Eco-efficient Downstream Processing of Biobutanol by Enhanced [617817]

Eco-efficient Downstream Processing of Biobutanol by Enhanced
Process Intensi fication and Integration
Iulian Patras ̧cu,†Costin Sorin B îldea,†and Anton A. Kiss *,‡,§
†University “Politehnica ”of Bucharest, Polizu 1-7, 011061 Bucharest, Romania
‡School of Chemical Engineering and Analytical Science, The University of Manchester, Sackville Street, Manchester, M13 9PL,
United Kingdom
§Sustainable Process Technology Group, Faculty of Science and Technology, University of Twente, PO Box 217, 7500 AE Enschede,
The Netherlands
ABSTRACT: The biobutanol stream obtained after the
fermentation step in the acetone −butanol −ethanol process has
a low concentration (less than 3 wt % butanol) that leads to highenergy usage for conventional downstream separation. Toovercome the high downstream processing costs, this studyproposes a novel intensi fied separation process based on a heat
pump (vapor recompression)-assisted azeotropic dividing-wallcolumn (A-DWC). Pinch analysis and rigorous processsimulations have been used for the process synthesis, design,and optimization of this novel sustainable process. Remarkably,the energy requirement for butanol separation using heatintegration and vapor recompression assisted A-DWC is reducedby 58% from 6.3 to 2.7 MJ/kg butanol.
KEYWORDS: ABE process, Azeotropic distillation, Dividing-wall column, Heat pumps
șINTRODUCTION
Biomass is a natural way of storing solar energy which can be
converted afterward into biofuels. In the quest for sustainable
production of renewable energy and chemicals, biore fineries
hold the promise for e ffectively 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, andseparation in order for biobutanol to be sustainable andeconomically pro fitable.
1Biobutanol was found to be an
efficient fuel with better physical and chemical properties as
compared to other biofuels. The selection of species, substrates,pretreatment, genetic engineering techniques, and variousdownstream 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., butyraldehydeand butyric acid) have mainstream applications in the solvent,polymer, fuel oxygenate, and specialty chemical markets.
Similarly, butanol dehydration products (butenes and buta-
diene) present essential opportunities in the hydrocarbon fueland synthetic elastomers markets.
3
In the acetone −butanol −ethanol (ABE) fermentation
process, biobutanol is obtained in diluted form, typically lessthan 3 wt % concentration (owing to the severe butanol toxicityto microorganisms). The key challenges in biobutanol
production emphasize the idea of improving the e fficiency ofthe ABE process by altering the upstream (e.g., pretreatment
and fermentation) and the downstream steps (product recoveryand puri fication) by various methods.
4The modi fication of
microorganisms by genetic engineering (to keep them alive and
active at high butanol concentrations) could increase theproductivity, yield, and concentration and thus reduce theproduction costs.
5However, this long-term goal is yet to be
realized, and even it becomes a reality, the product separationand puri fication will still remain a critical challenge.6The other
approach is the development of more e fficient downstream
processes for butanol recovery. This is not a singular issue, asthe need for cost-e ffective 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.9A higher butanol
yield is achieved by using anaerobic bacteria as Clostridium
acetobutylicum10,11and Clostridium beijerinckii .12But the ABE
concentration can be further increased by in situ productrecovery (ISPR) methods such as gas stripping technology to4.5 wt % acetone, 18.6 wt % butanol and 0.9 wt % ethanol.
13
Many techniques are available for ABE separation, for example,distillation, reverse osmosis, adsorption, liquid −liquid extrac-
tion, pervaporation, and others.14−17Process alternatives based
Received: January 21, 2018
Revised: February 24, 2018
Published: February 28, 2018Research Article
pubs.acs.org/journal/ascecg Cite This: ACS Sustainable Chem. Eng. 2018, 6, 5452 −5461
© 2018 American Chemical Society 5452 DOI: 10.1021/acssuschemeng.8b00320
ACS Sustainable Chem. Eng. 2018, 6, 5452 −5461

on advanced distillation technologies have been also reported −
seeFigure 1 .18More 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.
