U.P.B. Sci. Bull., Series B, Vol. 82, Iss. 3, 2020 ISSN 1454-2331

SEPARATION OF ACETONE-BUTANOL-ETHANOL

MIXTURE BY A HYBRID PROCESS

Iulian PATRAȘCU1, Marilena NICOLAE

2, Costin Sorin BÎLDEA

3

Biobutanol is considered an alternative biofuel which can be obtained

through acetone-butanol-ethanol (ABE) fermentation process with low

concentration (<3%wt.). The aim of this work is to design an energy-efficient hybrid

separation system by combining distillation with liquid-liquid extraction. In this

way, the most plentiful component (water) can be removed with minimum costs. Two hybrid separation sequences are designed (one uses mesitylene and the other 2-

ethyl-1-hexanol as separation agent), optimized and economically evaluated. The

total annual cost and the energy requirements are reduced with 25% and 34%

respectively by using dividing wall column and heat integration.

Keywords: Downstream processing, liquid-liquid extraction, distillation,

dividing-wall column, heat integration

1. Introduction

The increase of global energy demand and the importance of

environmental safety are two important aspects that lead nowadays to a possible

depletion and price rise of fossil fuels. Therefore, biofuels are environmentally

friendly and significantly reduces the gas emissions. Biobutanol is considered a

bio-derived fuel with high energy content (32 MJ/kg butanol) that can be

produced in the acetone-butanol-ethanol (ABE) fermentation process. Lately,

biobutanol gained interest over ethanol thanks to its characteristics as low water

miscibility, low flammability and corrosivity, and being able to replace gasoline

in car engines [1]. Biobutanol can be produced on an industrial scale from

lignocellulosic feedstocks like corn stover, wheat straw, corn fiber, barley straw,

switchgrass or wood residue [2]. The best butanol concentration was obtained

using microorganisms such as Clostridium Acetobutylicum and Clostridium

Beijerinckii. However, through the fermentation process, butanol cannot exceed

3% wt. in broth, because butanol inhibits the activity of microorganisms [3]. The

fermentation process requires genetic engineering to increase the concentration of

1 PhD. Student, Dept. of Chemical Engineering, University POLITEHNICA of Bucharest,

Romania, e-mail: patrascu.iulian90@yahoo.ro 2 S.l., Petroleum Processing and Environmental Engineering Department, Petroleum-Gas

University of Ploiesti, Romania, e-mail: nicolae_marilena@yahoo.com 3 Prof., Dept. of Chemical Engineering, University POLITEHNICA of Bucharest , Romania,

e-mail: sorin.bildea@upb.ro

34 Iulian Patrașcu, Marilena Nicolae, Costin Sorin Bîldea

butanol, which will reduce the energy demand in downstream processing and will

provide an economical process of ABE separation [4].

The separation of diluted ABE mixture can be achieved through different

separation technics: distillation, liquid-liquid extraction, adsorption, gas

extraction, reverse osmosis, perstraction, flash vacuum pervaporation and hybrid

separation [1]. Distillation is a separation technique largely used in industry, with

high potential of process intensification and heat integration. The separation of

ABE mixture by distillation is an energy intensive method (14.7 to 79.05 MJ/kg

butanol) [1]. However, a new separation sequence was published, where the

energy requirement was reduced to only 2.7 MJ/kg butanol by using vapor

recompression (VRC) and heat integration in an azeotropic dividing-wall column

(A-DWC) [5,6].

This work presents two hybrid separation sequences which combine

liquid-liquid extraction with conventional distillation columns. The liquid-liquid

extraction is performed with two different solvents (mesitylene and 2-ethyl-1-

hexanol). Both separation sequences are designed and optimized for a minimum

total annual cost (TAC). The most efficient hybrid sequence is further subject to

process intensification through dividing-wall column technology and heat

integration. Therefore, a new hybrid separation process is designed which features

34% energy and 25% TAC savings compared to the conventional hybrid

separation sequence.

