Design and Control of an Extractive Distillation System for Benzene/Acetonitrile Separation Using Dimethyl Sulfoxide as an EntrainerShengkai Yang,*, Yujie Wang,*, Guangyue Bai,*, and Yong Zhu

School of Chemistry and Chemical Engineering, Henan Institute of Science and Technology, Xinxiang, Henan, 453003, ChinaSchool of Chemistry and Chemical Engineering, Key Laboratory of Green Chemical Media and Reactions, Ministry of Education,Henan Normal University, Xinxiang, Henan, 453003, China

ABSTRACT: A method for benzene/acetonitrile separation using extractive distillation is presented, and dimethyl sulfoxide(DMSO) is selected as the entrainer. Steady state and dynamic simulations for this process are implemented by commercialsimulators (Aspen Plus and Aspen Dynamics). On the basis of global economic optimization, a design with optimized operationconditions for this process is developed. Two single temperature control structures with xed reux ratio or xed reux-to-feedratio and a dual-end temperature control structure for entrainer recovery column are presented to evaluate the dynamicperformance with feed ow rate and feed composition disturbances, and the last structure is quite eective.

1. INTRODUCTION

2-Picoline is an important pharmaceutical intermediate and araw organic chemical material. Catalyzed by organic cobalt, ethyneand acetonitrile are used to produce 2-picoline presented byQin et al.,1 and the synthesis process achieved wide-plant prod-uction scale in 2006. In the synthesis method, organic cobaltis dissolved in pure benzene solution as catalytic agent, andsimultaneously, benzene is also one of the byproducts. After2-picoline and some other byproducts are separated from reactantmixture, there are many benzene/acetonitrile mixtures left, and itis very necessary to separate them for recycle use.According to the ref 2, benzene (normal bp 80.09 C) and

acetonitrile (normal bp 81.60 C) at 1 atm form a minimum-boiling azeotrope with azeotropic temperature at 73 C andazeotropic composition at 52.9 mol % benzene. Figure 1A givethe Txy curve for the benzene/acetonitrile system.In this paper, we attempt to use an extractive distillation

process to separate benzene/acetonitrile. Among some suitablecandidate solvents, dimethyl sulfoxide (DMSO) is elected asentrainer according to the values of relative volatility. Thethermodynamics model Wilson is used in the simulation cal-culation. Using DMSO as entrainer, an optimized design for theextractive distillation process is developed using total annualcosts (TAC) as an objective function from many alternatives.Finally, appropriate control structures were introduced, andtheir dynamic performances were evaluated.

2. STEADY STATE DESIGN

2.1. Entrainer Selection. Since the entrainer is a key factorin extractive distillation, more attention should be paid toits election. A criterion for entrainer selection is through thecomparison of relative volatilities in the presence of dierententrainers.3 The higher the relative volatility, the easier theseparation. The relative volatility is dened as

=

K

Ki js i s

j s,

,

, (1)

where i,js is the relative volatility of component i and compo-

nent j in the presence of entrainer s, Ki,s is the innite dilution

K value for a trace of species i in the entrainer, and Ki,s is the

innite dilution K value for a trace of species j in the entrainer.Of all candidate solvents used for the separation of the

benzene and acetonitrile azeotrope mixture, the followingthree entrainers are studied in this paper: dimethyl sulfoxide(DMSO), phenol (PhOH), and sulfolane (SFO). Their Kvalues at innite dilution and the relative volatility are listed inTable 1. Because DMSO does not lead into further azeotropein the system and shows a higher relative volatility (2.139),it is chosen as entrainer in the simulation. This large relativevolatility is helpful to gain an economical separation sequence.Figure 1BC gives Txy diagrams for benzene/DMSO and

acetonitrile/DMSO mixtures, respectively. The former plays animportant role in the top of the extractive column, and thelatter applies in the solvent-recovery column.2.2. Thermodynamic Model Used in the Simulation.

