Energy Saving Designs for Separation of a Close Boiling 1 2 Propanediol and Ethylene Glycol Mixture

8/18/2019 Energy Saving Designs for Separation of a Close Boiling 1 2 Propanediol and Ethylene Glycol Mixture

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Energy-Saving Designs for Separation of a Close-Boiling 1,2-Propanediol and Ethylene Glycol Mixture Yi-Chun Chen,† Shih-Kai Hung,† Hao-Yeh Lee,‡ and I-Lung Chien* , ††Department of Chemical Engineering, National Taiwan University, Taipei 10617, Taiwan‡Department of Chemical Engineering, National Taiwan University of Science and Technology, Taipei 10607, Taiwan

ABSTRACT: Separation of 1,2-propanediol and ethylene glycol is an important task for coproduction of these two compoun via hydrogenolysis of glycerol for the purpose of utilizing this biodiesel byproduct. These two compounds exhibit close-b behavior which requires many stages of a regular column and also large energy consumption to meet stringent product puspeci cations. In this paper, several alternative designs are investigated in order to save steam cost and also total annual cothis separation task. Alternative designs considered include multieff ect distillation, heterogeneous azeotropic distillation, andextractive distillation. Signicant reductions of 38.3% in steam cost and 30.6% in total annual cost as compared to the regulacolumn can be obtained by a design owsheet via an extractive distillation system using triethylene glycol as entrainer. Method

for further improving the economics of this extractive distillation system have also been investigated. A simple worthwimproved design is to preserve the energy from the hot entrainer to preheat the fresh feed via a feed-effluent heat exchanger. With this simple improvement, a further 12.8% reduction in steam cost can be made as compared to the original extracdistillation system.

1. INTRODUCTIONGlycerol as a low-cost byproduct of the biodiesel industry can beutilized as a feed material to produce valuable products. One of the economically attractive reaction routes is to coproduce 1,2-propanediol (1,2-PDO) and ethylene glycol (EG) via hydro-genolysis of glycerol.1

− 4 Thus, the energy-saving design of separating these two components is an important research topic

worthy of detailed investigation.1,2-PDO and EG are close-boiling compounds with medianrelative volatility close to unity (1.34 at 1 atm). These twocompounds are difficult to be separated using a regulardistillation column. A common industrial application forseparating a close-boiling mixture is acetic acid dehydration. Itis customary to add an ester into the system as light entrainer forthis separation via heterogeneous azeotropic distillation. Designand control of the acetic acid deh y dration process were studiedextensively in the open literature.5− 11 For the separation of otherclose-boiling mixtures, Jongmans et al.12,13 proposed using ionicliquid as entrainer in an extractive distillation s ystem to separateethylbenzene and styrene. Lek-utaiwan et al.14 proposed using acomputer-aided molecular design (CAMD) approach for thesolvent selection to separate ethylbenzene and mixed xylene.Long and Lee15 discussed a retrot design using a thermally coupled distillation scheme for the separation of heptanes andtoluene using aniline as entrainer. The other ways for separatingclose-boiling mixtures in the open literature make use of themelting point diff erences of the mixtures in either a distillationfreezing design conguration or a hybrid distillation/meltcrystallization process.16

− 19

In Chapter 7 of a book by Doherty and Malone20 alternativeenergy-efficient distillation congurations for the separation of ternary A/B/C systems with no azeotrope were discussed.However, those design congurations cannot directly be appliedin this study because a two-column sequence was required in the

base case of ternary A/B/C systems. For the separation of close boiling systems, the base case contains only one distillationcolumn. For energy-saving design of one column, anothermethod mentioned in the open literature was to use a vaporrecompression scheme so that the vapor from the overhead of this column is compressed as the heat pumping uid to supportthe energy required for its reboiler in the bottom of the samecolumn. However, this scheme requires adding an expensivecompressor to the process and thus is not studied in this paper.

To the best of our knowledge, there are only two studiesinvestigating more energy-efficient ways to separate 1,2-PDOand EG. A U.S.patent21 listedabout 30 potential entrainers to aidthe separation via heterogeneous azeotropic distillation.However, only information on the enhancement of the relative volatility for these two compounds was given by adding apotential entrainer. To make the separation work viaheterogeneous azeotropic distillation, the entrainer shouldform at least one additional minimum-boiling azeotrope witheither of the two original compounds. Also, the new azeotropeshould be heterogeneous so that a decanter can be designed.Thisdesign conguration for this mi xture will be discussed on section

