Comparison of Reactive Distillation and Dual Extraction Processes for

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Comparison of Reactive Distillation and Dual Extraction processes for the separation of acetone, butanol and ethanol from fermentation broth Antti Juhani Kurkijärvi, Kristian Melin, and Juha Lehtonen Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.5b03196 • Publication Date (Web): 29 Jan 2016 Downloaded from http://pubs.acs.org on January 31, 2016

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Industrial & Engineering Chemistry Research

1

Comparison of Reactive Distillation and Dual

2

Extraction processes for the separation of acetone,

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butanol and ethanol from fermentation broth

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Antti J. Kurkijärvi*, Kristian Melin, and Juha Lehtonen

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School of Chemical Technology, Department of Biotechnology and Chemical Technology, Aalto

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University, POB 16100, 00076 Aalto, Finland.

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Butanol, extraction, separation, downstream processing

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Two processes, Reactive Distillation (RD) and Dual Extraction (DE), were presented and

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compared for the separation and purification of acetone, butanol and ethanol (ABE) from

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fermentation broth. The Reactive Distillation produces all the extraction solvents needed from

11

the fermentation products, while the Dual Extraction utilizes extremely effective, but non-

12

biocompatible solvents in extraction. In this work these processes were simulated using Aspen

13

Plus. The sugars consumed in the ABE fermentation were produced using SO2-ethanol-water

14

(SEW) pulping from lignocellulosic biomass. According to the simulations DE was more energy

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efficient of these two processes with energy consumption of 4.97 and 6.12MJkg-1 for all products

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and butanol, respectively. RD consumed 11.02 and 15.98MJkg-1 for all products and butanol,

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respectively. According to the economic analysis, the total investment costs were very similar for

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both of the processes. The DE process, with its slightly higher sales revenue, proved to be a more

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economical option. The economic analysis also showed that 83.3 to 94.9% of the total operating

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costs were caused by the price of lignocellulosic material and its pretreatment. Therefore,

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significant advances have to be made in the pretreatment of lignocellulosic material, for it to be

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noteworthy option for production of fermentable sugars.

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1 Introduction

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Currently, practically all butanol is produced from fossil sources. In this process propene is

25

hydroformylated with synthesis gas (CO and H2) to produce butyraldehyde, which is then

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hydrogenated to butanol. However, with growing energy demands, limited resources of fossil

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fuels and ever growing environmental concerns, the interest in producing platform chemicals and

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fuels from renewable biomass has increased1. Butanol has the potential to replace ethanol as the

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bio component in gasoline due to its better fuel properties: higher energy content, lower

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volatility, less hygroscopic nature and a better compatibility with older combustion engines and

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existing fuel infrastructure2. Due to these reasons and the high petroleum price, numerous

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biobutanol plants started operating in China during 2007 and 20083. Similarly in 2012 and 2013

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ethanol plants in the United States were retrofitted to biobutanol production by Gevo4 and

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Butamax5. In 2013 Abengoa Bioenergy announced that it plans to start commercial-scale

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production of butanol in 20156 and Green Biologics announced plans to start production of

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butanol and acetone with genetically manipulated microbial strains in the United States in 20167.

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However, it is not clear whether the subsequent drop in crude oil price has affected these plans.

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The main biochemical route to biobutanol production is often referred to as the acetone-butanol-

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ethanol (ABE) fermentation. It converts biomass to butanol by fermentation with bacteria of the

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Clostridium spp. The ABE process has a long industrial history and it is easily the most


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established production route to biobutanol2. However, in the western countries ABE

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fermentation became economically unfavorable in the 1950's8. Its key challenges are low

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productivity, sensitivity to lignin and energy intensive product recovery2.

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The Thermochemical route to biobutanol production starts with the gasification of biomass into

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synthesis gas, which is cleaned from impurities and catalytically converted to butanol and other

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alcohols via aldol condensation reactions9. This route is still in development stages and it has not

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been commercially demonstrated. The thermochemical route can handle wider range of

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feedstocks than the biochemical route, as synthesis gas can also be produced from lignin9. Also

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the product separation is easier, because the product concentrations are high in the product

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mixture. Its disadvantages include low CO conversion, high temperature and pressure

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requirements and highly exothermic reactions, which make the temperature control difficult2.

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Also the process is much more complex than the biochemical route, as it contains reactors for

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biomass pretreatment, indirect steam gasification, synthesis gas purification and mixed-alcohol

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synthesis, where the biochemical route only contains a fermenter or fermenters2.

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In addition to the thermochemical process, other novel methods are also developed. For example

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Abengoa Bioenergy is developing a catalytic ethanol to butanol process6,10. However, not much

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is known about this process. It involves catalytic condensation of ethanol to produce butanol

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through Guerbet reaction11. It could be speculated that the selectivity towards butanol in the

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process has to be very high for the process to be competitive against fossil butanol. If the butanol

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was to be used as a traffic fuel, its price should be competitive against ethanol, which is very

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difficult to achieve if it is the starting material of the process. On the other hand, buyers can be

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speculated to pay a little premium for biobutanol and government paid subsidies for example in

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the United States will also help.


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It should be noted that even though processes for biobutanol production exist, they are all

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economically challenging. To improve the economics, many novel product purification methods

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have been suggested in the literature, but to our best knowledge none of them have been utilized

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on industrial scale. Some suggested methods, with their corresponding energy requirements, are:

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24.2 MJ kg-1 for steam stripping12, 18.4 MJ kg-1 for traditional, continuous distillation13, 13.8 MJ

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kg-1 for gas stripping12, 13.3 MJ kg-1 for oleyl alcohol extraction13, 8.2 MJ kg-1 for adsorption-

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desorption12, and 4.8 MJ kg-1 for mesitylene extraction13. Other methods have also been

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suggested: pervaporation14, perstraction15, critical fluid extraction16, adsorption-desorption12,

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hollow-fiber reactors17, reverse osmosis18, liquid membranes19, salt-induced phase separation20

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and continuous flashing21.

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Currently significant capital and overall production costs are involved in the processes required

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to produce fermentable sugars from lignocellulosic material2. Despite this, lignocellulosic

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biomass has the potential to become a low-cost substrate for biobutanol production3.The

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processes, where lignocellulosic biomass is converted to fermentable sugars are commonly

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called pre-treatment. One such method is the SEW pulping, which can be considered a hybrid

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between acidic sulfite and organosolv pulping. It utilizes easily evaporable components, ethanol

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and SO2, for the fractionation of lignocellulosic material. When compared with conventional

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pulping methods, the chemical recovery is simplified, which results in lower capital costs, while

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producing high yield of fermentable sugars22. In SEW pulping biomass is fractionated into fibers

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and spent liquor. The fibers can be utilized after enzymatic hydrolysis, and the spent liquor,

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which contains the dissolved sugars, can be utilized as a feedstock for ABE fermentation after

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purification. In recent years, extensive work has been done to optimize the SEW pulping process

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for fermentative sugars production23. In this work, the product mixture from the SEW process


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was used as a feed of the ABE fermentation, but the pretreatment process was not considered in

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detail.

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The main aim of this paper is to present and compare two processes for the recovery of

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lignocellulosic based ABE components from dilute fermentation broth. The methods included in

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this study were Reactive Distillation (RD) and Dual Extraction (DE). The RD method utilizes the

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intermediate acids from the fermentation to produce the extraction solvents used in ABE

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recovery. Production of esters from carboxylic acids and alcohols by reactive distillation is

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relatively well-known technology; for example, the production of ethyl acetate from acetic acid

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and ethanol24, or the reactive distillation of pyrolysis oil with butanol to reduce its carboxylic

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acid content25. Furthermore, simultaneous butanol fermentation and lipase-catalyzed butyl

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butyrate production has been suggested as an alternative recovery method for butanol, since

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butyl butyrate is much easier to separate from aqueous environment than butanol26. The DE

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method utilizes non-biocompatible, but highly effective extraction solvents in ABE recovery.

