Energy-Efficient Design of an Ethyl Levulinate Reactive DistillationProcess via a Thermally Coupled Distillation with External HeatIntegration ArrangementFelicia Januarlia Novita,† Hao-Yeh Lee,*,‡ and Moonyong Lee*,†

†School of Chemical Engineering, Yeungnam University, Dae-dong 712-749, Republic of Korea‡Department of Chemical Engineering, National Taiwan University of Science and Technology, Taipei 10607, Taiwan

*S Supporting Information

ABSTRACT: In this paper we describe an energy-efficientprocess design using a reactive distillation (RD) of theesterification reaction for the synthesis of ethyl levulinate(LAEE). Two designs, a neat design and an excess design,were examined to find the best configuration for the process.In the neat design, there are two columns with equal molarfeeds to produce LAEE. An additional column is required inthe excess design to separate LAEE from the unreactedreactant (levulinic acid, LA). Later in the process, thisunreacted LA is recycled back into the RD column. Comparedto the neat design, the excess design showed superiority for theesterification process in terms of the total annual cost (TAC)and energy requirement. Therefore, three energy integration configurations, external heat integration (HI), thermally coupleddistillation (TCD), and a combination of TCD and HI, were investigated in this excess design. A side reboiler was implementedin the water removal column to overcome the limitation of the temperature difference for heat transfer. The simulation resultsshowed that the configuration of TCD with HI was the best one, which saved 32.9% and 10.5% of the energy and TAC,respectively, compared to the conventional RD with three columns without HI.

1. INTRODUCTION

Levulinic acid (LA) is a keto acid that can be produced by theacid-catalyzed dehydration and hydrolysis of hexose sugars.1 In2012, BioMetics Inc. established the biorefine process via diluteacid hydrolysis to produce LA from cellulosic feedstocks,including paper mill waste, wood waste, and agriculturalresidues, in 50−70% yield.2−4

LA produces many derivatives of ester levulinate (EL), whichhas been proposed for use as a fuel additive.5 Those ELs havetheir own characteristics.6 As seen in Table S1 (SupportingInformation), among various ELs, ethyl levulinate (LAEE) hasa high blending octane number, (RON + MON)/2, in whichRON is the research octane number and MON is the motoroctane number, of 107.5 and a high oxygen content (33 wt %),making LAEE better than other ester levulinates for gasolinefuel additives.7 Grove et al.8 reported the properties of LAEEand other ester blends with diesel fuel. They found that LAEEcould be separated from diesel fuel at temperatures below 0 °C.However, there is a limit to how much LAEE can be blendedinto diesel fuel due to its very low cetane number.9 Lake andBurton10 compared test mixtures of ultralow sulfur diesel(ULSD) fuel, biodiesel, and LAEE in a turbocharged enginewith those of ULSD and 20 vol % biodiesel (B20). The resultsrevealed that the LAEE blend has the lowest particulate matter(PM) emissions. The LAEE blend also eliminates the small

increase in nitrogen oxides (NOx) that have been observed forB20 in many studies.11 LAEE is also used as a flavor additive insome foods and a fragrance additive in cosmetics.12 As a resultof these applications, LAEE has attracted considerable attentionas a valued chemical that positively impacts the market. Asreported by Grand View Research,13 the demand for LAEE inthe market is expected to reach $11.8 million by 2022.EL can be produced in two reaction steps. LA is dehydrated

to angelica lactone (AL), which is then mixed with an alcoholto form levulinate.14 However, the two reaction steps make theprocess complicated and expensive. Therefore, another method,in which LA reacts directly with alcohol to form levulinate, wasconsidered. Unfortunately, the conventional esterificationprocess has an equilibrium limitation, which leads to highoperating costs, due to the separation of unreacted reactantsand products. To overcome those problems, reactive distillation(RD), where two important operations, reaction andseparation, occur in a single unit, can be used for theesterification process.15 RD can be more economical for thebiobased chemical manufacturing of esters than traditional

Received: February 15, 2017Revised: April 24, 2017Accepted: May 4, 2017Published: May 4, 2017