2. Problem statement

The ABE mixture obtained from the fermentation process contains a large

amount of water. Separation of water by distillation requires a large amount of

energy. The liquid-liquid extraction technique can solve this problem by

eliminating the most plentiful component (water) without energy costs. This

technique requires a good solvent with low viscosity, different density than water,

high selectivity for butanol and which does not form azeotropes with the

components from the mixture [4, 7]. There are several solvents used for ABE

recovery e.g. oleyl alcohol, n-hexyl acetate, mesitylene and 2-ethyl-1-hexanol [4,

7, 8, 9]. The most energy-efficient hybrid processes are obtained using mesitylene

(4.8 MJ/kg butanol) and 2-ethyl-1-hexanol (9.37 MJ/kg butanol) [4, 9]. However,

these two studies neglect the impurities (acetic acid and butyric acid) present in

the ABE mixture. According to azeotropic data predicted by Aspen Plus,

mesitylene forms a high boiling point azeotrope with butyric acid (Table 1).

Moreover, this azeotrope can accumulate in the solvent recycle and must be

removed, leading to a high economic penalty. This work considers a feed stream

mixture of 4.5%wt. acetone, 18.6%wt. butanol, 0.9 %wt. ethanol, 75.9 %wt. water

and ppm butyric acid and acetic acid, which can be recovered from fermentation

Separation of acetone-butanol-ethanol mixture by a hybrid process 35

by gas stripping [8]. The constraints of downstream processing are a production

rate of 40 kt/years butanol at 99.4 %wt. purity and water removal at 99.8 %wt..

Both mesitylene and 2-ethyl-1-hexanol are considered as potential solvents.

Two hybrid processes with conventional distillation columns are compared

hereafter, in terms of energy requirement and total annual cost. The best hybrid

separation sequence is further studied for total annual cost reduction by using

dividing-wall column technology and energy minimization by heat integration.

Table 1

Azeotropes

Temp (°C) Type Acetone Butanol Ethanol Acetic

Acid

Butyric

Acid Water Mesitylene

2-Ethyl-

1-Hexanol

95.91 Heterogeneous - 0.42 - - - 0.58 - -

93.97 Heterogeneous - 0.25 - - - 0.42 0.33 - 78.15 Homogeneous - - 0.96 - - 0.04 - -

99.18 Homogeneous - - - 0.36 - 0.64 - -

96.05 Heterogeneous - - - 0.20 - 0.42 0.38 -

99.82 Homogeneous - - - - 0.15 0.85 - -

96.61 Heterogeneous - - - - 0.01 0.53 0.46 -

154.62 Homogeneous - - - - 0.37 - 0.63 -

96.61 Heterogeneous - - - - - 0.53 0.47 -

99.34 Heterogeneous - - - - - 0.85 - 0.15

3. Modeling approach

The design and optimization of each separation sequence is performed

with Aspen Plus simulation software using NRTL property method to model the

non-ideality of the liquid phase. The binary parameters between butyric acids and

solvents are estimated by UNIFAC method. Fig. 1 shows that mesitylene forms an

azeotrope with butyric acid (left), meanwhile 2-ethyl-1-hexanol does not (right).

154

156

158

160

162

164

166

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

T /

[°C

]

Butyric acid mass fraction

160

165

170

175

180

185

190

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

T /

[°C

]

Butyric acid mass fraction

Fig. 1. T-xy diagram for solvent - butyric acid (mesitylene – left, 2-ethyl-1-hexanol – right)

The conceptual design is based on the following approach:

- remove the most plentiful component (water) by liquid-liquid extraction;

36 Iulian Patrașcu, Marilena Nicolae, Costin Sorin Bîldea

- recover the solvent with high purity;

- perform last the most difficult separation (butanol purification);

- purify the water product by removing the light components (acetone,

ethanol)

The liquid-liquid extraction column is designed for counter current flow.

The minimum solvent flow rate is found by using equations (1) – (2), where XE,

XS, XF and XR represent the molar fraction of butanol in the extract, solvent, feed

and raffinate, respectively; E, S, F and R are the molar flow rates of the same

streams.

RF

SE

XX

XX

S

R

, (1)

RXSXFXEX RSFE , (2)

log

1

/

/1log

KXX

KXX

NSR

SF

, (3)

The molar fractions of the butanol in the raffinate and extract are obtained

from the liquid – liquid equilibrium diagrams (Fig. 2) calculated by the Aspen

Plus simulator, using the NRTL property model. The mass balance of the

extraction column can be calculated using the equation (2). The theoretical

number of stages is given by the equation (3), where K is the slope of the

equilibrium line and ε is the extraction factor. Due to different selectivities of the

solvents for butanol, a minimum of 14525.5 kg/h mesitylene or 1273.7 kg/h 2-

ethyl-1-hexanol are required for liquid-liquid extraction.

