Eco-efficient Downstream Processing of Biobutanol by Enhanced

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Eco-efficient downstream processing of biobutanol by enhanced process intensification and integration Iulian Patrascu, Costin Sorin Bildea, and Anton Alexandru Kiss ACS Sustainable Chem. Eng., Just Accepted Manuscript • DOI: 10.1021/ acssuschemeng.8b00320 • Publication Date (Web): 28 Feb 2018 Downloaded from http://pubs.acs.org on March 2, 2018

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ACS Sustainable Chemistry & Engineering

Eco-efficient downstream processing of biobutanol by enhanced process intensification and integration Iulian Patraşcu,1 Costin Sorin Bîldea,1 Anton A. Kiss 2,3* 1

University “Politehnica” of Bucharest, Polizu 1-7, 011061 Bucharest, Romania

2

School of Chemical Engineering and Analytical Science, The University of Manchester,

Sackville Street, Manchester, M13 9PL, United Kingdom. E-mail: [email protected] 3

Sustainable Process Technology Group, Faculty of Science and Technology, University of

Twente, PO Box 217, 7500 AE Enschede, The Netherlands KEYWORDS: ABE process, azeotropic distillation, dividing-wall column, heat pumps

ABSTRACT. The biobutanol stream obtained after the fermentation step in the ABE process has a low concentration (less than 3 %wt butanol) that leads to high energy usage for conventional downstream separation. To overcome the high downstream processing costs, this study proposes a novel intensified separation process based on a heat pump (vapor recompression) assisted azeotropic dividing-wall column (A-DWC). Pinch analysis and rigorous process simulations have been used for the process synthesis, design and optimization of this novel sustainable process. Remarkably, the energy requirement for butanol separation using heat integration and vapor recompression assisted A-DWC is reduced by 58% from 6.3 to 2.7 MJ/kg butanol.

ACS Paragon Plus Environment

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INTRODUCTION Biomass is a natural way of storing solar energy which can be converted afterwards into biofuels. In the quest for sustainable production of renewable energy and chemicals, biorefineries hold the promise for effectively converting biomass into biofuels and bioproducts. Among them, biobutanol is a very promising biofuel but one must consider carefully the correct planning of lignocellulosic feedstock, fermentation, and separation in order to be sustainable and economically profitable.1 Biobutanol was found to be an efficient fuel with better physical and chemical properties as compared to other biofuels. The selection of species, substrates, pre-treatment, genetic engineering techniques and various downstream processing techniques have been evaluated in a recent review paper.2 In addition to its role as biofuel, biobutanol can be used to produce a wide range of chemicals. For example, n-butanol and its primary derivatives retaining the oxygen (e.g. butyraldehyde and butyric acid) have mainstream applications in the solvent, polymer, fuel oxygenate, and specialty chemical markets. Similarly, butanol dehydration products (butenes and butadiene) present essential opportunities in the hydrocarbon fuel and synthetic elastomers markets.3 In the acetone-butanol-ethanol (ABE) fermentation process, biobutanol is obtained in diluted form, typically less than 3 %wt concentration (owing to the severe butanol toxicity to microorganisms). The key challenges in biobutanol production emphasize the idea of improving the efficiency of the ABE process by altering the upstream (e.g. pretreatment and fermentation) and the downstream steps (product recovery and purification) by various methods.4 The modification of microorganisms by genetic engineering (to keep them alive and active at high butanol concentrations) could increase the productivity, yield and concentration and thus reducing the


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production costs.5 However, this long term goal is yet to be realized, and even it becomes a reality, the product separation and purification will still remain a critical challenge.6 The other approach is the development of more efficient downstream processes for butanol recovery. This is not a singular issue, as the need for cost effective separations is essential for sustainable biorefineries.7,8 The low concentration of butanol obtained by fermentation leads to high energy requirements for downstream processing, in the range of 14.7-79.05 MJ/kg butanol.9 A higher butanol yield is achieved by using anaerobic bacteria as Clostridium acetobutylicum

10,11

and Clostridium

beijerinckii.12 But the ABE concentration can be further increased by in-situ product recovery (ISPR) methods such as gas stripping technology to: 4.5 %wt acetone, 18.6 %wt butanol and 0.9 %wt ethanol.13 Many techniques are available for ABE separation, e.g. distillation, reverse osmosis, adsorption, liquid-liquid extraction, pervaporation and others.14-17 Process alternatives based on advanced distillation technologies have been also reported – see Figure 1.18 More insights into the appropriate selection and design of fluid separation processes (applicable also to biofuels) have been reported in a recent review paper.19 This work proposes a new biobutanol downstream process based on combining azeotropic distillation in a dividing wall column (DWC) with vapor recompression technology.20,21 Remarkably, the azeotropic DWC integrates three distillation columns into one unit with enhanced thermodynamic efficiency, and further reduces the primary energy used for separation by employing heat pumping and energy integration.