Article

pubs.acs.org/IECR

© 2017 American Chemical Society 7037 DOI: 10.1021/acs.iecr.7b00667Ind. Eng. Chem. Res. 2017, 56, 7037−7048

reactor−separator systems because a higher reaction conversioncan be achieved with lower energy consumption, which meansthat the RD column can overcome the equilibrium conversionlimitations present in batch or flow reactors.12,16 Thus far, morethan 1000 papers on RD have been published andapproximately 236 reaction systems that can be built withRD configurations have been reported.17 Sharma andMahajani18 summarized more than 100 important industrialreactions and highlighted the importance of RD in the industry.Tung and Yu19 examined the design of quaternary RDprocesses based on the relative volatilities of the reactantsand products. They concluded six configurations of the RDprocesses as provided in Table S2 (Supporting Information).From Table S2, type I processes have reactants or productswith a huge difference in volatility ranking. Type II processeshave reactants or products with closer relative volatilities. If therelative volatility ranking is organized in an alternating way forthe reactants and products, the process is classified as type III.The esterification system in this study included four

components and one azeotrope. Table S3 (SupportingInformation) lists the boiling point ranking20 of the purecomponents and the azeotrope. The lightest and heaviest of thesystem are the reactants EtOH and LA, respectively. Bothproducts are in the middle ranking of those components.Therefore, the esterification of LA and EtOH can be classifiedas type IR.Sharma and Mahajani highlighted the importance of RD in

industry in their book chapter18 and also mentioned lowconversion of the reaction and high operating costs in the batchor flow reactor. Therefore, this study does not demonstrate thetraditional reactor−separator systems for the LAEE productionprocess. From the literature review of the LAEE productionprocess, no detailed modeling study of the synthesis of LAEEvia the RD configuration has been reported. In addition, mostlythe RD studies for heavy esters in the open literature are type IPor IIIP.

16 However, the LAEE production process is categorizedas type IR. Therefore, the main contributions of this study are tolaunch the RD design for the type IR process and to implementthe RD column to produce LAEE.Distillation is a popular method used in separation processes

and accounts for 40−70% of the costs of a chemical plant.21

Therefore, several technologies have been used to reduce theenergy requirements. One of the technologies that has attracteda great deal of attention is thermally coupled distillation(TCD). Approximately 30% of energy consumption can besaved using TCD compared to the conventional arrangement.22

Several studies have been carried out to combine an RDcolumn and TCD23−27 to obtain more benefits in terms ofenergy savings. Thermally coupled reactive distillation (TCRD)is more economically promising than the conventional RDprocess. However, the application of TCRD is often limited dueto the complexity in the design and control of theconfiguration. On the basis of the location of the interconnect-ing vapor−liquid streams, there are three types of TCRDsystems for ternary separations: (1) TCRD side rectifier, whereone reboiler is eliminated and replaced with an interconnectingvapor liquid stream and the reactive zone is in the top part ofthe column, (2) TCRD side stripper, where one condenser iseliminated and replaced with an interconnecting vapor liquidstream and the reactive zone is in the bottom part of thecolumn, and (3) TCRD fully coupled, where one condenserand one reboiler are eliminated and replaced with theinterconnecting vapor liquid stream in each part and the

reactive zone is in the middle of the column. Both the TCRDside rectifier and side stripper systems have a tendency to offerthe most thermodynamically efficient designs compared to theTCRD fully coupled system.22,24 External heat integration (HI)is also another promising method to improve energy efficiency,providing significant economic benefits from process intensi-fication (PI). Using the heat of the overhead vapor stream inone column to the side or bottom reboiler of the other columnsis a simple way to implement HI.24

Another main contribution of this study is implementation ofthe external HI and TCD into the best design of the LAEEproduction process to take additional advantages from theTCD and HI structures. In this study, a commercial simulator,Aspen Plus, was used to conduct a rigorous simulation. Theequilibrium stages were assumed in all column simulations. Thetotal material, component, and energy balances were calculatedin each column tray via the RadFrac module in Aspen Plus. Forthe optimal design of the LAEE production process, thedesigners’ domain knowledge and engineering trade-off wereemphasized.

2. REACTION KINETICS AND THERMODYNAMICDATA2.1. Reaction Kinetics. The esterification reaction of LA

and ethanol (EtOH) in eq 1 is a reversible reaction that usesAmberlyst 39 as a catalyst. The kinetic data were correlatedwith the nonideal quasi-homogeneous model to determine thekinetic parameters.

+ ⇔ +C H O C H OH C H O H Ok

k5 8 3

(A)2 5

(B)7 12 3

(C)2

(D)2

1

(1)

A, B, C, and D are LA, EtOH, LAEE, and water, respectively.The kinetic equation of the above reaction was obtained fromTsai.28 The rate expression of the quasi-homogeneousesterification kinetic model is

− = −⎛⎝⎜

⎞⎠⎟r k a a

a aKA 1 A BC D

a (2)

Therefore, the equilibrium constant (Ka) of the reversiblereaction can be expressed as

=Kkka

1

2 (3)

The units of the kinetic equation (−rA) are kmol/m3·s, whereai is the liquid activity of each component. The reaction rateconstants (k1 and k2) of the esterification reaction are expressedas a function of temperature as follows:

=−⎛

⎝⎜⎞⎠⎟k A

E

RTexp1 f

0,f

(4)

=−⎛

⎝⎜⎞⎠⎟k A

E

RTexp2 r

0,r

(5)

In eqs 4 and 5, Af and Ar are the pre-exponential factors of theforward and backward reactions, respectively. E0,f and E0,r arethe activation energies of the forward and backward reactions,respectively. The relationship between the reaction heat (Δhf)and activation energy is

Δ = −h E Ef 0,f 0,r (6)

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The kinetic rate of this esterification process can be obtainedusing the above equations and the kinetic model parameters inTable S4 (Supporting Information), where R is the universalgas constant (R = 8.314 kJ/kmol·K) and T (K) is the absolutetemperature. Note that the units of the kinetic modelparameters in Table S4 need to be converted to fit therequirements of the Aspen Plus simulator. The catalyst densityused in this study was assumed to be 770 kg/m3.2.2. Thermodynamic Data. The studied system has four

components, including two reactants (LA and EtOH) and twoproducts (LAEE and H2O). The nonrandom two-liquid(NRTL) model was used to calculate the liquid activitycoefficient (γi) of each component, as written in eq 7.

∑γτ

ττ

=∑

∑+

∑

−∑

∑

=

= = =

=

=

⎡⎣⎢⎢

⎤⎦⎥⎥

G x

G x

G x

G x

G x

G x

ln ij ji ji j

k ki k j

ij j

k kj k

ijm mj mj m

k kj k

1nc

1nc

1

nc

1nc

1nc

1nc

α τ τ= − = + + +G a b T e nT f Texp( ) /ij ij ij ij ij ij ij ij

α τ= + − = =c d T G( 273.15 K) 0 1ij ij ij ii ii

(7)

The NRTL model requires a set of parameters for eachbinary component system. Since this process has fourcomponents, six sets of binary parameters are needed for thisquaternary system. The water and ethanol pairing parameterswere obtained from the Aspen Plus built-in database, and theremaining parameters refer to the regressed result from Resk’spaper.12 The Hayden−O’Connell (HOC) model27 was used toconsider the association in the vapor phase. The HOC equationconsistently manages the solvation of polar compounds anddimerization of the vapor phase in mixtures containingcarboxylic acids.20 Tables S5 and S6 (Supporting Information)list the regressed NRTL parameters and HOC parameters,respectively, used in data regression for the four relatedcomponents.

3. PROCESS DESIGN AND OPTIMIZATIONThere are two important design approaches that can beconsidered when implementing RD processes: a neat design

and an excess design. In the neat design the reactants (A and Bcomponents) have equal molar feeds to satisfy the reactionstoichiometry. In contrast, in the excess design, there is asurplus of one of the reactants (for example, component A)that is fed along with the second fresh feed (component B). Asa consequence, an additional column is required in the excessdesign to separate the unreacted reactant from the products.The amount of excess reactant will be determined from theamount of recycle. Reactant B is the “limiting reactant” in theRD column, and its conversion is high. The conversion ofreactant A in the RD column is not high. In the excess design,the excess ratio of the two reactants is an important designvariable. The reaction conversion can be improved using ahigher excess ratio. This also decreases the number of reactivestages required in the RD column under the same LAEEproduction rate. On the other hand, the energy requirementswill be higher. In contrast, a larger number of reactive stages inthe RD column are required for a low excess ratio, but therecovery of residual reactant has lower energy consumption.Therefore, there is a trade-off between those variables. Bothneat and excess designs were compared in the followingsections.

3.1. Neat Design of Two Columns. First, following theapproach of Lin,29 a system with a low chemical equilibriumconstant (Ka) is favored with an excess reactant design.Therefore, a neat design was implemented because thisesterification system has a high Ka value (Ka = 32.18 at 353.5K) from the reaction kinetics section. Therefore, if the neatdesign is to be applied to this system, a heavy reactant, LA, isentered into the reflux drum of the RD column and a lightreactant, EtOH, is fed into the second stage, which is theoptimal feed location for EtOH. The RD column is operatedunder total reflux. The entire upper section of this column is areactive section. Rectifying trays are not needed in the RDcolumn because the light reactant, EtOH, goes into the uppersection of the RD column, which is the low-temperature zone,and the upper section of the column should serve as a reactivetray. The reaction takes place in the RD column, and thebottom stream is then fed into the next distillation column toseparate the two products. The reaction conversion in the RDcolumn was set to 95.10% so both products (H2O and LAEE)could meet the 95 mol % specification in the next column.Following the design procedure, Figure 1 shows the result of

Figure 1. Flow diagram of the optimal neat design case. X denotes the mole fraction [XEtOH, XH2O, XLAEE, XLA].