BUTANOL

0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

0.9

0.8

0.7

0.6

0.5

0.4

0.3

0.2

0.1

BUTANOL

0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

0.9

0.8

0.7

0.6

0.5

0.4

0.3

0.2

0.1

Fig. 2. Liquid – liquid equilibrium line of the butanol – water – solvent

Separation of acetone-butanol-ethanol mixture by a hybrid process 37

4. Process design

4.1. Design of the process using mesitylene as extraction solvent

The diluted ABE mixture and the solvent are fed on the top tray and on the bottom

tray of the liquid-liquid extraction column, respectively. The raffinate containing

water, acetone and ethanol is further sent for purification in the distillation column

COL-3. The extract contains mainly solvent, butanol, acetone and ethanol.

However, the liquid-liquid extraction selectivity in not 100%, and a small amount

of water remains in the extract. The first distillation column (COL-1) serves to

solvent recovery. Due to the presence of the butyric acid - mesitylene azeotrope, a

side stream is necessary to avoid the accumulation of acids in the solvent recycle

stream. The second distillation column (COL-2) separates the butanol in bottom

and acetone - ethanol as distillate. A side stream is also required for COL-2 to

recycle a small amount of butanol-water azeotrope. In the third column (COL-3),

the distillate of the COL-2 and the raffinate are fed together for water purification.

The side stream of COL-1 is fed in the fourth column (COL-4) to avoid solvent

loss. The distillation columns are optimized for a minimum of total annual cost, as

it will be described in a later section. Fig. 3 presents the mass balance and key

design parameters and Fig. 4 shows the composition profiles of the distillation

columns.

696 kg/hAcids: 3.3 %wt

S: 96.6 %wt

COL-1

COL-2

COL-3

EXTR

AC

TIO

N

26945 kg/hA: 4.5 %wtB: 18.6 %wtE: 0.9 %wtW: 75.9 %wtAcids: ppm

Feed

Make-up58.01 kg/h

Mixer

S: 99.99 %wt51321 kg/h

4997 kg/hB: 99.4 %wtButanol

20334 kg/hw: 99.99 %wtWater

357 kg/hA: 3.7 %wtE: 68.3 %wtW: 25.6 %wtEthanol

1209 kg/hw: 99.4%wtAcetone

Mixer

Cooler

Diam = 3.47 mRR = 3.91QR = 11342 kW

Diam = 1.56 mRR = 17.19QR = 5431 kW

Diam = 1.14 mRR = 7.45QR = 2137 kW

E: 0.8 %wtW: 99.1 %wt20544 kg/h

A: 2.1 %wtB: 9.1 %wtE: 0.1 %wtW:0.4 %wtS:88.3 %wtAcids: ppm58861 kg/h1

15

1

2

37

460.14 kg/hB: 71.3 %wtW: 28.6 %wt

2

1

8

A: 17.8 %wtB: 77.8 %wtE: 1.1 %wtW: 3.2 %wt6843 kg/h

A: 89.6 %wtE: 5.7 %wtW: 4.3 %wt1356 kg/h

31

2

1

27

54

41

36

QC = -6240 kW

QC = -2004 kW

QC = -3333 kW

-4545 kW

40

16

Mixer

COL-4

620.59 kg/hS: 99.99 %wtSolvent

Waste 75.66 kg/hAcids: 30 %wt

S: 70 %wt

Diam = 0.17 mRR = 3.63QR = 34.89 kW

7

2

1

35

QC = -34.62 kW

Mixer

Fig. 3. Separation sequence with mesitylene as solvent

38 Iulian Patrașcu, Marilena Nicolae, Costin Sorin Bîldea

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 5 10 15 20 25 30 35 40 45

Mas

s fr

acti

on

Stage

COL - 2

Water

ButanolAcetone

Ethanol

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 5 10 15 20 25 30 35 40

Mas

s fr

acti

on

Stage

COL - 3WaterAcetone

Ethanol

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 5 10 15 20 25 30 35 40

Mas

s fr

acti

on

Stage

COL - 4

Mesitylene

Butanol

Butyric Acid

Fig. 4. Composition profiles of the distillation columns (mesitylene)

4.2. Design of the process using 2-ethyl-1-hexanol as extraction solvent

The separation sequence showed in Fig. 5 carries out the liquid-liquid

extraction using 2-ethyl-1-hexanol as solvent. The ABE mixture is fed in the

liquid-liquid extraction column on the top tray and 2-ethyl-1-hexanol is fed on the

bottom tray. The raffinate (water with some ethanol and acetone) is sent to the

column COL-3 for water purification. Because the liquid-liquid extraction process

does not achieve 100% selectivity, a small amount of water is found in the extract.