3


Blower

Azeotrope

Cooler Feed Organic COL1

Mixer

Acetone

COL4

Decanter Butanol

COL3

Ethanol COL2

Aqueous

Water

Azeotrope Blower

Azeotrope

Cooler

Feed

Organic COL1

Mixer

Hex2 Decanter Acetone Butanol

Aqueous

Hex3 COL4 DWC

Hex1 Azeotrope

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

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Water

Ethanol

Figure 1. Process flowsheet of ABE downstream separation sequences: conventional distillation (top) and intensified process using DWC and heat integration (bottom)


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PROBLEM STATEMENT Taking into account the energy density of butanol (36 MJ/kg) the use of classic distillationdecanter method for butanol recovery is clearly too demanding with energy requirements of 14.5-79.5 MJ/kg butanol.9 But this value could be drastically reduced when the advanced distillation technologies are combined with in-situ product recovery (ISPR) techniques.13 An optimized conventional separation sequence using three distillation columns along with one decanter (Figure 1, top) requires 6.3 MJ/kg butanol (excluding COL-4 which separates the acetone and ethanol byproducts).18 When applying heat integration to the conventional sequence, up to 18% energy savings are possible. But using heat integration and combining two columns into a dividing-wall column (Figure 1, btm), the energy requirement was reduced to 4.46 MJ/kg butanol.18 This is already a major reduction (29%) of the energy requirements, but can the energy savings be pushed any further? To achieve more savings, this work proposes a novel enhanced process that makes use of process intensification (azeotropic dividing-wall column) and process integration techniques (e.g. energy integration and vapor recompression heat pump). By using a highly integrated azeotropic dividing-wall column (A-DWC) assisted by vapor recompression (VRC) technology the energy savings can be increased significantly as described hereafter. APPROACH AND METHODOLOGY The plant capacity considered here is 40 ktpy butanol. To account for a realistic composition of the ABE mixture such as reported in literature,13 impurities are also taken into account hence the mixture which must be efficiently separated contains: 4.5 wt.% acetone, 18.6 wt.% butanol and 0.9 wt.% ethanol, 0.1 wt.% CO2, 0.08 wt.% butyric acid and 0.04 wt.% acetic acid. The required product purities are butanol 99.4 wt.% and water 99.8 wt.%. Note that in Figure 1 the last


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column (COL-4) separates only the light components (acetone and ethanol) and does not contribute to the energy requirement for butanol separation. Therefore, the process designed in this work will deliver an acetone – ethanol mixture (96 %wt.) without attempting to split it into high-purity components. The feed stream and all the products are at 25 °C. Regarding the design and operational constraints, when dividing wall technology is employed each side of the dividing wall must have the same number of trays and the temperature difference between the two sides should not exceed 20 °C. Moreover, the use of vapor recompression is limited by the maximum temperature of the compressed vapor, which is 150 °C. Phase equilibrium The process is simulated in Aspen Plus, using non-random two-liquid (NRTL) as appropriate property model. Table 1 lists the boiling points of the components, while Figure 2 illustrates the T-xy diagram of the binary mixture butanol-water. Notably, both hetero- and homogeneous azeotropes are formed in this aqueous system, and this further complicates the separation. Table 1. Boiling points of the pure chemicals and azeotropes involved in the butanol recovery

Component

Boiling point ºC (at 1.013 bar)

Acetone

56.14

Homogeneous azeotrope: Ethanol (95.63% wt) / Water

78.15

Ethanol

78.31

Heterogeneous azeotrope: n-Butanol (42% wt) / Water

95.91

Water

100

n-Butanol

117.75

Acetic acid

118.01

n-Butyric acid

163.28


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120 115

Heterogeneous azeotrope 42 %wt. butanol

110

T / [°C]

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


105 100 95

92 %wt. Water

90 0

0.1

0.2

0.3

67 %wt. Butanol 0.4

0.5

0.6

0.7

0.8

0.9

1

Butanol mass fraction Figure 2. T-xy diagram for the binary mixture butanol-water. Conceptual design Figure 3 presents the conceptual design of the process. First, a distillation sequence is suggested (Figure 3, top) based on several heuristics, as follows:22 •

Remove first the most plentiful component: the pre-fractionator column COL-1 removes a large amount of water as bottom product. This reduces both the investment and operation costs of the sub-sequent units. The distillate contains acetone, ethanol and water-butanol mixture with close to azeotropic composition. COL-1 can work as a stripper (feed on one of the top trays), as its main function is to remove the light components such that high purity water is obtained in the bottom stream, without having a tight specification on the distillate.

•

Lights out first: from the distillate of COL-1, the lightest components (acetone and the


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ethanol-water azeotrope) are removed as distillate of column COL-2. A rectifying section is necessary to ensure high purity of the distillate, while the specifications of the bottom stream are not stringent (see below). •

Perform the most difficult separation last: removal of acetone, ethanol and a part of water in COL-2 is compatible with this heuristic, as the bottom product of COL-2 is a water-butanol azeotropic mixture, the most difficult to be separated into high purity products.

