0 0.5 1.0 1.5 2.0 2.5

Perm

eate

Wat

er m

ole

fract

ion

J tot

(kg/

(h.m

2 ))

Absolute feed pressure (bar)

Total fluxPermeate composition

Figure 6.2 Pressure dependence of total permeation flux and of

permeates composition for the system ethanol/water. ( 321.15 , T K= , 0.183 0.001W Fx = ± , 20 permP mbar= ).

6.3.2.2 Evaluation of mass transfer limitations in the boundary layer

The effect of concentration polarization on the pervaporation process was studied for

water/ethanol mixtures, at constant feed water concentration and operating temperature,

varying the feed flowrate. Analyzing Figure 6.3, it is possible to conclude that for feed

flowrates higher than 19 L/h, the total flux remains constant, indicating absence of mass

transfer resistance from the bulk liquid phase to the feed-membrane interface. Therefore, the

remaining pervaporation experiments were performed at a feed flowrate of 20 L/h.

0.70

0.75

0.80

0.85

0.90

0.95

1.00

10 13 16 19 22 25

J tot

(kg/

(h.m

2 ))

Feed flowrate (L/h)

Figure 6.3 Total permeation flux as a function of feed flowrate ( 321.15 , T K= , 0.182 0.006W Fx = ± , 25 permP mbar= ).


6.3.3 Detailed Studies

6.3.3.1 Water/Ethanol System

The influence of feed composition at different operating temperatures on the total permeation

flux and on the permeate composition is shown in Figure 6.4 for the water/ethanol pair.

0.45

1.45

2.45

3.45

4.45

5.45

6.45

0.0 0.1 0.2 0.3 0.4 0.5 0.6

J tot

(kg/

(h.m

2 ))

Feed water mole fraction

T=321.15 K

T=336.15 K

T=344.15 K

0.95

0.96

0.97

0.98

0.99

1.00

0.0 0.1 0.2 0.3 0.4 0.5 0.6

Perm

eate

wat

er m

ole

facr

tion


T=344.15 K

T=321.15 K

T=336.15 K

Figure 6.4 Pervaporation performance for water/ethanol mixtures. (a) Influence of feed water mole fraction on total permeation flux at different operating temperatures; (b) Influence of feed water mole fraction on permeate composition at different operating temperatures.

As can be seen, the total permeation flux increases with water concentration (linearly) and

with the temperature in the feed (Figure 6.4a). From Figure 6.4b, it is also observed an

increase in water permeate mole fraction with the feed water concentration, but little

influence is noticed with the temperature.

6.3.3.2 Water/Ethyl lactate System

Like before, in the water/ethyl lactate system, it is observed an increase on the total flux with

the increase in the feed water mole fraction, except when the system was kept at 321.15 K,

where this effect of feed water composition on the total flux is not so significant (see Figure

6.5a). The temperature effect on the total flux is once again relevant, increasing with the

operation temperature, for the same feed composition, it is observed an increase around

167 % in the total flux. This is justified by the fact that the permeation driving force increases

with temperature. Figure 6.5b, shows that the water mole fraction in the permeate is almost

independent (varies from 0.996 to 0.999) from the feed water concentration and temperature,

considering the studied range.

(a) (b)


2.15

3.15

4.15

5.15

6.15

7.15

8.15

0.35 0.45 0.55 0.65 0.75

J tot

(kg/

(h.m

2 ))


T=321.15 KT=334.15 KT=343.15 K

0.990

0.992

0.994

0.996

0.998

1.000

0.35 0.45 0.55 0.65 0.75

Perm

eate

wat

er m

ole

fract

ion


T=321.15 KT=334.15 KT=343.15 K

Figure 6.5 Pervaporation performance for water/ethyl lactate mixtures. (a) Influence of feed water mole fraction on total permeation flux at different operating temperatures; (b) Influence of feed water mole fraction on permeate composition at different operating temperatures.

6.3.3.3 Water/Lactic acid System

Two pervaporation experiments were performed for the water/lactic acid mixtures with a feed

water molar fraction of 0.77, at two different temperatures (321.15 K and 343.15 K).

Experimentally, the permeate stream was only composed by water and, therefore, no more

experiments were performed for this mixture.

6.3.3.4 Membrane performance evaluation

For both ethanol/water and ethyl lactate/water systems, the flux increases with feed water

composition and temperature. It was shown that the flux depends strongly on the temperature

like already mentioned for dehydration in similar membranes (Casado et al., 2005; ten Elshof

et al., 2003). In the case of water/lactic acid mixtures, no lactic acid was detected in the

permeate side and no more measurements were made regarding this separation.

In the Figure 6.6 the process separation factor for the pairs water/ethanol and water/ethyl

lactate is plotted as a function of the temperature. It can be seen that the membrane presents

good selectivity towards water in both pairs, but the selectivity to water is higher in mixtures

with ethyl lactate most probably due to the higher molecular size of ethyl lactate compared

with that of ethanol. One indicative measure of the molecular size is given by the gyration

radius and their values presented in Table 6.1 support that hypothesis. However, according to

that, it was expected the permeation of lactic acid through the membrane. Naturally, this is

(a) (b)


not the only factor affecting selectivity; the lower lactic acid vapor pressure compared to all

other species (Table 6.1) significantly reduces the permeation driving force for this

component. Moreover, the affinity between species and membrane should be considered as a

secondary factor (the molecular diameter is the most important). This affinity is directly

related with the dipole moment of each component (Table 6.1): small dipole moment might

complicate the penetration of molecules into the hydrophilic silica layer. Comparing lactic

acid with ethyl lactate, the molecular sizes are similar, but the first has the lowest dipole

moment, which in addition to the lowest vapor pressure, leads to infinite membrane

selectivity towards water in lactic acid mixtures.

Table 6.1 Physical properties of the different species (water, ethanol, ethyl lactate and lactic acid) (Delgado et al., 2008).

Molecule Radius of gyration ( A ) Dipole moment (Debyes)

Water 0.615 1.8497 Ethanol 2.259 1.6908

Ethyl lactate 3.622 2.4000 Lactic acid 3.298 1.1392

0

200

400

600

800

1000

1200

320 325 330 335 340 345 350

Sepa

ratio

n fa

ctor

Temperature (K)

water/ethanol

water/ethyl lactate xW,F = 0.494 ± 0.008

xW,F = 0.435 ± 0.006

Figure 6.6 Influence of the temperature on the separation factor for the binary

mixtures water/ethanol and water/ethyl lactate.


6.3.4 Parameter Estimation

The pervaporation process design requires the knowledge of the permeance of each

component as function of temperature (Equation 6.4). However, there is evidence that the

feed water content affects permeance (Delgado et al., 2009) and therefore, both factors are

needed to be accounted for.

6.3.4.1 Permeance temperature dependence

Experimental permeance for each species can be calculated using the experimental

information found so far in Equation 6.3, simultaneously with the activity coefficients

calculated by the UNIQUAC model. Those values are presented in Figure 6.7.

-3.0

-2.5

-2.0

-1.5

-1.0

-0.5

0.0

2.85 2.90 2.95 3.00 3.05 3.10 3.15

ln Q

mem

b, w

1000/T (K-1)

-9.0

-8.5

-8.0

-7.5

-7.0

-6.5

-6.0

2.85 2.90 2.95 3.00 3.05 3.10 3.15

ln Q

mem

b, E

th

1000/T (K-1)

-9.0

-8.0

-7.0

-6.0

-5.0

-4.0

-3.0

-2.0

2.90 2.95 3.00 3.05 3.10 3.15

ln Q

mem

b, E

L

1000/T (K-1) Figure 6.7 Temperature dependence of species permeances (mol/(min.dm2.bar)) and

linear fittings: (a) water; (b) ethanol; (c) ethyl lactate.

From the slope and intercept of the linear regression, the activation energy of permeation and

the pre-exponential factor are determined for each species, like presented in Table 6.2.

(a) (b)

(c)


Although, the permeation flux increases with temperature due to the increasing driving

forces, the membrane permeability decreases with the temperature like the negative values of

the activation energies of permeation indicates (Table 6.2). As mentioned previously, the

activation energy is a sum of the diffusion activation energy and the heat of adsorption on the

membrane, indicating that the permeation of water, ethanol and ethyl lactate are governed by

the adsorption. Similar results were found for water/ethanol permeation through silica

microporous membranes (Ma et al., 2009).

Table 6.2 Pervaporation parameters of Equation 6.4 and the mean relative deviation (MRD).

Component Qmemb,0 (mol/(min.dm2.bar)) Eperm (kJ/mol) MRD (%) Ethanol 1.41×10-7 -22.60 14.51

Ethyl lactate 1.12×10-4 -10.42 24.13 Water 1.78×10-3 -13.93 13.94

The error introduced considering composition invariant permeances can be calculated from

the mean relative deviation (MRD), defined by Equation 6.5.

exp

, ,exp

exp ,exp

1 100%i calc i

n i

J JMRD

n J

⎛ ⎞−⎜ ⎟= ×⎜ ⎟⎝ ⎠∑ (6.5)

In order to decrease the deviations between theoretical and experimental fluxes, parameter

estimation considering also composition dependence will be addressed in following sub-

section.

6.3.4.2 Permeance temperature and water content dependence

The strategy developed is based on the linearization of equation 6.4 at each composition.

Considering water/ethanol and water/ethyl lactate mixtures, the permeances of water as a

function of temperature, at a constant feed water mole fraction, are shown in Figure 6.8 and

Figure 6.9, respectively. Except for an ethyl lactate aqueous solution with 68 % of water

content, where the activation energy is positive, in all other situations the activation energy of

permeation is negative has previously determined. This might indicate that, for water/ethyl


lactate mixtures, the water permeance at high feed water concentrations is controlled by

diffusion instead of adsorption.

-1.90

-1.70

-1.50

-1.30

-1.10

-0.90

-0.70

2.85 2.90 2.95 3.00 3.05 3.10 3.15

ln Q

mem

b, W

1000/T (K-1)

-1.70

-1.60

-1.50

-1.40

-1.30

-1.20

-1.10

2.85 2.90 2.95 3.00 3.05 3.10 3.15ln

Qm

emb,

W

1000/T (K-1)

-1.50

-1.40

-1.30

-1.20

-1.10

2.90 2.95 3.00 3.05 3.10 3.15

ln Q

mem

b, W

1000/T (K-1)

-1.60

-1.50

-1.40

-1.30

-1.20

2.85 2.90 2.95 3.00 3.05 3.10 3.15

ln Q

mem

b, W

1000/T (K-1) Figure 6.8 Temperature dependence of water permeances (mol/(min.dm2.bar)) and

linear fittings (water/ethanol mixtures): (a) , 0.105 0.001W Fx = ± ; (b) , 0.183 0.003W Fx = ± ; (c) , 0.298 0.006W Fx = ± (d) , 0.492 0.005W Fx = ± .

(a) (b)

(d)(c)


-1.50

-1.40

-1.30

-1.20

-1.10

-1.00

2.90 2.95 3.00 3.05 3.10 3.15

ln Q

mem

b, W

1000/T (K-1)

-1.14

-1.12

-1.10

-1.08

-1.06

-1.04

-1.02

-1.00

2.90 2.95 3.00 3.05 3.10 3.15

ln Q

mem

b, W

1000/T (K-1)

-1.18

-1.16

-1.14

-1.12

-1.10

2.95 3.00 3.05 3.10 3.15

ln Q

mem

b, W

1000/T (K-1) Figure 6.9 Temperature dependence of water permeances (mol/(min.dm2.bar)) and linear

fittings (water/ethyl lactate mixtures): (a) , 0.435 0.006W Fx = ± ; (b) , 0.564 0.015W Fx = ± ; (c) , 0.680 0.003W Fx = ± .

The pre-exponential factors as well as the activation energies of permeation calculated from

Figure 6.8 and Figure 6.9 are presented in Table 6.3. Independently of binary mixture

considered, the representation of those parameters as function of the water feed mole fraction

is shown in Figure 6.10.

Table 6.3 Pre-exponential factor and activation energy for different feed water contents.

Parameters Water/ethanol mixtures Water/ethyl lactate mixtures

Water molar (%) 10.5 18.3 29.8 49.2 43.2 56.4 68.0

Eperm,W (kJ/mol) -30.24 -17.56 -17.52 -7.84 -13.56 -4.31 3.46

Qmemb0,W (mol/min.dm2.bar) 4.62×10-6 4.55×10-4 4.47×10-4 1.47×10-2 2.00×10-3 7.12×10-2 1.13

(a) (b)

(c)


0.0

0.2

0.4

0.6

0.8

1.0

1.2

0.0 0.2 0.4 0.6 0.8

Qm

emb,

0(m

ol/(m

in.d

m2 .

bar))


-35-30-25-20-15-10-505

10

0 0.2 0.4 0.6 0.8

E per

m(k

J/m

ol)

Feed water mole fraction Figure 6.10 Influence of feed water mole fraction on: (a) pre-exponential factor;

(b) activation energy of permeation.

Analysing Figure 6.10, it is possible to conclude that the pre-exponential factor and water

permeance dependencies on feed water mole fraction are well described by exponential and

linear functions, respectively, leading to the following expressions:

6 2,0 1.967 10 exp(18.64 ) (mol/(min.dm . ))memb WQ x bar−= × (6.6)

, 50.377 32.326 (KJ/mol)perm w WE x= − (6.7)

These expressions were combined in a final empirical correlation that describes the

permeance of water as a function of the temperature and water feed content:

6 2,

50377 323261.967 10 exp(18.64 )exp (mol/(min.dm . ))Wperm W W

xQ x barRT

− −⎛ ⎞= × −⎜ ⎟⎝ ⎠

(6.8)

The error between experimental and theoretical permeances calculated by Equation 6.8 is of

10.35 % (MRD), which is about 26 % smaller than the one obtained when just the

temperature influence was considered. Although derived from binary mixtures, expression 6.8

can be applied to estimate the water permeance in quaternary mixtures, since it depends only

on water content no matters the remaining constituents. This is in agreement with

experimental pervaporation results using binary and quaternary mixtures involved in the ethyl

lactate synthesis, since for the same temperature and water feed content, the value of water

permeance is almost the same for binary or quaternary mixtures (Delgado et al., 2009).