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This process resembles conventional extraction processes, where the extraction solvents are

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recycled using distillation. The main exception is that two extraction columns are used instead of

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one. The purpose of the second extraction is to remove the remains of the non-biocompatible

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solvent from the broth so that it can be safely recycled back to the fermenter. The overall goal of

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this research is to present and evaluate these new, cost effective methods for ABE recovery, and

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thus promote the utilization of these promising technologies.

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2 Materials and methods

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2.1 Process simulations


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All simulations in this work were performed using Aspen Plus software (Version 8.4). The Non-

109

Random Two-Liquid (NRTL) and Universal QuasiChemical (UNIQUAC) activity coefficient

110

models were used in the modelling. To describe liquid-liquid equilibria (LLE), new experimental

111

interaction parameters were regressed for NRTL from data measured in this study. Otherwise

112

default parameters from Aspen Plus were utilized throughout the work.

113

In all simulations, the extraction columns were operated adiabatically at approximately 37ºC and

114

1bar. The heat exchangers were assumed to operate with a 5ºC temperature difference between

115

the incoming cold stream and the outgoing hot stream. The distillation columns consisted of 20

116

ideal stages and the pressure profiles inside the columns were from 1.3 to 1.0bar. The feed plates

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of the distillation columns were selected by matching the concentrations of the feeds to the

118

concentration profiles inside the columns and the reflux ratios were minimized so that the desired

119

product purities were achieved. The purge flow from the aqueous fermentation broth loop was

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adjusted so that the sugar concentration of the fermentable sugars (glucose, mannose and 70% of

121

xylose) in the fermenter feed was constantly at a concentration of 50gl-1. Like extraction units,

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the flash unit was also assumed to behave adiabatically at 1bar.

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2.1.1 Feed from the SEW Process

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SEW-process is one method to produce fermentable sugars from lignocellulosic biomass such as

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spruce chips22. In SEW process lignocellulosic biomass is fractionated to pulp (cellulose), C5

126

monosugars and lignin by ethanol, water and SO2. Furthermore, the pulp can be hydrolyzed

127

enzymatically to C6 monosugars with 20% solid concentration as presented by Sklavounos1. In

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this work, the output stream from the SEW process is used as the feed for the ABE fermentation.

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Its composition is presented in Table 1.


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Table 1. Feed to the ABE fermentation process

Mass flow

Concentration

(kgh-1)

(gl-1)

Water

2731.6

885.5

Acetic Acid

7.4

2.4

Ethanol

6.9

2.2

Glucose

481.2

156.0

Galactose

25.2

8.2

Mannose

111.0

36.0

Xylose

59.4

19.3

Arabinose

11.9

3.9

Component

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The sugar concentration of the ABE fermentation should not exceed 60gl-1 as the substrate

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inhibition begins to restrict the sugar utilization of the microbes27. As can be seen from Table 1,

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the sugar concentration of the feed stream is too high to be used in fermentation directly.

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Therefore, before feeding it to the fermenter, it is diluted with the recycled aqueous extraction

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raffinate, which contains only a low amount of sugars and fermentation products.

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2.1.2 ABE fermentation

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The ABE fermentation in this work is assumed to occur in a continuously operated, unmixed

138

fermenter8.The reactions of the different sugars to ABE products, intermediate acids, hydrogen,

139

carbon dioxide and water are based on experimental data8,28. The once-though conversion of the

140

fermentation yielded approximately 0.36g of ABE and acids per gram of fermentable feed. The

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total conversion of sugars was 92.8wt-%, which produced a total solvent yield of 0.30gg-1. If the

142

acids are also considered, this yield becomes 0.32gg-1. Sklavounos1 demonstrated that in a

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continuous process, a solvent yield of 0.25gg-1 was achieved with spruce hydrolysate without


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recycling. Since in this study most of the unutilized sugars are recycled, a somewhat higher value

145

of 0.30gg-1 can be expected even at industrial scale.

146

In this work the following yields were assumed per gram of fermented sugar: 0.0614g, 0.200g,

147

0.0155g, 0.0404g, 0.0169g, 0.0103g, 0.2970g and 0.3585g for acetone, butanol, ethanol, acetic

148

acid, butyric acid, hydrogen, carbon dioxide and water, respectively8,28. The different sugars

149

were assumed to have different reactivities in the reactor. Glucose and mannose were assumed to

150

be completely fermentable as only 70% of xylose was assumed to be consumable by the

151

microbe22. The remaining sugars, galactose and arabinose, were assumed to be completely non-

152

fermentable, and they were removed from the process with the aqueous purge22. Experiments

153

carried out with isotopically labelled components showed that 15-55% of acetic acid and 2-85%

154

of butyric acid converts to butanol and acetone during the solventogenic phase8. In this work it

155

was assumed that 25% of acetic and butyric acids present in the reactor feed react further to

156

acetone and butanol in the solventogenic phase of the ABE fermentation. With these

157

assumptions, the fermenter output without recycling was 44.6, 138.4, 20.3, 26.2 and 8.6kgh-1 of

158

acetone, butanol, ethanol, acetic acid and butyric acid, respectively.

159

2.2 Selecting extraction solvents

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2.2.1 Reactive Distillation method

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The ABE fermentation products can react with each other to produce higher boiling solvents.

162

From these reaction products butyl butyrate was found to be the best extraction solvent for ABE.

163

It also has low water solubility and according to Santos and Castros29, it can be used as a natural

164

flavoring agent in food products. Therefore, butyl butyrate was selected to be the main extraction

165

solvent in the RD method. During butyl butyrate production some side products are also formed,


8

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mainly butyl acetate with small amounts of ethyl butyrate and ethyl acetate. Their water

167

solubilities (4.930, 6.831 and 80gl-1 32, respectively) are high when compared to butyl butyrate, and

168

therefore these products are not ideal extraction solvents. This means that a significant amount of

169

these light esters would be lost to the aqueous phase during extraction. Luckily, this is not a

170

problem as the formation of these side products in the RD process is minimal. Furthermore,

171

Jenkins33 stated that butyl butyrate, butyl acetate and ethyl butyrate were promising bio

172

components for aviation fuels, diesel and gasoline. This means that the excess extraction solvents

173

produced by the RD process could possibly be sold as a mixture, without further purification

174

steps.

175

2.2.2 Dual Extraction method

176

Kurkijärvi34 stated that alcohols and alkanes with at least nine carbons would be the best solvents

177

for the DE process, because smaller components presented azeotropic behavior with water.

178

However, in that study only straight chained components were tested. Cyclic and branched

179

alcohols offer the significant advantage that some of them are azeotrope free. As a rule of thumb

180

alcohols with lower boiling points have higher water solubility and higher extraction capacity for

181

ABE. Therefore, it was possible to find extraction solvents that are more effective than the

182

components used in our previous works34,35. It should be noted that only the components

183

included in the Aspen default databanks were considered as potential extraction solvents.

184

The solvents considered for the first extraction were selected using the following criteria. As

185

stated by Kurkijärvi34, the first solvents should be an alcohol. To facilitate an easy solvent

186

recovery, the component should be azeotrope free and the boiling points should not be within the

187

boiling point range of the ABE, which is 56-117.4°C. Alcohols with boiling points lower than


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56°C have very high water solubility, and are for that reason not easily applicable as extraction

189

solvents. A higher boiling alcohol was needed, but to minimize the energy consumption, the

190

boiling point should be as low as possible. The azeotrope search was performed with Aspen Plus

191

and the used activity coefficient model was UNIQUAC. The boiling points were taken directly

192

from the databanks of Aspen, and the lowest boiling alcohol that met all the set criteria was 2-

193

methyl-1-hexanol. It was therefore selected to be the solvent in first extraction. In this work 2-

194

methyl-1-hexanol will be referred to as methyl hexanol.