4 of this paper. Another paper22

proposed recovering 1,2-PDOand EG from aqueous solution via reaction with acetaldehyde toform acetals in a reactive distillation column. Since these twoacetals are much easier to be separated than the original twocompounds, two additional reactive distillation columns areneeded for hydrolysis of both acetals back to 1,2-PDO and EGSince the overall process is rather complicated including a rstreactive-distillation column to produce two acetals, a second

Received: October 4, 2014Revised: March 22, 2015 Accepted: March 31, 2015

Article

pubs.acs.org/IECR

© XXXX American Chemical Society A DOI:10.1021/ie503922f Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

http://localhost/var/www/apps/conversion/tmp/scratch_4/pubs.acs.org/IECRhttp://dx.doi.org/10.1021/ie503922fhttp://dx.doi.org/10.1021/ie503922fhttp://localhost/var/www/apps/conversion/tmp/scratch_4/pubs.acs.org/IECR


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section should be even more than that of the stripping section. We studied a regular column operating at vacuum condition. Theresult con rmed the prediction from the Txy plot that theeconomics of the vacuum column is much worse than that of anatmospheric column.

In all simulation studies for this paper, Aspen plus RADFRACsimulations (Version 8.4) in the aspenONE Engineering Suite were conducted. In the optimization study, total annual cost(TAC) is minimized to determine the optimal total number of stages and feed location. The column pressure is set to beatmospheric for convenience of operation. The formulas forcalculating the total annual cost were adopted mainly fromLuyben.23 Table 2 summarizes the formulas for TACcalculations.

The nal optimal design of this regular column is shown inFigure2 with Table3 itemizedTACtermsof this regular column.It isobserved that the optimal NTis 121,and optimal NFis at the68th stage. The numbering convention is to assume thecondenser as the rst stage and reboiler as the last stage. Notethat as predicted, many stages are required for this difficultseparation. In our Aspen Plus RADFRAC simulation, a sieve tray was assumed for the calculations of tray pressure drop via the

“ tray rating” feature. This case will be used as a base case forcomparison purposes with other alternative designs.

3. MULTIEFFECT DISTILLATIONChiang and Luyben24 ,25 investigated ve multieff ect heat-integration congurations for the potential of saving energy ina distillation column. The concept of heat integration is tocombine the condenser of a high-pressure (HP) column and thereboiler of a low-pressure (LP) column into one process-to-process heat exchanger to save energy. They used a methanol− water system as an example and found that three designcon gurations can provide signicant energy savings for all cases with diff erent feed concentrations as compared to the case with aregular column. The design conguration that gave the bestenergy savings is the one that featured heat integration in the

Table 2. Basis of Economics and Equipment Sizing

column diameter (D) Aspen tray sizingcolumn length (L) NT trays with 2 ft spacing plus 20% extra lengthcolumn and other vessel(D and L are in meters)

capital cost 17 640(D)1.066(L)0.802

condensers (area in m2

)heat-transfercoefficient

0.852 kW/K ·m2

diff erentialtemperature

re ux-drum temperature− 315 K

capital cost 7296(area)0.65

reboilers (area in m2)heat-transfercoefficient

0.568 kW/K ·m2


steam temperature− base temperature (Δ T > 20 K)


heat exchangers, liquid-to-liquid (area in m2)

heat-transfercoefficient

0.852 kW/K ·m2


LMTD of (inlet and outlet temperature diff erences)


coolers (area in m2)heat-transfercoefficient

0.852 kW/K ·m2


LMTD of (inlet or outlet temperature −315 K)


energy costHP steam $9.88/GJ (41 barg, 254 ° C)MP steam $8.22/GJ (10 barg, 184 ° C)LP steam $7.78/GJ (5 barg, 160 ° C)cooling water $0.354/GJ

electricity $16.9/GJTAC (capital cost/payback period) + energy costpayback period 3 years

Figure 2. Optimal design owsheet for a regular column at 1 atm.

Table 3. Itemized TAC Terms of Regular Column and Multieff ect Systems

multieff ect distillation(LSR)

multieff ect distillation(LSF)

con gurations

regularcolumn(1 atm) C1 C2 C1 C2

installed capital costfor column (1000$)

307.826 100.238 181.227 63.785 182.679

installed capital costfor reboiler (1000$)

64.223 62.092 123.695 106.203 49.814

installed capital costfor condenser(1000 $)

17.915 12.403 11.552

steam cost (1000$/y)

154.689 97.306 91.388

cooling water cost(1000 $/y)

5.047 2.866 2.570

total reboiler duty (kw)

543.6 342.0 321.2

total steam cost(1000 $/y)

154.689 97.306 91.388

TAC (1000 $/y) 289.724 260.058 231.969

Industrial & Engineering Chemistry Research Article

DOI: 10.1021/ie503922f Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

C

http://dx.doi.org/10.1021/ie503922fhttp://dx.doi.org/10.1021/ie503922f


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reverse direction to process ows. They abbreviated this designcon guration as light-split-reverse (LSR) with feed entering intoa LP column and then a bottom stream of LP column feedinginto a second HP column.