Therefore, it can be stated that the water transport is barely affected by the presence of the

other components supporting the validity of Equation 1.8. For ethanol and ethyl lactate, no

(a) (b)


relation was found between their permeances in the membrane and the feed contents. Thus,

the parameters of Table 6.2 were used to determine the permeances of these species at

different operating temperatures and water feed contents.

6.3.4.3 Estimation of the boundary layer mass transfer coefficient (kbl)

In pervaporation processes, in addition to the permeation resistance in the membrane there is

also the resistance in the boundary layer (concentration polarization). This can be

conveniently represented by the resistance in series model, where the overall resistance to

transport is the sum of boundary layer and membrane resistances. The resistance in series

equation for pervaporation was derived by Wijmans and collaborators (1996):

0

,

, ,

1 1 i i mol ib

ov i memb i F

p Vk Q av

γ= + (6.9)

in which ,ov ik is a global membrane mass transfer coefficient 2(mol/(s.m . ))bar , that combines

the resistance due to the diffusive transport in the boundary layer with the membrane

resistance, Vmol,i is the molar volume of component i 3( / )m mol , bbl FK av= where Fv is the

feed liquid velocity ( / )m s . Equation 6.9 will be used to determine the mass transfer

resistance in the boundary layer.

The pervaporation data experimentally obtained for water/ethanol mixtures using different

feed flowrates was used to plot the inverse of the global membrane mass transfer coefficient

(1/ ovk ) of water and ethanol as a function of 1/ bFv , as shown in Figures 6.11 and 6.12,

respectively. The value of parameter b was chosen in order to keep the data points on a

straight line and also minimizing the difference between the inverse of the intercept of the

line and the experimental permeance measured in absence of mass transfer limitations on the

boundary layer (being the best fit for b=0.97), while the parameter a was obtained from the

slope of this line. By this analysis it was possible to obtain the following expressions to

determine the boundary layer mass transfer coefficients for both species (ethanol and water):

0.975, ( / ) 7.27 10 ( / )bl W FK m s v m s− ⎡ ⎤= × ⎣ ⎦ (6.10)

0.977, ( / ) 3.56 10 ( / )bl Eth FK m s v m s− ⎡ ⎤= × ⎣ ⎦ (6.11)


1.4

1.5

1.6

1.7

1.8

1.9

2.0

5 6 7 8 9 10

1/k o

v,W

((m2 .b

ar.s

)/mol

)

1/vFb (m/s)-b

Figure 6.11 ,1/ ov Wk versus 1/ bFv , b=0.97.

600

700

800

900

1000

5 6 7 8 9 10

1/k o

v,Et

h((m

2 .ba

r.s)/m

ol)

1/vFb (m/s)-b Figure 6.12 ,1/ ov Ethk versus 1/ b

Fv , b=0.97.

6.4 Modelling

6.4.1 Batch Pervaporation Model

The mathematical model developed to describe the behaviour of the batch pervaporation

membrane (BPM) considers:

- Plug flow for the bulk fluid phase;

- Total feed volume inside the tank and the retentate velocity variations (inside membrane)

due to permeation of components;

- Concentration polarization, where the resistance due to the diffusive transport in the

boundary layer is combined with the membrane resistance in a global membrane resistance;

- Non-isothermal operation due to heat consumption for species vaporization;

- Temperature polarization.

Following these assumptions the BPM model equations are:

Feed tank mass balance to component i

( ),, ,

f iret ret i f f i

d VCQ C Q C

d t= − (6.12)


where t is the time variable, V is the volume of the feed thank, fQ is the flowrate fed to the

membrane modules and retQ is the flowrate at the end of the membrane modules, fC and retC

are the liquid phase concentration fed to the membrane and at the end of the membrane

modules, respectively.

Feed volume variation

ret fdV Q Qd t

= − (6.13)

Retentate mass balance to component i

( ),, 0ret iret im i

vCCA J

t z∂∂

+ + =∂ ∂

(6.14)

where z is the axial coordinate at the membrane modules, v is the superficial velocity, mA is

the membrane area per unit membrane modules volume and iJ is the permeate molar flux of

species i, through the membrane, defined as:

0,k ( )i ov i i i i i permJ x p y Pγ= − (6.15)

where ,kov i is the global membrane mass transfer coefficient, that combines the resistance due

to the diffusive transport in the boundary layer with the membrane resistance (Wijmans et al.,

1996):

0,

, ,

1 1 i i mol i

ov i memb i bl

p Vk Q K

γ= + (6.16)

in which Vmol,i is the molar volume of component i and blK is the boundary layer mass transfer

coefficient. For laminar flow and Graetz number ( 2 /( )int md v D L ) much greater than one, the

mass transfer coefficient for transport in the boundary layer, blk , is determined by the Lévêque

correlation (Lévêque, 1928):

0.330.33 0.33 int1.62 Re dSh Sc

L⎛ ⎞= ⎜ ⎟⎝ ⎠

(Re 2300)< (6.17)


where int /bl mSh k d D= , int /Re d vρ η= and /( )mSc Dη ρ= are the Sherwood, Reynolds and

Schmidt numbers, respectively, mD is the solute diffusivity in the boundary layer, intd is the

inside diameter of the membrane, L is the membrane length, ρ is the density and η is the

viscosity. The accuracy of the Lévêque correlation to estimate the mass transfer coefficient in

the laminar regime has been validated experimentally by several works (Crowder and

Cussler, 1998; Wickramasinghe et al., 1992). The prediction of the solute diffusivity was

made using the Perkins and Geankoplis method (Perkins and Geankoplis, 1969). Further

details concerning its calculation can be found in Chapter 4.

The molar fraction of component i on the vapor phase (permeate side), iy , is defined as:

1

ii n

ii

JyJ

=

=

∑ (6.18)

Fluid velocity variation in the membrane feed side calculated from the total mass balance

,1

n

m i mol ii

dv A J Vdz =

= − ∑ (6.19)

where n is the total number of components.

Retentate heat balance

( ), ,, ,1 1

0n n

p i p iret i ret i m F mi i

T TC C vC C A h T Tt z= =

∂ ∂+ + − =

∂ ∂∑ ∑ (6.20)

where ,p iC is the liquid heat capacity of component i, T is the absolute temperature in the

feed side of the membrane, mT is the membrane temperature and Fh is the heat transfer

coefficient in the liquid boundary layer.

Membrane heat balance

int int2 2 2 2

1int int int int

( )( / 2) ( / 2)

nV

F m i ii

d dh T T H Jr r r r

δδ δ =

+− = Δ

+ − + − ∑ (6.21)


where intr is the internal radius of the membrane, δ is the membrane thickness and ViHΔ is the

heat of vaporization of species i. The heat transport coefficient was estimated by the Sieder-

Tate correlation, valid for laminar pipe flow (Welty et al., 2008):

0.140.33int1.86 Re Pr b

w

dNuL

ηη⎛ ⎞⎛ ⎞= ⎜ ⎟⎜ ⎟

⎝ ⎠ ⎝ ⎠ (6.22)

where int /FNu h d λ= , Pr /pCη λ= are the Nusselt and Prandtl numbers, respectively, bη and

wη are the viscosity of the liquid in the feed and in the membrane wall, respectively and λ is

the thermal conductivity.

Initial boundary conditions:

, , ,00 : f i ret i it C C C= = = (6.23a)

0V V= (6.23b)

FT T= (6.23c)

0 : Fz T T= = (6.24a)

Fv v= (6.24b)

, ,ret i f iC C= (6.24c)

where subscripts 0 and F refer to initial state and membrane feed conditions, respectively.

6.4.2 Pervaporation Membrane Reactor Model

A mathematical model was also developed to describe the behaviour of the tubular

pervaporation membrane reactor that similarly to the BPM model considers:

- Concentration polarization, where the resistance due to the diffusive transport in the

boundary layer is combined with the membrane resistance in a global membrane resistance;

- Temperature polarization;


Additionally, it also takes in account:

- Axial dispersion flow for the bulk fluid phase;

- External and internal mass transfer for adsorbable species combined in a global particle

resistance;

- Non-isothermal operation due to heat effects during species vaporization and reaction

(slightly endothermic);

- Velocity variations due to permeation of components and adsorption/desorption rates;

- Constant column length and packing porosity.

Following this assumption the PVMR model equations are:

Bulk fluid mass balance to component i:

( ) ( ), ,(1 ) 3ii i m

L i i p i ax T ip

u CC x AK C C D C Jt z r z z

εε ε

∂ ⎛ ⎞∂ ∂− ∂+ + − = −⎜ ⎟∂ ∂ ∂ ∂⎝ ⎠

(6.25)

where iLK , is the global mass transfer coefficient of the component i, ε is the bed porosity,

axD , and u are the axial dispersion coefficient and the interstitial velocity, respectively, pr is

the particle radius, pC is the average particle concentration and all the other variables were

already defined. The determination of the global mass transfer coefficient ( LK ) is presented

in detail in Chapter 4 and the axial dispersion coefficient ( axD ) is estimated from the

empirical correlation (Butt, 1980) valid for liquids in packed beds:

0.480.2 0.011RePeε = + (6.26)

in which int / axPe d u D= and int /Re d uρ η= are the Peclet and Reynolds numbers,

respectively.

The permeate flux of membrane is defined by Equation 6.15 (BPM model).

Interstitial fluid velocity variation calculated from the total mass balance

( ) ( ),, ,1 1

1 3 n nm

p iL i mol i i ii ip

Adu K V C C Jdz r

εε ε= =

−= − − −∑ ∑ (6.27)


where Vmol,i is the molar volume of component i and n is the total number of components.

Pellet mass balance to component i

( ) ( )ipbii

pip

pipiiLp

Crt

qt

CCCK

r ,,

,, 1)1(3

ερν

εε−

−∂∂

−+∂

∂=− (6.28)

where iν is the stoichiometric coefficient of component i , bρ is the bulk density, pε the

particle porosity, iq is the average adsorbed phase concentration of species i in equilibrium

with ipC , , and r is the kinetic rate of the chemical reaction relative to the average particle

concentrations in the fluid phase. The reaction rate and adsorption isotherms are those

determined in Chapters 3 and 4, respectively.

Retentate heat balance

( ), ,1 1

0n n

m bp i p ii i F m r

i i

AT TC C uC C h T T H rt z

ρε ε= =

∂ ∂+ + − +Δ =

∂ ∂∑ ∑ (6.29)

where rHΔ is the reaction enthalpy, determined in Chapter 3. It must be considered that the

heat of adsorption was neglected since while some species are being adsorbed, others are

being desorbed and there is compensation on the heat released or absorbed, respectively.

The equation of the membrane heat balance is the one represented in the BPM model

(Equation 6.21).

Initial and Danckwerts boundary conditions

0=t : , ,0p ii iC C C= = , ,0i iq q= and FT T= (6.30)

0=z : ,0

ii ax i F

z

CuC D uCz =

∂− =

∂ (6.31a)

0u u= (6.31b)

FT T= (6.31c)

cLz = : 0i

z Lc

Cz =

∂=

∂ (6.31d)



The viscosities, heat capacities, thermal conductivities and heats of vaporization of pure

components were calculated for each temperature by the expressions presented in Appendix

B. The mixture heat capacity was estimated assuming linear mole fraction averages, while the

mixture density and thermal conductivity were estimated assuming linear mass fraction

averages. The methods used to predict binary and multi-component mixture viscosities were

presented in detail in Chapter 4. Details on numerical solution are the same presented in

Chapter 5.


6.5.1 Batch Pervaporation

Within the mathematical model proposed for the batch pervaporation, the mass transfer

coefficient in the boundary layer is, for all the species, estimated through the Lévêque

correlation. However, its accuracy should be validated experimentally. In Section 6.3.4.3, the

boundary layer mass transfer coefficients, for water and ethanol, were determined from the

experimental data (Equations 6.10 and 6.11) and will be applied in the BPM model. The

water permeation fluxes were calculated either by using the experimental mass transfer

coefficient in the boundary layer (Equations 6.10 and 6.11) or using the Lévêque correlation

(Equation 6.17). As can be observed in Figure 6.13, there is a good agreement between the

water permeation flux estimated using the experimental mass transfer coefficients with the

one using the Lévêque correlation, proving its accuracy in the prediction of the boundary

layer mass transfer coefficients. Besides, small deviations are observed between the

experimental fluxes and the ones determined by the BPM model.

In order to validate the BPM non-isothermal model, the dehydration of ethanol and ethyl

lactate aqueous solutions was studied. The evolution of water mole fraction in the retentate

stream is in agreement with the experimental data: for both ethanol and ethyl lactate cases the

retentate composition after 15 and 10 minutes, respectively, is similar to the theoretical

values, as shown in Figure 6.14. In terms of temperature drop, experimentally it was not

observed any temperature variation for ethanol dehydration, while it has dropped 2ºC for the

ethyl lactate experiment. According to the simulation it would be expected a decrease of 2

and 4ºC for ethanol and ethyl lactate dehydrations, respectively. This deviation might be due

to experimental error of the thermocouple used (accuracy of 2.2ºC).


0.0

0.5

1.0

1.5

2.0

2.5

0.0 0.5 1.0 1.5 2.0 2.5

J W(m

ol/(h

.dm

2 )

JW,exp (mol/(h.dm2)

Estimated mass transfer coefficient (Lévêque correlation)

Experimental mass transfer coefficient

Figure 6.13 Experimental water permeation flux as a function of water permeation flux

determined by the BPM model with the Kbl estimated (Lévêque correlation) and determined from experimental data.

