Downstream processing and in situ product recovery ... · Downstream processing and in situ product recovery techniques of isopropanol and butanol production from biomass Diogo Miguel

Downstream processing and in situ product recovery techniques of

isopropanol and butanol production from biomass

Diogo Miguel Silva Felgueira

Thesis to obtain the Master of Science Degree in Chemical Engineering

Integrated Master in Chemical Engineering

Supervisors:

Prof. Dr. Filipe José da Cunha Monteiro Gama Freire

Dr. Vincent Coupard

Dra. Eszter Toth

Dra. Helena Gonzalez Penã

Examination Committee

Chairperson: Prof. Dr. Carlos Manuel Faria de Barros Henriques

Supervisor: Prof. Dr. Filipe José da Cunha Monteiro Gama Freire

Members of the Committee: Prof. Dr. Maria Cristina de Carvalho Silva Fernandes

October 2015

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Acknowledgements

Firstly, I would like to express my sincere gratitude to my internship supervisors at IFP Energies

Nouvelles, Dra. Eszter Toth and Dra. Helena Gonzalez Penã, as well as my project director Dr. Vincent

Coupard for all the help provided during my internship. Their patience, motivation and immense

knowledge was an invaluable contribute to this work. In fact, without their guidance I would dare

affirm this work couldn’t have been possibly. For that they truly have my most sincere gratitude.

I would also like to thank my IST supervisor Prof. Dr. Filipe Gama Freire for is supervision and

patience, especially in the writing of the report. Besides my advisor, I would like to thank the rest of

my thesis committee: Prof. Dr. Carlos Henriques and Prof. Dr. Maria Fernandes, for accepting the

invitation to my thesis committee and providing their time for its examination.

My sincere thanks also goes to all my friends that have supported me during this internship, as well

as during this years at IST. I’m grateful for having met such unique people and I wish them a

successful future, carrer and personal wise. My special gratitude goes to Ana, Catarina, David, João,

Joana, Solange and Tiago who have been my IFP buddies and also André, Ricardo, Filipa and Rita for

their most sincere friendship. Lastly I would also like to thank Bárbara and Catarina their deepest

support during this six months and the last years.

I thank all my family for their support and for believing in me, even when I do not. For this I specially

mention my brothers and my cousin for being the examples anyone could desire and I hope to have

made them proud.

Moreover I dedicate this work to my mother, for all the sacrifices she had to make and to my

stepfather, without which, all this efforts wouldn’t have been possible. They will always have my

deepest gratitude.

“Believe you can and you’re halfway there.”

Theodore Roosevelt

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Abstract

This works considers a preliminary study of butanol and isopropanol production from biomass.

Therefore the downstream recovery of the fermentation broth was evaluated while having a bigger

focus in in situ product recovery techniques to achieve a higher fermentation productivity as well as

more concentrated process streams.

For that purpose an analysis of the effect of achieving more concentrated fermentation broth in the

downstream recovery has been made and three different distillation schemes for this separation

have been developed. Regarding the in situ product recovery techniques, four different process

schemes were obtained by varying the separation method between gas stripping, liquid-liquid

extraction and condensation. Different solvents were used and for each a possible configuration for

its regeneration was evaluated, when conceivable.

The results obtained are promising but there is still much more work needed to achieve final results

for an economically feasible project. A higher understanding of the advantages of each technology

for in situ product recovery has been possible and some of the process optimizations indicate the

improvement that can still be made. From this works a preliminary choice of the in situ recovery

technology to focus further studies can be decided.

There are still many possible improvements to this work that may focus in further optimization of the

in situ recovery technologies, in the downstream recovery process and as well in the economic

evaluation of the project in terms of the CAPEX, OPEX and raw material costs reduction achievable.

Keywords: n-Butanol, Isopropanol, IBE fermentation, Biomass, Clostridium

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Resumo

Este trabalho consiste num estudo preliminário da produção de isopropanol e butanol através de

biomassa. Para tal foi feita uma breve análise do processo de separação e purificação dos álcoois do

meio de fermentação. No entanto este trabalho centra-se em técnicas de recuperação de produto

“in situ” para obter uma maior productividade e correntes de processo mais concentradas.

Três esquemas de distilação foram estudados para o processo de separação e purificação dos álcoois

de modo a avaliar o impacto de correntes de processo mais concentradas no consumo energético do

esquema. Relativamente a técnicas “in situ”, quatro esquemas de processo foram establecidos, entre

os quais variam as unidades de separação entre gas stripping, extacção líquido-líquido e

condensação. Deste modo, diferentes solventes foram ponderados e, para cada, um processo de

regeneração foi avaliado, quando possível.

Os resultados obtidos são promissores mas será necessário um estudo mais intenso de modo a obter

resultados finais sobre a viabilidade económica do projecto. Um estudo inicial das vantagens de cada

tecnologia de recuperação “in situ” foi feito e algumas optimizações de processo foram sugeridas.

Deste trabalho a escolha de um processo para um estudo mais profundo poderá ser decidida.

Existem ainda muitas melhorias possíveis para o projecto para uma melhor optimização do mesmo e

é também necessário obter valores de avaliação económica do projecto, relativamente a custos fixos

e custos operatórios.

Palavras chave: n-Butanol, Isopropanol, Fermentação IBE, Biomassa, Clostridium

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Table of Contents 1 Introduction ..................................................................................................................................... 1

1.1 Motivation ............................................................................................................................... 1

1.2 Objectives ................................................................................................................................ 2

1.3 Thesis outline .......................................................................................................................... 2

2 State of art ....................................................................................................................................... 3

2.1 Downstream recovery ............................................................................................................. 5

2.2 In situ/In stream recovery techniques .................................................................................... 7

2.2.1 Gas stripping .................................................................................................................. 10

2.2.2 Steam stripping ............................................................................................................. 11

2.2.3 Adsorption ..................................................................................................................... 11

2.2.4 Extraction liquid-liquid .................................................................................................. 12

3 Methodology ................................................................................................................................. 15

3.1 Downstream .......................................................................................................................... 15

3.1.1 DSLC ............................................................................................................................... 17

3.1.2 DSHC .............................................................................................................................. 18

3.1.3 DSDZ .............................................................................................................................. 19

3.2 Coupling techniques .............................................................................................................. 20

3.2.1 FCC ................................................................................................................................. 21

3.2.2 FWWC & FW5 ............................................................................................................... 27

3.2.3 FLLR ................................................................................................................................ 27

3.3 Solvent Regeneration ............................................................................................................ 29

3.3.1 2-butyl-octanol .............................................................................................................. 29

3.3.2 Vegetable oil .................................................................................................................. 32

3.4 Energy Consumption ............................................................................................................. 34

4 Results and Discussion .................................................................................................................. 35

4.1 Downstream .......................................................................................................................... 35

4.1.1 DSLC ............................................................................................................................... 35

4.1.2 DSHC .............................................................................................................................. 36

4.1.3 DSDZ .............................................................................................................................. 36

4.1.4 Comparison ................................................................................................................... 37

4.2 Fermenter concentration evolution ...................................................................................... 37

4.3 Gas Stripping ......................................................................................................................... 42

4.4 Extraction Methods ............................................................................................................... 42

4.4.1 FCC ................................................................................................................................. 43

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4.4.2 FWWC & FW5 ................................................................................................................ 43

4.4.3 Solvents ......................................................................................................................... 45

4.5 Schemes 1/2/3 optimization ................................................................................................. 46

4.5.1 Gas stripping .................................................................................................................. 46

4.5.2 Condenser temperature ................................................................................................ 47

4.6 FLLR........................................................................................................................................ 48

4.7 Final results ........................................................................................................................... 50

5 Conclusions .................................................................................................................................... 53

6 References ..................................................................................................................................... 55

7 Annexes ......................................................................................................................................... 58

7.1 Annex 1 - Constants ............................................................................................................... 59

7.2 Annex 2 – Simulation results ................................................................................................. 61

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List of figures Figure 1. Fermentation traditional process. ............................................................................................ 5

Figure 2. Schematic diagram of a two column distillation + decanter system for production of 99.9 wt

% 1-butanol from binary butanol-water mixtures [19]. .......................................................................... 6

Figure 3. Fermentation In situ recovery process [18]. ............................................................................ 7

Figure 4. Fermentation In stream recovery process [18]. ....................................................................... 8

Figure 5. Schematic diagram of a gas stripping process with product recovery by condensation. ...... 10

Figure 6. Schematic diagram of an extraction liquid-liquid process with solvent regeneration. ......... 12

Figure 7. In situ product recovery general process scheme. ................................................................. 15

Figure 8. Ternary diagram for butanol-ispropanol-water system (mol/mol), obtained in PROII using

SRKSV11 thermodynamic package. ....................................................................................................... 16

Figure 9. Distillation Scheme for low concentrations (DSLC). ............................................................... 17

Figure 10. Distillation Scheme for high concentrations (DSHC). ........................................................... 18

Figure 11. Distillation Scheme for the demixing zone (DSDZ). .............................................................. 19

Figure 12. Scheme for the first coupling technique process (FCC) ....................................................... 21

Figure 13. Scheme for fermenter mass balance. .................................................................................. 24

Figure 14. Scheme for the fourth coupling technique (FLLR)................................................................ 27

Figure 15. Ternary diagram for 2B1O- butanol -water system at 1 bar, obtained in PROII using UNIFAC

thermodynamic package. ...................................................................................................................... 30

Figure 16. Ternary diagram for 2B1O- butanol -water system at 0.01 bar, obtained in PROII using

UNIFAC thermodynamic package. ......................................................................................................... 31

Figure 17. Scheme for the 2BO1 regeneration by distillation. .............................................................. 32

Figure 18. Ternary diagram for Triolein- butanol -water system at 0.005 bar, obtained in PROII using

UNIFAC thermodynamic package. ......................................................................................................... 33

Figure 19. Scheme for vegetable oil regeneration. ............................................................................... 34

Figure 20. Graphic for the energy consumption in function of the feed compositions for the three

different distillation schemes. ............................................................................................................... 37

Figure 21. Curves of fermenter concentrations during the simulation for the FCC scheme. ............... 38

Figure 22. Curves of fermenter concentrations during the simulation in batch for the standard IBE

fermentation. ........................................................................................................................................ 39

Figure 23. Butanol production and recovery evolution during FCC fed batch simulation. ................... 40

Figure 24. Curves of fermenter concentrations during FCC fed batch simulation. ............................... 40

Figure 25. Batch process total energy consumption for FCC scheme. .................................................. 41

Figure 26. Fed batch process total energy consumption for FCC scheme. ........................................... 41

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Figure 27. Gas stripping extraction capacity for FCC scheme. .............................................................. 42

Figure 28. First condenser recovery capacity for FCC scheme. ............................................................. 43

Figure 29. First condenser recovery capacity for FCC scheme. ............................................................. 43

Figure 30. Recovery capacity for a standard FWWC scheme process unit. .......................................... 44

Figure 31. Fermenter concentration evolution in the organic phase. .................................................. 49

Figure 32. Fermenter concentration evolution in the aqueous phase. ................................................ 49

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List of tables

Table 1. Energy requirement of butanol recovery systems (N/A: not available data) ......................... 14

Table 2. Initial parameter conditions for coupling techniques simulation. .......................................... 20

Table 3. Work conditions for the condenser in FCC process scheme. .................................................. 26

Table 4. Results obtained for the energy consumptions of the DSLC. .................................................. 35

Table 5. Results obtained for the energy consumptions of the DSHC. ................................................. 36

Table 6. Results obtained for the energy consumptions of the DSHC. ................................................. 36

Table 7. Results for total recovery factors in for scheme FWWC & FW5 for both alternatives. .......... 44

Table 8. Results for 2BO1 regeneration process. .................................................................................. 45

Table 9. Results for vegetable oil regeneration process. ...................................................................... 46

Table 10. Results for FCC process scheme for different gas stripping initiation.. ................................. 47

Table 11. Results for FWWC process scheme for different condenser temperature.. ......................... 48

Table 12. Results for FLLR process scheme. .......................................................................................... 50

Table 13. Results for fed batch process schemes. ................................................................................ 51

Table 14. Results for batch process schemes. ....................................................................................... 52

Table A. 1. Kinetic constants used for the fermentation model ........................................................... 59

Table A. 2. Inhibition function parameters values. ............................................................................... 59

Table A. 3. Values for the Sig function parameters. .............................................................................. 60

Table A. 4. Values for mass transfer coefficients for each solvent used. .............................................. 60

Table A. 5. Property values considered in the model. .......................................................................... 60

Table A. 6. General simulation results for the standard batch and in situ recovery techniques for the

FCC process scheme. ............................................................................................................................. 61

Table A. 7. General simulation results for the FWWC process scheme. ............................................... 63

Table A. 8. General simulation results for the FW5 process scheme .................................................... 64

Table A. 9. General simulation results for the FLLR process scheme .................................................... 65

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Nomenclature 𝐴 Area [𝑚2]

𝐶 Concentration [𝑔 𝑙⁄ ]

C1 Condenser 1

C2 Condenser 2

𝐶𝑚𝑎𝑥 Concentration of inhibition [𝑔 𝑙⁄ ]

𝐶𝐴𝐵,𝑐𝑟𝑖𝑡𝑖𝑐 Butyric acid critic concentration [𝑔 𝑙⁄ ]

𝐶𝐵𝑢𝑂𝐻𝑏 Butanol concentration to begin extraction [𝑔 𝑙⁄ ]

𝐺 Gas mass [𝑔]

ℎ𝑜𝑙𝑑 − 𝑢𝑝 Gas hold-up

𝑖𝑛𝑏 Inhibition function

𝑘𝑖 Kinetic parameter where i ranges between 1 and 12 [ℎ−1]