•

Use liquid-liquid split to cross the distillation boundary induced by a heterogeneous azeotrope: The butanol-water azeotrope is cooled and separated in an organic and an aqueous phase. The aqueous phase can be split into water (heavy product) and butanol-water azeotrope (light product). As column COL-1 already performs this function, the aqueous phase is sent there. This also ensures that any acetone and ethanol which are left in the COL2 bottoms are recycled to a location which still allows their separation to a product stream (preventing therefore accumulation). The composition of the organic phase allows separation into butanol (heavy product) and butanol-water azeotrope (light product). This is achieved in column COL-3, which can also work as a stripper (feed on the top stage, no tight specification on the top product). The butanol-water azeotrope is sent to the decanter.

Next, one can consider energy coupling by combining the condensers of COL-1 and COL3 with the reboiler of COL-2. This leads to the flowsheet shown in Figure 3 (left). Note that the reboilers of COL-1 and COL-3 provide the vapors required for column COL-2. Finally, it can be observed that the boiling points of the COL-1 and COL-3 distillate streams are almost the same, corresponding basically to the boiling point of the butanol-water azeotrope. Moreover, because the bottom streams also have similar boiling points (water: 100 °C, butanol: 117.7 °C), the


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temperature profiles along COL-1 and COL-3 are expected to be similar. Therefore, columns COL-1 and COL-3 can be integrated into one section provided with a dividing wall. Thus, all three distillation columns can be combined into a single unit, as shown in Figure 3 (right).

AEW

COL2

AEW + Butanol water

Butanol water

Cooler

Butanol water

Feed

Decanter Organic

COL3 COL1

Butanol

Water

Aqueous

Condenser

AEW

AEW COL2

Butanol Water

AEW + Butanol Water

Cooler Azeotrope Feed

Butanol Cooler Water Organic

COL1

Decanter

DWC-R

Feed

Aqueous

DWC-L

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


Organic Decanter

COL3

Aqueous

Water

Butanol

Water

A-DWC

Butanol

Aqueous

Figure 3. Conceptual design of butanol separation: (top) sequence based on conventional distillation; (left) thermally coupled column (right) DWC equivalent configuration


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Process optimization The process shown in Figure 3 (right) was optimized using the total annual cost (TAC) as the objective function to be minimized (according to standard industrial practices):

TAC = OPEX +

CAPEX Payback period

(1)

More details about the equipment and utilities cost are provided in the Appendix. For the TAC minimization, the following decision variables and restrictions are considered:18 •

Number of stages in the distillation column (both sides). The same number of stages on each side of the dividing wall was considered in the optimization, as this is the normal constructive solution for large diameter trayed columns. However, the number of stages could be different on the two sides, particularly when (structured) packing is used.

•

The design specifications for purity of product stream (e.g. butanol product 99.4 %wt., water by-product 99.8 wt.%).

•

Max. 0.1 kg/h butanol in water-product (water quality) by changing the vapor flow rate

•

Butanol purity min. 99.4 %wt. (butanol quality) by manipulating distillate flow rate.

•

Max. 30 kg/h water in distillate (AEW quality), obtained by manipulating the reflux ratio.

•

CO2 recovery in distillate is obtained manipulating the side flowrate.

•

The feed mixture, organic phase and aqueous phase are preheated at 97 °C. This reduces the energy requirement in reboilers, then by heat integration the hot streams will provide the necessary heating for cold feed streams.


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Energy integration Pinch analysis provides understanding of the energy targets and subsequent design of the optimal heat exchange network (HEN). The procedure was applied according to the literature.23 A minimum temperature difference of 10 K was used. No additional correction factor was used, as only 1-pass shell & tube heat exchangers were considered – but other correction factors may be needed for other types of heat exchangers. The coefficient of performance (COP) is used for evaluating the feasibility of using a heat pump, while also accounting the additional costs and the payback time. COP = Q / W = 1 / η = Tc /(Tr − Tc ) > 10

(2)

where Q = reboiler duty, W = work provided, η = Carnot efficiency, Tr = reboiler temperature, and Tc = condenser temperature.24 If the Q/W ratio is lower than 5, using a heat pump (HP) brings no advantages, but if it exceeds 10 then a HP should be considered. The maximum energy savings are given by:19

Max.savings (%) = 100 ⋅

Qreb − HR ⋅ Qreb / COP Qreb

(3)

where HR is the heat ratio (thermal to electrical), with a typical ratio: 1 MWe = 2.5 MWth. Then, the investment and operation costs of the complete process are evaluated based on the correlations provided in the Appendix, to show the advantages of this design.22 Environmental impact The potential environmental impact was evaluated in Aspen Plus using Carbon Tracking to calculate the CO2 emissions. The fuel source considered is natural gas and the CO2 emission


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factor data source used is the US Environmental Protection Agency Rule of ‘E9-5711’ (CO2 EUS) proposed in 2009. The standard used for the Global Warming Potential is USEPA (2009) with a carbon tax of 5 $/ton (this can be updated to any particular year). The following equations were used for calculating the CO2 emissions:

 Q   C%  [CO2 ]emissions =  fuel   α  NHV   100 

 Q proc Q fuel =  λ  proc

(4)

 (TFTB − T0 )  ( hproc − 419 ) (TFTB − Tstack ) 