334

335

336

337

0.175

0.177

0.179

0.181

0.183

0.185

0.187

0.189

0 2 4 6 8 10 12 14

Tem

pera

ture

(K)

Wat

er m

olar

frac

tion

Time (min)

Feed (Experimental)Retentate (Experimental)Retentate (BPM model)Temperature (BPM model)

338

339

340

341

342

343

344

345

0.535

0.539

0.543

0.547

0.551

0.555

0.559

0 2 4 6 8 10

Tem

pera

ture

(K)

Wat

er m

olar

frac

tion

Time (min)

Feed (Experimental)Retentate (Experimental)Retentate (BPM model)Temperature (BPM model)

Figure 6.14 Evolution of water composition on retentate and temperature predicted by

the non-isothermal BPM model: a) dehydration of 81 % of ethanol in aqueous solution at 336 K; b) dehydration of 44 % of ethyl lactate in aqueous solution at 344 K.

6.5.2 Pervaporation Membrane Reactor

The performance of the pervaporation membrane reactor unit, packed with A15-wet resin in

the lumen side of the tubular membranes (where the active layer of the silica membranes is

placed) was evaluated by simulation. It was considered that the resin was initially saturated

with ethanol and then it was fed with a mixture of ethanol and 85 wt. % lactic acid aqueous


solution. The permeation fluxes were estimated using the expressions previously obtained in

section 6.3.3 and a summary of the parameters used in the simulations is presented in Table

6.4.

Table 6.4 Parameters used in the simulations.

Parameters

Feed temperature 323.15 K Feed composition , , W,F0.3546; 0.3424; x =0.3030Eth F LA Fx x= =

Feed flowrate 1.0 mL/min Permeate pressure 10 mbar Bed porosity 0.424 Bulk density 374.0 g/L Length of the bed 100 cm Internal diameter 0.7 cm

The PVMR unit was evaluated considering isothermal and non-isothermal operation and its

performance was compared with the one of a fixed bed reactor (FBR) in the same operational

conditions. The concentration histories at the end of the PVMR and FBR, considering

isothermal operation, are shown in Figure 6.15. It can be observed that the enhancement

introduced by the water removal through the membranes is significant for the ethyl lactate

production, where a 69 % lactic acid conversion was achieved in the PVMR unit, at the

steady-state, while in the FBR the conversion obtained was just 29 %.


02468

1012141618

0 20 40 60 80 100

C (m

ol/L

)

Time (min)

EthanolLactic acidEthyl lactateWater

02468

1012141618

0 20 40 60 80 100

C (m

ol/L

)

Time (min)


Figure 6.15 Concentration histories at the reactor outlet, considering isothermal operation: a) PVMR; b) FBR.

02468

1012141618

0 20 40 60 80 100

C (m

ol/L

)

Time (min)


02468

1012141618

0 20 40 60 80 100

C (m

ol/L

)

Time (min)


Figure 6.16 Concentration histories at the reactor outlet, considering non-isothermal

operation: a) PVMR; b) FBR.

Considering, alternatively, a non-isothermal operation (see Figure 6.16b), the lactic acid

conversion at steady state of the FBR is about 29 %, much better than 11 % found for the

PVMR (Figure 6.16a). This low performance of the PVMR is due to the temperature drop

noticed in the bulk, around 36ºC, while only 2ºC was noticed for the FBR (Figure 6.17). This

slightly decrease in the FBR is due to the heat needed to the reaction only (endothermic

reaction), while for the PVMR there is a large heat consumption to vaporize the species.

Since the flowrate is very small, the heat capacitance provided by the liquid stream is low and

therefore its temperature is drastically decreased, leading to lower reaction kinetic rates,

higher mass transfer resistances and lower water permeation fluxes through the membranes.


280

290

300

310

320

330

0 20 40 60 80 100 120 140

T (K

)

Time (min)

FBRPVMR

Figure 6.17 Temperature histories at the outlet of the FBR and of the PVMR.

Clearly, to operate the PVMR for efficient ethyl lactate production, it is necessary to provide

the heat required for vaporization and reaction, operating in isothermal conditions. As stated

before, the silica membranes used in this work have the selective layer coated inside the

membrane tube. Therefore, the most suitable way to supply heat to the retentate liquid is

introducing a heated sweep gas in the permeate side instead of using vacuum. However, when

the membrane selective layer is coated on the external side of the membrane tube (shell side),

the heat can be supplied through an appropriate heated solution re-circulated through jacketed

modules; in this situation, the heat is rapidly transferred to the retentate liquid stream.

Assuming that the PVMR unit could operate isothermally replacing vacuum by a heated

sweep gas, it is possible to determine by simulation the number of membranes connected in

series needed to maximize the lactic acid conversion of the same feed processed before. For

an operating temperature of 50ºC (same conditions from Table 6.4), it is possible to achieve

93 % of lactic acid conversion and 84 % of ethyl lactate purity when using a 250 cm long

membrane reactor (5 tubular membranes, effective area of about 550 cm2), as it can be seen in

Figure 6.18.


02468

1012141618

0 40 80 120 160 200 240

C (m

ol/L

)

Time (min)


Figure 6.18 Concentration histories at the outlet of the PVMR for a bed

length of 250 cm at 323.15 K.

Trying to further enhance the separation performance, as well as the reaction rate, the PVMR

temperature was increased to 70ºC. In this case, it was obtained a lactic acid conversion of

98 % and an ethyl lactate purity of 96 %, as shown in Figure 6.19. From the analysis of the

internal concentration profiles on the retentate side, shown in Figure 6.20, it is concluded that

using only two membranes it is possible to get about 90 % of lactic acid conversion, but only

82 % of ethyl lactate purity; being necessary 3 more membranes to maximize both lactic acid

conversion and ethyl lactate purity.

02468

1012141618

0 40 80 120 160 200 240

C (m

ol/L

)

Time (min)


Figure 6.19 Concentration histories at the outlet of the PVMR at 343.15 K

(L=250 cm).


0

2

4

6

8

10

0 0.2 0.4 0.6 0.8 1

C (m

ol/L

)

Z


Figure 6.20 Concentration profiles at steady state at 343.15 K.

These results demonstrate that the PVMR has a great potential for the ethyl lactate

production, where high conversions and purities can be obtained, when this reactor is

operated under isothermal conditions.

6.6 Conclusions

Commercial hydrophilic silica membrane from Pervatech BV (The Netherlands) was

evaluated for the dehydration of ethanol, ethyl lactate and lactic acid, in the temperature range

varied from 48ºC to 72ºC. First pervaporation studies indicate that the membranes have no

major imperfections since the total flux and water selectivity is barely affect by absolute feed

pressure. The influence of hydrodynamic conditions on the membrane polarization was

analyzed, and for velocity values higher than 0.14 m/s polarization effects are eliminated. The

effect of feed temperature and composition on the pervaporation performance was evaluated

by batch experiments. It was concluded that the microporous silica membranes have high flux

and high selectivity for water, while ethanol and ethyl lactate permeation is reduced and lactic

acid does not permeate at all. In summary, the permeances for all species through the

microporous silica membranes, as a function of temperature and feed water content, are

described by the following equations: 3

7,

22.60 101.41 10 expmemb EthQRT

− ⎛ ⎞×= × ⎜ ⎟

⎝ ⎠2/( min)mol dm bar


, 0memb LAQ = 2/( min)mol dm bar

34

,10.42 101.12 10 expmemb ELQ

RT− ⎛ ⎞×

= × ⎜ ⎟⎝ ⎠

2/( min)mol dm bar

6,

50377 323261.967 10 exp(18.64 )exp Wmemb W W

xQ xRT

− −⎛ ⎞= × −⎜ ⎟⎝ ⎠

2/( min)mol dm bar

A mathematical model was developed for the batch pervaporation membrane, considering (i)

plug flow for the bulk fluid phase; (ii) total feed volume inside the tank and retentate velocity

inside the membrane variations due to permeation of components; (iii) concentration

polarization, where the resistance due to the diffusive transport in the boundary layer is

combined with the membrane resistance in a global membrane resistance; (iv) non-isothermal

operation due to heat consumption for species vaporization; and (v) temperature polarization.

The batch pervaporation membrane model was validated experimentally and, therefore, it was

extended to the integrated pervaporation membrane reactor packed with the catalyst

Amberlyst-15wet. This model was used to evaluate the performance of the PVMR in both

isothermal and non-isothermal conditions. It was concluded that non-isothermal operation

worsens the performance of the PVMR, being even worse than that obtained in the fixed bed

reactor. In isothermal conditions, the PVMR is a very attractive solution leading to a lactic

acid conversion of 90 % and ethyl lactate purity of 82 %, for the PVMR set-up designed in

this work (100 cm of length, 2 membrane modules) operating at 70ºC. For near lactic acid

depletion (98 % conversion), it is produced ethyl lactate with 96 % purity if the membrane

modules has 250 cm (5 membranes in series), as indicated by the model predictions.

6.7 Notation

mA membrane area per unit membrane modules volume (m2/m3)

C liquid phase concentration (mol/m3)

fC liquid phase concentration in the thank (mol/m3)

pC average particle concentration (mol/m3particle)

pC liquid heat capacity (J/(mol.K))

retC liquid phase concentration in the retentate (membrane feed side) (mol/m3)


axD axial dispersion coefficient (m2/s)

intd internal diameter of the membrane (m)

mD solute diffusivity in the boundary layer (m2/s)

DE activation energy of diffusion (J/mol)

permE activation energy of permeation (J/mol)

Fh heat transfer coefficient in the liquid boundary layer (W/K)

sHΔ heat of adsorption (J/mol)

VHΔ heat of vaporization (J/mol)

rHΔ enthalpy of reaction (J/mol)

J permeate flux (mol/(m2s))

totJ total permeate flux (mol/(m2s))

blK boundary layer mass transfer coefficient (m/s)


ovk global membrane mass transfer coefficient (mol/(m2sPa))

L column length (m)


Nu Nusselt

P permeability coefficient (mol/(m.s.Pa)) 0p saturation pressure (s)

permP total pressure on the permeate side (s)

Pr Prandtl


particle (mol/m3)

Qmemb permeance (mol/(m2sPa))

,0membQ pre-exponential factor (mol/(m2sPa))

fQ flowrate fed to the membrane modules (m3/s)

retQ flowrate at the exit of the membrane modules (m3/s)

r rate of reaction (mol kg-1 s-1)

R gas constant (J/(mol.K))


intr internal radius of the membrane (m)


Re Reynolds

Sc Schmidt

Sh Sherwood

t time variable (s)

T temperature in the feed side of the membrane (K)

mT membrane temperature (K)

u interstitial velocity (m/s)

V volume of the feed thank (m3)

v superficial velocity (m/s)

molV molar volume (m3/mol)

w mass fraction

ix liquid molar fraction of component i in the feed side

iy molar fraction in the vapor phase of component i

z axial coordinate at the membrane modules (m)

Greek letters

iγ activity coefficient

ε bulk porosity



bρ bulk density (kg/m3bed)

ρ fluid density (kg/m3)

μ viscosity (Pa.s)

δ membrane thickness (m)

λ thermal conductivity (W/(m.K))

thickness of the selective layer of the membrane (m)

α process separation factor

Subscripts





Perm relative to permeate




W relative to water

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7. PermSMBR – A New Hybrid Technology

Abstract. In this Chapter, a new technology, the Simulated Moving Bed Membrane Reactor

(PermSMBR), was presented and applied for the esterification of lactic acid with ethanol. Its

conception was a result of ethyl lactate process re-intensification by integrating a permeable

membrane reactor with a simulated moving bed reactor, since the products separation by

selective adsorption is enhanced by continuous water removal through the membrane, shifting

the equilibrium till high lactic acid conversion with high ethyl lactate purity. It was

demonstrated that the PermSMBR technology applied for the ethyl lactate synthesis can

reduce the desorbent consumption in 62 % and increase the ethyl lactate productivity in 33 %

when compared with the SMBR and in 98 % when compared with reactive distillation for the

same purity and conversion requirements.

Adapted from: V. M. T. M. Silva, Pereira C. S. M. and A. E. Rodrigues, "Reactor de membranas adsorptivo de

leito móvel simulado, novo processo híbrido de separação e respectivas utilizações”, (PT 104496, patent

pending, 2009)

172 CHAPTER 7. PermSMBR – A New Hybrid Technology

7.1 Introduction

The paradigm of chemical engineering process is changing. Traditional processes (where the

reactor is followed by separation units in order to recover the desirable product, to remove the

by-product and to recycle the unconverted reactants to the reactor) are being replaced by

integrated processes where reaction and separation occur in the same device. These integrated

processes are of considerable interest, mainly for equilibrium-limited reactions where the

continuous removal of at least one reaction product shifts the equilibrium in order to increase

conversion and reduce by-product formation. The term multifunctional reactor is often used

to embrace reactive separations technology, which main advantages are higher yields,

reduction of energy requirements, decrease of solvents consumption and lower capital

investments. Typical examples of equilibrium-limited reactions include:

Esterification: R´-COOH + HO-R = RĆOOR + H2O

Acetalization: R´-CHO + 2 HO-R = RĆH-(OR)2 + H2O

Ketalization: R´R´ĆO + 2 HO-R = R´R´Ć(OR)2 + H2O

In the state of the art, the custom multifunctional reactors used for that type of reactions are:

reactive distillations, reactive extractions, membrane reactors and chromatographic reactors.