șPROBLEM STATEMENT
Taking into account the energy density of butanol (36 MJ/kg),the use of a classic distillation −decanter method for butanolrecovery is clearly too demanding with energy requirements of
14.5−79.5 MJ/kg butanol.9But 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 threedistillation columns along with one decanter ( Figure 1 , top)
requires 6.3 MJ/kg butanol (excluding COL-4 which separates
the acetone and ethanol byproducts).
18When heat integration
is applied to the conventional sequence, up to 18% energy
savings are possible. But when heat integration was used and
two columns were combined into a dividing-wall column(Figure 1 , btm), the energy requirement was reduced to 4.46
MJ/kg butanol.
18This 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
Figure 1. Process flowsheet of ABE downstream separation sequences: conventional distillation (top) and intensi fied process using DWC and heat
integration (bottom).ACS Sustainable Chemistry & Engineering Research Article
DOI: 10.1021/acssuschemeng.8b00320
ACS Sustainable Chem. Eng. 2018, 6, 5452 −54615453

novel enhanced process that makes use of process intensi fica-
tion (azeotropic dividing-wall column) and process integrationtechniques (e.g., energy integration and vapor recompressionheat pump). By using a highly integrated azeotropic dividing-wall column (A-DWC) assisted by vapor recompression (VRC)technology, the energy savings can be increased signi ficantly as
described hereafter.
șAPPROACH AND METHODOLOGY
The plant capacity considered here is 40 ktpy of butanol. Toaccount for a realistic composition of the ABE mixture such asreported in the literature,
13impurities are also taken into
account; hence, the mixture which must be e fficiently separated
contains 4.5 wt % acetone, 18.6 wt % butanol, 0.9 wt % ethanol,0.1 wt % CO
2, 0.08 wt % butyric acid, and 0.04 wt % acetic acid.
The required product purities are butanol 99.4 wt % and water99.8 wt %. Note that in Figure 1 the last column (COL-4)
separates only the light components (acetone and ethanol) anddoes not contribute to the energy requirement for butanolseparation. Therefore, the process designed in this work willdeliver an acetone −ethanol mixture (96 wt %) without
attempting to split it into high-purity components. The feedstream and all the products are at 25 °C. Regarding the design
and operational constraints, when dividing wall technology isemployed each side of the dividing wall must have the samenumber of trays and the temperature di fference between the
two sides should not exceed 20 °C. Moreover, the use of vapor
recompression is limited by the maximum temperature of thecompressed vapor, which is 150 °C.
Phase Equilibrium. The process is simulated in Aspen
Plus, using nonrandom two-liquid (NRTL) as an appropriateproperty model. Table 1 lists the boiling points of the
components, while Figure 2 illustrates the T−xydiagram of
the binary mixture butanol −water. Notably, both hetero- and
homogeneous azeotropes are formed in this aqueous system,and this further complicates the separation.
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
prefractionator column COL-1 removes a large amountof water as bottom product. This reduces both theinvestment and operation costs of the subsequent 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 toptrays), as its main function is to remove the light
components such that high purity water is obtained inthe bottom stream, without having a tight speci fication
on the distillate.
Lights out first: from the distillate of COL-1, the lightest
components (acetone and the 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 speci fications of the bottom stream are not
stringent (see below).
Perform the most di fficult separation last: removal of
acetone, ethanol, and a part of water in COL-2 iscompatible with this heuristic, as the bottom product ofCOL-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 splitinto 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
COL-2 bottoms are recycled to a location which stillallows 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 speci fication 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 requiredfor column COL-2. Finally, it can be observed that the boilingpoints of the COL-1 and COL-3 distillate streams are almostthe 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 temperature pro files along COL-1 and
COL-3 are expected to be similar. Therefore, columns COL-1and COL-3 can be integrated into one section provided with adividing wall. Thus, all three distillation columns can becombined into a single unit, as shown in Figure 3 (right).Table 1. Boiling Points of the Pure Chemicals and
Azeotropes Involved in the Butanol Recovery
componentboiling point °C
(at 1.013 bar)
acetone 56.14
homogeneous azeotrope:
ethanol (95.63 wt %)/water78.15
ethanol 78.31
heterogeneous azeotrope:
n-butanol (42 wt %)/water95.91
water 100
n-butanol 117.75
acetic acid 118.01n-butyric acid 163.28
Figure 2. T−xydiagram for the binary mixture butanol −water.ACS Sustainable Chemistry & Engineering Research Article
DOI: 10.1021/acssuschemeng.8b00320
ACS Sustainable Chem. Eng. 2018, 6, 5452 −54615454

Process Optimization. The process shown in Figure 3
(right) was optimized using the total annual cost (TAC) as theobjective function to be minimized (according to standardindustrial practices):
=+ TAC OPEXCAPEX
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 thenormal constructive solution for large diameter trayed
columns. However, the number of stages could bedifferent on the two sides, particularly when (structured)
packing is used.