(5)

where α = 3.67 is the ratio of molar masses of CO2 and C; NHV (net heating value) is 48900 kJ/kg for natural gas; C% (carbon content) is 0.41 kg/kg; Qproc is the heat duty required by the process and duty provided by the steam (kW); λproc is the latent heat of steam delivered to the process (kJ/kg); hproc is the enthalpy of steam delivered to the process (kJ/kg); T0 is the ambient temperature; TFTB (K) and Tstack (K) are the flame and stack temperature, respectively. RESULTS AND DISCUSSION Azeotropic dividing-wall column Figure 4 shows the mass and energy balance around the azeotropic dividing wall column, together with the main design parameters. The column has a total of 45 stages, 13 stages for the fractionation section and 32 stages for stripping sections. Butanol and water are the bottom products, while acetone and ethanol with some water (AEW) are obtained as distillate. The ABE feed and the aqueous phase recycled from the decanter are fed on 1st stage of the stripping section (14th stage of A-DWC), which separates water as bottom product. The liquid flowing down the column is routed to the right stripping section. From the 13th stage, a mixture close to


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the azeotropic composition is withdrawn as side stream, cooled, and sent to liquid-liquid separation. The organic phase is recycled on 2nd stage of the right stripping section (15th stage of A-DWC), from which butanol is obtained as bottom product. The aqueous phase is recycled to the left stripping section. 59.43°C, 1 bar

Qcond = –7790 kW 1

AEW 31.7°C, 1 bar 1515.87 kg/h 80.16%w A 16.14%w E 01.98%w W 01.72%w CO2

Aqueous 97°C, 1.13 bar

Azeotrope

Qc = –1634 kW

13 98.1°C, 1.12 bar

Feed

1

97°C, 1.13 bar

97°C, 1.14 bar 15

Qh = 5777 kW

DWC-R

25°C, 1.13 bar 26971 kg/h 04.50%w A 18.54%w B 00.91%w E 75.83%w W 00.04%w Acetic Acid 00.08%w Butiric Acid 00.10%w CO2

DWC-L

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


Diam = 1.22 m Qreb = 3012 kW

Qh = 814 kW

40°C, 1.12 bar 17549.5 kg/h 00.63%w A 42.89%w B 00.74%w E 55.69%w W 00.05%w Acids

Organic

Reflux ratio = 27.02 Boilup ratio = 2.49 Diam = 2.09 Qreb = 1971 kW

38.3°C, 1.12 bar 9072.94 kg/h 00.72%w A 76.32%w B 00.93%w E 21.95%w W 00.08%w Acids

Decanter

Qh = 661 kW

32 45

Butanol Water 25°C, 1.44 bar 20425 kg/h 99.86%w H2O 00.2%w Acids

25°C, 1.44 bar 5030.14 kg/h 99.4%w B 00.5%w W 00.1%w Acids

A-DWC 110.1°C, 1.44 bar

Qc = –2098 kW

127.9°C, 1.44 bar

Qc = –471 kW

Aqueous 38.3°C, 1.12 bar 8476.56 kg/h 00.53%w A 07.11%w B 00.53%w E 91.80%w W 00.03%w Acids

Figure 4. Azeotropic dividing-wall column (without heat integration) Figure 5 illustrates the internal vapor and liquid flowrates along the A-DWC unit, on both sides. While all stages are balanced in term of the liquid-vapor load, stage 14 (right side) has the lowest amount of liquid due to the side stream withdrawal, but this is compensated by the return of organic liquid from the decanter on stage 15. There is no liquid that comes from the fractionation section to the left stripping section; the preheated feeds work as “reflux” because is fed on first stage of this section. This configuration gives the lowest energy requirement in this left side. On


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the right side, the 14th stage is compulsory for the separation of butanol-water azeotrope. A minimum of one stage was kept between the side product stream and organic phase feed, which leads to a minimum energy requirement in right reboiler.

Figure 5. Internal vapor and liquid flowrates in the A-DWC unit (left and right side of the wall). Figure 6 plots the temperature and mass composition profiles along the A-DWC unit. It is worth noting that the temperature difference between the two sides of the wall is less than 20 K hence there is no need for special insulating measures. The bottom of the column shows two temperatures due to the use of two reboilers, for the water and butanol products respectively. In terms of composition, the profiles confirm the high purity of the bottom products, while a ternary mixture (AEW) is obtained as distillate stream. On the 13th stage of the A-DWC the composition of the heterogeneous azeotrope (butanol-water) can be observed, while the aqueous and organic phases can be seen on the 14th stage for the left and right sides of the dividing wall, respectively.