The extract is fed to the distillation column (COL-1). The solvent is

recovered with high purity as bottom product and recycled. The distillate is fed to

the column COL-2, which delivers high-purity butanol as bottom product, acetone

and ethanol with small amounts of water as distillate, and a side stream containing

butanol and water. The distillate of COL-2 is sent to COL-3 for water purification.

The side stream is sent to a decanter, from which the aqueous phase is recycled to

the extraction column and the organic phase is fed to a lower tray of COL-2.

The distillation columns are optimized for a minimum of total annual cost,

as it will be described in a later section. The mass balance and key design

parameters are presented in Fig. 5. The composition profiles of distillation

columns are shown in Fig. 6.

Note that this process (Fig. 5) is more promising: it does not involve the

formation of an azeotrope containing the solvent, thus fewer columns are required

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 5 10 15 20 25 30 35 40 45 50 55

Mas

s fr

acti

on

Stage

COL - 1

Water

Butanol

Mesitylene

Acetone

Butyric Acid

Separation of acetone-butanol-ethanol mixture by a hybrid process 39

compared to the previous process (Fig. 3). For this reason, in the next section the

energy minimization and total annual cost reduction by process intensification and

heat integration will be studied.

COL-1

COL-2

COL-3

EXT

RA

CT

ION

DEC26945 kg/hA: 4.5 %wtB: 18.6 %wtE: 0.9 %wtW: 75.9 % wt

Acids: ppmFeed

Make-up19.43 kg/h

Mixer

S: 99.99 %wt1387 kg/h

4999 kg/hB: 99.4 %wtButanol

20264 kg/hw: 99.99 %wtWater

450.32 kg/hA: 3.0 %wtE: 54.2 %wtW: 38.8 %wtEthanol

1209 kg/hw: 99.4 %wtAcetone

Mixer

Cooler

Diam = 1.65 mRR = 0.15QR = 3100 kW

Diam = 1.2 mRR= 10.53QR = 4158 kW

Diam = 1.68 mRR= 35.28QR = 4489 kW

A: 3.2 %wtE: 0.7 %wtW: 95.9 %wt21293 kg/h

A: 6.1 %wtB: 58.8 %wtE: 1.1 %wtW: 17.7 %wtS: 16.0 %wtAcids: 0.2 %wt

8686 kg/h1

15

1

2

15

1627.9 kg/hB: 6.8 %wt

W: 93.1 % wt

2

1

9

7

20

7299 kg/hA: 7.2 %wtB: 70.0 %wtE: 1.3 %wtW: 21.1 % wtAcids: 0.3 %wt

B: 78.3 %wtW: 21.4 % wt4617 kg/h

A: 82.4 %wtE: 14.3 %wtW: 3.3 %wt639.72 kg/h

21

2

1

2931

38

62

38

QC = -2018 kW

QC = -4051 kW

QC = -2040 kW

209 kW

Cooler620 kW

Fig. 5. Separation sequence with 2-ethyl-1-hexanol as solvent

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 5 10 15 20 25 30 35 40

Mas

s fr

acti

on

Stage

COL - 1

Water

Butanol2 Ethyl 1 Hexanol

Acetone Butyric Acid

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 5 10 15 20 25 30 35 40 45 50 55 60 65

Mas

s fr

acti

on

Stage

COL - 2

Water

Butanol

Acetone

Ethanol

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 5 10 15 20 25 30 35 40

Mas

s fr

acti

on

Stage

COL - 3WaterAcetone

Ethanol

Fig. 6. Composition profiles of the distillation columns from Fig. 5 (2-ethyl-1-hexanol)

40 Iulian Patrașcu, Marilena Nicolae, Costin Sorin Bîldea

4.3. Process design alternative (2-ethyl-1-hexanol as extraction solvent)

In this section we propose a novel design alternative of separation

sequence described in Fig. 5. The novel sequence uses the same extraction

solvent, 2-ethyl-1-hexanol. The process flowsheet is presented in Fig. 7. This

design uses a single dividing-wall (DWC) column to integrate the solvent

recovery column (COL-1) and butanol purification column (COL-2). As in the

previous flowsheet (Fig. 5), a butanol-water mixture is withdrawn as side stream

and submitted to liquid-liquid separation (DEC), from which the organic phase is

returned to DWC and the aqueous phase is recycled to extraction. Another side-

stream recovers the high-butanol butanol (99.4 %wt.). The dividing-wall column

is simulated in Aspen Plus as a prefractionator (PF) and a main column (DWC).