According to the refs 2 and 4, the best t was obtained by theWilson model for the binary isobaric VLE data composed ofthe benzene/acetonitrile and acetonitrile/DMSO system. Noisobaric VLE experimental data were found for the benzene/DMSO system. The predicted azeotropic temperature andazeotropic composition for benzene/acetonitrile at 0.43 atmusing the Wilson model are 73.46 C and 68.64 wt %(53.50 mol %), respectively, which were in very good agree-ment with experimental data.In this simulation, the Wilson model was used to describe

the nonideality of the liquid phase while the vapor phase wasassumed to be ideal. The Wilson model parameters of thethree pairs were taken from Aspen Plus, and all other physicalproperty model parameters were taken from the built-in valuesin Aspen Plus.

Received: March 15, 2013Revised: July 20, 2013Accepted: August 14, 2013

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2.3. Residue Curve Map (RCM) for the System. TheRCM of ternary system can be used to design and distinguishbetween feasible and infeasible sequences and can be mappedby Aspen Plus. Using the Wilson model, the RCM of thebenzene/acetonitrile/DMSO ternary system is drawn andshown in Figure 2. It can be seen that the benzene/acetonitrileazeotrope is the unstable node, DMSO is the stable node, andboth benzene and acetonitrile are the saddles. No distillationboundary exists to the system, which is an ideal situation forselection of an extractive distillation process. Material balancelines on behalf of the extractive distillation process were alsodrawn in Figure 2. It is noticed that M, denoting the mixture ofthe raw material and entainer, can be separated into D1 and B1on extractive distillation column, and B1 can be separated intoD2 and B2 on entainer recovery column. This means that thefeed F may be separated into relatively pure two products withthe aid of the entrainer. To balance the tiny loss of entrainer inboth D1 and D2 streams, a small makeup stream of DMSOshould be added (not displayed in Figure 2).2.4. Process Design and Economic Analysis. 2.4.1. Proc-

ess Design. In this paper, the raw material used is from severallocal plants and mainly composed of benzene and acetonitrilewith a little 2-picoline. After pretreatment, 2-picoline was re-moved from the raw material completely. The composition wasanalyzed by gas chromatography. The extractive distillationprocess was simulated with the following data: the feed was amixture made up of 65.5 wt % of benzene and 34.5 wt % of

Figure 1. Txy diagram for (A) benzene/acetonitrile or (B) benzene/DMSO at 0.43 atm and (C) acetonitrile/DMSO at 0.33 atm.

Figure 2. Residue curve map at 0.43 atm for the benzene/acetonitrile/DMSO system.

Table 1. Results of the Entrainer Selection for an ExtractiveDistillation System

benzene(1)/acetonitrile(2) using the Wilson model

solvent Tb/K K1,s K2,s

1,2

DMSO 464.00 27.12 12.68 2.139phenol 454.99 19.50 10.88 1.792sulfolane 560.45 79.40 53.17 1.493

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Figure 3. Eect of RR1 and entrainer ow rate S in extractive column (NT = 47) on (A) benzene and (B) acetonitrile.

Figure 4. Optimal process owsheet for extractive distillation.

Figure 5. Extractive distillation column C1: (A) temperature prole;(B) composition prole.

Figure 6. Entrainer recovery column C2: (A) temperature prole;(B) composition prole.

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acetonitrile, with a ow rate of 3500 kg per hour with 7200working hours annually. Aspen Plus and Aspen Dynamicsare used to do the steady state and dynamic simulations. TheWilson activity model is chosen to predict the vapor liquidequilibrium in the simulator. To ensure the products areavailable for recycling, the two product specications are set tobe as follows: the acetonitrile impurity in the benzene productis not more than 1 wt %, and the acetonitrile product has apurity of 99.9 wt %. Here, Aspen notation of numbering stagesfrom the top, with stage 1 being the ux drum and the last stagebeing the rebolier, is adopted.For an extractive distillation column (denoted as C1), when

the total stages, operating pressure, and feed location are xed,there are three design degrees of freedom: reux ratio (RR1),

entrainer ow rate (FE), and reboiler heat duty (QR1). Sincealmost all the benzene that is present in the C1 bottom product

Figure 7. Basic control structure (S1) with xed reux ratio for C1 and C2 columns.