The design speci fications for purity of product stream
(e.g., butanol product 99.4 wt %, water byproduct 99.8
wt %).
Maximum 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.
Maximum 30 kg/h water in distillate (AEW quality),
obtained by manipulating the re flux ratio.
Figure 3. Conceptual design of butanol separation: (top) sequence based on conventional distillation; (left) thermally coupled column; (right)
DWC equivalent con figuration.ACS Sustainable Chemistry & Engineering Research Article
DOI: 10.1021/acssuschemeng.8b00320
ACS Sustainable Chem. Eng. 2018, 6, 5452 −54615455

CO 2recovery in distillate obtained by manipulating the
sideflow rate.
Feed mixture, organic phase, and aqueous phase
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.
Energy Integration. Pinch analysis provides understanding
of the energy targets and subsequent design of the optimal heatexchange network (HEN). The procedure was appliedaccording to the literature.
23A minimum temperature di ffer-
ence of 10 K was used. No additional correction factor wasused, as only 1-pass shell and tube heat exchangers wereconsidered, but other correction factors may be needed forother types of heat exchangers. The coe fficient of performance
(COP) is used for evaluating the feasibility of using a heatpump, while also accounting the additional costs and thepayback time.
η == = − >QW T T T COP / 1/ /( ) 10cr c (2)
where Q= reboiler duty, W= work provided, η= Carnot
efficiency, Tr= reboiler temperature, and Tc= condenser
temperature.24If the Q/Wratio is lower than 5, using a heat
pump (HP) brings no advantages, but if it exceeds 10 then aHP should be considered. The maximum energy savings aregiven by
19
=−· QQ
Qmax savings (%) 100HR /COPreb reb
reb (3)
where HR is the heat ratio (thermal to electrical), with a typical
ratio: 1 MW e= 2.5 MW th.
Then, the investment and operation costs of the complete
process are evaluated based on the correlations provided in theAppendix , to show the advantages of this design.22
Environmental Impact. The potential environmental
impact was evaluated in Aspen Plus using Carbon Tracking tocalculate the CO 2emissions. The fuel source considered is
natural gas, and the CO 2emission factor data source used is the
US Environmental Protection Agency Rule of “E9-5711 ”(CO 2
E-US) 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). Thefollowing equations were used for calculating the CO
2
emissions:
α = ⎜⎟⎛
⎝⎜⎞
⎠⎟⎛
⎝⎞
⎠Q C[CO ]NHV%
1002 emissionsfuel
(4)
λ=−−
−⎛
⎝⎜⎜⎞
⎠⎟⎟ QQ
hTT
TT( 419)()
()fuelproc
procprocFTB 0
FTB stack (5)
where α= 3.67 is the ratio of molar masses of CO 2and C;
NHV (net heating value) is 48900 kJ/kg for natural gas; C%
(carbon content) is 0.41 kg/kg; Qprocis the heat duty required
by the process and provided by the steam (kW); λprocis the
latent heat of steam delivered to the process (kJ/kg); hprocis the
enthalpy of steam delivered to the process (kJ/kg); T0is the
ambient temperature; TFTB(K) and Tstack(K) are the flame and
stack temperature, respectively.
șRESULTS AND DISCUSSION
Azeotropic Dividing-Wall Column. Figure 4 shows the
mass and energy balance around the azeotropic dividing wallcolumn, together with the main design parameters. The columnhas 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 theaqueous phase recycled from the decanter are fed on first stage
of the stripping section (14th stage of A-DWC), whichseparates water as bottom product. The liquid flowing down
Figure 4. Azeotropic dividing-wall column (without heat integration).ACS Sustainable Chemistry & Engineering Research Article
DOI: 10.1021/acssuschemeng.8b00320
ACS Sustainable Chem. Eng. 2018, 6, 5452 −54615456

the column is routed to the right stripping section. From the
13th stage, a mixture close to the azeotropic composition iswithdrawn as side stream, cooled, and sent to liquid −liquid
separation. The organic phase is recycled on the second stage ofthe right stripping section (15th stage of A-DWC), from whichbutanol is obtained as a bottom product. The aqueous phase isrecycled to the left stripping section.