𝑘 Kinetic parameter for compounds mass transfer [(𝑔𝑜𝑟𝑔 𝑙𝑜𝑟𝑔⁄ )/(𝑔𝑎𝑞 𝑙𝑎𝑞⁄ )]

𝐾 Kinetic parameter [𝑔 𝑙⁄ ]

𝐾𝐺 Kinetic parameter for global mass transfer [ℎ−1]

𝐾𝑆 Affinity constant for the carbonic subtract [𝑔 𝑙⁄ ]

𝑙 Liquid mass [𝑔]

𝑚 Inhibition degree parameter

𝑀𝑔𝑎𝑠 Renewal gas mass [𝑔]

𝑀𝑀 Molecular weight [𝑔/𝑚𝑜𝑙]

𝑄𝐸𝑥𝑡 Solvent flowrate [𝑚3 ℎ⁄ ]

𝑄𝑔𝑎𝑠 Gas flowrate [𝑘𝑔 ℎ⁄ ]

𝑟 Solvent and gas ratio in wash stage entry

𝑠 Solvent mass [𝑔]

𝑠𝑖𝑔 Sig function

𝑆 Substrate concentration [𝑔 𝑙⁄ ]

𝑣𝑔𝑎𝑠 Gas velocity [𝑚 𝑠⁄ ]

𝑉 Volume [𝑚3]

𝑉𝑓𝑒𝑟𝑚0 Initial fermenter volume [𝑚3]

𝑡 Time [𝑠]

𝑇 Total mass [𝑔]

𝑥 Liquid fraction

𝑋 Biomass concentration [𝑔 𝑙⁄ ]

𝑦 Gas fraction

𝑧 Solvent and organic phase volume ratio in the fermeter

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Greek

𝜌𝑤 Water density [𝑔 𝑚3⁄ ]

µ Specific velocity of microorganisms production [ℎ−1]

µ𝑚𝑎𝑥 Maximum velocity of microorganisms production

𝜙 Fixed parameter for sig function

Indices

𝐴𝐴 Acetic acid

𝐴𝐵 Butyric acid

𝐵𝑢𝑂𝐻 Butanol

𝑐𝑜𝑛𝑑 Condenser

𝐸𝑡𝑂𝐻 Ethanol

𝑓𝑒𝑟𝑚 Fermenter

𝑖 Actual time step

𝑖 + 1 Next time step

𝐼𝑝𝑂𝐻 Isopropanol

𝑁2 Nitrogen

𝑠𝑜𝑙𝑣 Solvent

𝑤 Water

Exponents

𝑎𝑞 Aqueous phase

𝑐𝑜𝑛𝑑 Relative to condensate

𝑖𝑛 Inlet

𝑜𝑟𝑔 Organic

𝑜𝑢𝑡 Outlet

𝑟𝑒𝑐 Relative to initial fermenter volume

𝑤𝑎𝑠ℎ Washing unit

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1 Introduction

This report describes the work done at IFP energies nouvelles under the subject of downstream

processing and in situ product recovery techniques of isopropanol and butanol production from

biomass. This work shall therefore include a brief analysis of the effect of concentrating the

fermentation product on its energy consumption for separation as well as a simplified model for in

situ product recovery process schemes. In the end the results obtained shall be compared between

the different schemes.

1.1 Motivation

Nowadays petroleum and other fossil fuels are the main source of energy. From them, petrochemical

industry supplies a wide range of products in sectors of automotive fuels, plastic material, rubbers,

cosmetics and pharmaceutics. Consequently the depletion of cheap fossil fuels and their negative

impact on the environment have been motivating future work on the use of sustainable sources of

energy and of bulk chemicals.

As a sustainable source of energy and bulk chemicals, biomass has become increasingly more

attractive. In fact in the last decades the use of biofuels for transportation from biomass has become

much more relevant.

Therefore several work has been done to identify methods to substitute current fossil fuel demands

by alcohols produced by biomass fermentation by microorganism. Most of the investigation focus in

the production of biobutanol by ABE (Acetone-Butanol-Ethanol) fermentation due to its easiness to

be blended with gasoline and having more interesting properties than ethanol

Nevertheless in this work a new alternative, the IBE (Isopropanol-Butanol-Ethanol) fermentation, will

be explored to produce not only butanol but also isopropanol. Isopropanol has a bigger market value

than acetone, being used in many industrial processes as a solvent and can also be blended with

gasoline.

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1.2 Objectives

An economical study of the IBE fermentation process and downstream processing is underway at

IFP Energy Nouvelles to improve the process results and obtain a well-structured and economically

feasible project. Therefore the aim of this work is to make preliminary calculations for the energy

process consumption and process streams. The calculations were implemented in VBA (Visual basic

for applications).

This work will then be useful for further development of the process, as also for choosing which in

situ product recovery technologies to be focused on in order to achieve better process parameters at

lower project cost.

1.3 Thesis outline

This report has been divided in three sections. Firstly there is the State of Art, where can be found a

brief description of the use of butanol and isopropanol, the IBE fermentation basics, the traditional

IBE fermentation process and the most important constraints to the project. Then the fermentation

broth separation, the downstream recovery, is detailed, as also the increasing interest in achieving

higher fermentation broth concentrations. Lastly the most important difference in situ/in stream

product recovery techniques are described.

Henceforward, in the second section, the methods applied for the downstream processing and in situ

product recovery techniques are explained for different process schemes. For that, solvents may be

used and therefore there is also a brief methodology on the solvent regeneration. Lastly, equations

for energy consumption are explained.

In a third section, the results are presented and discussed by means of energy consumption,

concentration and extraction profiles as well as recurring to comparison by final results tables.

The report ends with a final conclusion where the remaining improvements are integrated and the

most important results are highlighted.

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2 State of art

The current reduction of petroleum resources and the impact to the environment these derived fuels

represent have culminated into an increasing interest in the use of biofuels. This biofuels should have

a lower process cost, be produced in large quantities from a renewable source and be able to be

directly used in existing motor vehicles and inserted in the petroleum distribution system. The

increase of usage of biofuels is also expected to reduce the oil market supply instabilities and the

greenhouse gas emissions [1].

As petroleum substitutes are included, as some of the most promising, isopropanol and butanol due

to being more advanced biofuels then ethanol, having bigger carbon chains and replacing better

gasoline.

Butanol is a great candidate because it can be obtainable from renewable sources, such as rice straw

and corn fiber, further it also as interesting properties, having an average energy density of 30%

more than ethanol, its less soluble in water, has a lower volatility, lower corrosivity and can be

blended with gasoline in any desired proportion. It can also be used as solvent for the anibiotics,

vitamins and hormones production, but most of it is used in the production of butyl acrylate and

methacrylate esters used in latex surfaces, coating, enamels and lacquers [2], [3], [4].

Isopropanol is also an important candidate, it is used in many industrial processes as solvent, oil,

lacquers and cosmetic products or as antiseptic agent and disinfectant for home, hospital and

industries, [5]. In the biofuel market it can be used as precursor of propylene, used for plastics

manufacture, and, with dehydration, to form Di-isopropyl ether (DIPE) which has been proposed by

Knifton et Day (1999) as a substitute for Methyl-tert-butyl-ether (MTBE), [6], an oxygenated additive

in gasoline blending to enhance the gasoline RON. Also it can be blended to gasoline due to its

heating value and other properties [2].

The process ABE (Acetone-Butanol-Etanol) is the conventional fermentation process for butanol

production and has been roughly studied in the last years, but this study will focus in the IBE process

(Isopropanol-Butanol-Ethanol) which is very similar to the ABE process but since it also produces

isopropanol instead of acetone it has an higher product value since it can be used directly as a

biofuel, which isn’t possible with acetone, being it considered un-desirable in the ABE process for its

poor fuel properties and corrosiveness to rubber engine parts [7].

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Both this fermentation processes are performed by microorganisms from the genus Clostridium and

the metabolic pathways are similar, also the ABE process is older and therefore more extensively

studied. Most known solventogenic microorganisms produce either acetone or isopropanol or

butanol, even so, the first report of butanol production by Clostidium bacteria was in 1861 by Louis

Pasteur, in 1905, Schardinger reported the production of acetone by Bacillus macerans and only in

1906 was the isopropanol production reported by Pringsheim [5].

ABE and IBE fermentations are known as solventogenic fermentations since its products are used as

solvents in various industrial processes. In fact ABE fermentation was almost uninterruptedly in

operation between 1910 and 1960 and the first plants were built in United Kingdom. Contrary to

nowadays the desired product from ABE fermentation during the two World Wars was acetone to be

used in the ammunitions manufacture. Butanol only started being desired with the development of

the automobile market, used in the manufacture of synthetic rubber and fast-dry car lacquer. While

this, IBE fermentation was being developed in Taiwan, but after the Second World War the

petrochemical industry developed and this fermentations market began to decline, ceasing to exist in

the beginning of the sixties in USA and Great Britain, while in Russia and China large scale plants

existed until the mid-1990s. Nowadays the ABE process as began to be introduced in China, Brazil

and South Africa [5].

Clostridium bacteria’s can use different carbon sources as fermentation substrates, including glucose,

xylose, cellobiose, sucrose and many others. Furthermore this bacteria’s have successfully produced

solvents from Jerusalem artichoke [8], liquefied corn starch [9] and starched package peanut [10],

[11].

The interest in studying the conversion of lignocellulosic biomass into biofuels by fermentation has

been increasing recently due to lignocellulosic biomass being one of the most abundant renewable

resources. This resource has a heterogeneous structure, being composed by cellulose (mainly

glucose), hemicellulose (glucose and pentoses) and lignin [12].A hydrolysis of the hemicellulose prior

to the fermentation is advised, obtaining a sugar solution of manly glucose and xylose and since

clostridium can ferment this substrates to produce solvents, these microorganisms have been gaining

more attention. Therefore there has been an extensively investigation in the production of solvents

from this lignocellulosic hydrolysates [13], [14], [15], [16] and [17].

An important characteristic of solvent-producing fermentations is the metabolic changes. Both ABE

and IBE fermentations show two phases, the acidogenesis during which acetic acid and butyric acid

are produced and the solventogenesis, the solvents production phase. During acidogenesis hydrogen

and carbon dioxide are also produced and with the acid increase the pH drops switching the

5

metabolism mechanic to the solventogenesis, where the acids are used as electron acceptors and

reassimilated by the bacteria to produce solvents. This dependency between both metabolisms

mean that as the solvents concentration increases lower will be the acid concentration [5].

The major setbacks for the production of biobutanol are due to the process high cost and the low

productivity, the fermentation inhibition threefold of butanol is very low, limiting its production to

usually only 2% (w/v) from the fermentation broth. The high process cost is mainly due to the high

energy required for the butanol recovery, but also because of the substrate cost and the low solvent

concentration [18].

The key for achieving an economical production is to increase the butanol concentration and

productivity, for this, studies focused in strain development, using inexpensive renewable

feedstocks, have been made.

The traditional way of ABE and IBE fermentation is “straightforward”. As shown in Figure 1, in the

fermenter the substrate is converted into the desired product until the critical concentration

(inhibition threshold) is reached or its own depletion (batch operation). Subsequently, the product is

recovered from the fermentation broth in a downstream recovery unit, usually composed of a train

of distillation towers. This primary alcohol removal is also called “end of pipe” removal.

Figure 1. Fermentation traditional process.

2.1 Downstream recovery

The separation of alcohol from the fermentation broth is normally done by distillation which is manly

controlled by the VLE (vapor-liquid equilibrium).

A way to measure the capability of the separation using VLE based techniques is the relative

volatility, since it is defined as the ratio between the vapor-phase concentrations ratio of the organic

compound and water and the liquid-phase concentrations ratio of the organic compound and water.

6

The distillation scheme efficiency can be measured as the energy necessary to separate the

compound with the desired composition per mass unit of alcohol recovered. To evaluate the viability

of the separation, its required energy may be compared with the product fuel value since it is an

alcohol. Also the normal behavior in a distillation scheme is to have an increase in energy required

with a decrease of alcohol purity in the feed stream [19].

ABE and IBE fermentations would have more promising results if the alcohols concentration by the

end of the batch cycle were higher. In fact the alcohol recovery by distillation from a fermentation

broth that dilute is highly energy intensive at low alcohol concentrations.

Indeed, butanol and water form an azeotrope at 55 % wt of butanol in the liquid phase at 1 atm.

Nevertheless due to the fact the solubility of butanol in water is around 7.7 % wt, two liquid phases

form at the azeotrope. The organic phase can contains about 80 % wt of butanol, while the aqueous

phase will be saturated in butanol, approximately 7.7 % wt, [19]. Therefore for the separation of

butanol an additional column will be necessary, following the distillation scheme in Figure 2.

Figure 2. Schematic diagram of a two column distillation + decanter system for production of 99.9 wt % 1-butanol from binary butanol-water mixtures [19].

7

Actually the butanol concentration during fermentation never reaches more than 1.5% wt and

consequently the high energy cost associated to the product recovery from such a low concentrated

broth turns the process economically unviable. Furthermore, associated to this large product

dilution, a huge volume of wastewater is generated. As that, there is a high process cost investment

due to the wastewater treatment.

It has been reported [19] that the energy required for producing 99% wt butanol using the previous

two-columns and decanter system from an aqueous stream at 20% wt butanol could decrease from

more than 30 to 4 MJ/ kg butanol produced, hence the energy consumption per unit of alcohol

recovered decreases as the concentration of alcohol at the end of the fermentation increases.

Therefore alternative technologies that are less energy expensive at low alcohol concentrations

would an interesting strategy to turn ABE and IBE fermentation in an economically viable process.