195

The solvents considered for the second extraction were cyclic and branched alkanes. In this case

196

the boiling points do not have to be outside the boiling point range of the ABE, as separating the

197

two extraction solvents from each other is of a higher priority. Two additional conditions were

198

set for the solvent: lack of azeotropes and liquid state at extraction conditions. It was assumed

199

that components with boiling points at least 10°C above extraction temperature are mainly in

200

liquid state in the extraction conditions. The azeotrope search was performed with Aspen Plus

201

and the used activity coefficient model was UNIQUAC. Cyclopentane, with its boiling point of

202

49.3°C was the first solvent to fulfill all the criteria and it was selected to be the second

203

extraction solvent.

204

2.3 Liquid-liquid equilibrium measurements

205

Butyl butyrate, methyl hexanol, cyclopentane, butanol, acetone, acetic acid and butyric acid were

206

from Sigma Aldrich and their purities were 99.9%. Ethanol with purity of 99.6% was purchased

207

from Altia oyj. The water was distilled, and ion exchanged. The purity of it was not analyzed.

208

To measure the LLE behavior of the selected solvents with ABE, an aqueous mixture containing

209

0.8, 0.4, 0.167, 0.1 and 0.1m-% of butanol, acetone, ethanol, acetic acid and butyric acid,


10

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210

respectively, was mixed with an organic phase consisting of one of the selected extraction

211

solvents. Because the DE utilizes two extraction solvents, another set of experiments were

212

carried out. The aqueous phase was otherwise identical to the first set, but it contained additional

213

0.1m-% of methyl hexanol. The organic phase in this experiment was cyclopentane. In both

214

cases three experiments were carried out with each extraction solvent: the organic to aqueous

215

phase mass ratios were 1:20, 1:10 and 1:6.67.

216

The experiments were carried out in 100ml glass bottles. The bottles were filled so that the gas

217

volume inside was small, approximately 5ml. This ensured that any errors caused by

218

vaporization would be minimized. All the experiments were carried out at 37.0°C with constant

219

mixing for at least 24 hours, after which the bottles were stored at 37.0°C. The temperature was

220

controlled within 0.05°C throughout the experiments. During these experiments the mixtures

221

were visually monitored to verify that no stable emulsions were formed. Both of the phases were

222

analyzed using a gas chromatograph with flame ionization detector (GC-FID). The organic

223

phases were further analyzed with a gas chromatograph with a mass spectrometer (GC-MS) to

224

determine their water content. The aqueous phase water content was calculated from mass

225

balance. To make the results more reliable all experiments were carried out in duplicate, all

226

samples were analyzed two times and the results from these analyses were averaged. The

227

distribution coefficients for the components are defined as mass fraction ratios in organic and

228

aqueous phases as shown in equation 1.

Di =

229

x i ,Org x i , Aq

(1)

230

Where Di is the distribution coefficient for component i; xi,Org is the mass fraction of component i

231

in organic phase and xi,Aq is the mass fraction of component i in aqueous phase. In practice,


11


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232

distribution coefficients were determined by plotting components mass fraction in organic phase

233

as a function of its mass fraction in aqueous phase with different organic to aqueous ratios. The

234

values of the distribution coefficients could then be observed as the slope of the graph.

235

2.3.1 Analysis methods

236

The GC-FID was a Hewlett Packard 6890 Series gas chromatograph, with a 60m long Zebron

237

ZB Wax plus column, with 0.25mm inner diameter and 0.25µm phase thickness. The carrier gas

238

was helium, the injection volume was 0.5µl and split ratio was 1:50. The temperature program

239

begun with a ten minutes hold at 40°C after which the temperature was elevated 10°Cmin-1 to

240

230°C followed by a hold of two minutes at that temperature.

241

The GC-MS was manufactured by Agilent Technologies: the gas chromatograph was 7890A and

242

the mass spectrometer was 5975C VL MSD. The used column was 30 m long Agilent

243

Technologies Innowax with 0.25mm inner diameter and 0.25µm phase thickness. The carrier gas

244

was helium, the injection volume was 0.5µl and split ratio was 1:50. The temperature program

245

begun with a 15 minutes hold at 40°C after which the temperature was elevated 7.5°Cmin-1 to

246

200°C followed by a hold of 15 minutes.

247

2.4 Process descriptions

248

2.4.1 Reactive Distillation method

249

The RD process with mass balances is presented in Figure 1. The sugar rich feed from the SEW

250

pre-treatment is mixed with the sugar lean raffinate from the second extraction. This forms the

251

feed to the fermenter, where the ABE is formed. The fermentation gases are washed with water

252

in a washing tower with 5 ideal stages to prevent losses of the more volatile components with the


12

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253

fermentation gases, and the resulting liquid phase was directed back to the fermenter, because it

254

was too diluted to be sent to the product purification stages.

255

From the fermenter the broth is fed to a concurrent Pre-Extraction, with 3 ideal stages. Here a

256

small amount of butyl butyrate was used to extract the main part of the butyric acid together with

257

butanol and a small part of ethanol, acetone and acetic acid. The organic phase from this Pre-

258

Extraction is fed to a Reactive Distillation column, where the mixture is esterified according to

259

reactions 1-4.

260

+ ↔ +

(1)

261

+ ↔ +

(2)

262

+ ↔ ℎ +

(3)

263

ℎ + ↔ ℎ +

(4)


13


10 231 30

Flow Temp 8 Flow 75 kg/h

1

2

Ads1

Temp 20 °C

13 Flow 983 kg/h Temp 122 °C

kg/h °C

Absorp

9 Flow

77

kg/h

Temp

32

°C

Flow 975 kg/h 15

Flow 3435 kg/h Temp 34 °C

11

Pre-Extraction

3

Flow 659 Temp

31

12 Flow 610 Temp

42

165°C

Reactive Distillation

ABE Reactor

Flow 933 kg/h Temp 101 °C 16

kg/h

Flow 42 kg/h

°C

14

Q1

kg/h °C

Flow

609

kg/h

Temp

176

°C

17 Flow

595 kg/h

Flow

25 14

kg/h

Temp

42 °C

Temp

42

°C

26 4

20

Flow 15207 kg/h

Flow 18439 kg/h

Flow 15 kg/h

Temp

30

°C

Temp

32

°C

Temp 45 °C

30

°C

Flow

2221

kg/h

Temp

123

°C

Flow

22 2072

kg/h

23

Temp

121

°C

61 kg/h 65 °C

Butanol

Q3 19 Flow 3069 kg/h

5

Temp

45

°C

76°C Q4

EtOH-Acetone

Flow 3243 kg/h

Solvent

18 Temp

21

Ads2

7

Main Extraction

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 14 of 45

Flow

6

kg/h

Temp

25

°C

Flow

27 36

kg/h

Temp

25

°C

Flow

19

kg/h

Temp

0

°C

28

86°C

Q2 6

264

Flow

3058

Temp

30

24 kg/h °C


29 Flow

130

kg/h

Temp

39

°C

265

Figure 1. The Reactive Distillation process. The heat fluxes of the heat integrations are 324.1,

266

35.5, 903.5 and 1.0MJh-1 for Q1, Q2, Q3 and Q4, respectively.

267

The reactions in the RD process were modelled using equilibrium assumptions. The reaction

268

equilibria were based on literature data for butyl butyrate36, butyl acetate37 and ethyl butyrate36.

269

The reaction equilibrium for ethyl acetate formation was calculated by the Gibbs energy

270

minimization method in Aspen. Because it is unlikely that the reactions would reach their

271

thermodynamical equilibria in the Reactive Distillation column, the values of the equilibrium

272

constants were calculated at a temperature, which was 10°C below the actual reaction


14

Page 15 of 45

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273

temperature for butyl butyrate and ethyl butyrate and 10°C above the reaction temperature for

274

butyl acetate. This resulted in equilibrium constants of 8.4, 8.0 and 9.8 for butyl butyrate, ethyl

275

butyrate and butyl acetate, respectively. When compared to the values calculated at the actual

276

reaction temperature, these values are reduced by 11.4, 12.8 and 2.1% for butyl butyrate, ethyl

277

butyrate and butyl acetate, respectively. It was estimated that this approach would describe the

278

behavior of the reactions in the Reactive Distillation column.