Figure 8. Improved control structure (S2) with xed reux-to-feed ratios for C1 and C2 columns.

Table 2. Temperature Controllers Tuning Parameters forthe Basic Control Structure

parameters TC1 TC2 TCHX

controlled variable T1,40 T2,10 Trecyclemanipulated variable QR1 QR2 QHXtransmitter range (K) 273443 273556 273370controller output range (GJ/h) 07.71 03.37 4.040ultimate gain 8.1037 2.9728 0.5200ultimate period (min) 6.0 4.8 2.4gain, Kc 2.5324 0.9290 0.1625integral time, (min) 13.2 10.56 5.28

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will go to the top of the entrainer recovery column (denotedas C2), to achieve the desired acetonitrile product, the massow rate of benzene in the bottom product of the C1 is held at1.2 kg/h (calculated from the material balance) by manipulat-ing the reboiler heat duty QR1.

2.4.2. Economical Optimization. It is known that, forthe extractive distillation process, an increase in the entrainerow rate will reduce the heat duty of the extractive distillationcolumn, but the heat duty of the entrainer recovery column willincrease, and a larger column diameter is expected. Thus, atrade-o between the extractive column costs (include both thexed capital costs and operating costs) and entrainer recoverycolumn costs needs to be made. Thus, when the theoreticalplates were xed, there exists an optimal entrainer ow ratethat minimizes TAC (total annual costs, dened in Luybensbook5). When the entrainer ow rate is larger than the optimalvalue, the costs of the entrainer recovery column become moredominant and TAC increase.It is a convention to use the TAC as the objective function

to be minimized, which includes annualized capital costs andoperating costs. The capital costs include the column shell,

reboiler, and condenser, and a payback period of 3 years isassumed. Small items such as reux drums, pumps, valves, andpipes are usually not considered due to their much lower costs.The operating costs include the steam and cooling water, andonly the former was reckoned in this simulation owing to itsmuch higher cost. The heat transfer area for the condensers andreboilers is determined using the overall heat transfer coecientand a dierential temperature driving force. Here, the overallheat transfer coecients and the calculation formula for theabove equipment are all taken from Luybens book.5

The Tray Sizing function in Aspen Plus is employed to sizethe column vessels, and a sieve plate is selected. The DesignSpec/Vary function was used to satisfy the product composi-tion. The weir height of 0.025 m is adopted in order to main-tain the pressure drop required, while other parameters use thedefault values. The stages are counted from top to bottom withthe condenser as the rst stage and the reboiler as the last stagefor both columns.The pressure selection is very important; lower pressure

leads to easier separation, but the pressure value is restricted bythe adopted cooling water temperature, simultaneously, due to

Figure 9. Dynamic responses for basic control structure (S1) with xed reux ratio for columns C1 and C2: feed ow rate disturbances. (Solid linesare denoted as +20% ow rate, and dashed lines are denoted as 20% ow rate.)

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the lower material temperature along with its pressurereduction. As the adopted temperature of the cooling water isconstant at 308.15 K and the temperature dierence is constantat 13.9 K for the column top condenser and heat exchangeraccording to Luybens book,5 the C2 column top temperature is322.05 K (i.e., the sum of cooling water temperature andthe temperature dierence of heat transfer) and its pressure iscalculated from Antoine extended formulation, namely,0.33 atm. Also, as the same cooling water and temperaturedierence was adopted, the material ow from C2 bottom isalso cooled to 322.05 K by the heat exchange (denoted as HX),namely, the recycling entrainer ow temperature input to C1column. According to the suggestion of Knight and Doherty,6

the temperature dierence of recycling entrainer ow temper-ature and C1 column top should be 515 K, and the tempera-ture of C1 column top was xed at 327.37 K (here the above-mentioned temperature dierence being at 5.32 K). Thus, theC1 column top pressure is calculated from Antoine extendedformulation to be 0.43 atm.