This could be achieved by applying an efficient technique that would allow the systematic removal of

the alcohols at a rate as fast as it is produced in the fermentation, improving the fermentation yields

and productivities in the reactor.

2.2 In situ/In stream recovery techniques

Optimization of the process has also been investigated with the use of in situ processes to remove

butanol. This technique consists in the removal of selective reaction products during fermentation.

Different approaches exist to develop an integrated product recovery system. This system consists of

a fermentation unit coupled to a product separation unit [20].

Two different setups can be designed:

- In situ recovery: the concentration step occurs inside the bioreactor, where the product is partially

separated, Figure 3. The alcohol-depleted fermentation broth never leaves the bioreactor.

Figure 3. Fermentation In situ recovery process [18].

8

- In stream recovery: fermentation and primary separation are carried out concomitantly but in two

distinct vessels. This implies the continuous pumping of the fermentation broth though another

column containing the selective phase or material, which allows separating a fraction of the product,

Figure 4. The alcohol – depleted stream is returned to the bioreactor.

This can be an advantage in different ways, like enriching the end product concentration leading to

lower downstream processing costs, improve the productivity by removal of the inhibition product,

reduce the process flows since it achieves higher product concentrations and for last it also increases

the product yield, reducing side reactions by the product removal [2].

In both previous process configurations, there is a decrease in the toxic effect of the alcohol on

bacterial cells due to a continuous removal of the product from the fermentation broth. In fact, if the

butanol is recovered as fast as it is produced the fermentation inhibition can be avoided and

therefore the microbial culture can achieve a higher product yield and productivity. Furthermore, by

alleviating product toxicity, higher substrate concentrations in the fermentation broth can be

considered (possibly in a fed-batch operation).

The advantages and disadvantages of both configurations are in fact according to the type of the

separation unit used, between adsorption, liquid-liquid extraction, stripping and more.

The integrated product recovery techniques of the alcohols from aqueous broth can be based on the

difference between physical or chemical properties of the different alcohols and water or on their

interaction with an auxiliary agent or material.

The end product enrichment depends on the selectivity of the ISPR (in situ process removal)

technique but processes with lower separation factors can still obtain similar or higher product

concentration in the concentrate, if operated at higher residual products concentrations. Therefore

the concentrate concentration is strongly connected to the product concentration in the

fermentation broth [2].

Figure 4. Fermentation In stream recovery process [18].

9

Many previous investigations using gas stripping had disappointing results since the butanol

concentration in the fermenter decrease to < 5 g.L-1 but according to Xue et al. (2012), [3], only by

controlling the gas stripping device so that it would only be “on” while the butanol concentration in

the fermenter was between 8 and 13 g.L-1 obtained concentrations of 151 g.L-1 of butanol, reaching a

total of 196 g.L-1 of ABE.

However as many process parameters are interrelated, in this case this increase in concentration will

lead to a lower solvent productivity than if it other ISPR techniques were used, as pervaporation [21].

The removal of inhibitory product will also increase the productivity since it relieves the cells

inhibition and therefore leads to a higher cell growth and consequentially a faster product formation

rate. The productivity from an ABE fermentation without ISPR is normally 0.53 g.L-1.h-1, [22] but

according to [21], a productivity of 1.16 g.L-1.h-1 can be obtained using gas stripping and applying a

fed-batch, also values as high as 2.5 were reported by Bankar et al (2012) with the combined effect

of cell immobilization and extraction, [23].

The process flow reduction is achieved by the utilization of more concentrated substrate. In

fermentations severely product-inhibited, like ABE and IBE, there is a tendency to use dilute

substrate solutions since production concentrations as low as 2 % wt can inhibit the reaction. Ezeji et

al. (2003) as reported the use of substrate concentrations up to 162 g.L-1 for ABE fermentation, [24],

and even the use of a fed-batch as concentrated as 500 g.L-1 in 2004, [25]. This higher substrate

concentrations will lead to less dilute product flows and therefore decreases the energy needed in

the downstream product separation.

For last, the yield defined as the total product produced dividing by the substrate consumed, can also

be improved with ISPR techniques since a higher concentration feed allows a greater turnover to

product than biomass formation, relives the microorganism stress by the inhibitory component and

decreases the intercellular maintenance. Yield enhancements as high as 18% have been

experimentally reported [24], [25], [26].

Among the most promising alcohol recovery techniques are gas stripping, steam stripping,

adsorption, liquid-liquid extraction and pervaporation.

10

2.2.1 Gas stripping

The option to use an inert gas as the stripping agent and operate at fermenter temperatures

contributes to turn this technique appealing because of its relative simplicity and incapacity to harm

the fermentation culture.

The principle of the process is the stripping agent injection at the fermenter base, parallel to a

countercurrent gas-liquid contactor. The transference of butanol is controlled by the vapor-liquid

equilibrium, so vapor pressure is an important parameter which is highly sensitive to the fermenter

temperature, resulting in a significantly increase of alcohol concentration in the gas after a

temperature mild increase, [19].

To recover the butanol in the gas phase a condenser is used to remove a richer liquid phase, [27],

[28], different units could also be used, such as adsorption, liquid-liquid extraction and membrane

separation, [29], [30]. A multi-step or fractional condensation system could also be used.

The alcohol concentration in the recycle gas must be maintained low enough to effectuate the

butanol removal required. The evaporation and condensation of water is a major parameter since its

energy requirement is 3.86 times that of butanol, [19], so the water title has to be carefully analyzed.

A possible solution is to condense the majority of water at higher temperatures and use a second

condenser at lower temperatures to condensate the alchools.

In the literature, [28], [24], [27], [15], [9] there are several examples of this study that assure gas

stripping as a possible in situ recovery technique, since the volatile products are removed selectively,

eliminating toxicity effects by them, higher sugar utilization of concentrate sugar substrates, greater

biomass concentration in the fermenter and enhanced productivity in overall.

Figure 5. Schematic diagram of a gas stripping process with product recovery by condensation.

11

2.2.2 Steam stripping

Steam stripping is very similar to gas stripping being the major difference the use of water vapor

instead of an inert gas. As in gas stripping the alcohol removed by steam is recovered in

condensation units but using steam the condenser operates at higher temperatures and complete

condensation of the vapor is possible.

By steam stripping energy usage may be reduced 32% relative to gas stripping, [19].

2.2.3 Adsorption

In adsorption there is a solid adsorbent material that will adsorb the alcohols in the fermenter.

Normally the extractant is contained in a packed-column contactor, serving as both adsorber and

desorber but the solid extractant is not moved through the system resulting in a cyclic process of

loading and unloading the adsorbent.

This solid must exhibit a sorption selectivity for alcohol greater than for water and high distribution

coefficients for sorption equilibrium resulting in a high separation factor, [19].

Hydrophobic zeolites are the most widely studied class of adsorbent for alcohol recovery, typically

with a ZSM-5 structure and a high silicon-aluminum elemental ratio, being, silicate-1, the best graded

alcohol-selective zeolite, containing no aluminum, [31], [32].

Polymeric resins as XAD resins and polyvinylpyridine and activated carbon are among other alcohol

selective adsorbents studied, [31], [33].

Silicate-1 can achieve, at room temperatures, a separation factor of 76 for ethanol, being even more

selective for butanol. However it’s important to keep in mind that the regeneration stream won’t

have an alcohol concentration as high, lowering the ethanol separation factor to somewhere like 27,

[34]. A possible explication is the existence of a liquid hold-up between particles within the sorbent

bed, this inefficiency can be reduced with a drying step prior to the regeneration phase, [34].

Nevertheless there will always exist some liquid in the bed and since this liquid is very dilute in

alcohol, it also leads to more evaporation energy spent.

Other worries with this technique include the fouling of the sorbent bed by cells and cell debris and

also the adsorption of other broth components, such as nutrients and the substrate itself.

12

2.2.4 Extraction liquid-liquid

For liquid-liquid extraction a liquid extractant is placed in contact with the fermentation broth so that

the alcohol is transferred to the extractant liquid, by mass transfer and as in the previous technology,

there must be a regeneration unit to remove the alcohols from the extractant liquid.

Figure 6. Schematic diagram of an extraction liquid-liquid process with solvent regeneration.

An extractant should have a selectivity for alcohol relative to water as high as possible, resulting in a

greater separation factor and a more concentrated product stream. The equilibrium distribution

coefficient is also important since it defines the extractant quantity necessary to remove the alcohol

desired.

Preferably there wouldn’t be any solubility between the extractant and the water, however it’s

expected to lose some extractant soluble in water since the extractant has a high affinity for alcohols,

which as to be carefully analyzed since it may create complications to the bioreactor, the products

and even the wastewater treatment facility. It should also be easily separated from the aqueous

phase, normally by gravity settling or continuous centrifuge, [35], [36].

To improve mass transfer in the extraction and regeneration phases a low extractant viscosity is

needed and according to the alcohol recovery process, a low or medium volatility may be required.

Fermenter

2nd Liquid

1st Liquid

Regen

eration

un

it Alcohols

13

A thoughtfully report on the interaction between the extractant and the bioreactor components

must be made, as also for the toxicity that it may represent in wastewater and air releases. Other

security criteria’s must also be achieved to ensure the process safety, as flammability, flash point and

reactivity with other process compounds, [35], [36].

For last, the cost of the liquid extractant must also be analyzed since, due to process losses, some

fresh addition will be needed periodically.

For the regeneration step the most common methods are: vacuum flash vaporization, distillation, gas

stripping and membrane-based pervaporation, also normally there is a condensation step and in the

case of butanol a decanter operation. In order to achieve the higher regeneration possible the

distribution coefficient for alcohol between a gas phase and the liquid must increase with increasing

temperature so that the tendency of the alcohol to leave the extractant phase is greater, [37].

The regeneration phase must be efficient enough to at least reduce the alcohol concentration to the

value which would be in equilibrium with the treated liquid concentration. In this way the extraction

a regeneration efficiencies are deeply related and as greater the efficiency, bigger the energy

demand and capital costs.

An advantage from this technology is the little thermal energy involved since in the extraction both

phases are liquid and the separation occurs by mass transfer, allowing to be used at fermenter

temperatures. The energy demand is due to the extractant regeneration phase and as a first

approximation may be considered as the energy required to evaporate the alcohol and water from

the extractant. The alcohol distribution coefficient will also affect the energy demand since higher

coefficients means a lower extractant-feed ratio, [19].

To sum up, ISPR systems may allow a more concentrated stream as feed for the distillation train, thus

reducing energy consumptions as well as investment and operational costs. Furthermore these

techniques can present some other advantages over the traditional “straight forward” scheme. This

may be the systematic removal of product which reduces microbial inhibition, and thus improves

solvent yield and productivity and the fermentation broth is more concentrated, generating less

wastewater.

14

Table 1. Energy requirement of butanol recovery systems (N/A: not available data)

Recovery system MJ/kg [38] MJ/kg [31] MJ/kg [39]

L-L extraction 8.8 7.7 N/A

Pervaporation 13.7 N/A 9

Gas Stripping 21.7 31 21

Adsorption 8.1 1.3 33

Steam stripping+distillation 24.1 N/A N/A

Among the several existing recovery techniques a system consisting in the fermentation coupled to

an extraction unit will be studied. In extractive fermentations, the end-product inhibition is reduced

by extracting toxic fermentation products out of the aqueous broth into an external agent.

15

3 Methodology

The general process scheme of this work aims the use of an in situ recovery technique that shall

extract the alcohols as long as they are produced. In addition, as shown in Figure 7, an extracting

agent shall be used, from the regeneration of this agent the alcohols will be recovered and

afterwards go through the traditional downstream recovery, a distillation train. Therefore in the

process there will exist different alcohols streams. They are, the fermenter broth remaining at the

end of the simulation and all the streams from the extracting agent regeneration, all this streams

shall be sent to downstream recovery units for further alcohol purification.

3.1 Downstream

PROII uses a reference state that differs accordingly to the method of units inputted. Using the SI

system the reference state is a temperature of 273.15 K and a pressure of 101.33 KPa.

The thermodynamic package used for the simulation is internal to IFPEN libraries and properly

developed for this compounds mixture (SRKSV11) and has integrated vapor-liquid equilibrium but

also liquid-liquid.

Regarding the simulation components, only IBE and water are considered in the columns scheme

feed.

Figure 7. In situ product recovery general process scheme.

16

Previous to the simulations, a study of the thermodynamic fluid phase equilibrium for the major

three compounds (water, butanol and isopropanol) was made recurring to an SIMSCI tool to draw

the ternary diagram. This component uses a PROII file, therefore recurring to the PROII

thermodynamics data it calculates the azeotropes of the mixture and traces its ternary diagram.

An analysis of the compound mixture thermodynamics was done assuming that ethanol fraction is

low enough to be excluded, where the gray line corresponds to the normal operating line of IBE

fermentation, that is a proportion BuOH:IpOH:EtOH of 6:3:1.

From the diagram it was concluded the existence of a two liquid phase’s zone where a possible

separation of the organic and aqueous phase may be advantageous. Also there are two azeotropes in

the system, butanol-water and isopropanol-water.

Thus from this diagram the feed concentrations were chosen to analyze the optimal distillation

scheme for that feed point, obtaining accordingly to the feed composition three different distillation

schemes. Also, all feed points analyzed are situated in the gray line, therefore assuming a normal

fermentation ratio between butanol and isopropanol.

Figure 8. Ternary diagram for butanol-ispropanol-water system (mol/mol), obtained in PROII using SRKSV11 thermodynamic package.