The second better design conguration with heat integrationin the forward direction to process ows is called light-split-forward (LSF). It is easily understandable that fresh feed isentered into a HP column and the bottom stream of this columnis then entered into the second LP column. The third betterdesign owsheet, perhaps most commonly used in industry, isthe one that split the feed into two feed streams and then go into both HP and LP columns. The industry often called the HPcolumn “ lower column” and the LP column “ upper column”

because the condenser of the HP column and the reboiler of theLP column are heat integrated. The conceptual design owsheetof these three multieff ect heat-integrated distillation columnsystems areshownin Figure3. We will develop designowsheetsfor the separation of 1,2-PDO and EG.

One of the plus sides of the LSR or LSF design congurationsis that theseparation eff ort of the rst column may not need tobe

that great because the bottom composition of this column is notstringently specied to be pure component. However, the totalnumber of stages of the second column is expected to be quitelarge because both outlet streams are nal products at highpurity. As for the feed-split (FS) design conguration, it isexpected that this design conguration is not suitable for theclose-boiling system because both HP and LP columns shouldhave stringent purity specications in the distillate and bottomstreams.

In this study, the top pressure of the LP column is set to beatmospheric, thesame as theregular column. Thetoppressure of the HP column is varied so that the temperature diff erence between the condenser temperature of the HP column to thereboiler temperature of the LP column cannot be too small. Weused the guideline in Chiang and Luyben24 to set thetemperature diff erence to be at about 20 K. The reason to setthe top pressure of the LP column at atmospheric pressure is because we want to avoid having tangent pinch behavior at thepure 1,2-PDO end when operating the low-pressure column at vacuum.

Figure 3. Conceptual design of three multieff ect distillation congurations.



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In the simulation study of the LSR design conguration inFigure 4 the bottom composition of the LP column is set so thatthe condenser duty of the HP column can be matched with thereboiler duty of the LP column. This is the case for obtaining themaximum energy savings with full heat integration. As for thetotal number of stages and feed location for both the HP and the

LP columns, they are determined by the short-cut distillationmodule in Aspen Plus (DSTWU). The reux ratio was set to be1.1 times the minimum reux ratio in the short-cut distillationcalculations. Of course, the total number of stages and feedlocationscan be designvariables to minimize total annualcost forfurther investigation. However, a time-consuming iterativesequential optimization procedure will be needed to obtain theoptimal design owsheet.

Note that the tray rating feature in the Aspen Plus RADFRACmodule was used to realistically calculate the pressure drop of both columns. It is observed that the bottom composition of the

rst LP column needs to be at 78.88 mol % 1,2-PDO so that thereboiler duty of this column is matched with the condenser duty of the second HP column. Using this heat-integrated design, the

overall reboiler duty can be saved from 543.6 kW of a regulacolumn (at 1 atm) to 342.0 kW (reboiler duty of HP column).Note that the operating pressure of the second HPcolumn has to be at 2.6 atm in order to meet the temperature diff erencerequirement for heat transfer. At this high pressure, the bottompressure of the HP column is at 3.39 atm with a resulting bottom

temperature at 516.27K.Althoughhigh-pressure steamat 527.15K can still beused toprovide the heat, the heat transfer areaof thereboiler will be larger, resulting in more capital cost of thireboiler. Table 3 also summarizes the itemized TAC terms forthis LSR design owsheet.

Figure 5 shows the LSF design owsheet of this separationsystem. It is observed that thebottom compositionof therstHPcolumn needs to be at 70.15 mol % 1,2-PDO so that thecondenser duty of this column is matched with the reboiler dutyof the second LP column. Using this heat-integrated design, theoverall reboiler duty can be saved from 543.6 kW of a regulacolumn (at 1 atm) to 321.2 kW. Note that the operating pressureof the rst HP column needs to be quite high (at 4.2 atm). This ismainly because the bottom composition of the LP column is at

Figure 4. Design owsheet of multieff ect light-split-reverse (LSR) conguration.

Figure 5. Design owsheet of multieff ect light-split-forward (LSF) conguration.



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pure EG, which gives a higher bottom temperature as comparedto thebottomtemperature of theLP column in theprevious LSR design conguration. Because of a high top pressure of the HPcolumn (although much fewer total number of stages), the bottom temperature of this HP column is still quite high at517.82 K. Table 3 also summarizes the itemized TAC terms forthis LSF design owsheet.

