Design and Control of an Extractive Distillation System for Benzene

Article pubs.acs.org/IECR

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

School of Chemistry and Chemical Engineering, Henan Institute of Science and Technology, Xinxiang, Henan, 453003, China School 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 commercial simulators (Aspen Plus and Aspen Dynamics). On the basis of global economic optimization, a design with optimized operation conditions for this process is developed. Two single temperature control structures with fixed reflux ratio or fixed reflux-to-feed ratio and a dual-end temperature control structure for entrainer recovery column are presented to evaluate the dynamic performance with feed flow rate and feed composition disturbances, and the last structure is quite effective. where αsi,j is the relative volatility of component i and component j in the presence of entrainer s, K∞ i,s is the infinite dilution K value for a trace of species i in the entrainer, and K∞ i,s is the infinite 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 following three entrainers are studied in this paper: dimethyl sulfoxide (DMSO), phenol (PhOH), and sulfolane (SFO). Their K values at infinite dilution and the relative volatility are listed in Table 1. Because DMSO does not lead into further azeotrope in the system and shows a higher relative volatility (2.139), it is chosen as entrainer in the simulation. This large relative volatility is helpful to gain an economical separation sequence. Figure 1B−C gives Txy diagrams for benzene/DMSO and acetonitrile/DMSO mixtures, respectively. The former plays an important role in the top of the extractive column, and the latter applies in the solvent-recovery column. 2.2. Thermodynamic Model Used in the Simulation. According to the refs 2 and 4, the best fit was obtained by the Wilson model for the binary isobaric VLE data composed of the benzene/acetonitrile and acetonitrile/DMSO system. No isobaric VLE experimental data were found for the benzene/ DMSO system. The predicted azeotropic temperature and azeotropic composition for benzene/acetonitrile at 0.43 atm using the Wilson model are 73.46 °C and 68.64 wt % (53.50 mol %), respectively, which were in very good agreement with experimental data. In this simulation, the Wilson model was used to describe the nonideality of the liquid phase while the vapor phase was assumed to be ideal. The Wilson model parameters of the three pairs were taken from Aspen Plus, and all other physical property model parameters were taken from the built-in values in Aspen Plus.

1. INTRODUCTION 2-Picoline is an important pharmaceutical intermediate and a raw organic chemical material. Catalyzed by organic cobalt, ethyne and acetonitrile are used to produce 2-picoline presented by Qin et al.,1 and the synthesis process achieved wide-plant production scale in 2006. In the synthesis method, organic cobalt is dissolved in pure benzene solution as catalytic agent, and simultaneously, benzene is also one of the byproducts. After 2-picoline and some other byproducts are separated from reactant mixture, there are many benzene/acetonitrile mixtures left, and it is 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 minimumboiling azeotrope with azeotropic temperature at 73 °C and azeotropic composition at 52.9 mol % benzene. Figure 1A give the Txy curve for the benzene/acetonitrile system. In this paper, we attempt to use an extractive distillation process to separate benzene/acetonitrile. Among some suitable candidate solvents, dimethyl sulfoxide (DMSO) is elected as entrainer according to the values of relative volatility. The thermodynamics model Wilson is used in the simulation calculation. Using DMSO as entrainer, an optimized design for the extractive distillation process is developed using total annual costs (TAC) as an objective function from many alternatives. Finally, appropriate control structures were introduced, and their dynamic performances were evaluated. 2. STEADY STATE DESIGN 2.1. Entrainer Selection. Since the entrainer is a key factor in extractive distillation, more attention should be paid to its election. A criterion for entrainer selection is through the comparison of relative volatilities in the presence of different entrainers.3 The higher the relative volatility, the easier the separation. The relative volatility is defined as αis, j =

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

K i∞ ,s K∞ j,s

(1) © XXXX American Chemical Society

A

dx.doi.org/10.1021/ie4008425 | Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Industrial & Engineering Chemistry Research

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Table 1. Results of the Entrainer Selection for an Extractive Distillation System benzene(1)/acetonitrile(2) using the Wilson model solvent