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1

130 RIGHT SIDE

Ethanol

0.8

110

LEFT SIDE

Butanol

Water

0.9

120

LEFT SIDE RIGHT SIDE

0.7

100

Mass fraction

Temperature / [°C]

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


90 80 70

0.6

Water

0.5 0.4 0.3

Acetone

0.2

60

Butanol

0.1

50

0

0

5

10

15

20

25

30

35

40

45

50

0

5

10

15

20

25

30

35

40

45

50

Stage

Stage

Figure 6. Temperature (left) and mass composition (right) profiles along the A-DWC unit. Energy-integrated heat-pump assisted distillation Figure 7 (top) shows the composite curves that reveal the energy targets. Heat integration may lead to important reduction of heating and cooling requirements. However, the energy savings are rather small compared to previously reported separation sequences (Figure 1). In particular, the vapor stream from the top of the A-DWC cannot be used for heat integration due to its low temperature. However, recompression to 5.8 bar (which requires 1646 kW) increases the temperature to 150 °C, which is useful for heat integration. More precisely, partial condensing the vapor stream, at about 116 °C, provides the heat (3012 kW) necessary to drive one reboiler of the A-DWC unit. Further condensation and sub-cooling to 54.5 °C makes available 5777 kW, which are used to preheat the feed stream. Figure 7 (bottom) shows that (for a Q/W ratio of 7.43) the vapor recompression heat pump helps to reduce the heating and cooling requirements by over 50% (equivalent to 2.7 MJ/kg butanol). This figure is impressive suggesting that applying heat pumping will be beneficial in this case. The hot utility requirements can be reduced from 7342 to 3309 kW, while the cold utility needs can be reduced from 7100 to 3067 kW (the rest being ensured by inter process streams transfer).


15


140

HEATING 7342 kW

Temperature[°C]

120 100 80 hot curve

60

cold curve pinch 54 °C

40 COOLING

20

7100 kW

0 0

5000

10000

15000

20000

Enthalpy [kW] 160

HEATING 3309 kW

140 120

Temperature[°C]

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

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100

hot curve

80

pinch 115.18 °C cold curve

60 40 COOLING 3067 kW

20 0 0

5000

10000

15000

20000

Enthalpy [kW]

Figure 7. Composite curve for simple heat integration (top) and heat pump assisted (bottom) Figure 8 shows the dependence on the pressure of the log-mean temperature difference (LMTD) in HEX1 and of the compressor power, as the key operating parameters of the heat pump. Note that the compressor outlet temperature is max. 150°C for safety reasons: at higher temperatures the system may fail from worn rings, acid formations, oil breakdown.25


16

1750

Compressor duty / [kW]

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


0.5

Compressor duty

1700

1650

1/LMTD

0.4

0.3

Tcomp,out=150°C 1600

1550

0.2

1/LMTD / [1/K]

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0.1

LMTD=7.21

1500

0 5.1 5.2 5.3 5.4 5.5 5.6 5.7 5.8 5.9

6

6.1 6.2 6.3

Pressure / [bar] Figure 8. Dependence on pressure of the compressor duty (used for vapor recompression) and of the log-mean temperature difference (LMTD) in the heat exchanger of the heat pump. Figure 9 shows the grid diagram used for the development of the heat exchanger network (HEN), based on Pinch analysis. The proposed HEN reduces the energy requirement for separation close to the calculated values in the composite curve and grand composite curve. The complete HEN includes: 1 heat pump (only its heat exchanger actually), 1 reboiler (Reb-R), 1 condenser (Cond), 3 coolers (Cool1, Cool2, Cool3), and 4 heat exchangers (Hex1, Hex2, Hex3, Hex4). The streams present are: the cold utility, the organic and aqueous phase streams, the initial feed (FEED), the right reboiler cold stream (REB-R), the left reboiler cold stream (REB-L), the vapor from top column (DIST.VAP.), the product streams (BUTANOL and WATER), the side product stream azeotrope (SIDE(Az)) and the hot utility (low pressure steam 6 bar).


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Figure 9. Grid diagram showing the process streams and the heat exchangers network Figure 10 presents the process flowsheet including vapor recompression and energy integration. The mass balance and key design parameters are also included, but some of them are identical to those presented in Figure 4 hence the same explanations will not be repeated here. Compared to Figure 4, a key difference is that the top vapor stream is compressed from 1 to 5.8 bar (in order to increase its temperature from about 60 °C to 150 °C), thus upgrading its thermal energy to provide heat to the (left) side reboiler (HEX1), then to preheat the diluted ABE feed (HEX2) and is eventually getting condensed. Additional heat is recovered by using the water product stream to preheat the aqueous and organic phases (HEX3, HEX4). The proposed heat integration for the new process design preheats the column-inlet streams to 97 °C. For this reason, every stream crosses a heat exchanger, being heated as follows: the initial feed stream is heated with the compressed vapors, while the aqueous and organic phase streams are heated by the hot water product stream. The temperature difference in each side of the heat exchangers exceeds 10 °C (a typical value used in industrial heat exchangers). It should be noted that in spite of the high degree of integration, a vapor recompression assisted


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dividing-wall column is still well controllable although some minor design modifications may be required – as demonstrated in the recent literature.26-28