In the above discussion of energy consumption for the LSR con guration (Figure 4) or LSF conguration (Figure 5),reboiler duty is used to represent the main energy consumption.In practical applications, a higher enough condenser temperature(as in previous cases in Figures 4 and 5) can be utilized togenerate low-pressure steam. For simplicity of analysis purposes, we did not consider this more complex scenario in this study.

Figure 6 shows the FS design owsheet of this separationsystem. Thefeed split ratio is adjustedso that the condenser duty of the HP column is matched with the reboiler duty of the LPcolumn. It is observed that savings of the reboiler duty is thehighest (from 543.6 kW of regular column to 283.8 kW).However, many stages are required for both columns. The othermain problem is that high-pressure steam cannot be used in thereboiler of HP column because the bottom temperature at534.99 K already exceeds the temperature of high-pressuresteam. The reason for the high temperature is the combinatory eff ects of pure EG composition at column bottoms of HPcolumn, and also many stages are required for this column. Sincethis FS designconguration is obviouslyinferior in terms of TAC

to the other two design congurations, the itemized TAC termsare not provided in Table 3. A nal note of this study is that the result for this close-boiling

system is qualitatively diff erent than the methanol/water systemstudied in Chiang and Luyben.25 In that paper, the LSR designcon guration gave the largest reboiler duty reduction ascompared to LSF or FS design congurations.

4. HETEROGENEOUS AZEOTROPIC DISTILLATIONSYSTEM

A U.S. Patent21 listed potential entrainers for the separation of 1,2-PDO and EG via a heterogeneous azeotropic distillationsystem. However, the important features of this kind of design to

aid the separation were not mentioned. We will use one of thesuggested entrainers, decane, as an example to explain thefeasibility of this separation method for a close-boiling mixtureFrom the RCM and LLE (at 40 °C) of this ternary system inFigure 7 it is noticed that two new azeotropes are formed with aminimum-boiling azeotrope of 1,2-PDO/decane at 158.00 ° C ataround 65 mol %. More importantly, this azeotrope isheterogeneous and can naturally be separated into two liquidphases in a decanter.

A conceptual design with material balance lines of thisheterogeneous azeotropic distillation system is also shown inFigure 7. The fresh feed is entered into a heterogeneousazeotropic distillation column. With the aid of a light-phaserecycle stream (mostly decane), a feasible separation of the top

Figure 6. Design owsheet of multieff ect feed-split (FS) conguration.



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vapor stream at (or near) the minimum-boiling azeotrope and bottom product to be pure EG at the highest boiling-pointtemperature of the left distillation region can be achieved. Figure7shows themerging of two inlet streams (freshfeed, F, andlight-phase recycle, L1) for the heterogeneous column and thensplitting into two outlet streams (top vapor, V, and the bottoms,B) representing the material balance lines of the dashed-lineenvelope at the design owsheet. This top vapor stream, after

condensation, can naturally be separated into two liquid phasesThe light phase (L1) is designed to be recycled back to theheterogeneous azeotropic column to continuously provideentrainer into this column. The heavy phase (L2) is designedto be drawn out of this column. Since from Figure 7 thecomposition of the heavy phase is not pure enough to meet 1,2-PDO product specication, a small product column is needed tofurther purify 1,2-PDO product. The distillate of this small

Figure 7. Conceptual design owsheet and material balance lines of the heterogeneous azeotropic distillation system.



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product column can be designed to be recycled back to thedecanter.

For the limiting case with top vapor right at the minimum- boiling azeotrope and the bottom right at pure EG, theow ratio(18.7:1.0) between the top vapor stream (V) and the bottomstream(B) can easily be estimated by the lever rule. It is observedthat a very large top vapor ow rate is needed for this separationsystem. This means that a very large recycle ow rate is expectedto be back to the heterogeneous azeotropic column. This willresult in large reboiler duty toboil up the large liquid traffic insidethis column.

Another serious drawback of this design can be observed fromthe paths of residue curves. It is observed that the residue curvesare all moved toward the EG and 1,2-PDO edge rst and thenmove toward pure EG. This means that the close-boilingproblem at the pure EG end cannot be circumvented viaheterogeneous azeotropic distillation.