Figure 5 illustrates the internal vapor and liquid flow rates
along the A-DWC unit, on both sides. While all stages are
balanced in terms of the liquid −vapor load, stage 14 (right
side) has the lowest amount of liquid due to the side streamwithdrawal, but this is compensated by the return of organicliquid from the decanter on stage 15. There is no liquid thatcomes from the fractionation section to the left strippingsection; the preheated feeds work as “reflux”because they are
fed on the first stage of this section. This con figuration gives the
lowest energy requirement in this left side. On the right side,the 14th stage is compulsory for the separation of the butanol −
water azeotrope. A minimum of one stage was kept betweenthe side product stream and organic phase feed, which leads toa minimum energy requirement in the right reboiler.
Figure 6 plots the temperature and mass composition pro files
along the A-DWC unit. It is worth noting that the temperaturedifference between the two sides of the wall is less than 20 K,
hence there is no need for special insulating measures. Thebottom of the column shows two temperatures due to the useof two reboilers, for the water and butanol products,respectively. In terms of composition, the pro files con firm the
high purity of the bottom products, while a ternary mixture(AEW) is obtained as distillate stream. On the 13th stage of theA-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 andright sides of the dividing wall, respectively.
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, theenergy savings are rather small compared to previously reportedseparation sequences ( Figure 1 ). In particular, the vapor stream
Figure 5. Internal vapor and liquid flow rates in the A-DWC unit (left
and right side of the wall).
Figure 6. Temperature (left) and mass composition (right) pro files along the A-DWC unit.
Figure 7. Composite curve for simple heat integration (top) and heat
pump-assistance (bottom).ACS Sustainable Chemistry & Engineering Research Article
DOI: 10.1021/acssuschemeng.8b00320
ACS Sustainable Chem. Eng. 2018, 6, 5452 −54615457

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, partiallycondensing the vapor stream, at about 116 °C, provides the
heat (3012 kW) necessary to drive one reboiler of the A-DWCunit. Further condensation and subcooling to 54.5 °C make
available 5777 kW, which is used to preheat the feed stream.
Figure 7 (bottom) shows that (for a Q/Wratio of 7.43) the
vapor recompression heat pump helps to reduce the heatingand cooling requirements by over 50% (equivalent to 2.7 MJ/kg butanol). This figure is impressive suggesting that applying
heat pumping will be bene ficial in this case. The hot utility
requirements can be reduced from 7342 to 3309 kW, while thecold utility needs can be reduced from 7100 to 3067 kW (the
rest being ensured by interprocess streams transfer).
Figure 8 shows the dependence on the pressure of the log-
mean temperature di fference (LMTD) in HEX1 and of thecompressor 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, and oil breakdown.
25
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 forseparation close to the calculated values in the composite curveand grand composite curve. The complete HEN includes 1 heatpump (only its heat exchanger actually), 1 reboiler (Reb-R), 1condenser (Cond), 3 coolers (Cool1, Cool2, Cool3), and 4heat exchangers (Hex1, Hex2, Hex3, Hex4). The streamspresent are the cold utility, the organic and aqueous phasestreams, the initial feed (FEED), the right reboiler cold stream(REB-R), the left reboiler cold stream (REB-L), the vapor fromtop column (DIST.VAP.), the product streams (BUTANOLand WATER), the side product stream-azeotrope (SIDE(Az))and the hot utility (low pressure steam 6 bar).
Figure 10 presents the process flowsheet including vapor
recompression and energy integration. The mass balance andkey 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 di fference is that the top vapor stream is compressed from
1 to 5.8 bar (in order to increase its temperature from about 60to 150 °C), thus upgrading its thermal energy to provide heat
to the (left) side reboiler (HEX1), then to preheat the dilutedABE feed (HEX2) and is eventually getting condensed.Additional heat is recovered by using the water product streamto preheat the aqueous and organic phases (HEX3, HEX4).The proposed heat integration for the new process designpreheats 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 thehot water product stream. The temperature di fference in each
side of the heat exchangers exceeds 10 °C (a typical value used
in industrial heat exchangers).