59.4°C, 1 bar

Cond

vapor

Qcond = –647 kW

1

Comp

liquid

54.5°C, 5.78 bar

31.7°C, 1 bar

AEW

P = 1646 kW Qc = –1634 kW 13 98.1°C, 1.12 bar

110.1°C, 1.44 bar

32 45 110.1°C, 1.44 bar

110.1°C, 1.44 bar

FL

97°C, 1.12 bar

Hex1

A = 491 m2 110.177°C, 1.44 bar

P2

A-DWC

38.3°C, 1.12 bar

P1

Organic Hex4

A = 74

Decanter Cool3

m2

Qc = –622 kW 51.0°C, 1.43 bar

25°C, 1 bar

51.0°C, 1.43 bar

Water

Reflux ratio = 27.02 Boilup ratio = 2.49 Diam = 2.43 m Hex3 Qreb = 1971 kW

Aqueous

DWC-L

97°C, 1.13 bar

Aqueous

vapor

vapor

149.6°C, 5.8 bar

97°C, 1.14 bar

DWC-R

15

38.3°C, 1.12 bar

Hex2

A = 301 m2

40°C, 1.12 bar

Azeotrope

Cool1

1

Feed

110.1°C, 1.44 bar

Feed

51.5°C, 1.43 bar

97°C, 1.13 bar

25°C, 1 bar

115.8°C, 5.79 bar

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


Cool2 110.1°C, 1.44 bar

A = 59 m2 127.9°C, 1.44 bar

Qc = –471 kW 25°C, 1 bar

Butanol

Figure 10. Process flowsheet of the new downstream separation process based on heat pump assisted A-DWC (heat integrated). PROCESS EVALUATION The total equipment cost was evaluated at 5250 k$/year, and the optimal operating cost is 1435 k$/year, including also the cooling of the products. Table 2 provides a summary of the costs. Table 3 presents a comparison of the energy efficiency, including the intensification factor as recently proposed and described.29 The energy required for heating without any heat integration and no heat pump assistance is 8.78 MJ/kg butanol. But the heat pump assisted A-DWC design requires only 2.7 MJ/kg butanol (amounting 58% less than the conventional separation


19


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sequence). Nonetheless, using a heat pump increases the capital cost due to the expensive compressor (1581.5 k$). Yet, considering the energy savings (1.69 MJ/kg butanol) evaluated at 1893 k$/year, the payback time of the heat pump is only 10 months. Table 2. Economic evaluation of the heat pump assisted A-DWC for butanol recovery Item description (unit) Shell / [10 US$]

DWC 718.1

Decanter 71.6

Coolers -

Exchangers -

Flash &Comp 1618.9

Trays / [103 US$]

94.6

-

-

-

-

Condenser / [103 US$]

266.7

-

1179.5

-

-

Reboiler / [103 US$]

497.8

-

-

803.4

-

Heating / [103 US$/year]

441.8

-

-

-

737.3

Cooling / [103 US$/year]

82.5

-

173.2

-

-

1050.0

23.9

566.4

267.8

1276.8

3

TAC / [103 US$/year]

Table 3. Energy efficiency comparison for the new vs previous butanol downstream processes

Downstream process type

Energy requirements (MJ/kg butanol)

Difference (%) vs. conventional

Intensification factor (IFenergy)

Conventional process (decanter + distillation)

6.30

0%

1.00

Dividing-wall column (DWC) distillation (heat integrated)

4.46

-29%

1.41

Azeotropic dividing-wall column (A-DWC)

8.78

+39%

0.71

Heat pump assisted A-DWC (heat integrated)

2.70

-58%

2.33

The sustainability of the ABE process for producing butanol has been reported using four metrics: energy efficiency, material efficiency, land use and costs.30 Here we look only at the


20

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downstream processing part of the ABE process. The CO2 emissions associated with the downstream process are estimated at 1429 kg/h (11.43 ktpy) when using a heat-integrated heat pump assisted A-DWC. The total net carbon tax has been evaluated at 7.87 $/h (62.96 k$/year). This figure could be further reduced (by 25 %) if the electricity used by the compressor of the heat pump comes from renewable sources (e.g. wind, solar, geothermal). CONCLUSIONS The biobutanol recovery from the ABE mixture obtained by fermentation can be efficiently achieved in only a few separation units: three classic distillation columns are combined in one azeotropic dividing-wall column (A-DWC) that is effectively coupled with a compressor for vapor recompression, and a decanter that is used for the liquid-liquid split of the heterogeneous azeotrope butanol-water. The novel downstream process proposed was successfully designed, optimized, and heat integrated using process intensification principles and process simulation. The process evaluation proves that the new downstream processing is economically feasible and sustainable. The energy requirement is drastically reduced by applying energy integration and vapor recompression technology. The investment cost of the process (with a 40 ktpy capacity) is 5250·103 US$, and the total operating cost is 1434·103 US$/year. Although the cost of the compressor used for heat pumping is rather high (1581.5·103 US$), the payback period is only 10 months. The highly integrated A-DWC system reduces the energy requirement for butanol separation to only 2.7 MJ/kg butanol, a reduction of 58% (equivalent to a 2.33 intensification factor) as compared to a conventional design.


21


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AUTHOR INFORMATION Corresponding Author Anton A. Kiss, School of Chemical Engineering and Analytical Science, The University of Manchester, Sackville Street, Manchester, M13 9PL, United Kingdom. [email protected] Author Contributions The manuscript was written through equal contributions of all authors. All authors have given approval to the final version of the manuscript.

ACKNOWLEDGMENT Financial support of the European Commission through the European Regional Development Fund and of the Romanian state budget, under the grant agreement 155/25.11.2016 (Project POC P-37-449, acronym ASPiRE) is gratefully acknowledged. AAK gratefully acknowledges the Royal Society Wolfson Research Merit Award. The authors also thank the reviewers for their insightful comments and suggestions.