Tb/K

K∞ 1,s

K∞ 2,s

α∞ 1,2

DMSO phenol sulfolane

464.00 454.99 560.45

27.12 19.50 79.40

12.68 10.88 53.17

2.139 1.792 1.493

2.3. Residue Curve Map (RCM) for the System. The RCM of ternary system can be used to design and distinguish between feasible and infeasible sequences and can be mapped by Aspen Plus. Using the Wilson model, the RCM of the benzene/acetonitrile/DMSO ternary system is drawn and shown in Figure 2. It can be seen that the benzene/acetonitrile azeotrope is the unstable node, DMSO is the stable node, and both benzene and acetonitrile are the saddles. No distillation boundary exists to the system, which is an ideal situation for selection of an extractive distillation process. Material balance lines on behalf of the extractive distillation process were also drawn in Figure 2. It is noticed that M, denoting the mixture of the raw material and entainer, can be separated into D1 and B1 on extractive distillation column, and B1 can be separated into D2 and B2 on entainer recovery column. This means that the feed F may be separated into relatively pure two products with the aid of the entrainer. To balance the tiny loss of entrainer in both D1 and D2 streams, a small makeup stream of DMSO should be added (not displayed in Figure 2). 2.4. Process Design and Economic Analysis. 2.4.1. Process Design. In this paper, the raw material used is from several local plants and mainly composed of benzene and acetonitrile with a little 2-picoline. After pretreatment, 2-picoline was removed from the raw material completely. The composition was analyzed by gas chromatography. The extractive distillation process was simulated with the following data: the feed was a mixture 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. B



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

Figure 4. Optimal process flowsheet for extractive distillation.

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

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



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

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

acetonitrile, with a flow rate of 3500 kg per hour with 7200 working hours annually. Aspen Plus and Aspen Dynamics are used to do the steady state and dynamic simulations. The Wilson activity model is chosen to predict the vapor liquid equilibrium in the simulator. To ensure the products are available for recycling, the two product specifications are set to be as follows: the acetonitrile impurity in the benzene product is not more than 1 wt %, and the acetonitrile product has a purity of 99.9 wt %. Here, Aspen notation of numbering stages from the top, with stage 1 being the flux drum and the last stage being the rebolier, is adopted. For an extractive distillation column (denoted as C1), when the total stages, operating pressure, and feed location are fixed, there are three design degrees of freedom: reflux ratio (RR1),

Table 2. Temperature Controllers Tuning Parameters for the Basic Control Structure parameters

TC1

TC2

TCHX

controlled variable manipulated variable transmitter range (K) controller output range (GJ/h) ultimate gain ultimate period (min) gain, Kc integral time, τ(min)

T1,40 QR1 273−443 0−7.71 8.1037 6.0 2.5324 13.2

T2,10 QR2 273−556 0−3.37 2.9728 4.8 0.9290 10.56

Trecycle QHX 273−370 −4.04−0 0.5200 2.4 0.1625 5.28

entrainer flow rate (FE), and reboiler heat duty (QR1). Since almost all the benzene that is present in the C1 bottom product D



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

reboiler, and condenser, and a payback period of 3 years is assumed. Small items such as reflux drums, pumps, valves, and pipes are usually not considered due to their much lower costs. The operating costs include the steam and cooling water, and only the former was reckoned in this simulation owing to its much higher cost. The heat transfer area for the condensers and reboilers is determined using the overall heat transfer coefficient and a differential temperature driving force. Here, the overall heat transfer coefficients and the calculation formula for the above equipment are all taken from Luyben’s book.5 The “Tray Sizing” function in Aspen Plus is employed to size the column vessels, and a sieve plate is selected. The “Design Spec/Vary” function was used to satisfy the product composition. The weir height of 0.025 m is adopted in order to maintain the pressure drop required, while other parameters use the default values. The stages are counted from top to bottom with the condenser as the first stage and the reboiler as the last stage for both columns. The pressure selection is very important; lower pressure leads to easier separation, but the pressure value is restricted by the adopted cooling water temperature, simultaneously, due to