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optimal neat design with a stoichiometric feed ratio of EtOHand LA. This neat design was optimized using an iterativeoptimization procedure for the five main design variablesoptimized: (1) number of reactive stages in the RD column(Nrxn,RD), (2) number of stripping stages in the RD column(Ns,RD), (3) EtOH feed location, (4) number of stages in C1(NC1), and (5) C1 feed stage (NF,C1). Figure 2 presents thesensitivity plots of those variables.In the design procedure, three assumptions were made to

design the RD column: (1) the downcomer occupies 10% ofthe tray area, (2) the liquid holdup is set as half-full of catalystAmberlyst 39 in each reactive stage (the catalyst density wasassumed to be 770 kg/m3), and (3) the weir height is 0.1 m,giving a reactive liquid holdup of 0.373 m3. A reflux drum isplaced on top of the reactive section, with a maximum kineticholdup taken to be 20 times that of the tray holdup.For the column operating pressure, it is well-known that, at

high pressure, a slight envelope in the XY diagram for vapor−liquid equilibrium (VLE) is not favorable30 because it willincrease the cost or duty for separation. Moreover, a higherpressure means a higher temperature in the column and higherutility costs due to hot oil or high-pressure steam beingrequired for heating. In contrast, operating under vacuummakes the separation easier. However, it requires additional,costly equipment to draw the vacuum and recover materialdrawn into the vacuum system, which means increasedcomplexity of the flow diagram.31 A greater cross-sectionalarea in the column is also needed to accommodate the highvolumetric flow of material when a column is operating undervacuum conditions. In addition, for an LAEE productionprocess, there is an upper limit on the operating temperaturedue to the Amberlyst 39 used as a catalyst. The maximum

operating temperature of the reactive section in the RD columnis 130 °C. On the basis of these design considerations, normalpressure (at 1 atm) was chosen as the most preferableoperating pressure of the columns. The temperature of thereaction section in the RD column could be kept under 130 °C(see Figure 8a), and cooling water could be used at normalpressure (at 1 atm).The RD column itself has 94 stages consisting of 81 reactive

stages and 12 stripping stages, not including the reboiler.Increasing the reactive stages would increase the reactiveholdup, requiring a further decrease in reboiler duty. On theother hand, extra stages require additional capital cost andcatalyst cost. A total of 81 reactive stages were found as a trade-off among the above advantages and disadvantages. Thereaction conversion in the RD column was set to 95.1% byadjusting the reboiler duty. Figure 3 shows the liquidcomposition profiles in the RD column. The compositionprofile is quite flat until stage 87. This suggests that thetemperature change is not significant in this part. Anotherobservation from the composition profile is the remixing effectin which the concentration of H2O increases as it approachesthe bottom of the column and then decreases between stages93 and 94. The remixing of H2O shows the waste of energy inthe column. As a result, a very high reboiler duty (17.17 Gcal/h) is needed in the RD column to produce LAEE.The bottom outlet of the RD column goes into the next

column, C1, to separate LAEE and H2O. The separation abilityof the C1 column will improve when the total number of stagesin C1 increases, which means that the energy requirement willdecrease; however, the capital cost increases significantly. FromFigure 2, the C1 column has 13 stages and the 11th stage is theoptimal feed location. LAEE was obtained at 95 mol % from the

Figure 2. Sensitivity plots of the optimization variables for the design case.

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bottom of the C1 column, and H2O as a byproduct wasobtained at 95 mol % from the overhead of the C1 column.The reboiler duty of the C1 column was adjusted to achieve thedesired LAEE purity. The reboiler duty of C1 was 2.42 Gcal/h.As a result, the total duty of the neat design was 19.59 Gcal/hwith an optimal total annual cost (TAC) of $5.36 × 106.3.2. Excess Design of Three Columns. As reported in the

previous section, a high TAC and total duty are required tosynthesize LAEE with the neat design; even the purity productis only 95 mol % for LAEE and byproduct, H2O. This meansthat the first approach from Lin29 is hardly applied in thisesterification system. Another approach from Lin,29 whichstates that the system with the lightest reactant, type IR, tendsto prefer the excess design, was examined because theesterification process in this study is categorized as type IR.Another reason to implement the excess design is that thecomposition profile in the RD column is poor using the neatdesign. As shown in Figure 3, a poor reactant compositiondistribution in the reactive zone results in a high concentrationof EtOH but a low concentration of LA. Recycling the LAreactant (HHK, heavier than heavy key) by adding more LA