It should be noted that despite the high degree of integration,
a vapor recompression-assisted dividing-wall column is still well
Figure 8. Dependence on pressure of the compressor duty (used for
vapor recompression) and of the log-mean temperature di fference
(LMTD) in the heat exchanger of the heat pump.
Figure 9. Grid diagram showing the process streams and the heat exchangers network.ACS Sustainable Chemistry & Engineering Research Article
DOI: 10.1021/acssuschemeng.8b00320
ACS Sustainable Chem. Eng. 2018, 6, 5452 −54615458

controllable although some minor design modi fications may be
requiredas demonstrated in the recent literature.26−28
șPROCESS EVALUATION
The total equipment cost was evaluated at 5250 k$/year, and
the optimal operating cost is 1435 k$/year, including also thecooling of the products. Table 2 provides a summary of the
costs.
Table 3 presents a comparison of the energy e fficiency,
including the intensi fication factor as recently proposed and
described.29The 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 2.7MJ/kg butanol (amounting to 58% less than the conventional
separation 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.
The sustainability of the ABE process for producing butanol
has been reported using four metrics: energy e fficiency, material
efficiency, land use, and costs.
30Here we look only at the
downstream processing part of the ABE process. The CO 2
emissions associated with th e 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 thecompressor of the heat pump comes from renewable sources(e.g., wind, solar, geothermal).
șCONCLUSIONS
The biobutanol recovery from the ABE mixture obtained byfermentation can be e fficiently 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 down-
stream process proposed was successfully designed, optimized,
Figure 10. Process flowsheet of the new downstream separation process based on heat pump assisted A-DWC (heat integrated).
Table 2. Economic Evaluation of the Heat Pump-Assisted A-DWC for Butanol Recovery
item description (unit) DWC decanter coolers exchangers flash and comp
shell/[103US$] 718.1 71.6 1618.9
trays/[103US$] 94.6
condenser/[103US$] 266.7 1179.5
reboiler/[103US$] 497.8 803.4
heating/[103US$/year] 441.8 737.3
cooling/[103US$/year] 82.5 173.2
TAC/[103US$/year] 1050.0 23.9 566.4 267.8 1276.8
Table 3. Energy E fficiency Comparison for the New vs
Previous Butanol Downstream Processes
downstream process typeenergy
requirements
(MJ/kg
butanol)difference
(%) vs
conventionalintensi fication
factor
(IF energy)
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% 2.33ACS Sustainable Chemistry & Engineering Research Article
DOI: 10.1021/acssuschemeng.8b00320
ACS Sustainable Chem. Eng. 2018, 6, 5452 −54615459

and heat integrated using process intensi fication principles and
process simulation.
The process evaluation proves that the new downstream
processing is economically feasible and sustainable. The energyrequirement is drastically reduced by applying energyintegration and vapor recompression technology. The invest-ment cost of the process (with a 40 ktpy capacity) is 5250 ×
10
3US$, and the total operating cost is 1434 ×103US$/year.
Although the cost of the compressor used for heat pumping israther high (1581.5 ×10
3US$), the payback period is only 10
months. The highly integrated A-DWC system reduces theenergy requirement for butanol separation to only 2.7 MJ/kgbutanol, a reduction of 58% (equivalent to a 2.33 intensi fication
factor) as compared to a conventional design.
șAPPENDIX
The total investment costs (CapEX) include the heatexchangers, coolers, heat pump, flash unit, distillation column,
and the decanter. A payback period of 3 years is used, with8000 h/year operating time. The total investment costs(CAPEX) include the heat exchangers, coolers, heat pump,flash (L-V), distillation column, and decanter. The cost of the
equipment can be estimated using standard cost correlationsand considering the Marshall & Swift equipment cost indexM&S = 1536.5 (in 2012), the latest value available in the openliterature.