Regarding to reactive distillation (RD) the best example is the methyl acetate synthesis

developed and patented by the company Eastman Kodak Company (Agreda and Partin,

1984). The entire process is carried out in a single column and represents one-fifth of the

capital investment and consumes one-fifth of the energy of the traditional process (reaction

followed by separation by distillation) (Krishna, 2002). However, there are some

disadvantages in the use of RD for systems that exhibit azeotropes formation and/or the

boiling points of the products are similar. Membrane reactors are widely used and typical

examples are pervaporation and permeation reactors, where the catalyst is in fluidized

(Alonso et al., 2001; Lee et al., 2006) or fixed bed (Lafarga and Varma, 2000; Zhu et al.,

1996). Several times, processes where the reactor and membrane are housed in separate units

in series or parallel are also regarded as membrane reactors (Datta and Tsai, 1998; Tsotsis et

al., 2007). Chromatographic reactors include fixed bed (FBR) (Pereira et al., 2009a; Silva

and Rodrigues, 2002), pressure swing adsorption (PSAR) and simulated moving bed reactors

(SMBR) (Kawase et al., 1996; Pereira et al., 2009b; Silva and Rodrigues, 2005). However,

from all the mentioned chromatographic reactors, the most common for process

intensification for the production of oxygenated products is the SMBR. The SMBR is well-


known equipment (Broughton and Gerhold, 1961) that consists of a set of interconnected

columns packed with an acid solid or a mixture of solids (catalysts and adsorbents). The

description of an SMBR unit is presented in Chapter 5. The SMBR has several advantages, as

the ones already mentioned for the reactive separations; however, it also has some

disadvantages, as for example, the difficulty in removing the more adsorbed species, which

implies high desorbent consumptions. Besides, in several applications, the feed is a mixture

of reactants with one of the products, which will influence the performance of the unit

resulting in low reactants conversion and products purities, low unit productivity and high

eluent/desorbent consumption. This is even worse for the case where the product in the feed

is the more retained one, since then the desorbent consumption will increase significantly. In

order to overcome these issues, it was developed a novel technology, the Simulated Moving

Bed Membrane Reactor (SMBMembR or PermSMBR), which comprises a reactor with two

different separation techniques (chromatography – Simulated Moving Bed (SMB) with a

selective permeable membrane – Pervaporation or Permeation) into a single device. It is a

clean and economic alternative to conventional processes and even competitive when

compared with the intensified processes. If reactive separations embrace the concept of

Process Intensification, the reactive hybrid separations (as PermSMBR) embody the Process

“ReIntensification” (Table 7.1).

Table 7.1 Chemical engineering process evolution.

Traditional Process Reactor + Separator

Distillation Adsorption Crystallization Membranes Extraction

Process Intensification Reactive Separations

• Reactive Distillation • Membrane Reactor • SMB Reactor

Process ReIntensification Reactive Hybrid Separations

• SMB Membrane Reactor

The PermSMBR technology is suitable mainly for oxygenated compounds as esters, acetals

and ethers, used as biofuels, solvents, flavours, among others. In this Chapter, the

PermSMBR will be exploited for the ethyl lactate synthesis. This new reactor leads to near

depletion of lactic acid and high ethyl lactate purity with higher concentration, and,

consequently, lower down streaming costs associated to the separation units.


7.2 Technical description of the PermSMBR technology

The PermSMBR consists of a set of columns with membranes connected in series and packed

with a solid, which could be a mixture of catalyst and selective adsorbent or a solid that acts

both as catalyst and as adsorbent. Typically, there are two inlets, feed and desorbent, and

three outlets, extract, raffinate and permeate. In Figure 7.1, there is schematic representation

of a PermSMBR unit, where a reaction of type A+B ↔ C+D is considered.

Figure 7.1 Schematic diagram of a PermSMBR unit with 4 sections and three columns per section.

In this case, the component A is used as reactant and desorbent, therefore it is introduced in

the system in the feed and desorbent streams. The other reactant B is used as feed. The

products formed (C and D) are removed from the PermSMBR in three different streams:

raffinate that is rich in the less strongly adsorbed product, C, extract, rich in the more strongly

adsorbed product, D, and total permeate that combines all the permeate streams and that

comprises mainly product D, for which the membranes are selective. All the inlet/outlet

streams, with exception of the permeate streams, are introduced/removed from the system

through ports positioned between the columns, and at regular time intervals, called the

switching time, this streams are switched for one column distance in direction of the fluid

flow. In this way, the countercurrent motion of the solid is simulated and its velocity is equal

to the length of a column divided by the switching time. A cycle is completed when the

number of switches is equal to a multiple of the total number of columns. The PermSMBR is

Liquid direction and ports switch A+B ⇔ C+D

Feed (A+B)

Raffinate (A+C)

Extract (A+D)

Desorbent (A)

Desorbent Extract

Raffinate Feed

Liquid direction and ports switch A+B ⇔ C+D

Feed (A+B)

Raffinate (A+C)

Extract (A+D)

Desorbent (A)

Desorbent Extract

Raffinate Feed


equipped with a rotary valve or a number of valves arranged in manner such that any feed

stream may be introduced to any column and any outlet stream may be withdrawn from any

column; it is also equipped with vacuum (or sweep gas flow) in order to withdraw the

permeate stream from each column. The vacuum (or sweep gas flow) can be turned on in all

the columns or just in some. Similarly to the SMBR, the position of the inlet/outlet streams

defines different sections existing in the PermSMBR unit, each one accomplishing a certain

function and containing a variable number of columns. The typical PermSMBR contains 4

sections, as represented in Figure 7.1 and Figure 7.2, where the section 1 is comprised

between the desorbent and extract nodes, the section 2 is comprised between the extract and

feed node; the section 3 is comprised between the feed and raffinate node, and the section 4 is

comprised between the raffinate and desorbent node. In section 1, the adsorbent is

regenerated by desorption of the more strongly adsorbed product (D) from the solid using the

desorbent (A); In sections 2 and 3, reactive sections, the products C and D are separated as

they are being formed. These products are continuously removed from the unit by adsorption

and also through the selective membranes and, therefore, the reaction will proceed beyond the

thermodynamic equilibrium, being possible to obtain 100 % of reactants conversion. In

section 4, before being recycled to section 1, the desorbent is regenerated by adsorption of the

less adsorbed product (C).

Section 1 Section 2 Section 3

XD F

PPP

Section 4

R

P

Figure 7.2 Schematic diagram of a PermSMBR unit with 4 sections: 2 inlet ports for

feed (F) and desorbent (D) streams; 2 outlet ports for extract (X) and raffinate (R) streams; and outlet permeate streams (P).

The PermSMBR can have different configurations depending on the number of streams

fed/removed from the unit. The total number of streams, with exception of the permeate

streams, corresponds to the total number of sections. For example, the PermSMBR unit can

be simplified to a unit of three sections: eliminating the extract stream, when the membrane is

selective to the more adsorbed product (Figure 7.3); or eliminating the raffinate stream, when

the membrane is selective to the less retained product (Figure 7.4).



feed (F) and desorbent (D) streams and 1 outlet port for raffinate (R) stream.


feed (F) and desorbent (D) streams and 1 outlet port for Extract (X) stream.

If necessary, the PermSMBR unit can also be more complex, having five or more sections.

For example, the schematic diagram of the process represented in Figure 7.5 is similar to the

one described for the Figure 7.2, but it has 5 sections since an additional feed stream (F2) is

introduced to the system. In this case, the regeneration of the solid and the desorbent is

accomplished in sections 1 and 5, respectively; and the complete conversion of reactants and

separation of the formed products occurs in section 2, 3 and 4. The feed F2 can comprise the

same reactants as the ones in feed F1, but in different proportions, or other reactants in order

to obtain the desired product. In the case where 3 products are formed, it will be necessary to

have other extra stream designed as R2 in Figure 7.6. Additionally, if it is observed a decrease

in the PermSMBR unit performance (decrease of products purity, reduction of reactants

conversion and/or loss of productivity, among others) due to problems related to

deactivation/poisoning of the catalyst/adsorbent and/or of the membranes, the PermSMBR

unit can be operated in order to correct those problems, making a bypass to each one of the

columns during a cycle to perform the necessary treatments (treatments with acids, solvents,

replacement of the solid and/or of the membrane, thermal treatments, …). For example, if the

desorbent is more expensive than the more adsorbed product, the treatment should by applied

to the last column of section I (Figure 7.7), since this column is still saturated with the most


adsorbed product, avoiding the unnecessary consumption of desorbent if other column of the

same section was selected. This procedure can also be performed in other sections in the most

appropriate column. Other ways of correction along time can be performed to the operational

variables of the PermSMBR unit, similarly to the procedure described in literature for the

SMB unit (Sá Gomes et al., 2007).

Section 1 Section 2 Section 3 Section 4

XD

Section 5

F1

P PPPP

RF2 Figure 7.5 Schematic diagram of a PermSMBR unit with 5 sections: 3 inlet ports for 2

feed streams (F1 and F2) and a desorbent stream (D), and 2 outlet ports for extract (X) and raffinate (R) streams.

Figure 7.6 Schematic diagram of a PermSMBR unit with 6 sections: 3 inlet ports for 2

feed streams (F1 and F2) and a desorbent stream (D), and 3 outlet ports for extract (X) and raffinate (R1 and R2) streams.

Figure 7.7 Schematic representation of the bypass to perform the

regeneration/activation of catalyst/adsorbent and/or of the membrane in section 1 of the PermSMBR unit.


7.3 PermSMBR mathematical model

Since the PermSMBR results from the integration of pervaporation membrane reactor and

simulated moving bed reactor, its model will combine the specificities of each model

previously described on Chapters 6 and 5, respectively; namely:

- axial dispersion flow for the bulk fluid phase;

- linear driving force (LDF) approximation for the inter and intra-particle mass transfer

rates;

- multi-component adsorption equilibrium;

- velocity variations due to adsorption/desorption rates and species permeation;

- constant porosity and length of the packed bed;

- membrane concentration polarization;

- isothermal operation.

In Figure 7.8 a schematic representation of the fluxes inside one membrane of the

PermSMBR is shown, where F is the molar flux in the feed side (retentate) and J is the

permeate molar flux through the membrane.

membrane

bPermeate

F|z

F|z+dz

z

z+dzJ|zJ|z

Feed Flow

Permeate Flow

Figure 7.8 Schematic representation of the fluxes inside one membrane of

the PermSMBR.


Following this assumption the PermSMBR model equations are:

Bulk fluid mass balance to component i in column k

( ) ( )2

,, , 2

1( ) 3ik ik k ik mp ikL ik ik ax k ik

p

C C u C AK C C D Jt z r z

εε ε−∂ ∂ ∂

+ + − = −∂ ∂ ∂

(7.1)

where ikC and ,p ikC are the bulk and average particle concentrations in the fluid phase of

species i in column k respectively, ikLK , is the global mass transfer coefficient of the

component i, ε is the bulk porosity, t is the time variable, z is the axial coordinate, kaxD , , and

ku are the axial dispersion coefficient and the interstitial velocity in column k, respectively,

pr is the particle radius, mA is the membrane area per unit reactor volume and ikJ is the

permeate flux of species i in column k.

The permeate flux ( iJ ) is defined as:

0ov,k ( )i i i i i permJ a p y P= − (7.2)

where ,ov ik is a mass transfer coefficient based on a partial vapour pressure driving force (see

Chapter 6), ia is the activity of component i in bulk, 0ip is the saturation pressure of

component i, permP is the total pressure on the permeate side and iy is the molar fraction of

component i in the vapour phase (permeate side) defined as:

1

ii n

ii

JyJ

=

=

∑ (7.3)

The determination of the global mass transfer coefficient ( LK ) is presented in detail in

Chapter 4 and the calculation of the axial dispersion coefficient ( axD ) is presented in previous

Chapter (Chapter 6).

Interstitial fluid velocity variation calculated from the total mass balance

( ) ( ),, ,1 1

1 3 n nk m

p ikL ik mol i ik iKi ip

du AK V C C Jdz r

εε ε= =

−= − − −∑ ∑ (7.4)


where Vmol,i is the molar volume of component i (60.87 mL/mol, 77.56 mL/mol,

118.44 mL/mol and 18.63 mL/mol at 50 ºC for ethanol, lactic acid, ethyl lactate and water,

respectively) and n is the total number of components.

Pellet mass balance to component i, in column k

( ) ( ),, ,,

31 ( )p ik ikp ik p ikp p L ik ik i p

p

qC K C C r Ct t r

ε ε υ ρ∂∂+ − = − +

∂ ∂ (7.5)

where ikq is the average adsorbed phase concentration of species i in column k in equilibrium

with ,p ikC , pε the particle porosity, iυ the stoichiometric coefficient of component i, pρ the

particle density and r is the chemical reaction rate relative to the average particle

concentrations in the fluid phase. The reaction rate and adsorption isotherms are those

determined in Chapters 3 and 4, respectively.

Initial and Danckwerts boundary conditions

0=t : , ,0p ikik ikC C C= = and ,0ik ikq q= (7.6)

0=z : , ,0

ikk ik ax k k ik F

z

Cu C D u Cz =

∂− =

∂ (7.7a)

,0k ku u= (7.7b)

cLz = : 0=∂∂

=Lcz

ik

zC (7.7c)


Mass balances at the nodes of the inlet and outlet lines of the PermSMBR:

Desorbent node: 1( 4, ) ( 1, 0)

4 4

DDi j z Lc i j z i

u uC C Cu u= = = == − (7.8a)

Extract (j=2) and Raffinate (j=4) nodes: )0,(),1( ==− = zjiLczji CC (7.8b)

Feed node: Fi

FziLczi C

uuC

uuC

2)0,3(

2

3),2( −= == (7.8c)

where,


Dsuuu += 41 Desorbent (Ds) node ; (7.9a)

Xuuu −= 12 Extract (X) node ; (7.9b)

Fuuu += 23 Feed (F) node ; (7.9c)

Ruuu −= 34 Raffinate (R) node ; (7.9d)

The ratio between the fluid interstitial velocity, uj, and the simulated solid velocity, Us,

(defined by the column length and switching time relation,*t

LUs = ) could be defined for

each section giving a new parameter:

Usu j

j =γ (7.9e)

The PermSMBR process performance and details on numerical solution are the same

presented in Chapter 5.