17

3.1.1 DSLC

This first scheme is the traditional downstream distillation of an IBE fermentation broth,

recommended for low concentration feedstocks and therefore forwardly called DSLC (Distillation

scheme for low concentrations). In a first step, a beer column is needed to remove most of the

water, corresponding also to the main OPEX cost. In a second column a commercial concentrated

solution in Isopropanol and Ethanol, averagely 800 g/l, can be obtained at the top, from the bottom

flows the rest of water and the butanol and due to the heteroazeotrope they form, have to be

separated in a heteroazeotropic column system.

The specs for this column scheme are:

Maximum of 100 ppm of Ethanol at beer column, column A, bottom outlet, since ethanol is

the component with higher solubility;

A recovery of at least 99.99% of isopropanol at the B column top outlet, while recovering

99.9% of the butanol at the bottom;

Maximum of 100 ppm of Butanol at the heteroazeotropic aqueous phase column C bottom

outlet;

99.9% Butanol at the heteroazeotropic organic phase column D bottom outlet.

Feed

Water

Isopropanol

Ethanol

Water

Butanol Water

Bee

r C

olu

mn

A B C D

Figure 9. Distillation Scheme for low concentrations (DSLC).

18

3.1.2 DSHC

Considering the possibility of achieving a much higher concentrated feed by coupling techniques in

the fermenter, a new distillation column arrangement may be obtained by suppressing the beer

column of the previous scheme, thus called the DSHC (Distillation scheme for high concentrations).

This is possible since the feed is further concentrated and so there is a lower water composition in

the stream. Therefore in the first column the commercial high concentrated solution in Isopropanol

and Ethanol can be obtained while a heteroazeotropic column system deals with the butanol and

water separation. The suppression of a column and the reduced water fraction in the feed may

contribute to a lower CAPEX and OPEX for this arrangement.


Pressure of 3 bar at column A to avoid two liquid phases distillation since the feed being

much more concentrated, has already an aqueous phase (saturated) and an organic phase;

A recovery of at least 99.99% of isopropanol at the column A top outlet, while recovering

99.9% of the butanol at the bottom;

Maximum of 100 ppm of Butanol at the heteroazeotropic aqueous phase column C bottom

outlet;

99.9% Butanol at the heteroazeotropic organic phase column D bottom outlet.

Figure 10. Distillation Scheme for high concentrations (DSHC).

Fee

d

Isopropanol

Ethanol

Water

Butano

l

Water

B C D

19

3.1.3 DSDZ

This arrangement can be used for dilute concentrations and even in the demixing ternary zone,

therefore referred as the distillation scheme for the demixing zone (DSDZ). By adding a zone of

liquid-liquid distillation in the second column and extracting the water rich heteroazeotrope phase

from the middle of the tower a compacted scheme is obtained. This water rich phase has a higher

organic concentration than the first column feed and therefore by adding it to the column, it will

decrease its OPEX cost.

With this arrangement the same streams from the other schemes are obtained from a low

concentration feed and with a possible lower energy consumption, but while adding complexity to

the distillation process, since the choice of the extracting plate will change the number of plates

under two liquid phases, and an increase in the tower CAPEX cost due to the need for internal plates

prepared to deal with a two liquid phase, but it also reduces the distillation scheme to only two

distillation columns.

Figure 11. Distillation Scheme for the demixing zone (DSDZ).

Feed

Water

Isopropanol

Ethanol

Water

Butanol

Bee

r C

olu

mn

A B

20


0.01% Butanol at the column B top outlet;

Existence of an extraction plate to remove a second liquid phase with an organics

concentration higher than fond in the feed.

0.01% Water at the column B bottom outlet.

3.2 Coupling techniques

As in situ alcohol removal techniques, the exploration of gas stripping and liquid-liquid extraction was

made. For that, 4 different process diagrams were studied with fixed parameters and a batch and fed

batch case were studied.

For each process diagram, a dynamic simulation was obtained in order to compare energy

consumption and product improvement over a common basis. For this, a simplified fermentation

kinetics model was used based on ABE fermentation literature. This model, whose aim is the dynamic

prediction of each metabolite evolution (except from fermentation gas) is developed in the next

pages and describes both phases of IBE fermentation (see section “State of Art“).

Recovery of alcohol thermodynamics and its energy consumption was achieved by using a

commercial software, Simulis by Prosim, FR, which as a developed thermodynamics package for IBE

components. In this software, just as Aspen or PROII, the main input is selecting the process

compounds, after that, Simulis can be called in excel recurring to a Simulis supplement called

“Calculators”. This “Calculators”, according to their name, can be used as excel functions to calculate

various thermodynamics calculations and therefore preview the process simulation.

To all simulation files the fermenter initial substrate, sugar, concentration was of 90 g/l and

fermenter with a volume of 500 m3 and an initial working volume of 400 m3. Every fixed conditions

are accordingly to the following table.

Table 2. Initial parameter conditions for coupling techniques simulation.

Initial Conditions

Substrate concentration (g/l) 90

Fermenter working volume (m3) 400

Biomass concentration (g/l) 0.1

Acid acetic concentration (g/l) 2.5

21

Batch dynamic simulations end when the substrate concentration reaches zero since by then the

alcohol production will no longer continue, except for remaining acids conversion.

For the fed batch case study, for comparison reasons, a fixed substrate consumption was made as

the substrate addition equivalent to a concentration of 500 g/l referred to the initial fermenter

volume. This substrate addition is activated when the sugar concentration in the fermenter reaches

10 g/l during the dynamic simulation and, in process terms, is made as an addition of an highly sugar

concentrated solution (500 g/l) in the fermenter to maintain a minimum of 10 g/l in the fermenter.

This feed rate was fixed as 5% of the initial fermenter volume.

Lastly general assumptions have also been made:

No CO2 formation;

No acids, biomass or substrate transfer to the gas or solvent;

No heat loss in the process;

No heat of reaction in the fermenter;

No formation entalpies;

No solution entalpies.

3.2.1 FCC

The first arrangement studied was the application of gas stripping in the fermenter using two

condensers in series to recover the alcohols stripped from the gas, denominated as FCC

(Fermenter-Condenser-Condenser). For this modelization nitrogen was used as the stripping media

while for the reaction kinetics, parameters from previous experimental kinetics studies were

Figure 12. Scheme for the first coupling technique process (FCC)

F

C1 C2 F – Fermenter

C - Condenser

22

available. With this the fermentation model could be easily constructed by writing the code for mass

balance.

The first step reaction is the formation of the butyl and acetic acids by sugar conversion.

𝑑𝑆

𝑑𝑡= −𝑘1µ𝑋−𝑘2

𝑆

𝑆 + 𝐾𝑆

𝐶𝐴𝐵

𝐶𝐴𝐵 + 𝐾𝐴𝐵𝑋−𝑘3

𝑆

𝑆 + 𝐾𝑆

𝐶𝐴𝐴

𝐶𝐴𝐴 + 𝐾𝐴𝐴𝑋 Equation 3. 1

Where 𝑆 is the sugar concentration in the fermenter (g/l), 𝑋 the biomass concentration (g/l), 𝐶𝐴𝐵 is

the concentration of butyric acid (g/l) and 𝐶𝐴𝐴 the acetic acid concentration (g/l). The remaining are

kinetic parameter that can be found in the Table A. 1. While that, µ, h-1, is the growth velocity

function that can be calculated as:

µ = µ𝑚𝑎𝑥

𝑆

𝑆 + 𝐾𝑆

(𝑖𝑛𝑏) Equation 3. 2

Therefore the grow velocity function depends on the maximum growth velocity, µ𝑚𝑎𝑥, h-1, the sugar

concentration and a function of inhibition, (𝑖𝑛𝑏).

𝑖𝑛𝑏 = 1 − (𝐶𝐴𝐵

𝐶𝐴𝐵𝑚𝑎𝑥)

𝑚𝐴𝐵

(𝐶𝐴𝐴

𝐶𝐴𝐴𝑚𝑎𝑥)

𝑚𝐴𝐴

(𝐶𝐵𝑢𝑂𝐻

𝐶𝐵𝑢𝑂𝐻𝑚𝑎𝑥 )

𝑚𝐵𝑢𝑂𝐻

(𝐶𝐼𝑝𝑂𝐻

𝐶𝐼𝑝𝑂𝐻𝑚𝑎𝑥 )

𝑚𝐼𝑝𝑂𝐻

(𝐶𝐸𝑡𝑂𝐻

𝐶𝐸𝑡𝑂𝐻𝑚𝑎𝑥 )

𝑚𝐸𝑡𝑂𝐻

Equation 3. 3

Regarding the inhibition function, it only depends on the concentration of the acids, 𝐶𝐴𝐵, 𝐶𝐴𝐴, and

organics, 𝐶𝐵𝑢𝑂𝐻, 𝐶𝐼𝑝𝑂𝐻, 𝐶𝐸𝑡𝑂𝐻 , in the fermenter. The remaining are all fixed parameters in the Table

A. 2.

The biomass concentration evolution also has to be calculated, being a function of the growth

velocity, µ, and the biomass concentration itself, 𝑋.

𝑑𝑋

𝑑𝑡= 𝑘1µ𝑋 Equation 3. 4

With all this the acids productions can then be calculated as:

𝑑𝐶𝐴𝐵

𝑑𝑡= 𝑘4𝑘1µ𝑋 − 𝑠𝑖𝑔 × 𝑘5

𝑆

𝑆 + 𝐾𝑆

𝐶𝐴𝐵

𝐶𝐴𝐵 + 𝐾𝐴𝐵𝑋 Equation 3. 5

𝑑𝐶𝐴𝐴

𝑑𝑡= 𝑘9𝑘1µ𝑋 − 𝑠𝑖𝑔 × 𝑘10

𝑆

𝑆 + 𝐾𝑆

𝐶𝐴𝐴


23

Where sig is a simple function to simulate a physiologic state of the cellule and is calculated as:

𝑠𝑖𝑔 =𝑚𝑎𝑥[0, 𝑎𝑏𝑠(𝐶𝐴𝐵 − 𝐶𝐴𝐵,𝑐𝑟𝑖𝑡𝑖𝑐)]

𝜙 + 𝑚𝑎𝑥[0, 𝑎𝑏𝑠(𝐶𝐴𝐵 − 𝐶𝐴𝐵,𝑐𝑟𝑖𝑡𝑖𝑐)] Equation 3. 7

The critic butyric acid, 𝐶𝐴𝐵,𝑐𝑟𝑖𝑡𝑖𝑐, from where henceforward the butanol and ethanol production shall

begin, and 𝜙 are parameters fixed accordingly to the constant simulation parameters table in the

Table A. 3.

With this equation it is avoided the mathematic discontinuity of simply implying:

𝑆𝑖𝑔 = 0 if 𝐶𝐴𝐵,𝑎𝑞 < 𝐶𝐴𝐵,𝑐𝑟𝑖𝑡𝑖𝑐 Equation 3. 8

𝑆𝑖𝑔 = 1 if 𝐶𝐴𝐵,𝑎𝑞 ≥ 𝐶𝐴𝐵,𝑐𝑟𝑖𝑡𝑖𝑐 Equation 3. 9

After the inclusion of the acid production calculations, follows the equations related to the alcohols

production.

𝑑𝐶𝐵𝑢𝑂𝐻

𝑑𝑡= 𝑠𝑖𝑔 × (𝑘6𝑘1µ𝑋 + 𝑘7

𝑆

𝑆 + 𝐾𝑆

𝐶𝐴𝐵

𝐶𝐴𝐵 + 𝐾𝐴𝐵𝑋) Equation 3. 10

𝑑𝐶𝐼𝑝𝑂𝐻

𝑑𝑡= 𝑠𝑖𝑔 × 𝑘11𝑘1µ𝑋 + 𝑘12

𝑆

𝑆 + 𝐾𝑆

𝐶𝐴𝐴


𝑑𝐶𝐸𝑡𝑂𝐻

𝑑𝑡= 𝑠𝑖𝑔 × 𝑘8𝑘1µ𝑋 Equation 3. 12

In all this equations there is a term related to the bacteria growth, recurring to the growth velocity,µ

and the biomass concentration, 𝑋. There is also another term that is dependent on the substrate

concentration and the acids concentration, in the case of butanol the butyric acid, 𝐶𝐴𝐵 and for the

isopropanol the acetic acid, 𝐶𝐴𝐴. Yet again, the remaining fixed parameters can be found in the Table

A. 1.

With all the previous equations the fermentation reactions can be simulated. Therefore what is

missing are the process calculations. For that, the gas flowrate can be calculated as follows if it is

considered that there is no N2 gas lost in the fermenter liquid. Also for the initial gas flow rate a vvm

of 2 liters of nitrogen per liter of water and per hour is assumed.

24

𝑄𝑔𝑎𝑠𝑖𝑛

𝑖+1= 𝑄𝑔𝑎𝑠

𝑖𝑛𝑖

∗𝑦𝑁2

𝑖𝑛

𝑖+1

𝑦𝑁2

𝑜𝑢𝑡

𝑖+1

Equation 3. 13

Where 𝑖 represents the time step, 𝑄𝑔𝑎𝑠𝑖𝑛 (kg/h) is the gas flowrate entering the fermenter, 𝑦𝑁2

𝑖𝑛 is the

nitrogen mass fraction in the entry gas and 𝑦𝑁2

𝑜𝑢𝑡 the nitrogen mass fraction at the fermenter outlet.

To calculate the flash equilibrium in the fermenter Simulis was used. For this calculation there is need

only for the feed composition and the fermenter pressure and temperature.

An important note is that all the gas injected won’t be in equilibrium with the liquid at the same

time, so a gas hold-up must be considered to establish the liquid-vapor equilibrium at each time step.