In the optimization process, the six design variables, namely,the total stages of the extractive distillation column (N1), thefresh feed location (NF1), the entrainer feed location (NFE), thetotal stages of the solvent recycle column (N2), the feedlocation of solvent recycle column (NF2), and the solvent feedrate (FE) should be determined by minimizing the total TAC.The procedure to search the optimal value of each designvariable is summarized below:

(1) Guess the six design variables, N1, NF1, NFE, N2, NF2, and FE.(2) Run the simulation until a normal result without any

wrong was achieved.(3) Calculate the total TAC. If the new calculated TAC was

smaller than the minimum obtained from previousrounds, the new design value was adopted, andotherwise, the search process should be stopped.

(4) Change the design variable based on the new calculatedTAC value and go back to step 2.

All the six design variables should be optimized according tothe procedure above. After a round search, namely, all the six

Figure 10. Dynamic responses for basic control structure (S1) with xed reux ratio for columns C1 and C2: feed composition disturbances. (Solidlines are denoted as +0.05 benzene mass composition, and dashed lines are denoted as 0.05 benzene mass composition.)

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design variables optimized once, the next round search shouldbe done until all the six design variables were not changed at acertain round. In the search process, step size for the theoreticalplate number and feed position were xed at 1 stage and theFE feed rate was xed at 100 kg/h. The search method isintuitive and eective but rather time-consuming in a plantwideoptimization.According to the optimal results, it is found that the optimal

mass ow rate of the entrainer (FE) is 9000 kg/h, and theoptimal total number of stages is 47 for the extractive distilla-tion column and 12 for the entrainer recovery distillationcolumn, respectively. The best feed position NFE is at the 4thstage, NF1 at the 31st stage for the extractive distillation column,and NF2 at the 7th stage for the entainer recovery column. Theoptimal reux ratios for C1 and C2 are 1.138 and 0.467, respec-tively. The nal minimum value of TAC is 451 150$/y.Extractive distillation systems have the interesting feature of

two dierent values of reux ratio that yield the same separa-tion. Figure 3A,B manifests a plot of benzene purity or aceto-nitrile impurity from C1 column top ow versus reux ratio atxed solvent rate, respectively. From Figure 3A,B, we can conclude

that the reux ratio of 1.138 is indeed the minimum of the twopossible reux ratios.Figure 4 gives the nal optimal ow sheet for this system,

with detailed steam information, heat duties, equipment sizes,and operating conditions at the steady state design conditions.Figures 5 and 6 show the temperatures and composition pro-les of C1 and C2 columns for the ow sheet, respectively.There is a rapid rise in the temperature for stage 4 and a rapidfall for stage 31 in Figure 5A, at which the entrainer and freshfeed are fed. It is obvious that stage 40 displays a fairly steepslope for the temperature and benzene composition proleswith the extractive distillation column shown in Figure 5B.In Figure 6, two steep slopes are found near the 4th and 10thstages. The prole distinguishing features indicate the propertemperature control point for the two columns.

3. CONTROL SYSTEM DESIGN3.1. Basic and Improved Control Structure. Single

temperature control structure is the rst to be considered dueto its less expensive cost. For the extractive distillation process,basic control structure, namely, xed reux ratio for C1 and C2

Figure 11. Dynamic responses for improved control structure (S2) with xed reux-to-feed ratio for columns C1 and C2: feed ow rate disturbances.(Solid lines are denoted as +20% ow rate, and dashed lines are denoted as 20% ow rate.)