7.4 PermSMBR geometrical specifications

The PermSMBR unit considered consists in 12 columns where each column has 13

commercial hydrophilic tubular membranes (Pervatech BV) in order to dehydrate the reaction

medium. Regarding to the position of the membrane separation layer, it is possible to have

different configurations. In the case of the silica membranes from Pervatech, the selective

layer is inside the tube and consequently, it was considered that these membranes were

packed with the resin Amberlyst 15-wet in the lumen side (inside the membrane tube). The

PermSMBR parameters were chosen in order to compare its performance to the one of

SMBR, since this technology will be the most competitive regarding to the PermSMBR. The

same mass of catalyst and effective area was considered. The porosity of the SMBR unit was

determined experimentally and presented in Chapter 5, while the PermSMBR porosity was

estimated in agreement with the experimental results in fixed bed columns with ratio between

tube diameter and particle diameter of 10 (Theuerkauf et al., 2006). The characteristics of the

columns are presented in Table 7.2.


Table 7.2 Characteristics of the columns for both SMBR and PermSMBR.

SMBR PermSMBR

Solid weight (A15) 47.6 g 47.6 g

Length of the bed (L) 23 cm 25.45 cm

Internal diameter (Di) 2.6 cm 0.7 cm

Bed porosity (ε ) 0.4 0.424

Bulk density (ρb) 390 kg/m3 374 kg/m3

7.5 Simulated Results

The ethyl lactate synthesis using the SMBR was previously studied (Chapter 5) and the best

performance obtained for that unit was under the following operation conditions: a feed of

lactic acid solution (85 wt. % in water), a desorbent of ethanol (99.5 wt. % in water), a

configuration of 3-3-4-2, a working temperature of 50ºC, a switching time of 2.1 min,

and 1 43.654 and 1.161γ γ= = . In this section the PermSMBR will be evaluated and compared

with the SMBR in order to demonstrate that the ethyl lactate production is enhanced by the

integration of perm selective membranes in the SMBR unit and, therefore, the same operation

conditions will be used. In order to keep the same 1γ and 4γ , the switching time ( *t ) was

changed to 2.323, since the PermSMBR column length is different of the one of the SMBR

unit (see Table 7.2). A summary of the operating conditions used is presented in Table 7.3.

Table 7.3 Operating conditions.

Operating conditions

Temperature 50 ºC Feed lactic acid (85 wt. % in water) Desorbent ethanol (99.5 wt. % in water) Configuration 3-3-4-2 Desorbent flow rate 58 mL/min Recycle flow rate 27 mL/min Switching time 2.323 min


Aiming to evaluate the geometrical parameters equivalence between the two units, the SMBR

was compared to the PermSMBR in absence of permeation (considering in the model zero

flux through the membranes) by using the parameters of Table 7.2 and operating conditions

of Table 7.3. The feed flow rate and extract flow rate used were 7 mL/min and 39 mL/min,

respectively. The performance parameters for both processes, presented in Table 7.4, are very

similar, as expected.

Table 7.4 Comparison between SMBR and PermSMBR*

SMBR PermSMBR*

PUX 99.92 99.94

PUR 96.19 97.85

X 99.51 99.37

PR (kgEL.Lresin-1.day-1) 14.64 14.63

DC (LEth/kgEL) 5.98 5.98

* Results obtained from the PermSMBR model considering zero flux through the membranes (equal to a SMBR).

7.5.1 Reactive/Separation Region: PermSMBR vs SMBR

The reactive/separation region is a feasible region (in the γ2 - γ3 plane), which determines the

operating conditions in sections 2 and 3, for given conditions on sections 1 and 4, to obtain

products with specific purity requirement. The reactive/separation regions were calculated for

both PermSMBR and SMBR processes, setting the conditions for complete regeneration on

sections 1 and 4, 1 4(3.654) and (1.161)γ γ , and imposing a 95 % criteria for extract and

raffinate purities and for lactic acid conversion. As it can be seen in Figure 7.9, the size of the

reactive/separation region using the PermSMBR technology is higher than the one for the

SMBR. This is justified by the fact that in the PermSMBR unit the water is removed not only

by the selective adsorption onto the resin but also by the selective removal through the

membrane; and therefore, the quantity of water removed from the system will be higher

allowing higher feed flowrates of lactic acid solution without the contamination of the

raffinate stream. It can be perceived that the PermSMBR allows working with higher feed

flowrates (12.1 mL/min) than the SMBR (8.8 mL/min) to obtain the same purity and

conversion requirements, and consequently higher productivity and lower desorbent

consumption are achieved using this new technology. This is corroborated by the


performance parameters obtained for both units working under the optimal operating

conditions, presented in Table 7.5. As it can be seen, the ethyl lactate synthesis on the

PermSMBR benefits its productivity in about 34 % and, additionally, decreases the desorbent

consumption in 28 % which will reduce downstream costs since the products are less diluted.

1.0

1.4

1.8

2.2

2.6

3.0

1.0 1.4 1.8 2.2 2.6 3.0

γ 3

γ2

PermSMBR

SMBR

Figure 7.9 Reactive/separation region for PermSMBR and SMBR processes

(PermSMBR switching time of 2.323 min and SMBR switching time of 2.1 min; remaining conditions of Table 7.3)

Table 7.5 Performance parameters of the SMBR and PermSMBR for the EL synthesis.

SMBR PermSMBR Improvement

PR (kgEL.Lresin-1.day-1) 18.06 24.19 33.94 %

DC (LEth/kgEL) 4.75 3.41 28.21 %

7.5.2 PermSMBR 3 zones

As stated before, the PermSMBR technology can be operated in different configurations,

depending on the products to be separated and on the membrane performance. The ethyl

lactate synthesis involves the formation of water, a by-product for which the membrane here

considered is selective; therefore, it might be convenient to simplify the PermSMBR unit

from 4 to 3 sections, eliminating the extract stream, which change the previous configuration

3-3-4-2 to 6-4-2. Setting the feed and desorbent flowrates at 8 / minFQ mL= and

25 / minDQ mL= , respectively, and using the remaining conditions of Table 7.3, the


PermSMBR performance can be significantly enhanced by adjusting the permeate pressure,

as shown in Table 7.6. Lactic acid conversion and ethyl lactate purity are the parameters most

significantly improved. The analysis of internal concentration profiles at the cyclic steady-

state, shown in Figure 7.10, allows better understanding the PermSMBR unit behaviour.

Table 7.6 Performance parameters for different permeate pressures.

Pperm 10 mbar 6 mbar 0 mbar

PUR (%) 92.85 96.15 99.63

X (%) 97.56 99.18 99.86

PR (kgEL.Lresin-1.day-1) 16.27 16.51 16.59

DC (LEth/kgEL) 2.00 1.96 1.95

Water, the most adsorbed component is formed from the esterification reaction but is also fed

to the system since an 85 % lactic acid aqueous solution is used. In order to have a high ethyl

lactate purity and high lactic acid conversion, water should be removed from section 2, and

desorbed from the resin in section 1. Since water is removed by pervaporation, the desorbent

flowrate can be reduced below the value used in the SMBR unit, without compromising the

resin regeneration. The reduction of the permeate pressure from 10 mbar (Figure 7.10a), to

6 mbar (Figure 7.10b) and 0 mbar (Figure 7.10c), increases the water permeation flux

(Equation 7.2), enhancing significantly the resin regeneration on section 1, which avoids

water to pass from section 1 to section 3 and increases the ethyl lactate purity. Since the

reaction is equilibrium limited, this decrease on water content near the raffinate port prevents

the ethyl lactate hydrolysis, increasing the lactic acid conversion, as consequence.

A SMBR unit to have the same raffinate purity, lactic acid conversion and ethyl lactate

productivity as the PermSMBR ( 6 permP mbar= ), would have to operate in the following

conditions: configuration of 3-3-4-2, * 2.1 mint = , 8 / minFQ mL= , 38 / minXQ mL= ,

58 / minDQ mL= and Re 27 / mincQ mL= . However, this would imply a desorbent

consumption of 5.20 LEth/kgEL, which is 62 % higher than the one needed for the PermSMBR.


0

2

4

6

8

10

12

14

16

18

0 3 6 9 12

Con

cent

ratio

n (m

ol/L

)Ethanol

Lactic acid

Ethyl lactate

Water

D F RP1 P2 P3 P4 P10 P11P7P5 P6 P8 P9 P12

0

2

4

6

8

10

12

14

16

18

0 3 6 9 12

Con

cent

ratio

n (m

ol/L

)

Ethanol

Lactic acid

Ethyl lactate

Water


0

2

4

6

8

10

12

14

16

18

0 3 6 9 12

Con

cent

ratio

n (m

ol/L

)

Ethanol

Lactic acid

Ethyl lactate

Water


Figure 7.10 Influence of permeate pressure on the concentration profiles at the middle

of the switching time at cyclic steady state for the PermSMBR with 3 sections and configuration 6-4-2: a) 10 permP mbar= ; b) 6 permP mbar= and c) 0 permP mbar= .

(a)

(b)

(c)


7.5.3 Comparison between PermSMBR, SMBR and RD technologies

The ethyl lactate synthesis using Amberlyst 15-wet as catalyst was successfully implemented

by means of RD processes; and the best unit performance was obtained using 88 wt. % lactic

acid solution, at 128ºC (bottom temperature), achieving 95 % of lactic acid conversion and

95 % of ethyl lactate purity (Asthana et al., 2005). From the previous studies on the SMBR

and PermSMBR processes for the ethyl lactate synthesis at 50ºC, the best ethyl lactate

productivity obtained for a criteria of 95 % of ethyl lactate purity and 95 % of lactic acid

conversion (see Table 7.7), is 48 and 98 % higher than that of the RD process, respectively.

However, in the SMBR and PermSMBR processes the excess of ethanol used to desorb water

is higher than in the case of RD, leading to higher ethanol consumption.

Table 7.7 Performance parameters for RD, SMBR and PermSMBR technologies.

RD+ SMBR PermSMBR

PR (kgEL.Lresin-1.day-1) 12.19 18.06 24.19

DC (LEth/kgEL) 2.17 4.75 3.41

+ (Asthana et al., 2005)

Another factor to take in consideration is that the RD operates at 128ºC, while the SMBR and

PermSMBR systems work at 50ºC; thus, the energy consumption will be higher on the RD

process. Therefore, comparison of technologies must be done in terms of economical

assessment.

7.6 Conclusions

In this Chapter a new technology was proposed, the PermSMBR, that consists in a SMBR

integrated with a permeable membrane reactor by using selective permeable membranes

inside the columns of the SMBR. The PermSMBR proved to be more effective than the

SMBR when applied to the ethyl lactate synthesis; higher ethyl lactate productivities and

lower desorbent consumptions were obtained for the same purity and conversion criterions


with this new reactor, since additionally to the driven force for product separation by selective

adsorption there is the selective membrane removal of water.

The 3 sections PermSMBR operated in configuration 6-4-2 was studied, and it was concluded

that it is not necessary the completely regeneration of the resin by water desorption in section

1 by introducing the sufficient amount of desorbent (ethanol), since water is also removed by

the permeable membranes, preventing the ethyl lactate contamination. It was stated a

decrease of 62 % in the desorbent consumption with this new technology using a 6 mbar

vacuum pressure when comparing with the SMBR to attain the same ethyl lactate purity,

productivity and lactic acid conversion.

When compared with the RD process, the PermSMBR requires higher ethanol consumption,

but allows a significantly ethyl lactate productivity increase of 98 %, operating at quite lower

temperatures.

Concluding, the PermSMBR is a very interesting technology, even compared with other

intensified processes, that allows complete reactants conversion, high productivity, high

purity, significant reduction of solvent consumption and consequently lowers downstreaming

costs associated to the separation units.

7.7 Notation

ia liquid-phase activity of component i in bulk side

mA membrane area per unit reactor volume (m2membrane/m3

bulk)

C liquid phase concentration (mol/L)

axD axial dispersion coefficient (m2/min)

DC desorbent consumption (L/mol)

iJ permeate flux of species i (mol/(m2min))


ovk global membrane mass transfer coefficient (mol/(m2sPa))

L column length (m)



0ip saturation pressure of component i (bar)

permP total pressure on the permeate side (bar)

PR raffinate productivity (kgEL/(L resin.day))

PUR raffinate purity (%)

PUX extract purity (%)


particle (mol/L)

Q volumetric flowrate (L/min)

r rate of reaction (mol kg-1 min-1)


t time variable (min)

*t switching time (min)

Us solid velocity (m/min)

u interstitial velocity (m/min)

,mol iV molar volume of species i (L/mol)

X lactic acid conversion

iy molar fraction in the vapor phase of component i


Greek letters

γ interstitial velocities ratio

ε bulk porosity




pρ particle density

Subscripts


j relative to section in SMBR (j = 1, 2, 3, 4)

k relative to column in SMBR





W relative to water




cRe relative to recycle


Superscripts





7.8 References

Agreda V. H. and L. R. Partin, "Reactive distillation process for the production of methyl acetate", U. S. Patent 4435595 (1984).

Alonso M., M. J. Lorences, M. P. Pina and G. S. Patience, "Butane partial oxidation in an externally fluidized bed-membrane reactor", Catal. Today 67(1-3): 151-157, 2001.

Asthana N., A. Kolah, D. T. Vu, C. T. Lira and D. J. Miller, "A continuous reactive separation process for ethyl lactate formation", Org. Process Res. Dev. 9(5): 599-607, 2005.

Broughton D. B. and C. G. Gerhold, "Continuous Sorption Process Employing Fixed Bed of Sorbent and Moving Inlets and Outlets", US Patent No. 2 985 589 (1961).

Datta R. and S.-P. Tsai, "Esterification of Fermentation-Derived Acids via Pervaporation", WO Patent No. 9823579 (1998).


Krishna R., "Reactive separations: more ways to skin a cat", Chem. Eng. Sci. 57(9): 1491-1504, 2002.

Lafarga D. and A. Varma, "Ethylene epoxidation in a catalytic packed-bed membrane reactor: Effects of reactor configuration and 1,2-dichloroethane addition", Chem. Eng. Sci. 55(4): 749-758, 2000.

Lee W. N., I. J. Kang and C. H. Lee, "Factors affecting filtration characteristics in membrane-coupled moving bed biofilm reactor", Water Res. 40(9): 1827-1835, 2006.