Therefore for each mass balance a renewal gas mass, 𝑀𝑔𝑎𝑠𝑖𝑛 , and its composition can be determined

in function of the hold-up and the fermenter volume 𝑉𝑓𝑒𝑟𝑚.

𝑀𝑔𝑎𝑠𝑖𝑛

𝑖+1=

ℎ𝑜𝑙𝑑 − 𝑢𝑝𝑖+1

(1 − ℎ𝑜𝑙𝑑 − 𝑢𝑝𝑖+1) ∗ 𝑉𝑓𝑒𝑟𝑚𝑖

Equation 3. 14

While the hold-up may be calculated by the Joshi and Sharma equation, a function of the gas

velocity, 𝑣𝑔𝑎𝑠 (m/s).

ℎ𝑜𝑙𝑑 − 𝑢𝑝𝑖+1 =𝑣𝑔𝑎𝑠𝑖+1

0.3 + 2𝑣𝑔𝑎𝑠𝑖+1

Equation 3. 15


𝑦𝐵𝑢𝑂𝐻𝑖𝑛

𝑦𝐼𝑝𝑂𝐻𝑖𝑛

𝑦𝐸𝑡𝑂𝐻𝑖𝑛

𝑦𝑤𝑖𝑛

𝑦𝑁2

𝑖𝑛

F

𝑄𝑔𝑎𝑠𝑜𝑢𝑡

𝑦𝐵𝑢𝑂𝐻𝑜𝑢𝑡

𝑦𝐼𝑝𝑂𝐻𝑜𝑢𝑡

𝑦𝐸𝑡𝑂𝐻𝑜𝑢𝑡

𝑦𝑤𝑜𝑢𝑡

𝑦𝑁2

𝑜𝑢𝑡

Figure 13. Scheme for fermenter mass balance.

25

This gas retention is normal to all bubble flow agitated reactions and associated to this, a level

expansions has to be maintained so that the total fermenter volume is lower than its volume,

otherwise the working volume must change.

Where the gas velocity is easily calculated considering the fermenter base area, 𝐴𝑓𝑒𝑟𝑚 in m2.

𝑣𝑔𝑎𝑠𝑖+1= 𝑄𝑔𝑎𝑠

𝑖𝑛𝑖+𝑖

/𝐴𝑓𝑒𝑟𝑚 Equation 3. 16

After this calculations, the total mass in the fermenter, 𝑇𝑓𝑒𝑟𝑚𝑒𝑛𝑡𝑒𝑟 (g), at any time may be calculated

as follows.

𝑇𝑓𝑒𝑟𝑚𝑖+1= 𝑄𝑔𝑎𝑠

𝑖𝑛𝑖+1

+ 𝑉𝑓𝑒𝑟𝑚𝑖∗ (𝐶𝐵𝑢𝑂𝐻𝑖

+ 𝐶𝐼𝑝𝑂𝐻𝑖+ 𝐶𝐸𝑡𝑂𝐻𝑖

) + 𝑉𝑓𝑒𝑟𝑚𝑖∗ 𝜌𝑤 Equation 3. 17

The fermenter broth density is assumed as the same of water, 𝜌𝑤.Its composition could then be

calculated for all components, Equation 3. 18 shows the example for butanol, where 𝑥𝐵𝑢𝑂𝐻𝑓𝑒𝑟𝑚

is the

butanol mass fraction in the fermenter and 𝑦𝐵𝑢𝑂𝐻𝑖𝑛 the butanol mass fraction in the gas entering the

fermenter.

𝑥𝐵𝑢𝑂𝐻𝑓𝑒𝑟𝑚

𝑖+1=


𝑖+1∗ 𝑦𝐵𝑢𝑂𝐻

𝑖𝑛𝑖+1

+ 𝑉𝑓𝑒𝑟𝑚𝑖∗ 𝐶𝐵𝑢𝑂𝐻𝑖

𝑇𝑓𝑒𝑟𝑚𝑖+1

Equation 3. 18

Simulis calculates the gas-liquid equilibrium from the fermenter composition and returns the gas

phase and liquid phase composition, resulted from the gas stripping.

Therefore the gas outlet, 𝑄𝑔𝑎𝑠𝑜𝑢𝑡 (kg/h), may be calculated as:

𝑄𝑔𝑎𝑠𝑜𝑢𝑡

𝑖+1= 𝑄𝑔𝑎𝑠

𝑖𝑛𝑖+1

∗𝑦𝑁2

𝑖𝑛

𝑖+1

𝑦𝑁2

𝑜𝑢𝑡

𝑖+1

Equation 3. 19

Then for the condensation steps the only necessary calculations are similar to the ones done until

now, that is, feed composition calculation to undergo a flash at condenser temperature and

pressure, which results in the liquid and gas composition obtained. The gas and liquid flow rate may

be calculated as Equation 3. 20 and Equation 3. 21.

𝐺𝑖+1 = 𝑇𝑐𝑜𝑛𝑑𝑖+1

(𝑦𝑁2

𝑖𝑛

𝑖+1− 𝑥𝑁2

𝑜𝑢𝑡

𝑖+1)

(𝑦𝑁2

𝑜𝑢𝑡

𝑖+1− 𝑥𝑁2

𝑜𝑢𝑡

𝑖+1)

Equation 3. 20

26

𝑙𝑖+1 = 𝑇𝑐𝑜𝑛𝑑𝑖+1− 𝐺𝑖+1 Equation 3. 21

Where 𝐺 is the gas mass flowrate from the condenser in g, 𝑇𝑐𝑜𝑛𝑑 the total mass in the condenser in

g, 𝑦𝑁2

𝑖𝑛 the mass fraction of nitrogen entering the condenser, 𝑥𝑁2

𝑜𝑢𝑡 the mass fraction of nitrogen in the

condensate and 𝑦𝑁2

𝑜𝑢𝑡 the mass fraction of nitrogen in the gas leaving the condenser.

The condenser conditions are in the following table.

Table 3. Work conditions for the condenser in FCC process scheme.

Condenser Temperature (°C) Pressure (bar)

C1 4 1.013

C2 -15 1.013

The alcohols recovered from the gas may then be calculated from the liquid flow rate and

composition, Equation 3. 22 shows the example of this calculation for butanol, being 𝑙𝐵𝑢𝑂𝐻 the mass

of butanol in the condensate (g) and 𝑥𝐵𝑢𝑂𝐻𝑜𝑢𝑡 the mass fraction of butanol in the condensate.

𝑙𝐵𝑢𝑂𝐻𝑖+1= 𝑙𝑖+1 ∗ 𝑥𝐵𝑢𝑂𝐻

𝑜𝑢𝑡𝑖+1

Equation 3. 22

Furthermore having the condensate streams determined two different concentration calculations for

each organic component may be made:

1. A concentration determined by the accumulation of alcohol recovered and referred to the

condensate volume, 𝐶𝐵𝑢𝑂𝐻𝑐𝑜𝑛𝑑 , useful to know the concentration obtained in the accumulated

condensates.

𝐶𝐵𝑢𝑂𝐻𝑐𝑜𝑛𝑑

𝑖+1=

∑ 𝑙𝑖+1 ∗ 𝑥𝐵𝑢𝑂𝐻𝑜𝑢𝑡

𝑖+1𝑖+10

∑ 𝑙𝑖+1 𝜌𝑤⁄𝑖+10

Equation 3. 23

2. A concentration determined by the accumulation of alcohol recovered and referred to the

fermenter initial volume, 𝐶𝐵𝑢𝑂𝐻𝑟𝑒𝑐 . This concentration is a method to compare the increase in

concentration obtained by fermentation coupling techniques to the normal fermentation

without coupling techniques. 𝑉𝑓𝑒𝑟𝑚0 refers to the inital fermenter volume.

𝐶𝐵𝑢𝑂𝐻𝑟𝑒𝑐

𝑖+1=

∑ 𝑙𝑖+1 ∗ 𝑥𝐵𝑢𝑂𝐻𝑜𝑢𝑡

𝑖+1𝑖+10

𝑉𝑓𝑒𝑟𝑚0

Equation 3. 24

27

3.2.2 FWWC & FW5

This schemes are not illustrated due to confidential reasons. Nevertheless the equations from the

previous scheme also apply for the fermentation kinetics. Also for this schemes there was an

increment of a new compound in the system, which will have to be taken in account in the former

equations for all mass balances and Simulis calculations.

Lastly, regarding the FWWC scheme, there exists a condenser unit for which the calculations made

are equivalent to the ones used for the FCC scheme and at atmospheric pressure while concerning

the temperature, simulations for both condenser temperatures from scheme FCC were done.

Lastly for both schemes, two different process alternatives were studied, corresponding to A and B.

3.2.3 FLLR

The last scheme replaces the gas stripping to total liquid-liquid extraction to promote the alcohol

production and extraction, where FLLR denominates Fermenter-Liquid-Liquid-Regeneration. The

reaction equations for this are the same as before but with the increment of the term for mass

transfer between the organic and liquid phases. For this scheme, as solvents, a vegetable oil and 2-

butyl-octanol were studied.

Reg

ene

rati

on

Un

it

F Alcohols

Figure 14. Scheme for the fourth coupling technique (FLLR).

28

Every kinetic parameter for the liquid-liquid extraction was previously obtained experimentally at

IFPEN labs, therefore this models does not recur to Simulis thermodynamic package. Also 𝐾𝐺 as a

value high enough to considerate that mass transfer governs.

Equation 3. 25 and Equation 3. 26, exemplifies the addition of the term of mass transfer in the

previous equation of butanol production in the fermenter, Equation 3. 10.

𝑑𝐶𝐵𝑢𝑂𝐻𝑎𝑞

𝑑𝑡= 𝑠𝑖𝑔 × (𝑘6𝑘1µ𝑋 + 𝑘7

𝑆

𝑆 + 𝐾𝑆

𝐶𝐴𝐵

𝐶𝐴𝐵 + 𝐾𝐴𝐵𝑋) − 𝐵𝑢𝑂𝐻𝑜𝑟𝑔 Equation 3. 25

𝐵𝑢𝑂𝐻𝑜𝑟𝑔 = 𝐾𝐺 × (𝑉𝑎𝑞 + 𝑉𝑜𝑟𝑔)

𝑉𝑎𝑞 × (𝑘𝐵𝑢𝑂𝐻 ∗ 𝐶𝐵𝑢𝑂𝐻

𝑎𝑞 − 𝐶𝐵𝑢𝑂𝐻

𝑜𝑟𝑔) Equation 3. 26

Where 𝐵𝑢𝑂𝐻𝑜𝑟𝑔 is the butanol mass transferred to the organic phase, and is a function of the

aqueous phase volume, 𝑉𝑎𝑞 (m3), the organic phase volume fixed as 10% of the aqueous phase,

𝑉𝑜𝑟𝑔 (m3), the butanol concentration in the aqueous phase, 𝐶𝐵𝑢𝑂𝐻𝑎𝑞

(g/l) and the butanol

concentration in the organic phase, 𝐶𝐵𝑢𝑂𝐻𝑜𝑟𝑔

(g/l). The remaining parameters were obtained in the

laboratory and its values may be found in Table A. 4.

𝑉𝑎𝑞𝑖+1 = 𝑉𝑎𝑞

𝑖 +𝑑𝐶𝑤

𝑎𝑞𝑖

𝑑𝑡× 𝑑𝑡 ×

𝑉𝑎𝑞𝑖

𝜌𝑤 Equation 3. 27

Where 𝐶𝑤𝑎𝑞is the water concentration in the aqueous phase (g/l) and 𝑑𝑡 the time step used in the

simulation, in hour units.

For the organic phase it was implemented new equations to the model. Equation 3. 28, shows the

example for the butanol concentration in the organic phase.

𝑑𝐶𝐵𝑢𝑂𝐻𝑜𝑟𝑔

𝑑𝑡= 𝐾𝐺 ∗

(𝑉𝑎𝑞 + 𝑉𝑜𝑟𝑔)

𝑉𝑜𝑟𝑔 ∗ (𝑘𝐵𝑢𝑂𝐻 ∗ 𝐶𝐵𝑢𝑂𝐻

𝑎𝑞 − 𝐶𝐵𝑢𝑂𝐻

𝑜𝑟𝑔) Equation 3. 28

For the alcohols extraction an extraction solvent flow rate, 𝑄𝐸𝑥𝑡 (m3/h) was calculated as:

𝑄𝐸𝑥𝑡𝑖+1= 𝑧 × (

𝐶𝐵𝑢𝑂𝐻𝑎𝑞

𝐶𝐵𝑢𝑂𝐻𝑏 )

2.5

× 𝑉𝑜𝑟𝑔 Equation 3. 29

29

Where 𝐶𝐵𝑢𝑂𝐻𝑏 is the needed concentration of butanol to initiate extracting the alcohols and 𝑧 is the

ratio between the extraction solvent flow rate and the organic phase volume, assumed as 0.001.

The water and the alcohols recovered from the solvent are calculated exactly as for a condenser in

the previous schemes, where 𝑙𝑖+1changes to 𝑄𝐸𝑥𝑡𝑖+1 , Equation 3. 22, and since this flow rate is

volumetric instead of composition, concentrations of the organic phase are used.

3.3 Solvent Regeneration

Preliminary simulations with PROII of the solvents regeneration were made. For this, the same

thermodynamics model was used and the components not in SRKS V11 pushed the need to use an

UNIFAC predictive model other than the SRKS without any coefficients.

To evaluate the feasibility of the separation of the solvent and the butanol the ternary diagram was

also studied.

3.3.1 2-butyl-octanol

As can be seen, the mixture obtained with the use of 2B1O has a ternary azeotrope at 93.37 °C,

therefore by distillation the separation of alcohols will always have this constraint point. In fact in a

first column the solvent could be obtained at the bottom while the ternary azeotrope at the top.