279

The esterification reactions are performed at the lowest stage of the Reactive Distillation column,

280

and water is removed from the reaction zone. It is adsorbed from the top product, thus making

281

the reflux flow of the column to be practically water free. The distillate contains acetone,

282

unreacted alcohols, light esters and butyl butyrate. The esters in this stream are used as the

283

makeup solvent for the main extraction. To prevent the accumulation of the light esters in the

284

Main Extraction solvent loop, some solvent is returned to the Pre-Extraction. The bottom product

285

of the Reactive Distillation column, which consists mainly of butyl butyrate, is returned to the

286

pre-extraction as the extraction solvent. It should be noted that this process can be operated so

287

that it produces more esters than is needed in the extractions. On the other hand, it can also be

288

operated in such manner that no excess extraction solvents are formed, thus maximizing the ABE

289

yield.

290

After the Pre-Extraction the fermentation broth goes to the Main Extraction, which contains 5

291

ideal stages. From here, the main part of the broth is recycled back to the fermenter. At this

292

stage, the broth contains mainly water, the unreacted sugars and low amounts of ABE. A purge

293

stream is separated from the broth, which is sent to waste water treatment done by for example

294

anaerobic digestion. This prevents the accumulation of unreactive sugars and other impurities to

295

the broth. The organic phase from the Main Extraction is fed to the Solvent column. Here the


15


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Page 16 of 45

296

butanol, ethanol and acetone are separated from the esters. In order to minimize the amount of

297

these esters in the distillate, the main part of the Butanol column bottom product is fed as reflux

298

stream to the top the Solvent column. A molecular sieve removes water from distillate of the

299

Solvent column, which at this point contains ABE, small amounts of esters and non-condensable

300

gases, which mainly consist of CO2. The butanol is separated from the lower boiling components

301

in the Butanol column, and the ethanol and acetone are then separated in the in EtOH-Acetone

302

column. In Figure 1, the dotted lines are the heat integration, which was used to minimize the

303

energy consumption of the process. In most columns the bottom product heats up the feed. The

304

Solvent column feed is also heated with the butanol product.

305

2.4.2 Dual Extraction method

306

The fermenter, washing of the fermentation gases, broth cycle and the purge are carried out

307

identically to the RD process. In DE the ABE is first extracted with highly effective, but non-

308

biocompatible solvent. To prevent the toxic effects of this solvent to the microbes, another

309

extraction is performed, which separates the non-biocompatible solvent from the broth before

310

recycling it back to the fermenter. The principle of DE is presented in more detail in our previous

311

works34,35. The DE process with mass balances is shown in Figure 2.


16

Page 17 of 45


8

Absorp


1

9 Flow Temp

54 20

13

kg/h °C

Flow Temp

2


ABE Reactor

316 93

kg/h °C

24

11 139°C

25 23

14 164°C 12 Flow 1870 kg/h Temp 42 °C

Flow 72 kg/h 60°C Temp 54 °C

Ads

Q1

15 155°C

Q4

21

161°C Flow 0.4 kg/h Temp 37 °C

18



Distillation

Q2

4



22


72°C

19 100°C

6

312

Flow Temp

20 2910 37

kg/h °C

Flow Temp

0.5 37


27

30

Decant

Flow 326 kg/h 80°C Temp 94 °C Q5

Temp Flow

146 80

kg/h °C

93°C Q6

AqDist

17

28

OrgDist

Q3


5

107°C


Flash

Extraction2

16

26

78°C



LightDist



Regeneration

3

Extraction1

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


104°C

29 kg/h °C

Flow 140 kg/h 85°C Temp 117 °C


313

Figure 2. The Dual Extraction process. The heat fluxes of the heat integration are 524.5, 44.4,

314

44.5, 0.6, 12.2 and 8.0MJh-1 for Q1, Q2, Q3, Q4, Q5 and Q6, respectively.

315

The organic stream from the Extraction1 is fed to the Regeneration column, which separates and

316

recycles the solvent from the ABE products. Due to the relatively high water solubility of the

317

extraction solvent, the distillate of the Regeneration column also contains significant amounts of

318

water and small amounts of non-condensable gases, mainly CO2. The Distillation unit separates

319

water-butanol mixture from the lower boiling components. The water-butanol mixture is fed to

320

Decanter, where an aqueous phase and butanol rich organic phase are formed. This phase

321

separation ensures that the OrgDist column operates above and the AqDist column operates

322

below water-butanol azeotropic point. This means that the bottom products from these columns


17


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Page 18 of 45

323

can be obtained as pure components, which would not be possible without the phase separation

324

in the Decanter. Furthermore it means that the distillates from these columns are near the water-

325

butanol azeotropic point. For that reason, the distillate from the OrgDist column is sent back to

326

the Decanter. The reason why the distillate from the AqDist column is not fed to the decanter is

327

the small amount of ethanol, which would accumulate in the Decanter-AqDist column loop. By

328

feeding the distillate from the AqDist column to the Distillation unit, both ethanol and butanol

329

can be recovered. The AqDist and the OrgDist columns have no condensers at all and thus the

330

feed streams act as reflux flows of the columns. The water from the light ethanol-acetone

331

fraction is first removed with adsorption column, and after this the ethanol and acetone are

332

separated in the LightDist column. In this column the purity of the distillate, in other words

333

acetone, is strongly dependent on the reflux ratio and thus the energy consumption of the

334

column.

335

The purpose of the second extraction is to separate the toxic component from the broth. To make

336

sure that the recycled broth is biocompatible, the amount of non-biocompatible solvent in the

337

raffinate was fixed to 20ppm. It can be assumed that this low solvent concentration does not have

338

any effect on the microbes in the ABE fermentation13. The organic stream from the second

339

extraction contains both extraction solvents, and the purpose of the Flash is to recycle these

340

solvents back to the extractions. The usage of flash drum instead of distillation column is

341

possible because the second solvent is a very low boiling component. This way very little

342

separation stages are needed to separate if from the high boiling solvent. It should be noted that

343

the ABE components are not recovered at all from the second extraction.

344

The only distillation column that has a liquid distillate in this process is the LightDist column.

345

All other columns produce their distillates in vapor phase. To reduce the energy consumption of


18

Page 19 of 45

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346

the process, heat integration was used. Similarly to Figure 1, the heat integration is marked with

347

dotted lines. The bottom product of the Regeneration column was used to heat up its feed and the

348

stream going to the Flash drum. Also the bottom products of the OrgDist, AqDist and LightDist

349

columns were used to heat up their own feeds.

350

2.4.3 Waste streams from proposed processes

351

The waste streams, the Purge flow and the Fermentation gases, are relatively identical in both

352

RD and DE processes. Therefore their utilization and potential further processing are also

353

assumed to be identical.

354

2.4.3.1 Fermentation gases

355

As a side product the ABE fermentation produces fermentation gases, which mainly consist of

356

carbon dioxide and hydrogen. In industrial scale this gas mixture can be used to generate heat

357

and power38. Another option would be to recover the hydrogen from the fermentation gases.

358

Approximately three moles of gases is formed per every mole of glucose consumed. The molar

359

ratio of carbon dioxide and hydrogen in this gas is approximately 3:2, and approximately 1.1t of

360

gases are formed per 1t of solvents39. Before the fermentation gases are processed any further,

361

the solvents present in this gas can be collected in an absorption tower2. The amount of solvent

362

recoverable by this method is about one ton per 100 000m3 of fermentation gas, which represents

363

about 1-2% of all the total solvents produced2,39.