Pereira C. S. M., V. M. T. M. Silva and A. r. E. Rodrigues, "Fixed Bed Adsorptive Reactor for Ethyl Lactate Synthesis: Experiments, Modelling, and Simulation", Sep. Sci. Technol. 44(12): 2721 - 2749, 2009a.

Pereira C. S. M., M. Zabka, V. M. T. M. Silva and A. E. Rodrigues, "A novel process for the ethyl lactate synthesis in a simulated moving bed reactor (SMBR)", Chem. Eng. Sci. 64(14): 3301-3310, 2009b.

Sá Gomes P., M. Minceva and A. E. Rodrigues, "Operation strategies for simulated moving bed in the presence of adsorbent ageing", Sep. Sci. Technol. 42(16): 3555-3591, 2007.



Theuerkauf J., P. Witt and D. Schwesig, "Analysis of particle porosity distribution in fixed beds using the discrete element method", Powder Technol. 165(2): 92-99, 2006.


Tsotsis T. T., M. Sahimi, B. Fayyaz-Najafi, A. Harale, B.-G. Park and P. K. T. Liu, "Hybrid adsorptive membrane reactor", US Patent No. 2007053811 (A1) (2007).


8. Conclusions and Suggestions for Future Work

This thesis focused on the development of a new efficient process to produce ethyl lactate

based in the esterification reaction between ethanol and lactic acid by using hybrid

technologies of reaction/separation based on the Simulated Moving Bed Reactor (SMBR) and

Pervaporation Membrane processes. Therefore, several topics were addressed from the

fundamentals (thermodynamic equilibrium and kinetics of reaction, multi-component

adsorption equilibria and pervaporation data), to the process development (modelling and

simulation) and intensification (simulated moving bed reactor, pervaporation membrane

reactor and an integration of both). The main results and conclusions are:

(i) Batch Reactor: Thermodynamic Equilibrium and Reaction Kinetics

To circumvent the lack of a suitable estimative for the thermodynamic equilibrium constant

as function of temperature, the equation ( )KTK 13.5159625.2ln −= , based on the

UNIQUAC method, and valid in the temperature range of 50-90ºC, was proposed.

The kinetic law was described by a Langmuir-Hinshelwood rate expression based on

activities, considering the surface reaction as the rate-controlling step. The following kinetic

law was proposed for the temperature range of 50-90 ºC, for the Amberlyst 15-wet catalyst:

( ) 2/ (1 )c Eth La EL W Eth Eth W Wr k a a a a K K a K a= − + + ; and the model parameters are

( )7( /( .min)) 2.70 10 exp 6011.55 / ( )ck mol g T K= × − , ( ))(/01.12exp19.15 KTKW = and

( ))(/63.359exp22.1 KTK Eth = .

(ii) Fixed Bed Adsorptive Reactor

The multi-component adsorption equilibria were determined from dynamic adsorption

experiments in absence of reaction, at 20 ºC and 50 ºC, using Amberlyst 15-wet as selective

194 CHAPTER 8 Conclusions and Suggestions for Future Work

adsorbent. The multi-component Langmuir adsorption isotherm was modified in order to

reduce the adjustable adsorption parameters from 8 (one molar monolayer capacity and one

equilibrium constant for each component) to 5 (one volumetric monolayer capacity for all

components and one equilibrium constant for each component):

( ) ( ),/ 1i V mol i i i j jq Q V K C K C= +∑ using the following parameters:

Component QV (ml/lwet solid) 20 ºC / 50 ºC

K (l/mol) 20 ºC / 50 ºC

Vmol (ml/mol) 20 ºC / 50 ºC

Ethanol 5.443 / 3.068 58.17 / 60.87

Lactic acid 4.524 / 4.085 74.64 / 77.56

Ethyl lactate 1.117 / 1.815 113.99 / 118.44

Water

390.0/383.5

15.353 / 7.055 18.08 / 18.63

(iii) Simulated Moving Bed Reactor

The ethyl lactate was produced in the Simulated Moving Bed Reactor pilot unit LICOSEP,

using Amberlyst 15-wet resin as catalyst and selective adsorbent. A mathematical model

considering external and internal mass-transfer resistances and variable velocity due to

change of liquid composition was developed to describe the dynamic behaviour of the SMBR

and it was validated by the experiments performed. The theoretical assessment of the SMBR

unit behaviour was performed ensuring complete regeneration of the resin (in section 1) and

desorbent (in section 4), by using the mathematical model to analyse the effect of SMBR

configuration, feed composition and switching time into the reactive/separation regions

or/and into the process performance at the optimal operating points. It was shown that the

SMBR is a very attractive technology for the production of ethyl lactate, since under

appropriate conditions the lactic acid conversion can be driven to completion and productivity

as high as 32 kgEL/(Lads.day) and purity of 95 % can be obtained.

(iv) Pervaporation Membrane Reactor

Pervaporation process using commercial Hydrophilic silica membranes from Pervatech was

evaluated for the ethyl lactate system aiming to contribute either for the separation of SMBR

extract stream (water/ethanol mixtures) or for the ethyl lactate process intensification by

continuous pervaporation membrane reactor (PVMR). First pervaporation studies indicate

that the membranes have no major imperfections since the total flux and water selectivity is


barely affected by absolute feed pressure. The influence of hydrodynamic conditions on the

membrane polarization was analyzed, and for velocity values higher than 0.14 m/s

polarization effects are eliminated. It was concluded that the permeances for all species as

function of temperature and feed water content, are described by the following equations for

the temperature range 48ºC - 72ºC: 3

7,

22.60 101.41 10 expmemb EthQRT

− ⎛ ⎞×= × ⎜ ⎟

⎝ ⎠2/( min)mol dm bar

, 0memb LAQ = 2/( min)mol dm bar

34

,10.42 101.12 10 expmemb ELQ

RT− ⎛ ⎞×

= × ⎜ ⎟⎝ ⎠

2/( min)mol dm bar

6,

50377 323262 10 exp(18.642 )exp Wmemb W W

xQ xRT

− −⎛ ⎞= × −⎜ ⎟⎝ ⎠

2/( min)mol dm bar

A mathematical model was developed for the batch pervaporation membrane (BPM),

considering (i) plug flow for the bulk fluid phase; (ii) total feed volume inside the tank and

retentate velocity inside the membrane variations due to permeation of components; (iii)

concentration polarization, where the resistance due to the diffusive transport in the boundary

layer is combined with the membrane resistance in a global membrane resistance; (iv) non-

isothermal operation due to heat consumption for species vaporization; and (v) temperature

polarization. The BPM model was validated experimentally and, therefore, it was extended to

the integrated pervaporation membrane reactor packed with the catalyst Amberlyst-15wet. In

isothermal conditions, the PVMR is a very attractive solution being possible to achieve near

lactic acid depletion (98 % conversion) and ethyl lactate with 96 % purity, at 70ºC.

(v) PermSMBR – A New Hybrid Technology

A new technology, the Simulated Moving Bed Membrane Reactor (PermSMBR), was

developed and applied for the esterification of lactic acid with ethanol. Its conception was a

result of process re-intensification by integrating a permeable membrane reactor with a

simulated moving bed reactor. The PermSMBR proved to be more effective than the SMBR

when applied to the ethyl lactate synthesis; higher ethyl lactate productivities and lower

desorbent consumptions were obtained for the same purity and conversion criterions with this

new reactor, since additionally to the driven force for product separation by selective

adsorption there is the selective membrane removal of water. It was demonstrated that the

196 CHAPTER 8 Conclusions and Suggestions for Future Work

PermSMBR technology can reduce the desorbent consumption in 62 % and increase the ethyl

lactate productivity in 33 % when compared with the SMBR and in 98 % when compared

with RD for the same purity and conversion requirements. Concluding, the PermSMBR is a

promising technology, even if compared with other intensified processes that allows complete

reactants conversion, high productivity, high purity, significant reduction of solvent

consumption and consequently lowers downstream costs associated to the separation units.

This thesis is focused on the ethyl lactate synthesis with a special contribution on its process

intensification. The SMBR technology exhibits higher productivity than that of the PVMR,

but has the disadvantage of requiring further separation units for recovery of ethanol from

both extract and raffinate streams. Integrating the benefits of SMBR and PVMR, a new

technology was invented, the PermSMBR. However, certain topics were not exploited and

others are unfinished and, therefore, can be deeper investigated. For that reason it is suggested

to explore, in a near future, the following research lines:

(i) Screening of new catalysts

Smopex 101 fibre has proved to be an efficient catalyst for acetalization and esterification

reactions; since it has negligible adsorption capacity is not suitable for chromatographic

reactor, namely the SMBR and PermSMBR, although would be valuable for the PVMR.

Another interesting catalyst is the Amberlyst 36-wet (A36), a strongly acidic ion exchange

resin, which proved to be more active than Amberlyst 15-wet (A15) for esterification

reactions; moreover, as adsorbent has selectivity to water since the swelling from dry to water

is 54 %, while A15 only swells 37 %. Therefore, A36 will enhance the performance of SMBR

and PermSMBR by improving both reaction kinetics and products separation.

(ii) Deactivation studies

It is known that catalyst and/or adsorbent ageing affects significantly the performance of

industrial processes, and might compromise the process feasibility. Ageing studies should

consider not only pro-analytical reactants but also industrial reactants (lactic acid from

fermentation broth and bio-ethanol). Its influence on the ion-exchange resin deactivation, in

terms of reaction kinetics and multi-component adsorption equilibria, should be addressed, in


order to understand its implications on the processes performance. Analysis of non-isothermal

effects onto the SMBR and PermSMBR unit performance should also be performed.

(iii) Process Integration and Economical Evaluation

The economical assessment of the ethyl lactate synthesis should be carried out considering its

integration with the subsequent separation units. Comparison between technologies (SMBR,

PVMR, PermSMBR, RD, ...) should be based on financial project evaluation using the Net

Present Value analysis, net annual profit or annual production cost. Alternatively, the

evaluation of the most sustainable technology should be based on the Life Cycle analysis.

Appendix A: Safety Data

1.1 Ethyl lactate

NFPA

1.1.1 General

Synonyms: lactic acid ethyl ester

Molecular formula: C5H10O3

CAS No: 97-64-3

1.1.2 Physical Data

Appearance: colourless liquid

Melting point: -26 ºC

Boiling point: 154 ºC

Vapour density: 4.07 (air = 1)

Vapour pressure (mmHg): 5 @ 30ºC (86ºF)

Density (g cm-3): 1.03

Flash point: 46 ºC (closed cup)

Explosion limits: 1.5 – 11.4%

Autoignition temperature: 400 ºC

Water solubility: appreciable

1.1.3 Stability

Stable. Combustible. Incompatible with strong oxidizing agents.

• Health: 3

• Flammability: 2

• Reactivity: 0

23 0

APPENDIX A. Safety Data A 2

1.1.4 Toxicology

Skin, eye and respiratory irritant.

Toxicity data:

Oral-mouse lethal dose 50% kill 2500 mg kg-1

Intraperitoneal – rat lowest published lethal dose 1000 mg kg-1

Subcutaneous – mouse lethal dose 50% kill 2500 mg kg-1

Risk phrases:

• R36 Irritating to eyes.

• R37 Irritating to respiratory system.

• R38 Irritating to skin.

1.1.5 Personal Protection

Minimize contact.

Safety phrases:

• S26 In case of contact with eyes, rinse immediately with plenty of water and seek

medical advice.

1.1.6 Hazard Summary

• Ethyl lactate can affect you when breathed in and may be absorbed through the skin.

• Prolonged contact can irritate the skin and eyes.

• Breathing ethyl lactate may cause dizziness, lightheadedness, and passing out.

1.1.7 How to determine if you are being expose

• Exposure to hazardous substances should be routinely evaluated. This may include

collecting personal and area air samples. Under OSHA 1910.20, you have a legal right to

obtain copies of sampling results from your employer. If you think you are experiencing

any work-related health problems, see a doctor trained to recognize occupational diseases.

Take this Fact Sheet with you.


1.1.8 Workplace exposure limits

No occupational exposure limits have been established for ethyl lactate. This does not mean

that this substance is not harmful. Safe work practices should always be followed.

• It should be recognized that ethyl lactate may be absorbed through the skin, thereby

increasing your exposure.

1.1.9 Ways of reducing exposure

Where possible, enclose operations and use local exhaust ventilation at the site of chemical

release. If local exhaust ventilation or enclosure is not used, respirators should be worn.

Wear protective work clothing.

Wash thoroughly immediately after exposure to ethyl lactate and at end of the workshift.

Post hazard and warning information in the work area. In addition, as part of an ongoing

education and training effort, communicate all information on the health and safety hazards of

ethyl lactate to potentially exposed workers.

1.1.10 Health hazard information

Acute Health Effects

The following acute (short-term) health effects may occur immediately or shortly after

exposure to ethyl lactate:

Prolonged contact can irritate the skin and eyes.

Breathing ethyl lactate may cause dizziness, lightheadedness and passing out.

Chronic Health Effects

The following chronic (long-term) health effects can occur at some time after exposure to

ethyl lactate and can last for months or years:

Cancer Hazard

According to the information presently available to the New Jersey Department of Health and

Senior Services, ethyl lactate has not been tested for its ability to cause cancer in animals.


Reproductive Hazard

According to the information presently available to the New Jersey Department of Health and

Senior Services, ethyl lactate has not been tested for its ability to affect reproduction.

Other Long-Term Effects

Ethyl lactate has not been tested for other chronic (long-term) health effects.

Medical Testing

There is no special test for this chemical. However, if illness occurs or overexposure is

suspected, medical attention is recommended.

Any evaluation should include a careful history of past and present symptoms with an exam.

Medical tests that look for damage already done are not a substitute for controlling exposure.

Request copies of your medical testing. You have a legal right to this information under

OSHA 1910.1020.