will go to the top of the entrainer recovery column (denoted as C2), to achieve the desired acetonitrile product, the mass flow rate of benzene in the bottom product of the C1 is held at 1.2 kg/h (calculated from the material balance) by manipulating the reboiler heat duty QR1. 2.4.2. Economical Optimization. It is known that, for the extractive distillation process, an increase in the entrainer flow rate will reduce the heat duty of the extractive distillation column, but the heat duty of the entrainer recovery column will increase, and a larger column diameter is expected. Thus, a trade-off between the extractive column costs (include both the fixed capital costs and operating costs) and entrainer recovery column costs needs to be made. Thus, when the theoretical plates were fixed, there exists an optimal entrainer flow rate that minimizes TAC (total annual costs, defined in Luyben’s book5). When the entrainer flow rate is larger than the optimal value, the costs of the entrainer recovery column become more dominant and TAC increase. It is a convention to use the TAC as the objective function to be minimized, which includes annualized capital costs and operating costs. The capital costs include the column shell, E



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

In the optimization process, the six design variables, namely, the total stages of the extractive distillation column (N1), the fresh feed location (NF1), the entrainer feed location (NFE), the total stages of the solvent recycle column (N2), the feed location of solvent recycle column (NF2), and the solvent feed rate (FE) should be determined by minimizing the total TAC. The procedure to search the optimal value of each design variable 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 previous rounds, the new design value was adopted, and otherwise, the search process should be stopped. (4) Change the design variable based on the new calculated TAC value and go back to step 2. All the six design variables should be optimized according to the procedure above. After a round search, namely, all the six

the lower material temperature along with its pressure reduction. As the adopted temperature of the cooling water is constant at 308.15 K and the temperature difference is constant at 13.9 K for the column top condenser and heat exchanger according to Luyben’s book,5 the C2 column top temperature is 322.05 K (i.e., the sum of cooling water temperature and the temperature difference of heat transfer) and its pressure is calculated from Antoine extended formulation, namely, 0.33 atm. Also, as the same cooling water and temperature difference was adopted, the material flow from C2 bottom is also cooled to 322.05 K by the heat exchange (denoted as HX), namely, the recycling entrainer flow temperature input to C1 column. According to the suggestion of Knight and Doherty,6 the temperature difference of recycling entrainer flow temperature and C1 column top should be 5−15 K, and the temperature of C1 column top was fixed at 327.37 K (here the abovementioned temperature difference being at 5.32 K). Thus, the C1 column top pressure is calculated from Antoine extended formulation to be 0.43 atm. F



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

that the reflux ratio of 1.138 is indeed the minimum of the two possible reflux ratios. Figure 4 gives the final optimal flow 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 profiles of C1 and C2 columns for the flow sheet, respectively. There is a rapid rise in the temperature for stage 4 and a rapid fall for stage 31 in Figure 5A, at which the entrainer and fresh feed are fed. It is obvious that stage 40 displays a fairly steep slope for the temperature and benzene composition profiles with the extractive distillation column shown in Figure 5B. In Figure 6, two steep slopes are found near the 4th and 10th stages. The profile distinguishing features indicate the proper temperature control point for the two columns.

design variables optimized once, the next round search should be done until all the six design variables were not changed at a certain round. In the search process, step size for the theoretical plate number and feed position were fixed at 1 stage and the FE feed rate was fixed at 100 kg/h. The search method is intuitive and effective but rather time-consuming in a plantwide optimization. According to the optimal results, it is found that the optimal mass flow rate of the entrainer (FE) is 9000 kg/h, and the optimal total number of stages is 47 for the extractive distillation column and 12 for the entrainer recovery distillation column, respectively. The best feed position NFE is at the 4th stage, NF1 at the 31st stage for the extractive distillation column, and NF2 at the 7th stage for the entainer recovery column. The optimal reflux ratios for C1 and C2 are 1.138 and 0.467, respectively. The final minimum value of TAC is 451 150$/y. Extractive distillation systems have the interesting feature of two different values of reflux ratio that yield the same separation. Figure 3A,B manifests a plot of benzene purity or acetonitrile impurity from C1 column top flow versus reflux ratio at fixed solvent rate, respectively. From Figure 3A,B, we can conclude

3. CONTROL SYSTEM DESIGN 3.1. Basic and Improved Control Structure. Single temperature control structure is the first to be considered due to its less expensive cost. For the extractive distillation process, basic control structure, namely, fixed reflux ratio for C1 and C2 G