364

The carbon dioxide could be separated from the hydrogen by cryogenic methods, washing with

365

water in basic conditions or by membrane processes. Jones and Woods8 suggests that this

366

separated carbon dioxide could be dried, purified and sold as a bulk gas, liquid CO2 or as dry


19


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Page 20 of 45

367

ice8,39. The hydrogen could be used in various ways, including hydrogenation,

368

hydrodeoxygenation, ammonia synthesis or used as fuel8,40. Moreira41 concluded that the ABE

369

fermentation gas would be ideal for methanol synthesis. It could also be converted to methanol

370

and formaldehyde39. Jones and Woods8 suggested that this gas mixture could be used in

371

production of methane by methanogenic bacteria or that the hydrogen could be used as a fuel in

372

fuel cell applications for electricity production. The usage of the hydrogen would naturally

373

depend on what kinds of processes are present nearby the ABE fermentation equipment.

374

2.4.3.2 Broth purge and molecular sieves

375

To avoid unwanted accumulation of components into the fermentation broth, a purge flow is

376

separated from the fermentation broth. Components that would accumulate in the process are for

377

example non-fermentable sugars, water and possible impurities from the pre-treatment. The

378

purge stream consists of mainly water, but some sugars and ABE products are also present:

379

generally the components for which the extraction yields are the lowest. This purge steam could

380

be sent to further treatment to recover those components or it could be sent to for example biogas

381

production. According to Chen42, the butyric and acetic acids can be converted into biogas. Only

382

at high concentration was acetic acid reported to be inhibitory to the anaerobic digestion process.

383

The removal of water is done by adsorbing it into molecular sieves. The energy consumption of

384

adsorption is caused by the regeneration, which is usually done by heating and potentially

385

applying vacuum to the adsorbent. Modelling of this kind of regeneration is not very

386

straightforward, so literature values were used. Kumar43 stated that regenerating water

387

adsorbents consumes from 1.94 to 6.62MJkg-1 water. An average of these two values, 4.28MJkg-

388

1

water, was used in this work to simulate the energy consumption of the adsorption units.


20

Page 21 of 45

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389

2.5 Economic evaluation of the processes

390

To study the economics of the processes at a commercially relevant scale, the capacities were

391

increased by a factor of 50, making the new capacities 50t/h of dry wood. First, the annual costs

392

and product revenues were calculated with this new capacity. After this, the heat exchanger areas

393

and total column weights were determined for both processes. The investment costs were

394

calculated based on reported price correlations for columns, heat exchangers, vessels and

395

pumps44. The purpose of this economic analysis was to study the differences in investment costs

396

of RD and DE process. Therefore, identical parts of the processes such as fermenter,

397

fermentation gas washing and treatment of the purge water were not considered. The economic

398

parameters used in the evaluation and their values are presented in Table 2.

399

Table 2. The parameters used in the economic analysis

Economic parameter

Value

Butanol price

1100eur/t45

Acetone Price

1100eur/t45

Ethanol Price

900eur/t45

Ester product price

1200eur/t27

Make-up methyl hexanol Price

1500eur/t45

Make-up cyclopentane

900eur/t45

Steam Price

32eur/MW46

Lang factor

4.7445

Annual operation of plant

8000h/a46

400

The chemical prices in Table 2 are the European bulk solvent prices of 201445. The lang factor is

401

used to calculate the total installed equipment cost based on purchased equipment cost.


21


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Page 22 of 45

402

2.5.1 Operating costs

403

The annual costs and revenues from products were calculated with equation 2 using the

404

economic parameters given in Table 2.

405

= ∑ −

406

Where S is sales income in euros, pi is yield of products (butanol, acetone, ethanol and ester

407

solution), ci is the value per unit given in Table 2 and fi is the amount of each input (steam,

408

make-up solvents etc.).

409

2.52. Investment cost

410

Based on the sum of the purchased equipment from Table S13 and S14, the total installed cost

411

was calculated according to Equation 3.

412

! = "#$%&'( )*+*

(2)

"#$

(3)

%&'&

413

Where TIC is the Total Installed Cost of equipment in euros, CPI is the Chemical Engineering

414

Plant Cost Index (CEPCI) for years 2015: 560.7 and 2010: 550.8, x is the exchange rate 0.91

415

euros/USD, l is the Lang factor and Etot is sum of calculated equipment costs in USD. Since the

416

price correlation were given for year 2010 the cost were transformed to present cost by the

417

CEPCI47.

418

In capital cost calculations only the process equipment that were different in RD and DE

419

processes were included in the calculation, since the point was comparison not absolute

420

investment cost. The distillation column diameters were sized based on the standard sizing tools

421

within Aspen plus. The sizing was done using tray columns with Glitch Ballast trays. The


22

Page 23 of 45

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422

distillation columns were sized assuming overall column efficiency of 70% and tray spacing of

423

61cm, with 6m extra height to accommodate the structures at top and bottom of the columns. The

424

reflux tanks of the distillation columns were sized for 15 minutes residence time. All pumps were

425

sized with 20% extra capacity, with 30 meter hydraulic head of water and an overall efficiency

426

of 40%. The water adsorption columns were sized for 1 week continuous operation and

427

duplicated so that one could always be in operation while the other one was in regeneration. The

428

adsorbent was assumed to adsorb 10% of its mass of water and the column was sized assuming

429

2300kgm-3 density and 70% filling fraction of the adsorbent49. The heat exchanger areas were

430

calculated in Aspen. For reboilers and condensers the heat exchanger areas were determined

431

based on the temperatures of the heating and cooling utilities with estimated overall heat transfer

432

coefficient of 850W/m2 K. Saturated steam at 200˚C and 140˚C were assumed to be used for the

433

heating purposes. The incoming cooling water temperature was 35˚C and outgoing temperature

434

55˚C, with the exception of the acetone-ethanol separation columns where 15˚C incoming and

435

45˚C outgoing temperatures were used. The extraction columns were assumed to be randomly

436

packed columns and sized with a flow of 45m3m-2h-1 48. The weights of the process units were

437

calculated assuming cylindrical shape and column wall thicknesses of 5-12mm depending of the

438

diameter of the unit44. All equipment were assumed to be made of carbon steel. The cost of water

439

adsorbents, fermenters, fermentation gas washing towers, valves and other process equipment

440

were not included in to the cost estimate, since they were assumed to be very similar on both

441

cases. It should be noted that the development stages of the RD and DE processes are low, and

442

therefore the exact sizes and materials of the equipment have not been determined. For this

443

reason, the investment cost estimate has a relatively low accuracy.

444

3 Results and discussion


23


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Page 24 of 45

445

3.1 Liquid-liquid equilibrium experiments

446

The reproducibility of the analysis was approximately ±5%. For this reason, it can be assumed

447

that evaporative losses and error from weighing are negligible when compared to the error from

448

analysis methods. The results from the experiments were added to Aspen plus to determine new

449

parameters for NRTL-model. The weighed sums of squares for the different regressions were

450

0.001, 0.062 and 0.022 for butyl butyrate, methyl hexanol and cyclopentane, respectively. Table

451

3 shows the experimental and simulated distribution coefficients after data regression.