1.1.11 Workplace controls and practices

Unless a less toxic chemical can be substituted for a hazardous substance, ENGINEERING

CONTROLS are the most effective way of reducing exposure. The best protection is to

enclose operations and/or provide local exhaust ventilation at the site of chemical release.

Isolating operations can also reduce exposure. Using respirators or protective equipment is

less effective than the controls mentioned above, but is sometimes necessary.

In evaluating the controls present in your workplace, consider: (1) how hazardous the

substance is, (2) how much of the substance is released into the workplace and (3) whether

harmful skin or eye contact could occur. Special controls should be in place for highly toxic

chemicals or when significant skin, eye, or breathing exposures are possible.

In addition, the following controls are recommended:

• Where possible, automatically pump liquid Ethyl Lactate from drums or other storage

containers to process containers.

Good WORK PRACTICES can help to reduce hazardous exposures. The following work

practices are recommended:

• Workers whose clothing has been contaminated by Ethyl Lactate should change into clean

clothing promptly.


• Contaminated work clothes should be laundered by individuals who have been informed of

the hazards of exposure to Ethyl Lactate.

• Eye wash fountains should be provided in the immediate work area for emergency use.

• If there is the possibility of skin exposure, emergency shower facilities should be provided.

• On skin contact with Ethyl Lactate, immediately wash or shower to remove the chemical.

At the end of the workshift, wash any areas of the body that may have contacted Ethyl

Lactate, whether or not known skin contact has occurred.

• Do not eat, smoke, or drink where Ethyl Lactate is handled, processed, or stored, since the

chemical can be swallowed. Wash hands carefully before eating, drinking, smoking or

using the toilet.

Personal Protective Equipment

WORKPLACE CONTROLS ARE BETTER THAN PERSONAL PROTECTIVE

EQUIPMENT. However, for some jobs (such as outside work, confined space entry, jobs

done only once in a while, or jobs done while workplace controls are being installed),

personal protective equipment may be appropriate.

The following recommendations are only guidelines and may not apply to every situation.

Clothing

• Avoid skin contact with Ethyl Lactate. Wear solvent-resistant gloves and clothing. Safety

equipment suppliers/manufacturers can provide recommendations on the most protective

glove/clothing material for your operation.

• All protective clothing (suits, gloves, footwear, headgear) should be clean, available each

day, and put on before work.

Eye Protection

• Wear indirect-vent, impact and splash resistante goggles when working with liquids.

• Wear a face shield along with goggles when working with corrosive, highly irritant or

toxic substances.


Respiratory Protection

IMPROPER USE OF RESPIRATORS IS DANGEROUS. Such equipment should only be

used if the employer has a written program that takes into account workplace conditions,

requirements for worker training, respirator fit testing and medical exams, as described in

OSHA 1910.134.

• Engineering controls must be effective to ensure that exposure to Ethyl Lactate does not

occur.

• Where de potential for overexposure exists, use a MSHA/NIOSH a approved supplied-air

respirator with a full facepiece operated in a pressure-demand or other positive-pressure

mode. For increased protection use in combination with an auxiliary self-contained

breathing apparatus operated in a pressure-demand or other positive-pressure mode.

1.1.12 Handling Storage

• Prior to working with Ethyl Lactate you should be trained on its proper handling and

storage.

• Ethyl Lactate is not compatible with OXIDIZING AGENTS( such as PERCHLORATES,

PEROXIDES, PERMANGANATES, CHLORATES, NITRATES, CHLORINE,

BROMINE and FLUORINE).

• Store in tightly closed containers in a cool, dark, well-ventilated area.

• Sources of ignition, such as smoking and open flames, are prohibited where Ethyl Lactate

is handled, used, or stored.

1.1.13 Fire Hazards

• Ethyl Lactate is a COMBUSTIBLE LIQUID.

• Use dry chemical, CO2, water spray or alcohol resistant foam extinguishers.

• CONTAINERS MAY EXPLODE IN FIRE.

• Use spray to keep fire-exposed containers cool.

• If employees are expected to fight fires, they must be trained and equipped as stated in

OSHA 1910.156.


1.1.14 Spills and emergencies

If Ethyl Lactate is spilled or leaked, take the following steps:

• Evacuate person’s not wearing protective equipment from area of spill or leak until clean-

up is complete.

• Remove all ignition sources.

• Ventilate area of spill or leak.

• Absorb liquids in vermiculite, dry sand, earth, or a similar material and deposit in sealed

containers.

• It may be necessary to contain and dispose of Ethyl Lactate as a HAZARDOUS WASTE.

Contact your Department of Environmental Protection (DEP) or your regional office of the

federal Environmental Protection Agency (EPA) for specific recommendations.

• If employees are required to clean-up the spills, they must be properly trained and

equipped. OSHA 1910.120(q) may be applicable.

1.1.15 First aid

Eye Contact

Immediately flush with large amounts of water. Continue without stopping for at least 15

minutes, occasionally lifting upper and lower lids.

Skin Contact

Remove contaminated clothing. Wash contaminated skin with soap and water.

Breathing

• Remove the person from exposure.

• Begin rescue breathing if breathing has stopped and CPR if heart action has stopped.

• Transfer promptly to a medical facility.


1.2 Lactic Acid

NFPA

1.2.1 General

Synonyms: 2-hydroxypropanoic acid, ethylideneactic acid, 1-hydroxyethanecarboxylic acid

Molecular formula: CH3CHOHCOOH

CAS No: 50-21-5

EC No: 200-018-0

1.2.2 Physical Data

Appearance: colourless to yellow liquid

Melting point: 18 ºC

Boiling point: 122 ºC @12mmHg

Specific gravity: 1.05

1.2.3 Stability

Stable. Combustible. Incompatible with strong oxidizing agents.

1.2.4 Toxicology

Eye or skin contact may cause severe irritation or burns.

Toxicity data:

Oral-rat lethal dose 50% kill 3730 mg kg-1

Subcutaneous – mouse lethal dose 50% kill 4500 mg kg-1

Risk phrases:

• R34 Causes burns.

• Health: 2

• Flammability: 1

• Reactivity: 0

12 0


1.2.5 Personal Protection

Safety glasses.

Safety phrases:

• S26 In case of contact with eyes, rinse immediately with plenty of water and seek

medical advice.

• S27 Take off immediately all contaminated clothing.

• S36 Wear suitable protective clothing.

• S37 Wear suitable gloves.

• S39 Wear eye / face protection.

1.2.6 Health Effects

• Inhalation: causes burns on contact with mucous membranes.

• Eye: Causes burns on contact with eyes.

• Skin: Causes burns on contact with skin.

• Ingestion: Harmful if swallowed. Can cause abdominal pain, diarrhea, and burns of

digestive tract.

• Routes of Entry: Inhalation, Ingestion or skin contact.

1.2.7 First Aid

• Inhalation: Remove to fresh air; give artificial respiration if breathing has stopped.

Get medical attention.

• Eye: Immediately flush eyes with large amounts of water for at least 15 minutes.

• Skin: Immediately flush thoroughly with large amounts of water. Remove

contaminated clothing and wash before reuse.

• Ingestion: Do not induce vomiting; get medical immediate attention.

1.2.8 Fire Hazard

• Extinguishing Media: water fog, foam, carbon dioxide, dry chemical


• Special Procedures: Wear self-contained breathing apparatus.

• Unusual Hazard: Emits acrid fumes when heated.

1.3 Ethanol

NFPA

1.3.1 General

Synonyms: ethyl alcohol, grain alcohol, fermentation alcohol, alcohol, methylcarbinol,

absolute alcohol, absolute ethanol, anhydrous alcohol, alcohol dehydrated, algrain, anhydrol.

Molecular formula: C2H5OH

CAS No: 64-17-5

1.3.2 Physical data

Appearance: colourless liquid

Melting point: -130 C

Boiling point: 78 C

Specific gravity: 0.789

Vapour pressure: 1.59

Flash point: 56 F

Explosion limits: 3.3% - 24.5%

Autoignition temperature: 683 F

Water solubility: miscible in all proportions

• Health: 0

• Flammability: 3

• Reactivity: 0

30 0

30 0


1.3.3 Stability

Stable. Substances to be avoided include strong oxidising agents, peroxides, acids, acid

chlorides, acid anhydrides, alkali metals, ammonia, moisture. Forms explosive mixtures with

air.

1.3.4 Toxicology

Causes skin and eye irritation. Ingestion can cause nausea, vomiting and inebriation; chronic

use can cause serious liver damage. Note that “absolute” alcohol, which is close to 100%

ethanol, may nevertheless contain traces of 2-propanol, together with methanol or benzene.

The latter two are very toxic, while “denatured” alcohol has substances added to it which

make it unpleasant and possibly hazardous to consume.

Risk phrases:

• R11 Highly flammable.

• R20 Harmful by inhalation.

• R21 Harmful in contact with skin.

• R22 Harmful if swallowed.

• R36 Irritating to eyes.

• R37 Irritating to respiratory system.

• R38 Irritating to skin.

• R40 Possible risk of irreversible effects.

1.3.5 Personal protection

Safety glasses. Suitable ventilation.

Safety phrases:

• S7 Keep container tightly closed.

• S16 Keep away from sources of ignition.

• S24 Avoid contact with skin.

• S25 Avoid contact with eyes.


• S36 Wear suitable protective clothing.

• S37 Wear suitable gloves.

• S39 Wear eye / face protection.

• S45 In case of accident or if you feel unwell, seek medical advice immediately (show

the label whenever possible.)

Appendix B: Thermodynamic Properties

1.1 Available Literature Data

The data presented in this section are from Yaws (Yaws, 1999), excepted when mentioned.

In Table B.1 some physical and thermodynamic properties of lactic acid, ethanol, ethyl lactate

and water are presented.

Table B.1 Basic properties of lactic acid, ethanol, ethyl lactate and water.

Properties Lactic Acid Ethanol Ethyl Lactate Water

Molecular weigh - M (g/mol) 90.079 46.069 118.133 18.015

Density- ρ (g/cm3) 1.209 0.789 1.031 1.027

Melting temperature - Tf (K) 289.95-291.15 159.15 248.25 273.15

Normal boiling temperature - Tb (K) 395.15 351.45 426.15-427.15 373.15

Critical temperature - Tc (K) 616.00 516.25 588.00 647.13

Critical pressure - Pc (bar) 59.65 63.84 38.60 221.20

Critical volume - Vc (cm3/mol) 216.9 166.9 354.0 57.1

Acentric factor - ω 1.035 0.637 0.793 0.344

1.1.1 Density

The modified form of the Rackett equation was selected for correlation of saturated liquid

density as a function of temperature.

n

cTT

L AB)1( −−

=ρ (B.1)

with )/( 3cmgLρ and T (K).

APPENDIX B. Thermodynamic Properties B2

Table B.2 Constants used for density calculation.

Constants Lactic Acid Ethanol Water Ethyl Lactate

A 0.39816 0.26570 0.34710 0.33372

B 0.26350 0.26395 0.27400 0.21190

n 0.28570 0.23670 0.28571 0.45530

Tmin (K) 291.15 159.05 273.16 247.15

Tmax (K) Tc Tc Tc Tc

1.1.2 Viscosity

The correlation for liquid viscosity as a function of temperature is given by Equation B.2.

210log DTCT

TBAL +++=η (B.2)

with )(cPLη and T (K).

Table B.3 Constants used for viscosity calculation.

Constants Ethanol Water Ethyl Lactate

A -6.4406E+00 -10.2158E+00 -20.0105E+00

B 1.1176E+03 1.7925E+03 3.2123E+03

C 1.3721E-02 1.7730E-02 4.1891E-02

D -1.5465E-05 -1.2631E-05 -3.2733E-05

Tmin (K) 240 273 247

Tmax (K) Tc 643 Tc

1.1.3 Vapour Pressure

The Antoine-type equation with extended term was selected for correlation of vapour

pressure as a function of temperature:

21010 loglog ETDTTC

TBAPvp ++++= (B.3)

with Pvp (mmHg) and T (K).


Table B.4 Constants used for vapour pressure calculation.

Constants Lactic Acid Ethanol Water Ethyl Lactate A -27.0836E+00 23.8442E+00 29.8605E+00 32.0863E+00

B -3.9661E+03 -2.8642E+03 -3.1522E+03 -2.9164E+03

C 2.0233E+01 -5.0474E+00 -7.3037E+00 -9.5666E+00

D -4.2176E-02 3.7448E-11 2.4247E-09 6.5114E-03

E 2.0310E-05 2.7361E-07 1.8090E-06 4.5645E-13

Tmin (K) 291.15 159.05 273.16 247.15

Tmax (K) Tc Tc Tc Tc

1.1.4 Liquid Heat Capacity

The correlation for heat capacity of liquid is a series expansion in temperature, given by

Equation B.4.

42 3 ETC A BT CT DTp += + + + (B.4)

with Cp (J/(Kmol.K)) and T (K), with exception for the ethyl lactate specie where Cp is given

in (J/(mol.K)).

Table B.5 Constants used for heat capacity calculation.

Constants Ethanol* Water* Ethyl Lactate Lactic acid*

A 1.0264E+05 2.7637E+05 -46.239E+00 6.1082E+04

B -1.3963E+02 -2.0901E+03 2.1823E+00 5.0343E+02

C -3.0341E-02 8.1250E+00 -5.9832E-03 ---------

D 2.0386E-03 -1.4116E-02 6.8683E-06 ---------

E --------- 9.3701E-06 --------- ---------

Tmin (K) 159.05 273.16 248.00 289.90

Tmax (K) 390.00 533.15 529.00 675.00

Error < 3% < 1% ----------- < 10%

* parameters taken from (DIPPR, 1998).


1.1.5 Heat of Vaporization

The correlation selected for the calculation of the heat of vaporization as a function of

temperature is given by Equation B.5 (DIPPR, 1998).

( )21V B CT DTr rTrAH + +⎡ ⎤−⎣ ⎦=Δ (B.5)

with /r cT T T= and VHΔ in (J/(Kmol).