reactant to the RD column can improve the reactantdistribution in the RD column in such a way that the reactiongoes toward to the production direction to increase the molefraction of LA in the RD column. A higher LAEE can beproduced due to the higher reaction conversion. From ananalysis of the composition profile by Tung and Yu,19 ifreactant B (HHK) is chosen to be in excess, the reactive zoneshould be placed in the upper section of the RD column, wherethe concentration of reactant A (LLK, lighter than light key) ishigher. Figure 4 shows the optimal flow diagram of the excessdesign, in which EtOH is considered as a limiting reactant. Theoverall flow diagram includes an RD column and two othercolumns. Two fresh feeds (EtOH and LA) are all entered intothe reflux drum of the RD column under the total refluxoperation. All of the upper section of this column was a reactivesection. All products and excess LA leave the bottom of the RDcolumn to the subsequent columns for further separation. Thedirect sequence rather than the indirect sequence is preferredfor further separation because less energy is required for thesame separation task. In the indirect sequence, the HHKcomponent will separate from the bottom stream of the firstseparation column while the upper stream will contain the LK(light key) and HK (heavy key) components. This separationtask requires considerable energy. Moreover, it still needs agreat deal of energy to separate the LK and HK components inthe last separation column. Negligible EtOH in the feed streamto the C1 column with no azeotrope in the ternary system ofwater, LAEE, and LA leads to easy separation. The lightproduct (H2O) of 99.5 mol % leaves the first distillationcolumn (C1) as a distillate. The heavy product (LAEE) of 99.5mol % is obtained from the distillate of the C2 column, whereasthe excess reactant (LA) of 24.7 kmol/h is withdrawn from thebottom of the same column and recycled back to the RDcolumn. The reboiler duties of the C1 and C2 columns werevaried to achieve the purity specifications of the products, H2O,and LAEE, while the distillate flow rate of the C1 column andthe bottom flow rate of the C2 column were fixed.Total annual cost (TAC) analysis was used to find the

optimal design. Equation 8 expresses the TAC, which consistsof the annual operating cost (AOC) and the annualized total

Figure 3. Liquid composition profiles of the RD column for the neatdesign case.

Figure 4. Flow diagram of the optimal excess design. X denotes the mole fraction [XEtOH, XH2O, XLAEE, XLA].

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capital cost (TCC), with a payback period (n) of 3 years. Thecalculation is based on the work of Douglas,32 ignoring thecosts of piping and pumps.

= +n

TAC AOCTCC

(8)

Seven variables were optimized on the basis of their apparentcompeting effect on the TAC of the process: (1) number ofreactive stages in the RD column (Nrxn,RD), (2) number ofstripping stages in the RD column (Ns,RD), (3) number ofstages in C1 (NC1), (4) C1 feed stage (NF,C1), (5) number ofstages in C2 (NC2), (6) C2 feed stage (NF,C2), and (7) recycleflow rate. Their feasible search boundaries were determinedfrom an extensive simulation. In this design flow diagram, allreactants should be fed to the RD column from the top. Threeassumptions were made to design the RD column: (1) thedowncomer occupies the 10% tray area, (2) the liquid holdup isset as half-full of catalyst Amberlyst 39 (assumed density 770

kg/m3) in each reactive stage, and (3) the weir height is 0.1 m,resulting in a reactive liquid holdup of 0.069 m3. As in the neatdesign case, the top of the reactive section for the excess designis a reflux drum, with a maximum kinetic holdup taken to be 20times that of the tray holdup.19

The optimal conditions were found using an iterativeoptimization procedure explained as follows: (1) Fix thestripping section number (Ns,RD). (2) Guess the stage numbersin the reactive trays (Nrxn,RD). (3) Go back to step 2 and changeNrxn,RD until TAC is minimized. (4) Go back to step 1 and varyNs,RD until TAC is minimized. (5) Fix the C1 column number(NC1). (6) Guess the C1 feed location (NF,C1). (7) Go back tostep 6 and change NF,C1 until TAC is minimized. (8) Go backto step 5 and vary NC1 until TAC is minimized. (9) Fix the C2column number (NC2). (10) Guess the C2 feed location(NF,C2). (11) Go back to step 10 and change NF,C2 until TAC isminimized. (12) Go back to step 9 and vary NC2 until TAC is

Figure 5. Results of optimization for the excess design case.