452

Table 3. Experimental and simulated distribution coefficients after data regression at 37°C, 1 bar

Butyl Butyratea

Methyl hexanola

Cyclo pentaneb

Experimental Simulated

Experimental

Simulated Experimental Simulated

DWater

0.005

0.005

0.06

0.06

0.0002

0.0002

DAcetone

2.60

2.60

2.44

2.45

0.37

0.37

DEthanol

2.46

2.46

3.89

3.89

1.00

1.00

DButanol

6.18

6.18

12.6

12.6

1.47

1.47

DAcetic acid

1.17

1.17

0.95

0.94

0.15

0.15

DButyric acid

4.22

4.22

10.2

10.2

0.78

0.79

DButyl butyrate

1651

1651

-

-

-

-

DMethyl hexanol -

-

143

150

671

664

DCyclo pentane

-

-

-

8286

8313

-

a

453 454

(aqueous composition 0.8, 0.4, 0.167, 0.1 and 0.1m-% of butanol, acetone, ethanol, acetic acid and butyric acid), b(aqueous phase contained additional 0,1m-% of methyl hexanol)

455

As can be seen from Table 3, the simulations and experimental results were in good agreement

456

and the new parameters for the NRTL model can be considered to be accurate and reliable. The

457

distribution coefficients presented in Table 3, are relatively large for ABE components and


24

Page 25 of 45

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458

minimal for water, as is to be expected. The values for acetone and acetic acid were however

459

slightly unexpected. It would have been more logical, if all the distribution coefficients for ABE

460

products were the largest with methyl hexanol, because of its higher polarity. But this was not

461

the case, as acetone and acetic acid are extracted as efficient7y with both butyl butyrate and

462

methyl hexanol. Differences in the extraction solvents are more apparent when their extraction

463

behavior is compared with ethanol, butyric acid and butanol. Extraction of ethanol is slightly

464

more efficient with methyl hexanol, but the real differences can be seen in the extraction of

465

butyric acid and butanol. For example, when using butyl butyrate over twice the amount of

466

extraction solvent is needed to achieve the same extraction efficiency for butanol when

467

compared with methyl hexanol. According to our knowledge, the selected extraction solvents

468

have not been suggested as extraction solvents for ABE components in the literature. For this

469

reason no literature values for these components were available for comparison.

470

3.2 Process simulations

471

3.2.1 Reactive Distillation

472

The sugars are extracted very moderately in the RD process and all extracted sugars are recycled

473

back to the extraction along with the extraction solvent. The intermediate acids, butyric and

474

acetic acid are extracted in both extractions. From the pre-extraction they are transferred to the

475

solvent production in the Reactive Distillation column and from the main extraction they are

476

mostly recycled back to extraction with the extraction solvent. Acetic acid goes partly to the

477

distillate and therefore ends up into the products. The light esters formed in the RD are

478

transferred to the main extraction and eventually majority of them leave from the process within

479

the aqueous purge flow. Small amount goes to the distillate of the Solvent column and thus to the


25


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Page 26 of 45

480

butanol product. Tables 4 and 5 presents the products and the product losses to purge flow as a

481

function of solvent flow in the main extraction. The organic to aqueous ratio for the pre-

482

extraction was 0.038 in all simulations.

483

Table 4. Product distributions of Reactive Distillation with different organic to aqueous ratios. mOrg:mAq (kgkg-1)

Acetone (kgh-1)

Butanol (kgh-1)

Ethanol (kgh-1)

Ester Product (kgh-1)

0.230

38.0

131.5

19.5

2.0

0.200

37.7

131.1

19.5

4.0

0.175

37.0

128.1

19.5

7.6

0.115

35.7

118.0

19.4

9.3

0.090

34.7

102.8

19.0

10.2

484

Table 5. Product losses into the waste water in the Reactive Distillation as function of extraction

485

solvent flow.

mOrg:mAq Acetone (kgkg-1) (kgh-1)

Butanol Ethanol (kgh-1) (kgh-1)

Acetic acid (kgh-1)

Butyric Butyl Butyl Ethyl Total acid butyrate acetate butyrate sugars (kgh-1) (kgh-1) (kgh-1) (kgh-1) (kgh-1)

0.230

7.7

0.4

2.0

8.3

3.2

2.0

10.3

0.1

101.5

0.200

7.8

0.7

2.3

8.3

3.0

2.0

10.3

0.1

101.4

0.175

8.9

1.6

2.9

8.3

1.2

2.0

10.4

0.1

101.5

0.115

10.0

10.9

3.9

8.9

0.8

2.1

10.3

0.1

101.8

0.090

11.2

24.3

4.4

9.3

0.6

2.1

11.2

0.1

101.8

486

As can be seen from Table 5, practically all ABE is extracted from the broth, when the solvent to

487

aqueous ratio exceeds the value of 0.175 in the Main Extraction. With higher solvent to aqueous

488

ratios, the broth entering the Pre-Extraction contains only the solvents that were formed in the

489

fermenter, and practically no recycled ABE. This is the case where there is the least amount of


26

Page 27 of 45

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490

solvents to be extracted in the pre-extraction. Therefore the ester to ABE ratio in the organic

491

phase from the pre-extraction is at its highest, and the mixture entering the Reactive Distillation

492

column is already close to the reaction equilibria and less esters are formed. When the solvent

493

amount in the main extraction is lowered, the ester to ABE ratio in the organic stream from the

494

pre-extraction is lower and more esters are formed in the Reactive Distillation column. The

495

product purities from the RD process were 94.7, 98.3 and 92.5wt-% for acetone, butanol and

496

ethanol respectively. The ester product had the following composition: 96.7, 2.0 and 0.6wt-% for

497

butyl butyrate, butyl acetate and butyric acid, respectively.

498

As can be seen from Table 5, with larger organic to aqueous ratios the main components in the

499

Purge stream, in addition to water and the unfermented sugars, are butyl acetate, acetic acid and

500

acetone. This is because the extraction efficiencies of acetic acid and acetone are the lowest, and

501

butyl acetate is transferred to the aqueous phase as it has higher water solubility than butyl

502

butyrate. With smaller organic to aqueous ratios the concentration of the main fermentation

503

product, butanol, increases in the broth and for that reason also in the Purge stream. In contrast to

504

other components, the amount of butyric acid in the Purge stream increases as a function of the

505

organic to aqueous ratio. The main extraction does not remove butyric acid from the broth, but it

506

does remove alcohols. So at higher organic to aqueous ratios there are less alcohols to be

507

extracted in the Pre-Extraction, but the amount of butyric acid is practically the same. This

508

reduces the amount of produced esters, which means it also reduces the conversion of the butyric

509

acid. As the unreacted butyric acid remains with the bottom product of the Reactive Distillation

510

column it is transferred back to the Pre-Extraction and thus more butyric acid remains in the

511

aqueous broth. The energy consumption of the RD process is shown in Table 6.


27


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 28 of 45

512

Table 6. Energy consumptions in the Reactive Distillation process as function of extraction

513

solvent flow in the main extraction. Reactive mOrg:mAq Distillation -1 (kgkg ) (MJh-1)

Solvent (MJh-1)

Butanol (MJh-1)

EtOHAcetone (MJh-1)

Molecular Total sieves (MJh-1) -1 (MJh )

0.230

362.5

1393.9

98.3

186.1

157.3

2198.2

0.200

362.5

1368.0

102.2

186.5

149.4

2168.6

0.175

376.9

1347.1

109.1

186.1

140.4

2159.6

0.115

496.8

1317.2

109.8

192.6

132.5

2248.9

0.090

559.8

1290.6

109.8

198.4

128.9

2287.4

514

As can be seen from Table 6, the Solvent column alone is responsible of over half of the total

515

energy consumption of the process. This is because a significant amount of reflux is needed to

516

minimize the amount of esters in the distillate. The net heat demand for the process was 2.16-

517

2.29GJh-1.

518

3.2.2 Dual Extraction

519

In the DE method, the sugars are extracted only moderately, if at all. And what little is extracted

520

is recycled back to the process with the bottom product of the Regeneration column. The

521

intermediate acids, acetic and butyric acid, are extracted more readily. The Regeneration column

522

recovers practically all butyric acid and it is recycled back to the process. Acetic acid is more

523

problematic; approximately two thirds of the extracted acetic acid will go in to the distillate and

524

thus in to the products. However, the extraction is not very effective towards acetic acid as only

525

approximately 10wt-% of acetic acid present in the fermentation broth is extracted. This means

526

that the overall fraction of acetic acid in the products remains small. Table 7 presents the amount

527

of products as a function of extraction solvent flow in the first extraction.