One other drawback can also be predicted by the ternary diagram in Figure 7. In a common heterogeneous azeotropiccolumn with moderate total stages, most likely the top vaporcomposition cannot be pushed down exactly to point V (at alimiting case with minimum-boiling temperature). This meanssome EG will remain in this top vapor stream. After condensingin a decanter, the heavy phase composition will not exactly be atL2 but with some EG. This will create a problem for the nextproduct column to obtain 1,2-PDO at high-purity specication because pure 1,2-PDO is not the stable node in the leftdistillation region of the ternary diagram.

All other suggested entrainers listed in this U.S. Patent21 havethe same problems as demonstrated by the decane system. Thisexample shows that with enough process understanding of theprinciple of separation method an unsuitable (although feasible)alternative design can be screened out without doing any furtherinvestigation via computer simulation. We cannot nd any otherentrainers in the open literature which can use the separationprinciple via heterogeneous azeotropic distillation for this close- boiling system without having the problems elaborated above.

5. EXTRACTIVE DISTILLATION SYSTEM Another separation method for a close-boiling system is to add aheavy entrainer into this close-boiling systemin order to enhancethe relative volatility between 1,2-PDO and EG. In this way, 1,2PDO can be drawn out of an extractive distillation column asdistillate while entrainer and EG can go out from the bottomstream. With another entrainer recovery column, the heavy entrainer can be puried at bottoms and then recycle back to theextractive distillation column. The EG product can go out fromthe distillate of this entrainer recovery column. We willinvestigate the energy-saving potential of this two-columnextractive distillation system.

5.1. Two-Colum n Extractive Distillation System.Doherty and Malone26 proposed to use homologues, polarity,and hydrogen-bonding tendencies for crude screening of candidate entrainers for extractive distillation. However, we were unable to use their solvent selection procedure to locate thespeci c candidates for this close-boiling system. The paper by Gmehling and Mo ̈llman27 listed commonly used heavy entrainers in industry for the purpose of extractive distillationFigure 8 displays the capability of relative volatility enhancemenof some limited heavy entrainers with a boiling point higher thanthat of EG. In the gure, the x axis is the ratio between theentrainer ow rate to the fresh feed ow rate. The y axis is the

relative volatility between 1,2-PDO and EG with addition of theentrainer into the system. The data can easily be collected by jusperforming a ash calculation using FLASH2 module in AspenPlus. The operating pressure is set at 0.1 atm for the main reasonfor still being able to use high-pressure steam at reboiler.

With no entrainer added into the system, the original relative volatility is close to 1.0 (at 1.231). By adding an equalowrate of triethylene glycol (TEG) or Sulfolane into the system, therelative volatility can greatly be enhanced to over 3.0. Notice thaanother entrainer (6-caprolactam) will reversely enhance therelative volatility of EG over 1,2-PDO. This means that at anextractive distillation column EG will be the distillate product busing 6-caprolactam as entrainer. From the gure, the enhance-ment of the relative volatility via 6-caprolactam is not as great a

Figure 8. Enhancement of relative volatility at 0.1 atm for 1,2-PDO to EG by adding several heavy entrainers.



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the other two entrainers; thus, this entrainer will not be used inthe design study.

Another observation is that there is a reversal of trend for therelative volatility enhancement of 1,2-PDO over EG at anentrainer/feed ratio over 1.5when using TEGas entrainer. Thus,the optimal entrainer/feed ratio is expected to be less than 1.5 because adding more entrainer into the system does not have any positive eff ect. A further conjecture can also be made with anoptimal ratio to be less than 1.0 because no signicant increasingof the relative volatility can be observed when this ratio isincreased from 1.0 to 1.5.

Of course, the separation of 1,2-PDO and the entrainer should be easy so that only few stages are required in the rectifying

section of the extractive distillation column. Also, the separationof EG and the entrainer should be easy as well so that only few stages are required in the entrainer recovery column. Figure 9 of Txy plots justies the above requirements for using either TEGor Sulfolane as entrainer.

A design owsheet for the two-column extractive distillationsystem by using TEG as entrainer will be established next. Thereare quite a few design variables for this system including thefollowing: TEG ow rate (or feed ratio); total stages (NT1) andfeed locations of fresh feed (NFF) and entrainer (NFE) of theextractive distillation column; total stages (NT2) and feedlocation (NF2) of the entrainer recovery column. The twocolumn pressures are set at 0.1 atm so that high-pressure steamcan still be used in the two reboilers. The outlet temperature of

the solvent cooler is set to be 10 K below the distillatetemperature of the extractive distillation column following thesuggestion by Doherty and Malone.20

In all simulation runs, top and bottom compositions of theextractive distillation column are set at 99.9 and 0.01 mol % 1,2PDO, respectively, and for the entrainer recovery column, topand bottom compositions are set at 99.9 mol % EG and 99.99mol % TEG, respectively. A very small entrainer makeup ow isdetermined tobalance theentrainer loss from two distillate outletstreams.