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column (denoted as S1), and its improved control structure,namely, xed reux-to-feed ratio (denoted as S2), are oftenadopted. Before starting the dynamic simulation, the plumblingsystem and major equipment sizes must be specied.The commonly used heuristic for reux drums and column bases

sizing is to provide 5 min of liquid holdup when half full. All thecontrol valves pressure drops are 3 bar with the valve half open atthe design ow rate. Then, the steady state owsheet is pressure-checked, and the Aspen Plus le is exported to Aspen Dynamics.The temperature at the 40th stage is adopted as the tem-

perature control point for the extractive column, and the tempera-ture at the 10th stage is used as the temperature control pointfor the extrainer recovery column. To the control schemes of S1and S2, the following control structure is proposed for theextractive distillation control system.

(1) Feed is ow-controlled (reverse acting).(2) Reux drum levels in both columns are held by manip-

ulating the ow of distillates (direct acting).(3) Base level in extractive distillation column is held by

manipulating the ow of the bottoms (direct acting).

(4) Base level in entrainer recovery distillation column isheld by manipulating the makeup DMSO ow rate(reverse acting).

(5) The total entrainer ow is in proportion to the feed ow.(6) The pressure in the two columns is controlled by

manipulating the heat removal rate in the condenser ofthe two columns (reverse acting).

(7) The reux ratios for C1 and C2 columns are xed for thebasic control structure (S1), and the reux-to-feed ratiosfor C1 and C2 columns are xed for the improved controlstructure (S2).

(8) Entrainer feed temperature is held by manipulating thecooler HX heat duty (reverse acting).

(9) The temperature for the 40th stage in the extractive distilla-tion column is controlled by manipulating the reboiler heatinput into the extractive distillation column (reverse acting).

(10) The temperatures in the 10th stage in the entrainer re-covery column are controlled by manipulating thereboiler heat input into the entrainer recovery column(reverse acting).

Figure 12. Dynamic responses for improved control structure (S2) with xed reux-to-feed ratio for columns C1 and C2: feed compositiondisturbances. (Solid lines are denoted as +0.05 benzene mass composition, and dashed lines are denoted as 0.05 benzene mass composition.)

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A noteworthy feature revealed here is that the base level inthe entrainer recovery distillation column is held by manipu-lating the makeup DMSO ow rate suggested by Grassi andLuyben.7,8 Figures 7 and 8 demonstrate the basic and improvedcontrol structures for this extractive distillation system, respectively.Proportional controllers are used for all liquid levels with

Kc = 2, and conventional PI controllers are used for all other

controllers. The proportional and integral (PI) settings ofthe top pressure control loops for both columns are setat Kc = 20 and I = 12 min. Three dead time elementsare inserted into the corresponding temperature controlloops with a dead-time of 1 min. Using the Tyreus-Luybentuning, relay-feedback tests are run on the temperaturecontrollers to determine ultimate gains and periods, and the

Figure 13. Dynamic responses for control structure (S3) with xed reux ratio for columns C1 and dual temperature control for column C2: feedow rate disturbances. (Solid lines are denoted as +20% ow rate, and dashed lines are denoted as 20% ow rate.)

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parameters are shown in Table 2 for S1 or S2 controlschemes.Now, the dynamic performance of the control structure is

evaluated by feed ow rate and composition disturbances.Figure 9 shows the dynamic responses of the basic controlstructure to positive and negative 20% step changes in feed owrate at t = 0.2 h. As you can see from Figure 9, benzene product

purity is held close its desired value, while acetonitrile productpurity is 0.9898 mass fraction at the new steady state in 8 h,having large deviation to its desired value. It is found that thetwo controlled tray temperatures were brought back to their setpoints and the reboiler duties of two columns reached newsteady state values in 5 h. Figure 10 shows the dynamicresponses of the basic control structure for positive and

Figure 14. Dynamic responses for control structure (S3) with xed reux ratio for columns C1 and dual temperature control for column C2: feedcomposition disturbances. (Solid lines are denoted as +0.05 benzene mass composition, and dashed lines are denoted as 0.05 benzene mass composition.)