22

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APPENDIX The total investment costs (CapEX) include the heat exchangers, coolers, heat pump, flash unit, distillation column, and the decanter. A payback period of 3 years is used, with 8000 hours/year operating time. The total investment costs (CAPEX) include the heat exchangers, coolers, heat pump, flash (L-V), distillation column, and decanter. The cost of the equipment can be estimated using standard cost correlations and considering the Marshall & Swift equipment cost index M&S = 1536.5 (in 2012) – the latest value available in the open literature.22

(

CHEX (US $) = ( M & S / 280 ) ⋅ ( 474.7 ⋅ A0.65 ) 2.29 + Fm ( Fd + Fp )

)

(6)

where A is the area (m2), Fm = 1 (carbon steel), Fd = 0.8 (fixed-tube), Fp = 0 (less than 20 bar). The heat transfer coefficient used to calculate the heat transfer area was U=0.585 kW/m2/K for reboiler and U=0.85 kW/m2/K for other heat exchangers. The design factor was taken as Fd = 1.35 for the reboilers and heat exchangers. Cshell (US $) = ( M & S / 280 ) ⋅ ( 957.9 ⋅ D1.066 ⋅ H 0.82 ) ⋅ ( 2.18 + Fc )

(7)

The cost of the columns shell was calculated considering the following correction factor: Fc = Fm·Fp

(8)

where the material factor is Fm = 1 (carbon steel), and the pressure factor: Fp = 1 + 0.0074 ⋅ ( P − 3.48 ) + 0.00023 ⋅ ( P − 3.48 )

2

(9)

The distillation columns diameter (D) were obtained by the tray sizing utility from Aspen Plus, while the height of the column was evaluated as:


23


H = 0.6·(NT-1) + 2; (expressed in meter)

Page 24 of 29

(10)

The cost of the trays was evaluated as follows:

Ctrays (US $) = NT ⋅ ( M & S / 280 ) ⋅ 97.2 ⋅ D1.55 ⋅ ( Ft + Fm )

(11)

where Ft = 0 (sieve trays) and Fm = 1 (carbon steel). The cost of the compressor was calculated by the following relationship:

Ccomp = ( M & S / 280) ⋅ (664.1P 0.82 Fc )

(12)

where P is the power (kW) and Fc =1 the correction factor (specific for a centrifugal motor). The heating and cooling costs considered are standard: LP steam (6 bar, 160 °C, $7.78/GJ), cooling water (25 °C, $0.72/GJ) and chilled water (5 °C, $4.43/GJ). The compressor power cost taken into account is 15.5 $ per GJ power.22


24

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ABBREVIATIONS ABE

Acetone-butanol-ethanol

A-DWC

Azeotropic dividing-wall column

CAPEX

Capital expenditures

DWC

Dividing-wall column

HP

Heat pump

HR

Heat ratio

IF

Intensification factor

LMTD

Log mean temperature difference

M&S

Marshall & Swift equipment cost index

OPEX

Operating expenditures

VRC

Vapor recompression

TAC

Total annual cost


25


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REFERENCES 1. Quiroz-Ramírez, J. J.; Sánchez-Ramírez, E.; Hernández-Castro, S.; Segovia-Hernández, J. G.; Ponce-Ortega, J. M. Optimal planning of feedstock for butanol production considering economic and environmental aspects, ACS Sust. Chem. Eng., 2017, 5 (5), 4018-4030. 2. Bharathiraja, B.; Jayamuthunagai, J.; Sudharsanaa, T.; Bharghavi, A.; Praveenkumar, R.; Chakravarthy, M.; Yuvaraj, D. Biobutanol - An impending biofuel for future: A review on upstream and downstream processing tecniques, Renew. Sust. Energy Rev., 2017, 68, 788-807. 3. Mascal, M. Chemicals from biobutanol: Technologies and markets, Biofuels, Bioprod. Bioref., 2012, 6, 483-493. 4. Garcia, V.; Päkkilä, J.; Ojamo, H.; Muurinen, E.; Keiski, R. L. Challenges in biobutanol production: How to improve the efficiency?, Renew. Sust. Energy Rev., 2011, 15, 964980. 5. Green, E. M. Fermentative production of butanol - The industrial perspective, Curr. Opin. Biotechnol., 2011, 22, 337-343. 6. Huang, H. J.; Ramaswamy, S.; Liu, Y-Y. Separation and purification of biobutanol during bioconversion of biomass, Sep. Purif. Technol., 2014, 132, 513-540. 7. Straathof, A. J. J. 2.57 - The proportion of downstream costs in fermentative production processes, Compr. Biotechnol. (Second Edition), 2011, 811-814. 8. Kiss, A. A.; Lange, J. P.; Schuur, B.; Brilman, D. W. F.; van der Ham, A. G. J.; Kersten, S. R. A. Separation technology - Making a difference in biorefineries, Biomass Bioenergy, 2016, 95, 296-309. 9. Kujawska, A.; Kujawski, J.; Bryjak, M.; Kujawski, W. ABE fermentation products recovery methods - A review, Renew. Sust. Energy Rev., 2015, 48, 648-661. 10. Pfromm, P. H.; Amanor-Boadu, V.; Nelson, R. Bio-butanol vs. bioethanol: A technical