We did not use a time-consuming iterative sequentialoptimization procedure to obtain the optimal design owsheet which minimizes TAC. Instead, we just want to demonstrate the

superiority of this design for the purpose of economics. FromFigure9 it appears the separation of EG and TEG should be easy.Thus, NT2 = 10 and NF2 = 5 was arbitrarily assumed for theentrainer recovery column. Also, for simplicity reasons, the totanumber of stages for the extractive distillation column was set to be 50 stages (still much less than the previous cases). With theabove assumptions, there areonly three designvariables left to bedetermined. They are TEG ow rate and feed locations of thefresh feed (NFF) and the entrainer (NFE).

Figure 10 shows the results of many simulation runs of totalreboiler duty vs NFF andNFE when entrainerowratewassetat ve diff erent values. Figure 11 summarizes the runs which

minimized total reboiler duty for each entrainer ow rate. Withincreasing entrainer ow rate, Qr1 is decreased rst but then

Figure 9. Txy plots of 1,2-PDO-TEG and EG-TEG or 1,2-PDO-sulfolane and EG-sulfolane at 0.1 atm.



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increased when this ow rate is over 5.0 kmol/h. This is mainly due to the reversal phenomena in Figure 8. Putting moreentrainer (feed ratio over 0.5) into the extractive distillationcolumn, thenegative eff ect of the overall feed rate increase is over

the positive eff ect of incremental relative volatility enhancement. As for Qr2, it is increased monotonically with the increase oentrainer owrate. It is found that the optimal entrainerow rateis at 4.5 kmol/h. This gives the entrainer/feed ratio at only 0.45.

Figure 10. Total reboiler duty vs NF and NE for ve FE for extractive distillation system.

Figure 11. Summarized plot of total and individual reboiler duties vs TEG ow rate.



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The optimal feed locations are at NFF = 20 and NFE = 4. Notethat this optimized case is for xing of NT1 = 50, NT2 = 10, andNF2= 5. For other valuesof these three design variables the totalreboiler duty can possibly be even lower. Alternatively, theseadditional three design variables can be determined by minimizing the TAC of the extractive distillation system.

Figure 12 shows the design owsheet of the extractivedistillation system. In comparison with the original base case inFigure 2 , the reboiler duty is reduced from 543.6 to 335.6 kW (asigni cant 38.3% reduction). The TAC calculated by usingformulas in Table 2 is also reduced from $2.897 × 105 to $2.011× 105 (30.6% reduction). The individual terms in the TACcalculationsarelisted in Table4. Note that tobemore accurate inthe TAC calculation, the costs associated with the vacuumsystem to generate operating pressure at 0.1 atm should also beincluded. However, by following the procedure outlined onpages 589−590 of Seider et al.’ s book,28 the annualized capitalcost of the liquid-ring vacuum pump and also the annualelectricity cost of this vacuum system can be neglected incomparison with other main terms in the TAC calculation.

Figure 13 shows the composition proles of the extractivedistillation column and entrainer recovery column. It is observed

that TEG is quickly depleted toward the column top; thus, veryfew stages are needed in the rectifying section of the extractivdistillation column. However, the depletion rate of 1,2-PDOtoward the column bottoms is relatively much slower. Thus,many more stages are needed in the stripping section of thiscolumn. This should be a unique characteristic of a close-boilingsystem because it is still difficult to separate 1,2-PDO and EGeven with the presence of entrainer in the stripping section. Another observation for the liquid composition prole of theextractive distillationcolumn is that a remixing eff ect is present atthe stripping section of this column. EG composition wentthrough a maximum inside the column and then dropped at thecolumn bottom.

The same study as above but just changing the entrainer fromTEG to Sulfolane obtains another design owsheet as in Figure14. The optimal entrainer ow rate is at 5 kmol/h. This gives anentrainer/feed ratio that is greater than the TEG system. Theoptimal feed locations are at NFF = 29 and NFE = 14. Note thatNT1, NT2, and NF2 are still xed at 50, 10, and 5, respectively.The larger rectifying section can be predicted by observing theTxy plot of 1,2-PDO and Sulfolane at pure 1,2-PDO end inFigure 9. A tangent pinch behavior is exhibited there. The

Figure 12. Design owsheet of extractive distillation system by using TEG as entrainer.