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minimized. (13) Vary the LA recycle flow rate until TAC isminimized.Figure 5 shows the result of the flow diagram optimization. A

total of 44 reactive stages gave the lowest TAC. Increasing thisnumber would add reactive holdup, which would further reducethe reboiler duty required. On the other hand, more stagesmean more capital and catalyst cost. A total of 44 reactivestages were found as a trade-off among the above advantagesand disadvantages. The separation ability of the columns willimprove when the number of total stages of both C1 and C2increase. This means that the energy requirement will decreasewhile the capital cost increases. A larger recycle flow ratesuggests a higher reactant concentration for the excess reactantand decreases the reactive stages required in the RD column,resulting in a higher recycle cost and energy requirement. Incontrast, a smaller recycle flow rate indicates a lower reactantconcentration for the excess reactant, which increases thenumber of reactive stages required in the RD column while itreduces the recycle cost and energy requirement.The optimal flow diagram shown in Figure 4 was finally

obtained from TAC analysis. As a result, the reactive section of44 stages and the stripping section of 7 stages were obtained inthe RD column (not including the reboiler stage). The numberof stages in the C1 column was 11, and the feed stage of the C1column was located at the 10th stage. The number of stages inthe C2 column was 18, and the optimal feed stage was found atthe 12th stage. The optimal excess ratio was 1.247 with a TACof $2.29 × 106 and a total reboiler duty of 7.77 Gcal/h.Figures 6 and 7 show the liquid composition and

temperature profiles of the RD, C1, and C2 columns in theoptimal excess design, respectively. The composition profile inFigure 7a indicates that the compositions of the productsimprove almost 1.2- and 2-fold for H2O and LAEE,respectively, compared to those of the neat design (from 0.25

to 0.3 mole fraction for H2O and from 0.05 to 0.1 mole fractionfor LAEE in the first stage of the RD column). The higherproduct purity with lower energy requirements in the excessdesign case results mainly from the improved reactantdistribution in the RD column. Of more importance, theTAC of the excess design was 57.3% lower than that of the neatdesign.The optimization procedure above can be applied to other

processes. Although it does not guarantee the global optimalsolution, the main advantage of this procedure is that thedesigners can analyze the sensitivity of the design variables andthe production phenomena on the basis of their knowledge.

3.3. Comparison of the Neat Design and ExcessDesign. Table 1 lists the design results of the neat and excessdesigns. Although there is one additional column in the excessdesign, the total duty (7.77 Gcal/h) and TAC ($2.29 × 106) of

Figure 6. Liquid composition profiles of the RD (a), C1 (b), and C2 (c) columns in the optimal excess design.

Figure 7. Temperature profiles of the RD (a), C1 (b), and C2 (c) columns in the optimal excess design.

Table 1. Comparison of the Design Results for the Neat andExcess Designs

neat design excess design

product purity (mole fraction) H2O 0.950 0.995LAEE 0.950 0.995

duty (Gcal/h) RD 17.17 3.68C1 2.42 2.37C2 1.72

total duty (Gcal/h) 19.59 7.77capital cost ($) RD 2.65 × 106 7.43 × 105

C1 3.76 × 105 2.39 × 105

C2 3.86 × 105

operating cost ($) RD 2.04 × 106 4.40 × 105

C1 2.92 × 105 2.77 × 105

C2 2.04 × 105

TAC ($) 5.36 × 106 2.29 × 106

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the excess design are significantly lower than those (total duty19.9 Gcal/h, TAC = $5.36 × 106) of the neat design. Theachievable product purity is also higher using the excess design(99.5 mol % LAEE and H2O). Consequently, the excess designwas chosen as the best design for this LA and EtOHesterification process. This design was adopted for a furthercomparison with other process configuration candidates interms of energy efficiency.

4. ENERGY INTEGRATION ARRANGEMENT OF THEEXCESS LAEE RD PROCESS

4.1. External HI Case. One of most common approachesfor HI between the adjacent columns is to utilize the latent heatof the overhead vapor stream of one column as the reboilerheat source of the other. For this type of HI, the temperature ofthe overhead vapor stream in one column should be sufficiently

higher than that of the bottom section in another column. Adouble effect configuration and a heat-pump-assisted config-uration are two typical options to secure the temperaturedifference required for this type of HI. On the other hand, theseconfigurations require either an adjustment of the pressure ofone of the columns or compression of the overhead stream ofone column, which might restrict the economics andimplementation.24 In this study, HI using a heat stream and/or a side reboiler, which requires neither a heat pump nor apressure change, was explored. The minimum approachtemperature (ΔTmin) chosen for HI was 10 °C.As shown in parts a and c of Figure 7, the reboiler

temperature of the RD column is approximately 105.0 °C,whereas the overhead temperature of the C2 column isapproximately 199.5 °C. This highlights the potential for HIwithout a pressure increase of C2 or a heat pump.

Figure 8. Heat integration between the RD and C2 columns.

Figure 9. Optimal flow diagram of the TCD between the RD and C1 columns. X denotes the mole fraction [XEtOH, XH2O, XLAEE, XLA].