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and economic assessment for corn and switchgrass fermented by yeast or Clostridium acetobutylicum. Biomass Bioenergy, 2010, 34, 515-24. 11. Qureshi, N.; Li, X. L.; Hughes, S.; Saha, B. C.; Cotta, M. A. Butanol production from corn ﬁber xylan using Clostridium acetobutylicum. Biotechnol. Progr., 2006, 22, 673-80. 12. Tashiro, Y.; Yoshida, T.; Noguchi, T.; Sonomoto, K. Recent advances and future prospects for increased butanol production by acetone-butanol-ethanol fermentation, Eng. Life Sci., 2013, 13, 432-445. 13. Xue, C.; Zhao, J-B.; Liu, F-F.; Lu, C-G.; Yang, S-T.; Bai, F-W. Two-stage in situ gas stripping for enhanced butanol fermentation and energy-saving product recovery, Bioresour. Technol., 2013, 135, 396-402. 14. Abdehagh, N.; Tezel, F. H.; Thibault, J. Separation techniques in butanol production: Challenges and developments, Biomass Bioenergy, 2014, 60, 222-246. 15. Liu, G.; Wei, W.; Jin, W. Pervaporation membranes for biobutanol production, ACS Sust. Chem. Eng., 2014, 2 (4), 546-560. 16. Sanchez-Ramirez, E.; Quiroz-Ramirez, J. J.; Hernandez-Castro, S.; Segovia-Hernandez, J. G.; Kiss, A. A. Optimal hybrid separations for intensified downstream processing of biobutanol, Sep. Purif. Technol., 2017, 185, 149-159. 17. Staggs, K. W.; Qiang, Z.; Madathil, K.; Gregson, C,; Xia, Y.; Vogt, B.; Nielsen, D. R. High efficiency and facile butanol recovery with magnetically responsive micro/ mesoporous carbon adsorbents, ACS Sust. Chem. Eng., 2017, 5 (1), 885-894. 18. Patrascu, I.; Bildea, C. S.; Kiss, A. A. Eco-efficient butanol separation in the ABE fermentation process, Sep. Purif. Technol., 2017, 177, 49-61. 19. Blahusiak, M.; Kiss, A. A.; Babic, K.; Kersten, S. R. A.; Bargeman, G.; Schuur, B. Insights into the selection and design of fluid separation processes, Sep. Purif. Technol., 2018, 194, 301-318.


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20. Kiss, A. A. Advanced distillation technologies - Design, control and applications, Wiley, Chichester, UK, 2013. 21. Kiss, A. A.; Infante Ferreira, C. A. Heat pumps in chemical process industry, CRCPress, Taylor & Francis Group, Boca Raton, US, 2016. 22. Dimian, A. C.; Bildea, C. S.; Kiss, A. A. Integrated design and simulation of chemical processes, 2nd Edition, Elsevier, Amsterdam, 2014. 23. Smith, R. Chemical process design and integration, 2nd Edition, Wiley-Blackwell, Chichester, UK, 2016. 24. Plesu, V.; Bonet, R. A. E.; Bonet, J.; Llorens, J. Simple equation for suitability of heat pump use in distillation, Comput. Aided Chem. Eng., 2014, 33, 1327-1332. 25. Luo, H.; Bildea, C. S.; Kiss, A. A. Novel heat-pump-assisted extractive distillation for bioethanol purification, Ind. Eng. Chem. Res., 2015, 54, 2208-2213. 26. Patrascu, I.; Bildea, C. S.; Kiss, A. A. Dynamics and control of a heat pump assisted extractive dividing-wall column for bioethanol dehydration, Chem. Eng. Res. Des., 2017, 119, 66-74. 27. Luyben, W. L. Control of an azeotropic DWC with vapor recompression, Chem. Eng. Process., 2016, 109, 114-124. 28. Luyben, W. L. Improved plantwide control structure for extractive divided-wall columns with vapor recompression, Chem. Eng. Res. Des., 2018, 123, 152-164. 29. Fernandez Rivas, D.; Castro-Hernández, E.; Villanueva Perales, A. L.; Van der Meer, W. Evaluation method for process intensification alternatives, Chem. Eng. Process., 2018, 123, 221-232. 30. Uyttebroek, M.; Van Hecke, W.; Vanbroekhoven, K. Sustainability metrics of 1-butanol, Catal. Today, 2015, 239, 7-10.


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For Table of Contents Use Only

VRC

A-DWC

Sustainable process Reduced investment Major energy savings Low CO2 emissions

Acetone Ethanol

ABE Water

Azeotrope Organic

Aqueous Water

Decanter

Water

Butanol

Aqueous

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


Flash

Process integration and intensification

Butanol

Downstream processing of butanol using vapor recompression assisted azeotropic dividing-wall column that allows major energy savings and low CO2 emissions


29

Eco-efficient Downstream Processing of Biobutanol by Enhanced

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