Table 4. Itemized TAC Terms of Various Extractive Distillation Systems

extractive distillation(TEG)

extractive distillation(SUL)

extractive distillation(TEG with FEHE)

extractive distillation(TEG with thermally

coupled)

con gurations C1 C2 C1 C2 C1 C2 C1 C2installed capital cost for column (1000, $) 187.293 26.433 201.068 27.948 189.249 21.685 187.640 43.3installed capital cost for reboiler (1000, $) 41.752 14.069 32.811 17.313 36.599 14.182 277.09installed capital cost for condenser (1000, $) 15.406 5.744 16.780 7.842 15.605 5.724 15.405 2.6installed capital cost for entrainer cooler (1000, $) 5.998 4.421 1.113 6.746installed capital cost for FEHE (1000, $) 20.154steam cost (1000, $/year) 80.592 14.908 84.006 24.401 68.200 15.052 91.367cooling water cost (1000, $/year) 2.292 0.548 2.614 0.885 2.339 0.545 2.293 0.220entrainer cooler cooling water cost (1000, $/year) 0.507 0.311 0.024 0.668

entrainer loss (1000, $/year) 3.363 14.420 3.363 3.363total reboiler duty (kW) 335.6 381.0 292.6 321.1total steam cost (1000, $/year) 95.501 108.407 83.252 91.367TAC (1000, $/year) 201.110 229.365 190.960 275.526



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individual terms in theTAC calculations arealso listedin Table4.Comparing with the TEG system, the total reboiler duty isincreased from 335.6 (Figure 12) to 381.0 kW (Figure 14). TheTAC is also increased from from $2.011 × 105 to $2.294 × 105.This demonstrates that TEG is a more eff ective entrainer thanSulfolane for this separation system.

5.2. Extractive Distillation System with Feed-E ffl uent-Heat-Exchanger. From the design owsheet in Figure 12 it isobserved that a solvent cooler is needed to drop down thetemperature of this stream from 493.4 to 388.6 K. With the

temperature of the fresh feed at 384 K, it is possible to add anextra feed-effluent-heat-exchanger (FEHE) to preserve theenergy of this high-temperature solvent stream. A simulation with this extra FEHE was conducted, and the result is shown inFigure 15. Note that the overall reboiler duty can further bereduced from 335.6 kW in the original two-column system to292.6 kW (a further 12.8% reduction). A quick calculation canalso been conducted tond the return of investment of this extraFEHE via savings of the steam cost to be 1.65 year. Theindividual terms in the TAC calculations for this simple

Figure 13. Liquid composition proles of the extractive distillation system using TEG as entrainer.

Figure 14. Extractive distillation design owsheet using Sulfolane as entrainer.



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improved design are also listed in Table 4. It is noted that TACcan further be reduced from $2.011 × 105 to $1.910 × 105.

5.3. Thermally Coupled Extractive Distillation System.Since from Figure 13 there is a remixing eff ect at the strippingsection of the extractive distillation column, another potentialretro t design can be made to thermally couple the two columnsin order to reduce this remixing eff ect and save energy. The way to thermally couple the two columns in Figure 12 is to combinetwo reboilers into one reboiler. A vapor sidedraw should bedesigned to draw-off from the second column to provide vaportraffic in the extractive distillation column. Since this alternativedesign is assumed to be a retrot from an existing owsheet as inFigure 12 , the total stages of the two columns are assumed toremain the same.

Figure 16 displays a thermally coupled system converted fromthe original extractive distillation system as in Figure 12. Notethat the rst column is now having 49 stages because the reboileris not counted for this column. An important change is to raisethe top pressure of the second column from 0.1 to 0.35 atm. Thereason for this change is to provide enough pressure drop for the vapor sidedraw so that a compressor can be avoided. Recall thathe tray rating feature in the Aspen Plus RADFRAC module waused to realistically calculate thepressure drop insidethecolumn.Thus, the top pressure change of the second column is necessaryso that the pressure at thefth stage of the second column can behigher than the pressure at the 49th stage of the rst column forthe vapor to ow.

In this design owsheet, the ow rate of vapor sidedraw is afree operating variable not used to maintain the composition of

Figure 15. Proposed design owsheet of the extractive distillation system with FEHE.

Figure 16. Design owsheet of the thermally coupled extractive distillation system with TEG as entrainer.