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Figure 8 shows the HI between the RD and C2 columns byimplementing a heat stream. In this HI, the latent heat of theoverhead stream in the C2 column was transferred to thebottom reboiler in the RD column. As a result, the duty of thebottom reboiler in the RD column could be reduced from 3.68to 1.98 Gcal/h. This HI saved 8.8% and 21.8% of the TAC andtotal energy requirement of the RD and C2 columns,respectively, compared to the conventional excess designwithout HI.The heat stream here is illustrated as a model of a heat

exchanger. Therefore, the area of the heat exchanger can becalculated using the following equation:

= ΔQ AU T (9)

where Q is the duty (Gcal/h), U is the overall heat transfercoefficient (kW/m2·K), A is the area of the heat exchanger(m2), and ΔT is the temperature difference between the C2

column condenser and the RD reboiler. The duty of this heatstream was 1.71 Gcal/h (1984.6 kW), the temperaturedifference was 94.5 K, and U (overall heat transfer coefficient)was taken from Luyben33 as 0.568 kW/m2·K. As a result, therequired heat exchange area of the heat exchanger wascalculated to be 36.98 m2.

4.2. Thermally Coupled Case. A thermally coupleddistillation (TCD) configuration has gained increasing popular-ity as an effective way to reduce the energy consumption byreducing the remixing effect in a distillation process.22 Asindicated in Figure 6a, an apparent remixing effect mostlyoccurs in the stripping section of the RD column. A partialTCD between the RD and C1 columns was investigated toimprove the energy efficiency by reducing the remixing effect inthe stripping section of the RD column. A partial TCD betweenthe C1 and C2 columns was not considered due to the slightremixing effect in the C1 column. Figure 9 shows the optimal

Figure 10. Effect of the reflux ratio of the RD column (a) and vapor side stream (b) for each recycle flow rate on the TAC.

Figure 11. Liquid composition profiles of the RD (a), C1 (b), and C2 (c) columns in the optimal TCD.

Figure 12. Temperature profiles of the RD (a), C1 (b), and C2 (c) columns in the optimal TCD.

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TCD configuration between the RD and C1 columns. TheTCD configuration was implemented in Aspen Plus using anequivalent TCD configuration with a side rectifier.To focus on the energy-saving effect according to the

configuration, the main structures (total number of stages andfeed location) of the columns in the TCD were the same asthose in the conventional configuration. In this way, all energysavings are credited to the thermal coupling, reflux ratio of theRD column, and recycle flow rate. Figure 10 shows theoptimization results of those three design variables. Theoptimal reflux ratio of the RD column was 1.701, and theoptimal recycle flow rate was 30 kmol/h. The reboiler duties ofthe C1 and C2 columns were 5.31 and 1.73 Gcal/h,respectively, under the optimal conditions of the vapor sidestream of 261 kmol/h. As a result, the total duties of the RDand C1 columns were reduced by approximately 0.73 Gcal/hby implementing thermal coupling between the RD and C1columns.

Figures 11 and 12 show the liquid composition andtemperature profiles of the RD, C1, and C2 columns in theproposed TCD, respectively. The composition at the topsection of an RD column profile was similar to that of theconventional RD design. On the other hand, the remixing effectin the bottom section of the RD column was reduced; 9.39% ofenergy consumption was saved in the TCD configuration witha total energy consumption of 7.04 Gcal/h. Around 2.0% of theTAC could be saved with this configuration compared to theconventional excess design without HI.

4.3. TCD with HI Configuration. As shown in Figure 9,the latent heat of the overhead of the C2 column was largeenough at approximately 1.70 Gcal/h. From parts b and c ofFigure 12, although the bottom temperature of the C1 columnis approximately 9.1 °C higher than the top temperature of theC2 column, a sharp decrease in temperature of the rectifyingsection in the C1 column results in a low enough temperatureto absorb the latent heat of the overhead of the C2 column

Figure 13. Optimal flow diagram of the proposed TCD with HI via a side reboiler. X denotes the mole fraction [XEtOH, XH2O, XLAEE, XLA].

Figure 14. Optimization result of the proposed TCD with HI case: effect of the reflux ratio of the RD column (a) and vapor side stream (b) for eachrecycle flow rate on the TAC.