The prefractionator has 30 theoretical stages and the main column has 45

theoretical stages. Note that the purpose of the column COL-3 (Fig. 5) is mainly

to purify the water. For a fair assessment of energy requirements (MJ/kg butanol),

the purification of acetone and ethanol should not be included in the analysis.

Therefore, the column COL-3 from the conventional hybrid separation sequence

(Fig. 5) is replaced by the two-product column (DC) which delivers high-purity

(99.8 %wt.) water as bottoms, and an acetone-ethanol mixture impurified with

small amounts of water as distillate.

EXTR

AC

TIO

N

1

15

8

10

2

15

30

7

1

30

40

45

2

38

14DWCPF

DC

1

1

26945 kg/hA: 4.5 %wtB: 18.6 %wtE: 0.9 %wtW: 75.9 %wtAcids: ppm

Feed

Make-up24.36 kg/h

DEC

QC = -3359 kW

Diam = 1.94 mRR= 30QR = 4376 kW

Diam = 0.8 mRR= 5QR = 1847 kW

QC = -1504 kW

QC = -125 kW

QC = -196 kWCooler

Cooler

Mixer

Mixer

Mixer

B: 73.8 %wtW: 22.4 % wt4669 kg/h

1675 kg/hB: 7.2 %wt

W: 90.4 % wt

8621.7 kg/hA: 6.2 %wtB: 59.4 %wtE: 1.13 %wtW: 18.03 %wtS: 14.9 %wtAcids: 0.18 %wt

5029.5 kg/hB: 99.4 %wtButanol

20449.8 kg/hw: 99.8 %wtWater

AEW1490 kg/hA: 81.55 %wtE: 16.42 %wtW: 2.02 %wt

A: 83.17 %wtE: 13.04 %wtW: 3.7 %wt620 kg/h

A: 3.2 %wtE: 0.7 %wtW: 95.9 %wt21293 kg/h

S: 98.5 %wtAcids: 1.5%wt1297.25 kg/h

Hex 1

Hex 2

A = 17.8 m2

A = 163.56 m2

25 ℃

25 ℃

93 ℃

95

℃5

8 ℃

60 ℃

25 ℃

25 ℃

197 ℃

132 ℃

147 ℃

106 ℃

107 ℃

96 ℃

52 ℃40 ℃

40 ℃

125 ℃

108 ℃

58 ℃

Dia

m=

1.1

4 m

B: 56.27 %wtW: 40.69 % wt6344 kg/h

Fig. 7. Heat integrated DWC flowsheet with 2-ethyl-1-hexanol as solvent

Separation of acetone-butanol-ethanol mixture by a hybrid process 41

40

60

80

100

120

140

160

180

200

0 5 10 15 20 25 30 35 40 45 50

Tem

pe

ratu

re /

[°C

]

Stage

PF

DWC

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 5 10 15 20 25 30 35 40 45

Mas

s fr

acti

on

Stage

Water

Butanol

Acetone

Ethanol

PF

DWC

Butanol

WaterBut. Ac.

2 Ethyl 1 Hexanol

2 Ethyl 1 Hexanol

Fig. 8. Temperature and composition profile heat integrated DWC (2-ethyl-1-hexanol)

In this design, heat integration is applied for minimization of the energy

requirement. Therefore, the extract is preheated to 60º before being fed to the

prefractionator (PF), using the DWC side stream. The raffinate is also preheated

from 25º to 95º using DC bottom stream, before being mixed with the DWC

distillate and fed to the DC column. The heat integration scheme reduces the

heating needs by 34%.

Fig. 7 shows the mass balance of the main streams together with the key

design parameters. Fig. 8 shows the temperature (left) and composition (right)

profiles of the DWC. The temperature difference between the two sides of the

dividing-wall column does not exceed 20º, as required by this technology. The

total annual cost is reduced by 25% compared to the conventional separation

system.