Table B.6 Constants used for heat vaporization calculation.

Constants Ethanol Water Ethyl Lactate Lactic acid

A 5.5789E+07 5.2053E+07 8.0260E+07 1.0436E+08

B 3.1245E-01 3.1990E-01 4.0930E-01 3.8548E-01

C --------- -2.1200E-01 --------- ---------

D --------- 2.5795E-01 --------- ---------

Tmin (K) 159.05 273.16 247.15 289.90

Tmax (K) 514.00 647.13 588.00 675.00

Error < 1% < 1% < 10% < 25%

1.1.6 Thermal Conductivity

The thermal conductivity was calculated by Equation B.6 (DIPPR, 1998).

42 3 ETA BT CT DTλ += + + + (B.6)

with λ (W/(m.K)) and T (K).

Table B.7 Constants used for thermal conductivity calculation.

Constants Ethanol Water Ethyl Lactate Lactic acid

A 2.4680E-01 -4.3200E-01 2.8358E-01 3.4850E-01

B -2.6400E-04 5.7255E-03 -3.5110E-04 -3.7085E-04

C --------- -8.0780E-06 --------- ---------

D --------- 1.8610E-09 --------- ---------

Tmin (K) 159.05 273.16 247.15 289.90

Tmax (K) 353.15 633.15 427.65 490.00

Error < 5% < 1% < 25% < 10%


1.2 Properties Estimation

1.2.1 Estimation of Liquid Viscosity

The lactic acid viscosity was estimated by the following expression (DIPPR, 1998):

4097.9exp 14.403 0.4407 ln( )LA TT

η ⎛ ⎞= − + − ×⎜ ⎟⎝ ⎠

(B.7)

with ( . )L Pa sη and T (K).

Table B.8 Viscosity of lactic acid.

T (K) 293.15 323.15 η (cP) 53.67 13.99

For the estimation of the viscosity of the remaining species (ethanol, ethyl lactate and water)

it was used the Equation B.2 presented in section 1.1.2.

1.2.2 Estimation of Vapour Pressure

For correlation of vapour pressure as a function of temperature the Antoine equation with

extended term was chosen (Equation B.3).

0.00

0.50

1.00

1.50

2.00

2.50

290 320 350 380

P vp

(atm

)

T (K)

Lactic acidEthanolWaterEthyl lactate

Figure B. 1 Variation of the vapour pressure of the compounds with the

temperature.


1.2.3 Estimation of Liquid Heat Capacity

The liquid heat capacity was estimated by the correlation given by Equation B.4 (see Figure

B.2).

0

50

100

150

200

250

300

350

290 310 330 350 370

Cp (J

/ (m

ol.K

))

T (K)

Ethanol WaterEthyl lactate Lactic acid

Figure B.2 Variation of the liquid heat capacity of the different species with the

temperature.

1.2.4 Estimation of Heat of Vaporization

The heat of vaporization was estimated by the correlation given by Equation B.5 (see Figure

B.3).

3.0E+04

4.5E+04

6.0E+04

7.5E+04

9.0E+04

290 310 330 350 370

ΑH V

(J /

mol

)

T (K)

EthanolWaterEthyl lactateLactic acid

Figure B.3 Variation of the heat of vaporization of the different species with the

temperature.


1.2.5 Estimation of Molar Volumes

The molar volumes for all components were estimated with Gunn-Yamada method (Reid et

al., 1987):

( ) ( )( )

R

RV

Tf

TfTV = (B.8)

where

( ) ( )21 1 HHTf ω−= (B.9)

4321 11422.102512.251941.133953.033593.0 rrrr TTTTH +−+−= (B.10)

22 04842.009045.029607.0 rr TTH −−= (B.11)

c

R

cr T

TorTTT = (B.12)

VR is the molar volume at the reference temperature TR (cm3/mol), ω is the acentric factor

and Tc is the critical temperature (K).

Gunn-Yamada method could be used just in the case when the liquid molar volume is known

at some temperature (reference temperature). The reference molar volumes are presented at

Table B.9.

Table B.9 Molar volumes at the reference temperature used for all calculations.

VR (cm3/mol) Reference Temperature TR (K) Lactic Acid Ethanol Ethyl Lactate Water

293.15 74.640 58.174 113.986 18.082

1.3 References

DIPPR, "Thermophysical Properties Database", (1998).

Reid R. C., P. J.M. and P. B.E., "The Properties of Gases and Liquids", McGraw-Hill, (1987).

Yaws C. L., "Chemical Properties Handbook", McGraw-Hill, (1999).

Appendix C: Calibration

1.1 Calibration

1.1.1 Pure Components

The calibration was realized by injecting different volumes of pure components at 20 ºC and

registering the respective peak area values. The molar volume, VM, of each species was used

to convert volume, V, in number of moles, n. In the case of lactic acid the volume was

converted to number of moles using the solution density (ρ), since lactic acid is not pure

(lactic acid 85 w/w % pure).

The response factor (fi) was defined as:

iii Afn = (C.1)

where in is the number of moles of component i (μmol) and iA is the area of component i (u.a.)

nEth = 5.879AEthR2 = 0.983

0

2

4

6

8

10

0 0.4 0.8 1.2 1.6

n Eth

(µm

ol)

AEth (u.a.)

Figure C.1

Table C.1 Ethanol

VM (cm3/mol) V (μL) n (μmol) A (u.a.)

0.1 1.7093 0.3394 0.1 1.7093 0.3410 0.2 3.4185 0.6408 0.2 3.4185 0.6385 0.3 5.1278 0.9178 0.3 5.1278 0.9208 0.4 6.8371 1.1555 0.4 6.8371 1.1796 0.5 8.5464 1.4011

58.17

0.5 8.5464 1.3564

APPENDIX C. Calibration C2

nLA = 6.994ALAR2 = 0.990

1

2

3

4

5

6

0.3 0.5 0.7 0.9

n LA

(µm

ol)

ALA (u.a.)

Figure C.2

nEL = 3.345AELR2 = 0.986

0

1

2

3

4

5

0 0.4 0.8 1.2 1.6

nEL(µmol)

AEL (u.a.) Figure C.3

nW= 11.826AWR2 = 0.988

0

5

10

15

20

25

30

0 0.5 1 1.5 2 2.5

n W(µ

mol

)

AW (u.a.) Figure C.4

Table C.2 Lactic acid ρ (g/cm3) V (μL) n (μmol) A (u.a.)

0.2 2.2375 0.3144 0.2 2.2375 0.3212 0.3 3.3562 0.4732 0.3 3.3562 0.4574 0.4 4.4750 0.6262 0.4 4.4750 0.6612 0.5 5.5937 0.7800

1.209

0.5 5.5937 0.8294

Table C.3 Ethyl lactate VM (cm3/mol) V (μL) n (μmol) A (u.a.)

0.1 0.8695 0.3074 0.1 0.8695 0.3020 0.2 1.7390 0.5713 0.2 1.7390 0.5721 0.3 2.6085 0.8174 0.3 2.6085 0.8202 0.4 3.4780 1.0289 0.4 3.4780 1.0371 0.5 4.3475 1.2496

113.99

0.5 4.3475 1.2407

Table C.4 Water

VM (cm3/mol) V (μL) n (μmol) A (u.a.)

0.1 5.5253 0.5409 0.1 5.5253 0.5434 0.2 11.0506 1.0200 0.2 11.0506 1.0287 0.3 16.5759 1.4712 0.3 16.5759 1.4729 0.4 22.1012 1.8591 0.4 22.1012 1.8350 0.5 27.6266 2.2369

18.08

0.5 27.6266 2.2602


1.1.2 Binary mixtures

Several standard binary mixtures with known concentration were prepared, analysed and the

molar fraction of each component was determined, accordingly to Equation C.2.

,i i

est in n

n

f Axf A

=∑

(C.2)

In order to introduce a correction factor, the estimated molar fraction (obtained from Equation

C.2) and the real molar fraction of each component were adjusted for each pair by a linear

fitting (see Figure C.5).

, ,Calc i est ix a bx= + (C.3)

xreal,EL = 1.1304xest,EL - 0.0177R2 = 0.9998

0

0.2

0.4

0.6

0.8

1

0 0.2 0.4 0.6 0.8 1

x rea

l,EL

xest,EL

Ethyl lactate/Ethanol

xreal,Eth = 1.0492xest,Eth - 0.018R2 = 0.9995

0

0.2

0.4

0.6

0.8

1

0 0.2 0.4 0.6 0.8 1

x rea

l,Eth

xest,Eth

Ethanol/Water

Figure C.5 Real molar fraction as function of estimated molar fraction of a specie in

different binary mixtures.

xreal,LA = 1.089xest,LA + 0.0118R2 = 0.9984

0

0.1

0.2

0.3

0.4

0 0.1 0.2 0.3 0.4

x rea

l,LA

xest,LA

Lactic acid/Water

xreal,EL = 0.9706xest,EL - 0.0044R² = 0.9993

0

0.2

0.4

0.6

0.8

0 0.2 0.4 0.6 0.8

x rea

l,EL

xest,EL

Ethyl lactate/Water


1.1.3 Multicomponent mixtures

Several standard multicomponent mixtures with known concentrations were prepared and

analysed. The real molar fraction and the estimated molar fraction (from Equation C.2) of

each component were adjusted by a linear fitting (Figure C.6).

xreal,EL = 0.8803xest,EL + 0.0031R² = 0.9993

0

0.2

0.4

0.6

0 0.2 0.4 0.6

x rea

l,EL

xest,EL

Figure C.6 Real molar fraction as function of estimated molar fraction of a specie in multicomponent mixtures.

The lactic acid molar fraction was calculated from mass balance: 1 ( )LA Eht EL Wx x x x= − + + .

1.2 Validation of Calibration

The molar fraction of a component in a binary mixture is calculated using Equation C.2 and

the equation obtained from the linear fitting of the respective pair (presented in Figure C.5).

For the case of multicomponent mixtures the procedure for the calculation of a specie molar

fraction is the same that the one used for binary mixtures; first it is used Equation C.2 and

then the equation obtained from the linear fitting of the specie in a multicomponent mixture

(presented in Figure C.6). For both cases, the concentration is then evaluated using the liquid

molar volumes (Equation C.4):

xreal,Eth = 0.9986xest,Eth - 0.016R² = 0.9986

0

0.2

0.4

0.6

0.8

1

0 0.2 0.4 0.6 0.8 1

x rea

l,Eth

xest,Eth

xreal,W = 0.9969xest,W - 0.0144R² = 0.999

0

0.2

0.4

0.6

0 0.2 0.4 0.6

x rea

l,W

xest,W


,

ii

n M nn

xCx V

=∑

(C.4)

In order to verify the analysis accuracy for the binary and quaternary mixtures several

samples with known concentration were prepared and analysed, as shown in Tables C.4-C.8.

Table C.4 Analysis of Ethanol/Ethyl lactate mixtures.

Ethyl lactate molar fraction Sample

Real Calculated Error (%) 1 0.1919 0.1935 0.82

2 0.5078 0.5075 -0.07

3 0.8968 0.8991 0.26

Table C.5 Analysis of Lactic acid/Water mixtures. Lactic acid molar fraction

Sample Real Calculated Error (%)

1 0.0818 0.0809 -1.12

2 0.3182 0.3213 0.98

3 0.4098 0.4146 1.17

Table C.6 Analysis of Ethanol/Water mixtures.

Ethanol molar fraction Sample

Real Calculated Error (%) 1 0.1560 0.1574 0.87

2 0.6025 0.5912 -1.88

Table C.7 Analysis of Ethyl lactate/Water mixtures.

Ethyl lactate molar fraction Sample

Real Calculated Error (%) 1 0.2789 0.2781 -0.27 2 0.6565 0.6622 0.87


Table C.8 Analysis of a quaternary mixture. Molar fraction Component

Real Calculated Error (%) Ethanol 0.3969 0.3981 0.29

Lactic acid 0.1865 0.1900 1.88 Ethyl lactate 0.2275 0.2233 -1.85

Water 0.1891 0.1886 -0.24

Appendix D: Binary adsorption experiments at 293.15 K

The breakthrough curves for the binary adsorption experiments performed at 293.15 K are

presented below.

(a)

0

10

20

30

40

50

60

0 40 80 120 160 200

Out

let c

once

ntra

tion

(mol

/L)

Time (min)

Water

Ethanol

Theoretical

(b)

0

10

20

30

40

50

60

0 10 20 30 40 50 60

Out

let c

once

ntra

tion

(mol

/L)

Time (min)

waterethanolTheorical

Figure C.1 Breakthrough experiments: outlet concentration of ethanol and water as a

function of time; Q = 5 mL/min; T = 293.15 K; (a) water displacing ethanol; Bottom up flow direction; (b) ethanol displacing water; Top-down flow direction.

APPENDIX D. Breakthrough curves at 293.15 K D 2

(a)

0

2

4

6

8

10

12

14

16

18

0 40 80 120 160 200

Out

let c

once

ntra

tion

(mol

/L)

Time (min)

Ethyl lactate

Ethanol

Theoretical

(b)

0

2

4

6

8

10

12

14

16

18

0 10 20 30 40

Out

let c

once

ntra

tion

(mol

/L)

Time (min)

ethyl lactateethanolTheorical

Figure C.2 Breakthrough experiments: outlet concentration of ethanol and ethyl

lactate as a function of time; Q = 5 mL/min; T = 293.15 K; (a) ethyl lactate displacing ethanol; Bottom up flow direction (b) ethanol displacing ethyl lactate; Top-down flow direction.

APPENDIX D. Breakthrough curves at 293.15 K D 3

0

10

20

30

40

50

60

0 20 40 60 80 100

Out

let c

once

ntra

tion

(mol

/L)

Time (min)

Water

Lactic acid

Theoretical

Figure C.3 Breakthrough experiments: outlet concentration of water and lactic acid

as a function of time; Q = 5 mL/min; T = 293.15 K; Lactic acid displacing water; Bottom up flow direction.

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