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any stream. This ow rate of 16 kmol/h is determined by minimizing the reboiler duty. Below this value, two productpurity specications (both 1,2-PDO and EG at 99.9 mol %)cannot be met. The liquid composition prole of this thermally

coupled system is shown in Figure 17. It canbe observed that thereduction of remixing eff ect from the original system is achieved.From Figure 16 it is observed that the overall reboiler duty can

be reduced from 335.6 kW in Figure12to321.1 kWin thisgure.Since in Figure 12 both reboilers need high-pressure steam,reboiler duty savings for the thermally coupled system candirectly become savings in the steam cost. The main drawback of this design owsheet can be observed from the bottomtemperature of the second column. This high temperature of 524.9 K (due to the required high top pressure of this column)gives a very small temperature diff erence to that of the high-pressure steam. The individual terms in the TAC calculations forthis thermallycoupled systemarealso listedin Table4.Itisnotedthat although steam cost can be saved by 4.3%, however, theextremely large capital cost of the reboiler is needed to make theTAC much larger than the original extractive distillation system.

The problem of a very small temperature diff erence for heattransfer at the reboiler can be lessened by lowering the toppressure of both columns. Also, more reduction of the reboilerduty may be possible by changing the two feed locations (NFEand NFF). However, from the study by Wu et al.,29 the controlperformance of this thermally coupled system is predicted to behampered because of combining tworeboilers, resulting in losingan important control degree of freedom. We did not pursuending a better thermally coupled system any further.

6. CONCLUSIONThis paper investigated alternative design methods of anindustrially relevant problem of separating a close-boilingmixture of 1,2-PDO and EG. The new ndings of this study are summarized below and should be extendable to other close- boiling systems.

(1) The best multieff ect distillation design for this separationsystem is a light-split-forward (LSF) conguration with40.9% reduction of the reboiler duty. This designcon guration is preferable compared to a light-split-reverse (LSR) conguration. Both congurations requirefewer total stages for the rst column as compared to thatof the second column.

(2) Another industrial commonly used feed-split (FS)con guration is not suitable for the close-boiling system because large total stages are still required for bothcolumns in the multieff ect distillation system. The high-pressure steam cannot be used at the reboiler of a high-pressure column.

(3) No suitable entrainer can be found for using heteroge-neous azeotropic distillation design. The candidateentrainers suggested by a U.S. Patent21 can be used in afeasible design owsheet. However, the internal recircu-lation rate in this separation system is very high to makethis design uneconomical. This inferior design can bepredictedby RCMs, LLE, andmaterial balance lines of thisseparation system.

(4) Triethylene glycol is a suitable heavy entrainer for thisseparation because of largely enhancing the relative volatility of 1,2-PDO to EG. However, the number of stages in the stripping section of the extractive distillation

Figure 17. Liquid composition proles of the thermally coupled system with TEG as entrainer.



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column still need to be quite large for a close-boilingsystem.

(5) The extractive distillation system for this separation usingtriethylene glycol as entrainer can reduce steam cost by 38.3% and also reduce TAC by 30.6% as compared to theregular column system. The TAC is also lower than the best multieff ect distillation design because of much fewerrequired total stages for the two columns in this extractivedistillation system.

(6) Further energysavings for theextractive distillationsystemcan beachieved byadding a heatexchanger to preserve theheat from a high-boiling entrainer stream to the fresh feedstream. The return of investment for this additional heatexchanger viasavings of thesteamcost was calculated to be1.65 year. The steam cost can be reduced by 46.2%, andTAC can be reduced by 34.1% as compared to the regularcolumnsystem. This is thenal proposeddesignowsheetfor this close-boiling separation system.

(7) Another retrot design by thermally coupling twocolumns in the extractive distillation system can beconducted for the purpose of further saving of reboilerduty. The top pressure of the second column needs to beadjusted to avoid using an additional compressor. Thereduction of reboiler duty is not as great to justify implementing this more complex design with one lessimportant control degree of freedom.

One additional remark for the cases when not introducingforeign component into the system is critically important,multieff ect distillation with LSF conguration is preferable forthis separation system to reduce reboiler duty. Further energy savings of this design conguration can be achieved by adding aheat exchanger to preserve the heat from the distillate product of the HP column to the fresh feed stream.

A nal remark of this study is that this study was based on a

NRTL thermodynamic model with the most binary parametersestimated by the UNIFAC group contribution method. Furtherphase equilibrium experiments of the ternary system of 1,2-PDO/EG/TEG should be conducted to have more condencein the proposed design owsheet developed in this study.

■ AUTHOR INFORMATIONCorresponding Author* Tel: +886-3-3366-3063. Fax: +886-2-2362-3040. E-mail:[email protected] .NotesThe authors declare no competing nancial interest.

■ ACKNOWLEDGMENTSResearch funding from the Ministry of Science and Technology of the R.O.C. under grant no. MOST 103-2221-E-002-257 isgreatly appreciated.

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Energy Saving Designs for Separation of a Close Boiling 1 2 Propanediol and Ethylene Glycol Mixture

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