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negative 0.05 benzene composition disturbances in the feed att = 0.2 h. From Figure 10, acetonitrile product purity is 0.9566,having large deviation to its desired value. Other controlvariables remaining, i.e., benzene product purity, the twocontrolled tray temperatures, and two columns reboiler duties,have almost similar responses as those in the feed ow ratedisturbance.The dynamic responses of improved control structure were

shown in Figure 11 to the feed ow rate positive and negative20% disturbance and in Figure 12 to the benzene compositionpositive and negative 0.05 disturbances in the feed at t = 0.2 h.The acetonitrile product purities are 0.9879 and 0.9902 to thesame feed ow rate and benzene composition disturbance inthe feed, respectively, and the benzene product purity has greatimprovement simultaneously (from 0.9882 to 0.9927 massfraction) to the negative 0.05 benzene composition disturbancein the feed than that in the basic control structure.3.2. Dual Temperature Control Structure. To the single

temperature control structure, the scheme of S2 has a better testresult to rate and component disturbance than S1, but to thedisturbances of positive 20% feed ow rate and positive 0.05feed benzene composition, deviations of the acetonitrileproduct purity from C2 top at the new steady state are stilltoo large in strict product requirements. To improve thecontrol performance, a dual temperature control structure(denoted as S3) was presented, in which xed reux ratio wasadopted for C1 column, and stage 4 and stage 10 werecontrolled by reux rate and by reboiler duty in the C2 column,respectively, along with the same other control strategymentioned above in single temperature control schemes. Thetest results were shown in Figures 13 and 14, from which theacetonitrile product purities have great improvement under thesame uctuation. The acetonitrile product purities are 0.9963 tothe feed ow disturbance and 0.9987 to the feed benzenecomposition disturbance, respectively. In actual production,which one to be adopted for S2 or S3, depends on the purity ofthe product requirements, or the trade-o between the addedvalue from improved purity of the two products and theincreased costs if S3 was adopted.

4. CONCLUSIONSDesign and control of an extractive distillation process forseparation of benzene/acetonitrile are investigated in our work.DMSO is chosen as a suitable entrainer by comparing the re-lative volatility. The Wilson model is used to calculate thethermodynamics properties. Using the total annual cost as theobjective function, the optimal design of the extractivedistillation process is presented. According to the simulationresults, it is found that the optimal mass ow rate of theentrainer (S) is 9000 kg/h, and the optimal total number ofstages is 47 for the extractive distillation column and 12 forthe entrainer recovery distillation column, respectively. Theoptimal solvent and raw feed position for extractive distillationcolumn are the 4th stage plate and the 31st stage plate,respectively, and the optimal feed position is the 7th stage platefor entrainer recovery column.Two single temperature control structures, i.e., basic and

improved control scheme, and a dual temperature controlscheme were presented and tested to the feed ow rate andcomposition disturbance. The improved control scheme withxed reux-to-feed ratio has slightly better tested results thanthe basic control scheme with xed reux ratio, while the dualtemperature control structure can maintain higher purity of the

two products with the same disturbance variations than theimproved control structure in 5h. Economic accounting decidesthe select of improved or dual temperature control scheme.

AUTHOR INFORMATIONCorresponding Authors*Tel: +86 13462354555. E-mail: skyang111@163.com.*Tel: +86 15903023407. E-mail: yujiewang2001@yahoo.com.*Tel: +86 13938729275. E-mail: baiguangyue@htu.cn.NotesThe authors declare no competing nancial interest.

NOMENCLATUREi,j = separation factor of component i and jC1 = extractive distillation columnC2 = entrainer recovery columnIDn = internal diameter for column n (m)Ki = K factor of component iNFE = feeding location for the entrainerNF = feeding location for the fresh feedNF2 = feeding location for the feed to column C2NTn = number of theoretical plates for column nQCn = condenser heat removal for column n (KW)QHX = heat duty of the heat exchanger HXQRn = reboiler heat input for column n (KW)Rn = reux ratio for column nT = absolute temperature (K)TAC = total annual costsTn,m = temperature for tray m in column n (K)

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