28

Page 29 of 45

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

528


Table 7. Product distributions for the Dual Extraction with different organic to aqueous ratios. Product streams mOrg:mAq (kgkg-1) Acetone Butanol Ethanol (kgh-1) (kgh-1) (kgh-1)

Purge stream Acetone Butanol (kgh-1) (kgh-1)

Ethanol (kgh-1)

Acetic Butyric acid acid -1 (kgh ) (kgh-1)

Total sugars (kgh-1)

0.150

36.4

141.6

20.9

9.1

0.1

1.4

13.3

6.1

101.5

0.125

35.8

141.1

20.5

9.8

0.8

2.2

13.7

6.1

101.7

0.100

35.2

140.5

20.1

10.4

1.4

3.0

14.1

6.1

102.0

0.075

34.5

123.2

20.0

11.1

18.7

3.4

14.8

6.0

102.4

0.050

33.7

105.8

19.8

11.7

35.9

3.8

15.5

6.0

103.0

529

As can be seen from Table 7, practically all butanol is extracted from the broth, when the flow of

530

extraction solvent is over 10wt-% of the aqueous flow in the first extraction. If the organic to

531

aqueous ratio is further increased, the extraction of butanol and ethanol are enhanced only

532

marginally, but the extraction of acetone is more affected. When using the proposed product

533

separation method for the ABE components, the mass based product purities are 95.6, 98.6, 99.2

534

and 99.7wt-% for acetone, butanol, ethanol and water, respectively. The compositions of the

535

Purge flow in DE process follow the trends of the RD. With larger organic to aqueous ratios the

536

main organic components in this stream are unfermented sugars, acetic acid and acetone, as these

537

are the components that are not extracted very effectively. With lower organic to aqueous ratios

538

the amount of butanol becomes the largest. This is because the fermentation produces is a lot

539

more butanol than acetone into the broth and if there is not enough extraction solvent to take the

540

butanol away, then its concentration in the Purge flow will increase. Table 8 presents the energy

541

consumptions of different distillation columns as a function of extraction solvent flow in the first

542

extraction.


29


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 30 of 45

543

Table 8. Energy consumptions at different of different organic to aqueous ratios in Dual

544

Extraction process

mOrg:mAq Regeneration Distillation OrgDist AqDist (kgkg-1) (MJh-1) (MJh-1) (MJh-1) (MJh-1)

Molecul LightDist Total ar sieves -1 (MJh ) (MJh-1) -1 (MJh )

0.150

748.8

63.0

69.8

59.4

41.4

12.6

995.0

0.125

665.6

61.2

68.8

57.2

43.6

12.6

909.0

0.100

582.1

59.4

67.3

55.1

7.7

11.9

821.5

0.075

477.4

52.9

65.2

47.9

41.4

10.4

695.2

0.050

373.0

46.4

63.4

41.0

37.1

8.6

569.5

545

As can be seen, the Regeneration column is by far the biggest energy consumer in this process.

546

On the other hand, of all the distillation columns in the process, it handles the largest process

547

flows, which explains the large energy consumptions. The total energy consumption of the DE

548

process is from 570 to 995MJh-1. The butanol and water products are at approximately 90°C,

549

when leaving the process. The temperature is not very high, but this heat could possibly be used

550

in some other parts of the biofuel plant.

551

4 Comparison of Reactive Distillation and Dual Extraction processes

552

4.1 Comparison based on energy consumption

553

What is common for both RD and DE processes is that they both utilize two extraction columns

554

in a row. The function of the main extraction is the same in both processes: to remove the ABE

555

components from the broth. However, the purpose of the secondary extraction column is quite

556

different in these processes. In RD the purpose is to extract the intermediate acids and some ABE

557

for the production of the extraction solvent. In DE the purpose of the additional extraction

558

column is to remove the non-biocompatible solvent from the broth.


30

Page 31 of 45

559

As can be seen from Tables 6 and 8, the total energy consumption goes down in both processes,

560

when the amount of extraction solvent is lowered. However, this also increases the product

561

losses to the Purge stream. Therefore, the energy consumption as a function of all products and

562

as a function of only butanol is drawn in Figure 3 for both processes. The All Products in Figure

563

3 are ABE and ABE plus esters for DE and RD, accordingly. It should be noted that the

564

intermediate acids, acetic and butyric acid, are not considered as products in Figure 3.

25 Reactive Distillation (Butanol) Reactive Distillation (All products) Dual Extraction (Butanol) Dual Extraction (All products)

20 MJ/kg(products)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


15 10 5 0 70

565

75

80 85 90 wt-% of products recovered

95

100

566

Figure 3. Energy consumptions of the processes as a function of products

567

As can be seen from Figure 3, the energy consumption of DE increases slightly when the amount

568

of extracted products is increased. The energy consumptions start to grow exponentially after 97

569

and 99wt-% for all products and for butanol, respectively. In RD the energy consumptions

570

decreases with increasing recovery until the amount of extracted products is 89 and 99wt-%, for

571

all products and for butanol, respectively. After this an exponential increase in energy

572

consumption starts. The exponential increase in energy consumption occurs, when almost all

573

products have been extracted. In this situation increasing the extraction solvent amount will

574

increase energy consumption, but the amount of products will remain practically unchanged. The


31


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Page 32 of 45

575

decreasing trend of the RD can be explained with the energy consumption of the largest energy

576

consumer in the process, the Solvent column, which is not very sensitive to the amount of

577

extraction solvent. This is because of the very high reflux ratio, which will keep the energy

578

consumption high even when there is only small amount of extraction solvent. This leads to a

579

situation, where increase in the amount of extraction solvent will increase the amount of

580

products, but the energy consumption will increase only slightly, and the ratio between energy

581

consumption and amount of recovered products will decrease. In Figure 3, the curves of the RD

582

are further apart from each other, than the curves of the DE. This is because the esters are heavy

583

as their mass consists partly of intermediate acids, which are not considered to be a product

584

otherwise. In addition, some of the butanol has reacted to esters further increasing the amount of

585

other products than butanol. As can be seen from Figure 3, the curves which contain only butanol

586

start the exponential growth closer to the 100% mark than the curves with All Products. This is

587

because butanol is easier to extract than for example ethanol, acetone or the smaller esters. The

588

DE process has smaller energy consumption than the RD process. At the operating point where

589

approximately 99wt-% of butanol is recovered to the products, the processes consumed 6.12 and

590

15.98MJkg-1 butanol for DE and RD, respectively. At this operating point 97.5 and 89.4wt-% off

591

all products were recovered for DE and RD processes, respectively.

592

In this work the heat exchangers were assumed to operate with 5°C temperature differences

593

between the ingoing cold streams and outgoing hot streams. This is relatively low value as 10°C,

594

is a commonly used in literature. This results in larger and therefore more expensive heat

595

exchangers, but at the same time it improves the energy efficiency of the process, which is

596

important as the energy efficiency is stated to be one of the most important aspects of the ABE

597

separation and purification process13. The conventional separation method for ABE, direct


32

Page 33 of 45

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


598

distillation, has been reported to consume 18.50MJkg-1 butanol13. When compared to this, the

599

energy requirements for both RD and DE are low. So far, the lowest energy consumptions

600

reported in the literature have been 3.6 and 5.8MJkg-1 butanol with DE and mesitylene

601

extraction, respectively13,34. These values are lower, but still in good agreement with the results

602

achieved in this work. However, these processes were not completely realistic, as neither of the

603

simulations contained purge flows, and the compositions of the broths were simplified as the

604

sugars and intermediate acids were omitted. In addition, in their work Kramer13 used ten ideal

605

steps for the extraction, which in industrial scale would require enormous equipment50. From

606

economic perspective it is unlikely that equipment of this size would be built. In this work, only

607

from 3 to 5 ideal stages were used. In addition Kramer13 estimated that the separation of the ABE

608

to pure components would consume 0.76MJkg-1 butanol. This value is not in agreement with the

609

results in this work: 1.26 and 1.8MJkg-1 butanol with RD and DE, respectively. One reason for

610

the difference might be that in their work Kramer13 used Ideal model to describe the vapor-liquid

611

equilibrium (VLE). This model does not take into account azeotropes, which reduces the energy

612

consumptions of the distillations significantly.