Nevertheless at atmospheric pressure the ternary composition is already low in 2B1O, having almost

the composition of the binary azeotrope between water and butanol, hence following to a

heteroazeotropic distillation. However a solvent waste would always exist.

30

As solution to have a lower waste of solvent, lower pressure was also tested. According to Figure 16,

by lowering the pressure the ternary azeotrope composition in solvent drops and therefore wastes

are greatly reduced.

Figure 15. Ternary diagram for 2B1O- butanol -water system at 1 bar, obtained in PROII using UNIFAC thermodynamic package.

31

Therefore a distillation scheme for the regeneration of the solvent can be proposed as follows in

Figure 17, where the ternary azeotrope shall be obtained at the top of the first column and by phase

separation purified while the remaining is sent to other column to achieve the main alcohol recovery.

Figure 16. Ternary diagram for 2B1O- butanol -water system at 0.01 bar, obtained in PROII using UNIFAC thermodynamic package.

32

3.3.2 Vegetable oil

Lastly, using vegetable oil there isn’t any ternary azeotrope at low pressure. Therefore the

regeneration can be easily obtained in a single distillation column. To this scheme fresh butanol was

also added to ease the separation, hence reduce the energy consumption. The use of a decanter for

using the distillation column top as feed turns possible to achieve streams with higher alcohol

concentration.

Feed

Solvent

Alcohol

s Alcohols

Solvent

Concentrated

Alcohols

Figure 17. Scheme for the 2BO1 regeneration by distillation.

33

Figure 18. Ternary diagram for Triolein- butanol -water system at 0.005 bar, obtained in PROII using UNIFAC thermodynamic package.

34

.

3.4 Energy Consumption The energy consumption may derive from different parcels, the compressor to maintain the gas

pressure, where an overall pressure loss of 0.2 bar was assumed, the condensers and lastly the

solvent regeneration towers reboilers.

For the compressor, recurring to simulis, the gas enthalpy is calculated before entering the

compressor, then the gas entropy is also calculated and lastly the gas outlet enthalpy is calculated

using as entry parameters the pressure increase and entropy and therefore assuming an isentropic

compression. Finally the difference between the gas enthalpies results in the energy consumption.

The condensers energy consumption is merely the difference between the gas enthalpy and outlet

enthalpies. This is easily done by Simulis, where it is only needed the entry composition, temperature

and pressure to automatically make a flash calculation to discover the resulting gas and liquid

enthalpies per kilogram. After this enthalpy shall be summed according to the vapor and liquid

corresponding fractions.

Therefore in all this calculations it was assumed an energy efficiency of 100 %.

Meanwhile for the solvent regeneration, the energy consumption is directly obtained in the PROII

simulation per kilogram of solvent and respecting a certain feed and product streams compositions.

Figure 19. Scheme for vegetable oil regeneration.

Feed

Solvent

Alcohol

s

Fresh

Butanol

Alcohol

s

Alcohol

s

35

4 Results and Discussion

Regarding this chapter structure, firstly the results for the downstream recovery of alcohol by the

distillation schemes described in 3.1. Afterwards, the study will focus in the results for the in situ

product recovery techniques, beginning by the fermenter concentration evolution during the

simulation, followed by the concentration factors achieved by gas stripping and each extraction

method used, then some process optimizations are discussed as well as the results for the FLLR

scheme. By the end, final results for each in situ recovery technique are compared.

4.1 Downstream

Concerning the results obtained for the downstream recovery of alcohols, for each distillation

scheme study, energy consumption values per column are shown. Also in the end a comparison

between the total energy consumption per kg of IBE of all the distillations schemes has been made.

4.1.1 DSLC

The DSLC scheme is more efficiently used for low concentrations feeds since the beer column

objective is to significantly remove water and therefore concentrate the alcohols in the top outlet.

For this arrangement 3 feed concentrations of IBE were used, 8 gIBE/l, 17 gIBE/l and 24 gIBE/l. This

concentrations follow the operating line in Figure 20, also the average concentration achieved in IBE

fermentation is 17 gIBE/l, meaning a final concentration of 10.2 gBuOH/l, 5.12 gIpOH/l and 1.7 gEtOH/l.

Table 4. Results obtained for the energy consumptions of the DSLC.

8 (gIBE/l) 17 (gIBE/l) 24 (gIBE/l)

Column 1 consumption (GJ/h) 327.16 178.18 133.80




Total consumption (GJ/TonIBE recovered) 25.48 15.42 12.47

As to expect the consumption of the beer column is the highest, ranging between 86 to 71% of the

total energy consumption. Therefore concentrating the feed will easily reduce the energy cost of this

scheme.

36

4.1.2 DSHC

This arrangement is meant to be used only for high concentrating feeds, therefore suppresses the

main separation cost, the beer column. Two feed concentrations were used 607 gIBE/l and 718 gIBE/l.

Table 5. Results obtained for the energy consumptions of the DSHC.

607 (gIBE/l) 718 (gIBE/l)

Column 1 consumption (GJ/h) 28.49 23.57



Total consumption (GJ/TonIBE recovered) 3.43 3.00

As supposed the suppression of the first column reduces greatly the energy consumption, compared

to the average of 17 gIBE/l, concentrating the feed may reduce the downstream energy consumption

between 78% and 81%.

4.1.3 DSDZ

The final arrangement is meant to be used for the mixing zone of the ternary and even the low

concentration zone, only the concentration point of 17 gIBE/l was used as a comparison method.

Table 6. Results obtained for the energy consumptions of the DSHC.

17 (gIBE/l)

Column 1 consumption (GJ/h) 183.89

Column 2 consumption (GJ/h) 32.42

Total consumption (GJ/TonIBE recovered) 14.30

This new arrangement can then reduce the energy consumption form the first scheme only by 7.2%,

since it adds complexity by adding two liquid phases column trays this scheme may also as well not

be as optimized as the ones before, therefore further optimization in this scheme is necessary before

concluding if it is a viable option and also test it for other concentration zones.

37

4.1.4 Comparison

As seen, concentrating the feed can drastically change the downstream processing and its energy

consumption. The next curve shows the effect of feed concentration and separation energy cost.

According to Figure 20, the energy consumption suffer a quick decrease just by having a little more

concentrated feed. In fact for the DSHC scheme both energy consumptions are very similar,

therefore energy consumptions seems to reach an asymptote around this value.

4.2 Fermenter concentration evolution

In this schemes the results for the fermenter concentration evolution are very similar, this because

the yield in the different alcohols won’t really vary much and the extraction of the alcohols is done by

nitrogen is present in this schemes, and therefore the specific tendency of extraction of each alcohol

is almost the same. The real difference is in the productivity, the energy consumption to recover the

alcohols from the gas and also the concentrations obtained in this recovered streams.

Figure 20. Graphic for the energy consumption in function of the feed compositions for the three different distillation schemes.

DSLC

DSDZ

DSHC

38

The fermenter concentration evolution is as follows in Figure 21, corresponding to the first scheme

accordingly to the Figure 12 where CAB is the concentration of butyric acid, CAA corresponds to

acetic acid concentration, CB, CIpOH, CEtOH to butanol, isopropanol, ethanol concentration,

respectively, S is the sugar concentration and finally the variables related directly to the coupling

technique are CB max, the total butanol produced dividing by the initial fermenter volume, Cond1

which is the butanol mass that remains in the fermenter plus the butanol recovered in the first

condenser, also dividing by the initial fermenter volume and for last Cond2 which parallel to Cond1

just sums the parcel of the butanol recovered in the second condenser.

Figure 21. Curves of fermenter concentrations during the simulation for the FCC scheme.

Analyzing the Figure 21 the total concentration of butanol that is achieved is of 18.6 g/l, only a

reduced fraction of butanol remains in the gas, not being removed from it as can be concluded from

the almost coincident curves of CB max and Cond2. The need for a second condenser is also

confirmed since the difference between curve Cond2 and curve Cond1 is very significant, in fact with

only one condenser merely a concentration of 11g/l of butanol was attained.

39

The butanol concentration in the fermenter, as the remaining alcohol concentrations, reaches a peak

concentration of 9.25 g/l, 4.85 g/l and 1.55 g/l, for butanol, isopropanol and ethanol, respectively.

From this point the alcohols concentrations in the fermenter will start decreasing since the alcohol

stripping velocity will be superior to its production, until reaching a sugar concentration of zero, from

where there is no production and only stripping, hence the total butanol production, CB max, is

maintained while the fermenter butanol concentration, CB, decreases.

Figure 22. Curves of fermenter concentrations during the simulation in batch for the standard IBE fermentation.

Comparing this results from that of a normal IBE batch fermentation, Figure 22, from introducing the

gas striping method from scheme FCC, an increase in production of 65%, 43% and 70% can be

expected for butanol, isopropanol and ethanol respectively.

Introducing the fed batch in the process Figure 23 and Figure 24 were obtained. This figures follows

the same symbology as before and are also from the first scheme.

40

Figure 23. Butanol production and recovery evolution during FCC fed batch simulation.

Analyzing only the gas stripping extraction and the alcohol recovery, CB max reaches a concentration

of 120.8 g/l in 124.6 hours. Comparing this result to the one without fed batch, its introduction

augments the productivity in butanol (mass per hour) in 110%. Observing for isopropanol and

ethanol there are augments in productivity of 76% and 79%, respectively.

Figure 24. Curves of fermenter concentrations during FCC fed batch simulation.

Observing the fermenter concentrations during the fed batch introduction there is an increase in

fermenter concentrations, in fact butanol reaches the concentration inhibition limit therefore the

41

Figure 25. Batch process total energy consumption for FCC scheme.

Figure 26. Fed batch process total energy consumption for FCC scheme.

results obtained are not optimized. This increase in fermenter concentrations can be explained as,

although there is an increase in fermenter volume, since fed batch is an aqueous solution of

500 gsugar/l, and hence the concentrations should decrease, subsequently there is also a reduction in

the vvm used because the nitrogen flow rate is constant. This reduction will also reduce the stripping

effect meaning less alcohol is extracted and therefore the alcohols that rest in the fermenter

accumulate and there’s concentration increases. In this condition can then be concluded that the

stripping rate is lower than the production rate.

To correct this efficiency problem a possible option would be an increase of the vvm with the

volume, which is the same as having a fixed vvm and a variable nitrogen flow rate, so that there

would be a feed of nitrogen to compensate the increment in volume. The disadvantage of this

resolutions is the difficulty of working at a variable gas flow rate experimentally.

In terms of energy for the FCC scheme the calculation is easily done as the energy consumption in

the condensers and compressor. This compressor maintains the gas pressure and a pressure loss of

0.2 bar was assumed in the whole process system.

In the batch process the energy

consumption is very high at the

beginning but with the development of

the reaction the alcohol concentration

increases and so does the alcohol

extraction, reducing the energy

consumption per mass of alcohol

produced. Obviously once the sugar

concentration reaches zero there is no

more production of alcohols, hence the

energy consumption increases, and

therefore there is no need to continue

the cycle.

Moving to a process with fed batch it is

possible to reach a minimum of energy

consumption and continue production

while maintaining the minimum energy

consumption.

42

Working in fed batch if assumed the case ends at 500 gsugar/l used, being the volume the initial

fermenter volume, then the energetic consumption is of 17.2 GJ/tonIBE produced, while the energetic

consumption for a batch case is 33.1 GJ/tonIBE produced. Therefore the introduction of fed batch can

reduce the energy consumption by 48%. Also the base case ends at 500 gsugar/l but the longer the fed

batch is used the lower the energy consumption will be, if imagining the fed batch will be used as a

continuous process and therefore the cycle never ends the energy consumption will be as low as 0.27

GJ/tonIBE produced since most of the energy cost is associated to the first 10 hours.

4.3 Gas Stripping

During gas stripping the alcohols will have different concentrations in the fermenter and therefore

the gas will extract more or less accordingly, hence it is interesting to plot this dependence.

Analyzing Figure 27 butanol is the

preferred alcohol extracted and

isopropanol the lowest. From this result it

is easier to understand why ethanol is the

only alcohol that doesn’t have its

concentration increasing in Figure 24

being the reason that ethanol is the

alcohol with the lowest production but

since its extraction is quicker than

isopropanol it is possible that for isopropanol the production rate is quicker than the extraction rate

while for ethanol the contrary occurs.

With nitrogen as stripping agent, butanol from the fermenter in the gas fraction is averagely 14.6

times more concentrated than in the fermenter, since there exists a quicker extraction in the

beginning, while for isopropanol and ethanol the corresponding factor is 14.6 and 11.0, respectively.

4.4 Extraction Methods To compare the different extraction processes it is also important to understand the concentration

factors so that different results may be explained.

Figure 27. Gas stripping extraction capacity for FCC scheme.

43

4.4.1 FCC The first condenser works at 4 °C and contrary to what was seen in the gas stripping extraction, for

low fractions, ethanol concentration can be higher to not only isopropanol but also butanol.

However if analyzed for all the fraction

range butanol is the preferable component.

In fact it is possible to recover in average,

85% of the butanol stripped, indeed for

butanol the recovery fraction is higher at

higher fractions. While for isopropanol and

ethanol recoveries of 62% and 66% can be

expected, in that order.

While that in the second condenser the

working temperature is of -15 °C thus a

higher concentrated factor is expected.

In terms of extraction tendency the same

affinity as in the first condenser is observed

but calculating, concentration factors

values of 1.2, 0.7 and 0.9 were observed for

butanol, isopropanol and ethanol,

respectively.

Hence in the end there has been, for butanol, in the first condensate a total concentration factor of

12.5 and for the second liquid steam of 1.2 relative to the residual butanol in the gas stream, which is

already two times more concentrated than in the first condenser feed, therefore a total

concentration factor of 33.6 can be expected in the second condenser.