22
=
++CA
FF F(US$) (M&S/280)(474.7 )
(2.29 ( ))HEX0.65
md p (6)
where Ais the area (m2),Fm= 1 (carbon steel), Fd= 0.8 ( fixed-
tube), Fp= 0 (less than 20 bar). The heat transfer coe fficient
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.
=+ CD H F (US$) (M&S/280)(957.9 )(2.18 )shell1.066 0.82
c
(7)
The cost of the columns shell was calculated considering the
following correction factor:
=FF Fcm p (8)
where the material factor is Fm= 1 (carbon steel), and the
pressure factor:
=+ − + −FP P 1 0.0074( 3.48) 0.00023( 3.48)p2
(9)
The distillation columns diameter ( D) were obtained by the
tray sizing utility from Aspen Plus, while the height of thecolumn was evaluated as
=− +H NT 0.6( 1) 2 (expressed in meter) (10)
The cost of the trays was evaluated as follows:
=+ CN D F F (US$) (M&S/280)97.2 ( )trays T1.55
tm (11)
where Ft= 0 (sieve trays) and Fm= 1 (carbon steel).
The cost of the compressor was calculated by the following
relationship:
= CP F (M&S/280)(664.1 )comp0.82
c (12)where Pis the power (kW) and Fc= 1 the correction factor
(speci fic 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
șAUTHOR INFORMATION
Corresponding Author
*E-mail: tonykiss@gmail.com .
ORCID
Anton A. Kiss: 0000-0001-5099-606X
Notes
The authors declare no competing financial interest.
șACKNOWLEDGMENTS
Financial support of the European Commission through the
European Regional Development Fund and of the Romanianstate budget, under the Grant Agreement 155/25.11.2016(Project POC P-37-449, acron ym ASPiRE) is gratefully
acknowledged. A.A.K. gratef ully acknowledges the Royal
Society Wolfson Research Merit Award. The authors alsothank the reviewers for their insightful comments andsuggestions.
șABBREVIATIONS
ABE Acetone −butanol −ethanol
A-DWC Azeotropic dividing wall columnCAPEX Capital expendituresDWC Dividing wall columnHP Heat pumpHR Heat ratioIF Intensi fication factor
LMTD Log mean temperature di fference
M&S Marshall & Swift equipment cost indexOPEX Operating expendituresVRC Vapor recompressionTAC Total annual cost
șREFERENCES
(1) Quiroz-Ramírez, J. J.; Sa ́nchez-Ramírez, E.; Herna ́ndez-Castro, S.;
Segovia-Herna ́ndez, J. G.; Ponce-Ortega, J. M. Optimal planning of
feedstock for butanol production considering economic and environ-
mental aspects. ACS Sustainable 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. Renewable Sustainable Energy Rev. 2017 ,68,
788−807.
(3) Mascal, M. Chemicals from biobutanol: Technologies and
markets. Biofuels, Bioprod. Biorefin. 2012 ,6, 483−493.
(4) Garcia, V.; Pa ̈kkila ̈, J.; Ojamo, H.; Muurinen, E.; Keiski, R. L.
Challenges in biobutanol production: How to improve the efficiency?
Renewable Sustainable Energy Rev. 2011 ,15, 964−980.
(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 , 811−814.10.1016/B978-0-08-088504-9.00492-XACS Sustainable Chemistry & Engineering Research Article
DOI: 10.1021/acssuschemeng.8b00320
ACS Sustainable Chem. Eng. 2018, 6, 5452 −54615460

(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. Renewable
Sustainable Energy Rev. 2015 ,48, 648−661.
(10) Pfromm, P. H.; Amanor-Boadu, V.; Nelson, R. Bio-butanol vs.
bioethanol: A technical 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. Prog. 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 Sustainable 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 Sustainable 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 ,
177,4 9−61.
(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.
(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 ; CRC-Press, 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 ed.; Elsevier: Amsterdam, The
Netherlands, 2014.
(23) Smith, R. Chemical Process Design and Integration , 2nd ed.,
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 ,119,6 6−74.
(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-Herna ́ndez, E.; Villanueva Perales,
A. L.; Van der Meer, W. Evaluation method for process intensificationalternatives. 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.ACS Sustainable Chemistry & Engineering Research Article
DOI: 10.1021/acssuschemeng.8b00320
ACS Sustainable Chem. Eng. 2018, 6, 5452 −54615461