613

4.2 Comparison based on economic analysis

614

According to Tables S13 and S14, the investment cost for the RD process would be

615

approximately 0.9 million euros higher than the investment cost of the DE process. Considering

616

the accuracy of the cost estimates, it can be stated that the processes are equally expensive and

617

neither of the processes have the advantage of lower investment costs. Therefore further

618

economic comparison is needed. In Table 9, the factors affecting the costs of the processes are

619

presented.


33


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

620

Page 34 of 45

Table 9. Cost structure of the processes

Process

Sugar feedstock (Meur/a)a

RD DE Difference (RD-DE)

Steam (Meur/a)

Make-up solvents (Meur/a)

110.0 110.0

7.7 2.9

0.0 0.7

0.0

4.8

-0.7

Annuity of Total Installed Cost of recovery (Meur/a) 2.4 2.3 0.1

Total cost of Production the recovery cost part (Meur/t (Meur/a) sugars)a 10.1 5.9

300.2 289.8

4.2

10.4

621

a

622

As can be seen from Table 9, the sugars are clearly the biggest expense. The cost of the sugars

623

alone was responsible for 83.3 to 94.9% of the total operating costs. Therefore, the energy

624

consumption or the other factors are playing only a small role in the total productions costs. The

625

total cost of the recovery part, which includes energy consumptions, make-up solvents and

626

annuities of the initial investments were between 5.9 and 10.1Meur/a, which is a relatively

627

normal price for plants of this type and size51. As can be seen from Table 9, DE process has

628

somewhat lower production costs than RD process. However, it should be noted that the price of

629

sugars is again dominating these costs, and it can be concluded that the energy consumption of

630

the ABE recovery is not very significant factor, if lignocellulosic sugars are used. Of course, if

631

cheaper ways to produce fermentable sugars from lignocellulosic biomass are found, then the

632

situation could be different. However, Table 9 does not take into account the income that comes

633

from the products. This was taken into account in Figures 4 and 5.

Sugar price 275eur/t


34

Page 35 of 45

40.0

Sales revenues (Meur/a)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


30.0

DE RD

20.0 10.0 0.0 0

50

100

150

200

250

300

350

-10.0 -20.0

634

Sugar price (eur/t)

635

Figure 4. Sales revenues of the RD and DE processes as a function of sugar price

636

As can been seen from Figure 4, the sales revenue was higher in the DE process. This was

637

mainly caused by the lower product recovery yields and higher steam consumption of the RD

638

process. Per 50 tons of dry wood the RD and DE processes produced 9610 and 9790kg of

639

products. The sales income from these products were 81.83 and 83.14Meur/a for RD and DE

640

processes, respectfully. As can be seen from Figure 4, the breakeven prices of sugars for the RD

641

and DE processes were 233 and 268eur/t, respectively. It should be noted that these are slightly

642

lower than the estimated price of sugars in Table 9. A pretreatment unit of this size was

643

estimated to produce sugars at a price of approximately 400eur/t46. This is significantly higher

644

than the breakeven prices of the RD and DE processes. It could be concluded that it is very

645

challenging to produce fermentable sugars from lignocellulosic biomass economically, and that

646

combining SEW with ABE is most likely not a very economical option at the moment.

647

The RD process has different products than the DE process. It was calculated that with ester

648

price of 4376eur/t the RD and DE processes would produce equal sales revenues. As can be seen

649

from Table 2, the average bulk price of 2014 for the ester product was 1200eur/t. It is unlikely


35


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 36 of 45

650

that its price would increase dramatically enough to make the RD equally or more profitable than

651

DE process. Therefore, it can be concluded that the DE process will produce higher sales

652

revenue than RD process with practically identical investment cost.

653

4.3 Comparison based on technological maturity

654

Neither RD nor DE process is completely mature technology. The main uncertainties in the RD

655

process are related to the Reactive Distillation column. Catalyst research is needed to derive

656

proper kinetical models for all reactions, and to verify that the catalyst remains active in the

657

reaction conditions. The novelty of the DE process is not within the unit operations, as they are

658

all well-known technology. Therefore, the biggest uncertainty regarding this process is in the

659

accuracy of the VLE and LLE models. In addition to this, experimental work is needed for both

660

processes to verify that the used solvents do not inhibit the ABE production. It could be

661

concluded that the DE is slightly less complicated process and that its technical maturity is

662

higher. However, technological maturity is still far for both of these processes.

663

5 Conclusions

664

Two processes, RD and DE, were presented and compared for the separation and purification of

665

ABE components from fermentation broth. In the simulations it was found that DE was more

666

energy efficient of these two processes as it consumed 4.97 and 6.12MJkg-1 for all products and

667

butanol, respectively. Similarly, the RD consumed 11.02 and 15.98MJkg-1 for all products and

668

butanol, respectively. These values are higher than the lowest values reported in literature, but

669

this is largely due to the more realistic approach of this work. Economic analysis showed that the

670

usage of lignocellulosic starting material represented from 83.3 to 94.9% of the total operating

671

costs. Significant advances have to be made in the pretreatment of lignocellulosic material, for it


36

Page 37 of 45

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


672

to be noteworthy option for biobutanol production. Furthermore, the economic analysis showed

673

that the total investment costs were almost identical for both processes. Also the total costs of the

674

product recoveries were estimated to be at reasonable level for these processes. The DE process,

675

with its higher sales revenue, proved to be somewhat more economical option of the two

676

processes. However, to truly see whether these processes have potential, their technical maturity

677

would have to be further increased.

678

Supporting Information

679

-Components considered for the first extraction solvent in the DE process

680

-Determination of experimental distribution coefficients

681

-Regressed NRTL interaction parameters

682

-Process flow data

683

-Equipment sizing and prizing for the economic analysis

684

This material is available free of charge via the Internet at http://pubs.acs.org.

685

AUTHOR INFORMATION

686

Corresponding Author

687

* Phone: +358 504589535. E-mail: [email protected]

688

Author Contributions

689

The manuscript was written through contributions of all authors. All authors have given approval

690

to the final version of the manuscript.

691

Funding Sources


37


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 38 of 45

692

The Financial support from Aalto University is gratefully acknowledged.

693

ACKNOWLEDGMENT

694

We greatly appreciate it that Aalto University patented the Reactive Distillation concept

695

presented in this article.

696

ABBREVIATIONS

697

ABE, acetone-butanol-ethanol; CO, carbon monoxide; H2, hydrogen; RD, Reactive Distillation;

698

DE, Dual Extraction; NRTL, Non-Random Two-Liquid; UNIQUAC, Universal QuasiChemical

699

activity coefficient; LLE, liquid-liquid equilibrium; GC-FID, Gas chromatograph with flame

700

ionizing detector; GC-MS, gas chromatograph with mass spectrometer detector; TIC, Total

701

Installed Cost; CECPI, Chemical Engineering Plant Cost Index; VLE, vapor-liquid equilibrium.

702

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703

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704

Water Spent Liquor from Spruce Chips/Softwood Biomass for ABE Fermentation. Ind. Eng.

705

Chem. Res. 2013, 52(11), 4351-4359.

706

(2) Okoli, C.; Adams, T.A. Design and Economic Analysis of a Thermochemical Lignocellulosic

707

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708

(3) Ni, Y.; Sun, Z. Recent progress on industrial fermentative production of acetone-butanol-

709

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710

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711

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712

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714

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715

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716

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717

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718

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719

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720

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721

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722

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723

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724

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725

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726

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727

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728

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729

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730

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731

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732


733

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734

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735

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738

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739

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740

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741

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742

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743

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744

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745

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747

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748

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749

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751

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752

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753

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754

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757

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For Table of Contents Only

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45

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