5. Process optimization

In both conventional separation sequences, the distillation columns are

optimized for a minimum of total annual cost (TAC).

periodpayback

CAPEXOPEXTAC

. , (4)

The equipment (CAPEX) and operating (OPEX) costs were evaluated for

a payback period of 3 years and 8000 h/year operating time [10,11]. The heating

utilities used are: for the solvent recovery columns (COL-1) - HP steam (42 bar,

254 °C, 9.88 $/GJ), for the butanol purification column (COL-2) and the water

purification column (COL-3) - LP steam (6 bar, 160 °C, 7.78 $/GJ). The cost of

cooling is $0.72/GJ. The total investment cost (CAPEX) includes the extraction

column, the cooler for solvent, the distillation columns, and the decanter. The cost

of the equipment was estimated using standard cost correlations (Marshall &

Swift equipment cost index M&S = 1536.5 in 2012):

42 Iulian Patrașcu, Marilena Nicolae, Costin Sorin Bîldea

0.65( $) & / 280 474.7 2.29HEX m d pC US M S A F F F (5)

1.066 0.82( $) & / 280 957.9 2.18shell cC US M S D H F (6)

1.55( $) & / 280 97.2trays T t mC US N M S D F F (7)

where A is the area (m2), Fm = 1 (carbon steel), Ft = 0 (sieve trays), Fd =

0.8 (fixed-tube), Fp = 0 (less than 20 bar), Fd = 1.35 (for reboilers), and for the

shell Fc = Fm∙Fp, 2)48.3(00023.0)48.3(0074.01 PPFp .

The optimization procedure for the distillation columns followed the next

steps: specify the number of stages; set design specifications in order to satisfy the

constraints of product purities; perform a sensitivity analysis (included in the

Aspen Plus software) to find the feed stage with the lowest energy requirement;

calculate the total annual cost. Fig. 9 and Fig. 10 present the total annual cost

(TAC) versus the number of stages, from which the optimum number of stages is

found, for each distillation column.

4025000

4075000

4125000

4175000

50 53 56 59

TAC

/ [

US$

/ye

ar]

No of stages

COL-1

950000

965000

980000

995000

38 40 42 44

TAC

/ [

US$

/ye

ar]

No of stages

COL-2

1740000

1760000

1780000

1800000

33 35 37 39

TAC

/ [

US$

/ye

ar]

No of stage

COL-3

30850

30900

30950

31000

32 34 36 38

TAC

/ [

US$

/yea

r]

No of stages

COL-4

Fig. 9. Optimization of distillation columns from Fig. 3 (mesitylene)

Separation of acetone-butanol-ethanol mixture by a hybrid process 43

1234000

1236000

1238000

1240000

35 37 39 41

TAC

/ [

US$

/ye

ar]

No of stages

COL-1

1867000

1868000

1869000

1870000

59 61 63 65

TAC

/ [

US$

/yea

r]

No of stages

COL-2

1355000

1361000

1367000

1373000

35 37 39 41

TAC

/ [

US$

/ye

ar]

No of stage

COL-3

Fig. 10. Optimization of distillation columns from Fig. 5 (2-ethyl-1-hexanol)

6. Economic evaluation

Table 2 shows the economic evaluation of the separation sequence which

uses mesitylene as solvent. The operating cost (OPEX) of this process is 5230·103

US$/year and the capital cost (CAPEX) is 6905.81 ·103 US$, which gives a total

annual cost of 6890.6·103 US$/year. The energy requirement for butanol

purification is 11.05 MJ/kg butanol. Table 2

Economic evaluation of the hybrid separation system (mesitylene)

Item description (unit) Extractor COL - 1 COL – 2 COL - 3 COL - 4 Total

Shell / [103 US$] 294.7 1246.5 281.4 363.4 30.72 2216.72

Trays / [103 US$] 25.5 196.2 27.1 38.3 1.22 288.32

Condenser / [103 US$] - 640.9 480.8 690.4 12.81 1824.91

Reboiler / [103 US$] - 904.2 515.3 704.5 20.25 2144.25

Exchangers / [103 US$] - - - - - 431.61

Heating / [103 US$/year] - 3185.9 478.9 1216.9 9.92 4891.91

Cooling / [103 US$/year] - 130.8 41.5 69.1 0.72 336.12

TAC / [103 US$/year] 106.7 4050.9 955.4 1746.7 30.9 6890.6

Table 3 presents the economic evaluation of separation sequence which

uses 2-ethyl-1-hexanol as solvent. The operating cost (OPEX) of this process is

44 Iulian Patrașcu, Marilena Nicolae, Costin Sorin Bîldea

3004.6·103 US$/year and the capital cost (CAPEX) is 5538.4·10

3 US$, which

gives a total annual cost of 4659.9·103 US$/year. The energy requirement for

butanol purification is evaluated at 6.76 MJ/kg butanol.