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through the simple implementation of a side reboiler. In thisconfiguration, the 10th stage was found to be the best locationfor the side reboiler implementation to give the minimumenergy requirement under the required ΔTmin constraint.Figure 13 presents the optimal flow diagram of the proposedTCD with HI via a side reboiler.From the HI method, the bottom reboiler duty of the C1

column could be reduced from 5.31 to 3.42 Gcal/h. The overallenergy requirement of the proposed TCD with HI case was5.21 Gcal/h under the optimal conditions of a vapor sidestream of 273 kmol/h. The net energy savings by the proposedconfiguration was 2.56 Gcal/h, which is equivalent to 32.9%energy savings compared to the conventional case. Note thatthe net savings is higher than the simple sum of those by theTCD configuration (0.73 Gcal/h) and the HI configuration(1.69 Gcal/h) due to the difference in recycle flow rates. Eachrecycle flow rate was found to be the optimal recycle flow rateof each configuration. According to eq 9, if a side reboiler dutyis illustrated as a heat exchanger, then the required heatexchange area of the heat exchanger is 64.20 m2. Figure 14presents the optimization variables of the proposed TCD withHI case. The optimal reflux ratio of the RD column was 1.558,and the optimal recycle flow rate was 30 kmol/h. The TAC ofthis configuration was $ 2.05 x106, which is equivalent to 10.5%savings compared to that of the conventional excess designwithout HI.4.4. Comparison of the Three Configurations. Table 2

lists the design results of the three configurations. The aim of

the proposed combined configuration is to achieve moreadvantages from both energy integrations. All three config-urations saved energy and TAC; however, as shown in Table 2,the TCD with HI case saved the most energy (32.9%) andTAC (10.5%). The TCD-alone case saved 9.4% and 2.0% ofenergy and TAC, respectively, and the HI-alone case saved21.7% and 8.8% of energy and TAC, respectively. The netenergy savings of the TCD with HI case was higher than thesimple sum of the TCD case and the HI case due to thedifference in recycle flow rates. Although the heat exchange areaof the heat exchanger in the TCD with HI case is larger thanthat in the HI case, the percent energy savings in the TCD withHI case was still higher. An additional reason for the highenergy savings in the TCD with HI case is that thermalcoupling can reduce the remixing effect. HI has the greatestinfluence here in terms of energy efficiency. The result showedthe potential of the TCD with HI configuration for enhancingenergy efficiency in the esterification between the LA andEtOH system.

5. CONCLUSIONSLAEE is an important blending component in biodiesel thatcauses modifications in the fuel properties. LAEE is also a goodcandidate as a gasoline fuel additive due to the high blendingoctane number and high oxygen content. A novel LAEE

production process using RD technology was proposed toovercome the equilibrium limitation of esterification. Theultimate goal of this design was to obtain an LAEE with 99.5%purity and maintain the purity of the byproduct (H2O) at99.5%. The effects of several key design variables, including therecycle flow rate, on the economical production of LAEE wereinvestigated. The results showed that the excess design is moreeconomical than the neat design by avoiding much of theenergy required to achieve the desired product in the RDcolumn in the neat design. A detailed optimal steady-statedesign was then found through TAC analysis. Finally, an energyintegration arrangement combining a TCD and external HI wasproposed to enhance the energy efficiency of the excess RDdesign configuration. The combined arrangement offered moreadvantages compared to the TCD- and HI-alone configurationsand achieved significant energy and TAC savings of 32.9% and10.5%, respectively. Overall, the proposed excess RD designwith the combined energy integration configuration is apromising option for the esterification process between theLA and EtOH system to produce high-purity LAEE in anenergy-efficient manner.

■ ASSOCIATED CONTENT*S Supporting InformationThe Supporting Information is available free of charge on theACS Publications website at DOI: 10.1021/acs.iecr.7b00667.

Fuel characteristics of each ester levulinate, process typesbased on the relative volatility ranking, boiling pointranking, kinetic model parameters, thermodynamicparameters, and total annual cost calculation (PDF)

■ AUTHOR INFORMATIONCorresponding Authors*Tel.: +886-2-2737-6613. E-mail: haoyehlee@mail.ntust.edu.tw.*Tel.: + 82-5-3810-2512. E-mail: mynlee@yu.ac.kr.

ORCIDMoonyong Lee: 0000-0002-3218-1454NotesThe authors declare no competing financial interest.

■ ACKNOWLEDGMENTSThis study was supported by the Basic Science ResearchProgram through the National Research Foundation of Korea(NRF) funded by the Ministry of Education (Grant2015R1D1A3A01015621) and by the Priority ResearchCenters Program through the NRF funded by the Ministry ofEducation (Grant 2014R1A6A1031189). We are also gratefulfor the support from the Ministry of Science and Technology ofTaiwan under Grant MOST 104-2221-E-011-149-MY2.

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Table 2. Comparison of the Conventional, TCD, HI, andTCD with HI Cases

conventional TCD HI TCD with HI

total duty (Gcal/h) 7.77 7.04 6.08 5.21recycle flow rate (kmol/h) 26 30 26 30energy savings (%) 9.4 21.7 32.9TAC savings (%) 2.0 8.8 10.5

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