4.4.2 FWWC & FW5

Since this schemes are confidential only a brief discussion has been made. For each scheme two

different process alternatives were used, designed as A and B. Also, while for the FCC scheme the

condenser can’t concentrate more than the gas stripping since water is easier to condense, in this

cases there is a higher tendency to remove alcohols than water, due to the process scheme nature.

From the FWWC scheme three recovered alcohol streams may be obtained, while for the FW5

scheme, five streams are recovered.

Figure 29. First condenser recovery capacity for FCC scheme.

Figure 28. First condenser recovery capacity for FCC scheme.

44

In Figure 30 the alcohols fraction at the outlet of a typical unit in this process schemes is shown in

function of the alcohols fraction in the feed to this unit. From this figure, concentration factors

increase compared to the ones obtained for the FCC scheme, being superior to one.

Similarly to before the different feed concentrations were also calculated. Therefore the total

concentration factor can be calculated for both process alternatives.

Table 7. Results for total recovery factors in for scheme FWWC & FW5 for both alternatives.

Stream number

Total Concentration factor

FWWC FW5

A B A B

1 45.3 67.2 45.3 67.1

2 18.4 54.8 19.6 54.8

3 1.1 6.3 7.3 45.3

4 - - 2.3 34.6

5 - - 0.8 26.0

In the B process alternative there is an increase in the concentration factor since there is more water

accumulated in the gas.

Figure 30. Recovery capacity for a standard FWWC scheme process unit.

45

4.4.3 Solvents

In terms of solvent regeneration costs the thermodynamic system for each solvent used was

analyzed.

Considering Figure 15, for the 2B1O, from a mixture of majorly solvent in a first separation the

ternary azeotrope shall be obtained at the column top outlet, while the pure solvent at the bottom.

In a second column an outlet with a composition similar to the binary azeotrope between water and

butanol is obtained.

Truly, according to Figure 17, the process streams obtained are improved recurring to phase

separation by a decanter.

Table 8. Results for 2BO1 regeneration process.

1 2 3 4 5

𝑸 𝑸𝒇𝒆𝒆𝒅⁄ 0.034 0.847 0.012 0.097 0.010

Temperature (°C) 81.6 163.4 40.0 91.6 102.7

Pressure (Bar) 0.57 0.1 0.1 1 1

𝒙𝑩𝒖𝑶𝑯 (m/m) 1.40×10-03 2.00×10-04 9.80×10-03 3.00×10-04 0.612

𝒙𝑰𝒑𝑶𝑯 (m/m) 1.03×10-03 0 7.56×10-03 0 0.164

𝒙𝑬𝒕𝑶𝑯 (m/m) 7.13×10-04 0 4.93×10-03 0 0.037

𝒙𝒘 (m/m) 0.997 0 0.978 0 0.184

𝒙𝟐𝑩𝑶𝟏 (m/m) 3.28×10-05 0.9998 1.80×10-05 0.9997 2.54×10-03

Regarding Figure 18, from a mixture of vegetable oil, in a first separation the azeotrope between

water and butanol is directly obtained. Once more recurring to the scheme used, the product

streams compositions are improved.

2 3

4

5

1

46

Table 9. Results for vegetable oil regeneration process.

1 2 3 4

𝑸 𝑸𝒇𝒆𝒆𝒅⁄ 0.406 0.012 0.581 0.001

Temperature (°C) 156.6 -5.8 -5.8 -5.8

Pressure (Bar) 0.005 0.005 0.005 0.005

𝒙𝑩𝒖𝑶𝑯 (m/m) 2.00×10-04 0.052 0.185 0.629

𝒙𝑰𝒑𝑶𝑯 (m/m) 0 0.022 0.263 0.115

𝒙𝑬𝒕𝑶𝑯 (m/m) 0 7.00×10-03 0.047 0.016

𝒙𝒘 (m/m) 0 0.914 0.505 0.239

𝒙𝑶𝒊𝒍 (m/m) 0.9998 0 0 0

4.5 Schemes 1/2/3 optimization

4.5.1 Gas stripping

In the FCC scheme, as in all schemes, there is a big energy waste in the initial time period of the

fermentation since the gas injection begins and there is no alcohol to be stripped, therefore

simulations where the gas only starts to be injected once butanol concentration in the fermenter as

reached 6 g/l were performed.

Results for the FCC scheme are in the following table:

1

4

2

3

47

Table 10. Results for FCC process scheme for different gas stripping initiation..

FCC process scheme

Batch Fed-Batch Fed-Batch C=6

Total mass produced

Butanol total (ton) 7.50 49.67 49.37

IpOH total (ton) 2.98 16.84 16.71

EtOH tot (ton) 1.16 6.40 6.38

Concentration total

(Initial fermenter volume)

Butanol total (g/l) 18.75 124.18 123.43

IpOH total (g/l) 7.44 42.10 41.77

EtOH tot (g/l)) 2.91 16.01 15.94

Concentration

(reactor volume)

Butanol (g/l) 3.65 10.16 10.14

IpOH (g/l) 2.93 5.94 5.92

EtOH (g/l)) 0.72 1.24 1.23

Energy consumption

compressor GJ 34.15 102.64 82.91

Energy consumption

condenser 1 GJ 318.27 981.74 800.34

Energy consumption

condenser 2 GJ 75.45 236.40 194.21

Gas rate (m3/h) m3/h 10567.49 10567.49 10567.49

Total Energy consumption

per kg IBE MJ/kg 36.76 18.11 14.87

Final Volume m3 300.46 428.68 486.14

Water added in fed batch 0.00 328.00 328.00

Fermentation time h 45.00 135.20 135.40

Consequently by initiating the gas stripping later the energy costs are reduced by 19.4%, total alcohol

production reduces logically, but not in a really significant value. Also there is more water in the

fermenter and so the concentration remaining in the fermenter is lower but it also means that the

recovered alcohol streams are more concentrated, since the same water is added and the initial

volume is the same.

4.5.2 Condenser temperature

In scheme FWWC results, condenser temperature was of -15 ºC, but a different temperature could

have been used, like for example 4 ºC as the first condenser in scheme 1. Reducing the condenser

efficiency may reduce the total alcohols recovery but as seen previously in this condenser butanol is

already at low concentration.

48

Table 11. Results for FWWC process scheme for different condenser temperature..

FWWC process scheme A

Fed-Batch

C=6 g/l

T=-15 ºC

Fed-Batch

C=6 g/l

T=4 ºC

Total mass produced

Butanol total (ton) 53,43 53,72

IpOH total (ton) 18,67 18,85

EtOH tot (ton) 6,22 6,17

Energy consumption compressor GJ 125,55 141,44

Energy consumption condenser GJ 1187,13 932,57

Gas rate (m3/h) m3/h 10567,49 10567,49

Total Energy consumption per kg IBE GJ/ton 30,84 28,96

Condenser Recovery mass

Butanol (ton) 2,48 1,06

IpOH (ton) 7,27 3,59

EtOH (ton) 1,70 0,61

Eau (ton) 262,23 236,21

Other units Recovery mass

Butanol (ton) 47,02 48,53

IpOH (ton) 9,91 13,20

EtOH (ton) 4,32 5,30

Eau (ton) 104,83 109,58

Final Volume m3 360,19 381,46

Water added 328,00 328,00

Fermentation time h 191,60 198,60

Increasing the condenser temperature total productivity decreases by 3%. While this compressor

energy consumption increases since the time cycle increases and condenser energy consumption

decreases. Energy savings are around 6% per kilogram of alcohol produced. Also since the

condensation process is less severe there is more water in the fermenter and therefore a lower

concentration in the fermenter and a higher in the condensate stream.

4.6 FLLR

The last scheme is the least explored. Only results for the process in batch conditions were obtained.

The fermenter concentration evolution is as follow in Figure 32. Where 2-butyl octanol was used as

the solvent and CB Ext is the contribution to CB max from the butanol recovered in the organic

phase. Results for simulation with vegetable oil are available in annex.

49

The results are similar to those obtained for other scheme processes. In fact CB max is slightly lower

than for other schemes, therefore there is less butanol per sugar consumed. Nevertheless its

productivity is higher than previously and the majority of butanol produced resides in the organic

phase.

To evaluate the organic phase, similar curves were traced to evaluate the simulation concentrations

obtained for the alcohols.

There is an enormous difference in the butanol concentration achieved compared to the other

alcohols, butanol achieves values over 30 g/l while isopropanol is always lower than 5 g/l.

Figure 32. Fermenter concentration evolution in the aqueous phase.

Figure 31. Fermenter concentration evolution in the organic phase.

50

Table 12. Results for FLLR process scheme.

FLLR

2BO1 Vegetable Oil

Total mass produced

Butanol total (ton) 6.83 6.85

IpOH total (ton) 2.81 2.80

EtOH tot (ton) 1.20 1.20

Concentration total (Initial fermenter volume)

Butanol total (g/l) 17.08 17.13

IpOH total (g/l) 7.02 7.01

EtOH tot (g/l)) 3.00 3.01

Concentration (reactor volume)

Butanol (g/l) 7.25 6.75

IpOH (g/l) 5.79 4.69

EtOH (g/l)) 2.76 2.06

Energy consumption solvent regeneration

GJ 81.51 60.04

Total Energy consumption / kg IBE GJ / ton 7.52 5.53

Total mass recovered

Butanol (ton) 2.81 3.99

IpOH (ton) 0.33 0.88

EtOH (ton) 0.06 0.36

Eau (ton) 1.61 0.82

Solvent (m3) 93.79 860.99

Final Volume m3 399.37 399.96

Water added 0.00 0.00

Fermentation time h 37.07 37.23

Comparing results, for a liquid-liquid extraction process vegetable oil is the better solvent, it leads to

almost the same productivity, same alcohol production and the alcohols are more concentrated in

the organic phase, all this while reducing energy consumption.

4.7 Final results

So far each process arrangement and its different optimizations has been discussed, therefore a final

comparison between schemes shall be done.

For this to compare the first three schemes, the results in normal fed batch conditions without any of

the previous optimization options discussed have been evaluated, where for the FWWC and FW5

schemes, A and B stands for different process alternatives.

51

Table 13. Results for fed batch process schemes.

FCC FWWC-A FWWC-B FW5-A FW5-B

Total mass produced

Butanol (ton)

49.67 53,28 54,02 52,89 54,77

IpOH (ton)

16.84 18,63 19,07 18,48 19,47

EtOH (ton)

6.40 6,26 6,27 6,16 5,98

Total (ton) 72.92 78.17 79.36 77.53 80.22

Productivity

Butanol (g/l/h)

0.92 0.70 0.67 0.72 0.64

IpOH (g/l/h)

0.31 0.25 0.24 0.25 0.23

EtOH (g/l/h)

0.12 0.08 0.08 0.08 0.07

Total

(g/l/h) 1.35 1.03 0.99 1.06 0.94

Yield

Butanol (g/gsugar)

0.25 0,27 0,27 0,26 0,27

IpOH (g/gsugar)

0.08 0,09 0,10 0,09 0,10

EtOH (g/gsugar)

0.03 0,03 0,03 0,03 0,03

Total

(gIBE/gsugar) 0.36 0.39 0.40 0.39 0.40


per IBE GJ / ton 18.11 33,50 33,02 17,27 10,43

Fermentation time

h 135.20 190,00 200,40 182,60 214,40

From this results, FW5-B has the highest butanol and isopropanol production while the lowest

productivity for both, while Scheme FCC has the highest productivity in all alcohols but the lowest

yield also. Schemes FWWC-A and FWWC-B fall behind schemes FW5 since energy costs are almost

the double, while productivities and yields are similar. Therefore Schemes FWWC may be deemed as

the possible worst process arrangement and annulled from the following results discussion.

Between the FCC scheme and the FW5-A, energy consumption is very similar, being the main

differences alcohol productivity and yield. Nevertheless, comparing to FW5-B, the scheme becomes

very interesting since it as a much lower energy consumption. In fact as previous settled, the FCC

scheme as a higher productivity since it also as a higher extraction rate, promoting the reactions in

the fermenter. While that since schemes FW5 are slower it also has an improved yield since there is

more time for the sugar to fully react.

52

To evaluate the liquid-liquid extraction scheme compared to the first and third schemes, the results

in normal batch conditions are applied and also without any of the previous optimization options.

Table 14. Results for batch process schemes.

Batch FCC FW5-A FW5-B FLLR 2B1O

FLLR Vegetable

Oil

Total mass produced

Butanol (ton)

4.49 7.50 7,53 7,51 6.83 6.85

IpOH (ton)

2.07 2.98 2,98 2,98 2.81 2.80

EtOH (ton)

0.69 1.16 1,16 1,15 1.20 1.20

Total (ton)

7.25 11.64 11.67 11.65 10.84 10.86

Productivity

Butanol (g/l/h)

0.33 0.42 0.42 0.41 0.46 0.46

IpOH (g/l/h)

0.15 0.17 0.17 0.16 0.19 0.19

EtOH (g/l/h)

0.05 0.06 0.07 0.06 0.08 0.08

Total

(g/l/h) 0.54 0.65 0.66 0.64 0.73 0.73

Yield

Butanol (g/gsugar)

0.12 0.21 0.21 0.21 0.19 0.19

IpOH (g/gsugar)

0.06 0.08 0.08 0.08 0.08 0.08

EtOH (g/gsugar)

0.02 0.03 0.03 0.03 0.03 0.03

Total

(gIBE/gsugar) 0.32 0.32 0.32 0.32 0.30 0.30


GJ / tonIBE 0 36.76 16,29 8.09 7.52 5.53

Fermentation time

h 33.60 45.00 44,40 45,80 37.07 37.23

This last coupling technique, FLLR, has the lowest total alcohol production but it actually has a higher

productivity than the first scheme. This point is very similar to what seen before, the balance

between producing more quickly with a lower yield or the contrary. What truly may distinguish the

liquid-liquid extraction from the previous schemes is the energy consumption. In fact its energy

consumption in batch conditions is already lower than most other arrangements in fed batch

conditions by at least 65%. Therefore, the FLLR scheme may be a much better coupling technique

option and must be deeply investigated and also at fed batch conditions.