Table 3

Economic evaluation of the hybrid separation system (2-ethyl-1-hexanol)

Item description (unit) Extractor COL - 1 COL – 2 COL - 3 Decanter Total

Shell / [103 US$] 188.4 406.5 595.4 281.5 25.9 1497.7

Trays / [103 US$] 15.0 44.4 74.1 27.2 - 160.7

Condenser / [103 US$] - 281.1 745.1 501.9 326.1 1854.2

Reboiler / [103 US$] - 454.5 918.5 595.6 - 1968.6

Exchangers / [103 US$] - - - - - 57.2

Heating / [103 US$/year] - 882.3 1006.1 931.9 - 2820.3

Cooling / [103 US$/year] - 41.8 84.0 42.3 12.9 184.3

TAC / [103 US$/year] 67.8 1235.9 1867.7 1358.3 130.2 4659.9

Table 4 shows the economic evaluation of heat integrated DWC separation

sequence which uses 2-ethyl-1-hexanol as solvent. The operating cost (OPEX) of

this process is 1788.9·103 US$/year and the capital cost (CAPEX) is 3438.35·10

3

US$, which gives a total annual cost (TAC) of 3029.8·103 US$/year. The energy

requirement for butanol purification is 4.46 MJ/kg butanol.

Table 4

Economic evaluation of the heat integrated DWC flowsheet (2-ethyl-1-hexanol)

Item description (unit) Extractor DWC DC COOL Decanter Total

Shell / [103 US$] 188.4 659 174.2 - 38.2 1059.8

Trays / [103 US$] 15.0 84.6 14.1 - - 113.7

Condenser / [103 US$] - 665.5 390.7 189 - 1245.2

Reboiler / [103 US$] - 576.5 350.29 - - 926.79

Exchangers / [103 US$] - - - - - 92.86

Heating / [103 US$/year] - 1245.2 413.9 - - 1659.1

Cooling / [103 US$/year] - 70.4 31.2 27.5 - 129.1

TAC / [103 US$/year] 67.8 1977.5 754.99 123.9 12.7 3029.8

7. Conclusions

The purification of butanol in a hybrid liquid-liquid extraction - distillation

system can significantly reduce the energy requirement and the total annual cost.

In this paper, two separation agents were used for liquid – liquid extraction

columns (mesitylene and 2-ethyl-1-hexanol). The acetone-butanol-ethanol (ABE)

mixture is normally recovered from the fermentation process with impurities as

Separation of acetone-butanol-ethanol mixture by a hybrid process 45

acetic acid and butyric acid. Considering these impurities in the feed mixture with

ABE, the downstream processing becomes difficult. This happens because of the

high-boiling azeotrope formed by mesitylene and butyric acid. However, two

conventional hybrid separation sequence were designed, optimized and

economically evaluated. The best results were obtained by using 2-ethyl-1-

hexanol as separation agent.

The conventional hybrid separation sequence which uses mesitylene as

separation agent is an energy intensive process (11.05 MJ/kg butanol with a TAC

of 6890.6·103 US$/year), due to the high amount of solvent used and the existence

of the mesitylene-butyric acid azeotrope. The solvent loss with this azeotrope is

around 424·103 kg/year which means 1695·10

3 US$/year loss. However, the

energy requirement when using 2-ethyl-1-hexanol as a solvent is much lower -

6.76 MJ/kg butanol with a TAC of 4659.9·103 US$/year. A new, alternative heat

integrated process has also been studied for the conventional separation sequence

which uses 2-ethyl-1-hexanol. This process design uses a dividing-wall column to

integrate two conventional columns in a single one. Thereby, the total annual cost

is reduced to 3029.8·103 US$/year and the energy requirement is minimized to

4.46 MJ/kg butanol.

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