Concluding, FW5-B scheme or the FLLR scheme using both solvents may be where a true solution for

IBE fermentation may lie.

53

5 Conclusions

A first study of the downstream recovery for different feed compositions and a model for in situ

coupling techniques has been developed.

Regarding the downstream processing study developed, three different distillation schemes were

established and results according to the bibliographic study were obtained in terms of great energy

reduction of separation only by concentrating the fermentation broth.

Concerning the model developed for the in situ recovering techniques four different process

schemes were studied and energy requirements similar to the shown in the state of art were

achieved. Also the impact of condenser temperatures and gas stripping initiation were evaluated.

Preliminary concentration potential of each process scheme was established.

Comparing the results for each process scheme the economic advantageous may be expected and

therefore the FWWC scheme was assumed as the less recommendable. While that the FLLR scheme

was deemed the most promising in spite of not having achieved results in fed batch conditions.

In terms of solvent selection, results using vegetable oil seem more promising in preliminary

calculations.

Nevertheless there are many improvements needed to achieve a final work and conclusion of the

impact of the many variables in the process to obtain the minimal project cost. For that the FLLR

scheme as to be simulated in fed batch conditions since is the most promising process and fed batch

conditions are desired. In the fermentation model CO2 formation should be included and this gas

should be used as extracting agent for gas stripping. The acids extraction by the extraction should

also be taken into account, as well as a more rigorous heat balance for each process scheme. All

process schemes can be further optimized in terms of the alcohols concentration range in which the

extraction agent is active.

Despite the model providing promising results it still takes much time to compute, hence further

optimization of the model is needed to be able to obtain optimizations results much quicker and

easier in order to provide a more rigorous study for the many process variables.

To grasp a real economic value distillation schemes can still be improved and also a deep study has to

be made focused in how to separate each process stream obtained in the in situ recovery technique

scheme. In fact different approaches can be made, all alcohol recovered streams may be mixed and

depending to its final concentration sent to the according distillation scheme or each stream can be

54

sent to different distillation schemes or even in the same distillation scheme but sent to different

columns according to its titer.

Last but not least the final evaluation of the project consists in the revenue obtained, thus a deeper

economic analysis as to be made to evaluate the total CAPEX, OPEX and raw material costs for each

process scheme and the possible separation by distillation alternatives.

55

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58


59

7 Annexes 7.1 Annex 1 - Constants

Table A. 1. Kinetic constants used for the fermentation model

Kinetics constants

Parameters units value

max h-1 0.15

KS 2.862

k1 14.3

k2 h-1 0.193

k3 h-1 0.286

KBA 2.862

KAA 1.818

k4 0.107

k5 h-1 0.39

k6 0.156

k7 h-1 0.36

k8 0.039

k9 0.081

k10 0.356

k11 0.033

Table A. 2. Inhibition function parameters values.

Inhibition function parameters


CBuOH,max g/l 11

mBuOH 4

CAB,max g/l 5

mAB 2.5

CAA,max g/l 12

mAA 1.5

CIpOH,max g/l 22

mIpOH 2.5

CEtOH,max g/l 44

mEtOH 1.5

60

Table A. 3. Values for the Sig function parameters.

Sig function parameters


CABcr g/l 0.5

0.001

Table A. 4. Values for mass transfer coefficients for each solvent used.

Solvents

Units 2-butyl-1-octanol Vegetable Oil

KG h-1 10.00 10

kBuOH (gorg/lorg)/(gaq/laq) 3.88 0.62

kIpOH (gorg/lorg)/(gaq/laq) 0.71 0.254

kEtOH (gorg/lorg)/(gaq/laq) 0.30 0.24

kw (gorg/lorg)/(gaq/laq) 0.019 1.07E-03

Table A. 5. Property values considered in the model.

Property values


MMB g/mol 74,1216

MMIpOH g/mol 46,0684

MMEtOH g/mol 60,095

MMH2O g/mol 18,0153

MMN2 g/mol 28,0134

MMDBE g/mol 130,228

MM2B1O g/mol 186,334

MMVO g/mol 885,43206

w kg/m3 983

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7.2 Annex 2 – Simulation results

Table A. 6. General simulation results for the standard batch and in situ recovery techniques for the FCC process scheme.

Standard case FCC

Batch Batch Fedbatch Fedbatch C=6

Total mass produced BuOH total (ton) 4,23 7,50 49,67 49,37 IpOH total (ton) 1,99 2,98 16,84 16,71 EtOH tot (ton) 0,65 1,16 6,40 6,38

Total concentration (relative to the initial aqueous

volume)

BuOH total (g/l) 10,58 18,75 124,18 123,43 IpOH total (g/l) 4,97 7,44 42,10 41,77 EtOH tot (g/l)) 1,62 2,91 16,01 15,94

Concentration (relative to fermenter volume)

BuOH (g/l) 10,58 3,65 10,16 10,14 IpOH (g/l) 4,97 2,93 5,94 5,92 EtOH (g/l)) 1,62 0,72 1,24 1,23

Energy consumption compressor GJ 0,00 34,15 102,64 82,91 Energy consumption condenser 1 GJ 0,00 318,27 981,74 800,34 Energy consumption condenser 2 GJ 0,00 75,45 236,40 194,21

Gas flowrate m3/h 0,00 10567,49 10567,49 10567,49 Total energy consumption MJ/kgIBE 0,00 36,76 18,11 14,87

Mass recovered (Condenser 1)

BuOH (ton) 0,00 3,76 29,92 29,45 IpOH (ton) 0,00 1,30 9,12 8,82 EtOH (ton) 0,00 0,53 3,55 3,50

w (ton) 0,00 89,81 270,41 218,59 gw/gorg gw/gorg 0,00 16,06 6,35 5,23

Mass recovered (Condenser 2)

BuOH (ton) 0,00 2,63 15,33 14,93 IpOH (ton) 0,00 0,79 5,13 4,97 EtOH (ton) 0,00 0,41 2,32 2,28

w (ton) 0,00 9,51 28,31 22,78 gw/gorg gw/gorg 0,00 2,48 1,24 1,03

Total stripping gw/gorg gw/gorg 0,00 10,54 4,57 3,77 Final Volume m3 400,00 300,46 428,68 486,14

Water additioned

0,00 0,00 328,00 328,00 Simulation time h 34,60 45,00 135,20 135,40

Acids Mass AB (ton) 0,20 0,40 0,16 0,18 AA (ton) 0,09 0,19 0,07 0,08

62


63

Table A. 7. General simulation results for the FWWC process scheme.

FWWC – A FWWC – B

Batch Fedbatch

Fedbatch C=6

Fedbatch C=6 + 4°C

Batch Fedbatch Fedbatch

C=6 Fedbatch C=6 + 4°C

Total mass produced BuOH (ton) 7,58 53,28 53,43 53,72 7,59 54,02 54,04 54,04 IpOH (ton) 2,99 18,63 18,67 18,85 3,00 19,07 18,98 18,98 EtOH (ton) 1,16 6,26 6,22 6,17 1,16 6,27 6,21 6,21

Total concentration (relative to the initial

aqueous volume)

BuOH (g/l) 18,95 133,20 133,56 134,30 18,97 135,05 135,10 135,10 IpOH (g/l) 7,48 46,57 46,68 47,13 7,49 47,68 47,46 47,46 EtOH (g/l)) 2,90 15,65 15,56 15,43 2,90 15,67 15,51 15,51

Concentration (relative to fermenter

volume)

BuOH (g/l) 4,31 10,82 10,70 10,63 4,35 10,76 10,72 10,72 IpOH (g/l) 1,67 4,14 4,04 5,28 1,70 4,29 4,06 4,06 EtOH (g/l)) 0,21 0,59 0,56 0,66 0,27 0,63 0,60 0,60

Energy consumption compressor

GJ 37,94 144,22 125,55 141,44 38,85 151,88 132,74 132,74

Energy consumption condenser

GJ 358,42 1363,19 1187,13 932,57 433,83 1712,29 1500,53 1500,53

Energy consumption Other units

GJ 148,69 1111,47 1102,43 1206,68 50,93 396,34 391,70 536,07

Gas flowrate m3/h 10567,49 10567,49 10567,49 10567,49 10567,49 10567,49 10567,49 10567,49 Total energy consumption MJ/kgIBE 46,45 33,50 30,84 28,96 44,59 28,48 25,56 27,38

gw/gSolv 3,52 13,55 11,85 12,37 0,82 3,31 2,93 3,81

Mass recovered (Condenser )

BuOH (ton) 0,32 2,49 2,48 1,06 2,39 19,23 19,03 12,98 IpOH (ton) 1,06 7,32 7,27 3,59 2,09 14,91 14,71 12,93 EtOH (ton) 0,31 1,71 1,70 0,61 0,81 4,61 4,55 3,02 Eau (ton) 79,87 301,74 262,23 236,21 106,38 415,39 362,94 421,61

Mass recovered (other units)

BuOH(ton) 6,02 47,39 47,02 48,53 3,95 31,65 31,30 40,26 IpOH (ton) 1,45 10,00 9,91 13,20 0,41 2,91 2,87 5,70 EtOH (ton) 0,79 4,36 4,32 5,30 0,27 1,47 1,44 2,77 Eau (ton) 31,07 119,89 104,83 109,58 7,25 29,44 25,99 33,90 Solv (ton) 8,82 8,85 8,85 8,86 8,84 8,88 8,88 8,90

gw/gorg gw/gorg 3,76 1,94 1,71 1,63 1,57 0,82 0,73 0,70 Final Volume m3 288,82 305,51 360,19 381,46 286,14 282,27 338,28 271,52

Water additioned m3 0,00 328,00 328,00 328,00 0,00 328,00 328,00 328,00 Simulation time h 50,00 190,00 191,60 198,60 51,20 200,40 201,40 253,00

Acids Mass AB (ton) 0,27 0,07 0,10 0,09 0,26 0,06 0,08 0,01 AA (ton) 0,14 0,00 0,01 0,01 0,14 0,00 0,00 0,00

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Table A. 8. General simulation results for the FW5 process scheme

FW5 – A FW5 – B

Batch Fedbatch Batch Fedbatch

Total mass produced BuOH (ton) 7,53 52,89 7,51 54,77 IpOH (ton) 2,98 18,48 2,98 19,47 EtOH (ton) 1,16 6,16 1,15 5,98

Total concentration (relative to the initial aqueous volume)

BuOH (g/l) 18,82 132,22 18,79 136,91 IpOH (g/l) 7,46 46,20 7,46 48,68 EtOH (g/l)) 2,90 15,41 2,89 14,94

Concentration (relative to fermenter volume) BuOH (g/l) 6,81 10,57 6,95 10,11 IpOH (g/l) 2,65 4,16 4,92 8,93 EtOH (g/l)) 0,42 0,49 1,12 0,89

Energy consumption compressor GJ 41,99 84,98 21,47 99,55

Energy consumption other units GJ 148,10 1253,99 72,76 737,31

Gas flowrate m3/h 10567,49 10567,49 10567,49 10567,49 Total energy consumption MJ/kgIBE 16,29 17,27 8,09 10,43

Mass recovered (other units)

Butanol (ton) 5,16 47,43 4,84 48,09 IpOH (ton) 2,06 16,33 1,06 13,56 EtOH (ton) 1,01 5,91 0,71 5,38 Eau (ton) 52,40 216,84 15,14 73,10

Solvent (ton) 21,87 21,89 22,10 22,16 gw/gorg gw/gorg 6,37 3,11 2,29 1,09

Final Volume m3 347,27 510,51 384,43 654,37 Water additioned m3 0,00 328,00 0,00 328,00 Simulation time h 44,40 182,60 45,80 214,40

Acids Mass AB (ton) 0,28 0,14 0,25 0,13 AA (ton) 0,14 0,05 0,12 0,05

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Table A. 9. General simulation results for the FLLR process scheme

FLLR – 2B1O FLLR – Vegetable Oil

Batch Batch

Total mass produced BuOH (ton) 6,83 6,85 IpOH (ton) 2,81 2,80 EtOH (ton) 1,20 1,20

Total concentration (relative to the initial aqueous volume)

BuOH (g/l) 17,08 17,13 IpOH (g/l) 7,02 7,01 EtOH (g/l)) 3,00 3,01

Concentration (relative to fermenter volume)

BuOH (g/l) 7,25 6,75 IpOH (g/l) 5,79 4,69 EtOH (g/l)) 2,76 2,06

Energy consumption Regeneration GJ 81,51 60,04 Gas flowrate m3/h 0,00 0,00

Total energy consumption MJ/kgIBE 7,52 5,53

Mass recovered in the extracting solvent

Butanol (ton) 2,81 3,99 IpOH (ton) 0,33 0,88 EtOH (ton) 0,06 0,36 Eau (ton) 1,61 0,82 Solv (m3) 93,79 860,99

gw/gorg gw/gorg 0,50 0,16 Final Volume m3 399,37 399,96

Water additioned m3 0,00 0,00 Simulation time h 37,07 37,23

Acids Mass AB (ton) 0,20 0,20 AA (ton) 0,35 0,37

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