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

column (denoted as S1), and its improved control structure, namely, fixed reflux-to-feed ratio (denoted as S2), are often adopted. Before starting the dynamic simulation, the plumbling system and major equipment sizes must be specified. The commonly used heuristic for reflux drums and column bases sizing is to provide ∼5 min of liquid holdup when half full. All the control valves’ pressure drops are ∼3 bar with the valve half open at the design flow rate. Then, the steady state flowsheet is pressurechecked, and the Aspen Plus file is exported to Aspen Dynamics. The temperature at the 40th stage is adopted as the temperature control point for the extractive column, and the temperature at the 10th stage is used as the temperature control point for the extrainer recovery column. To the control schemes of S1 and S2, the following control structure is proposed for the extractive distillation control system. (1) Feed is flow-controlled (reverse acting). (2) Reflux drum levels in both columns are held by manipulating the flow of distillates (direct acting). (3) Base level in extractive distillation column is held by manipulating the flow of the bottoms (direct acting).

(4) Base level in entrainer recovery distillation column is held by manipulating the makeup DMSO flow rate (reverse acting). (5) The total entrainer flow is in proportion to the feed flow. (6) The pressure in the two columns is controlled by manipulating the heat removal rate in the condenser of the two columns (reverse acting). (7) The reflux ratios for C1 and C2 columns are fixed for the basic control structure (S1), and the reflux-to-feed ratios for C1 and C2 columns are fixed for the improved control structure (S2). (8) Entrainer feed temperature is held by manipulating the cooler HX heat duty (reverse acting). (9) The temperature for the 40th stage in the extractive distillation column is controlled by manipulating the reboiler heat input into the extractive distillation column (reverse acting). (10) The temperatures in the 10th stage in the entrainer recovery column are controlled by manipulating the reboiler heat input into the entrainer recovery column (reverse acting). H



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Figure 13. Dynamic responses for control structure (S3) with fixed reflux ratio for columns C1 and dual temperature control for column C2: feed flow rate disturbances. (Solid lines are denoted as +20% flow rate, and dashed lines are denoted as −20% flow rate.)

A noteworthy feature revealed here is that the base level in the entrainer recovery distillation column is held by manipulating the makeup DMSO flow rate suggested by Grassi and Luyben.7,8 Figures 7 and 8 demonstrate the basic and improved control 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 of the top pressure control loops for both columns are set at Kc = 20 and τ I = 12 min. Three dead time elements are inserted into the corresponding temperature control loops with a dead-time of 1 min. Using the Tyreus-Luyben tuning, relay-feedback tests are run on the temperature controllers to determine ultimate gains and periods, and the I



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

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

parameters are shown in Table 2 for S 1 or S 2 control schemes. Now, the dynamic performance of the control structure is evaluated by feed flow rate and composition disturbances. Figure 9 shows the dynamic responses of the basic control structure to positive and negative 20% step changes in feed flow rate at t = 0.2 h. As you can see from Figure 9, benzene product J



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

negative 0.05 benzene composition disturbances in the feed at t = 0.2 h. From Figure 10, acetonitrile product purity is 0.9566, having large deviation to its desired value. Other control variables remaining, i.e., benzene product purity, the two controlled tray temperatures, and two columns’ reboiler duties, have almost similar responses as those in the feed flow rate disturbance. The dynamic responses of improved control structure were shown in Figure 11 to the feed flow rate positive and negative 20% disturbance and in Figure 12 to the benzene composition positive and negative 0.05 disturbances in the feed at t = 0.2 h. The acetonitrile product purities are 0.9879 and 0.9902 to the same feed flow rate and benzene composition disturbance in the feed, respectively, and the benzene product purity has great improvement simultaneously (from 0.9882 to 0.9927 mass fraction) to the negative 0.05 benzene composition disturbance in 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 test result to rate and component disturbance than S1, but to the disturbances of positive 20% feed flow rate and positive 0.05 feed benzene composition, deviations of the acetonitrile product purity from C2 top at the new steady state are still too large in strict product requirements. To improve the control performance, a dual temperature control structure (denoted as S3) was presented, in which fixed reflux ratio was adopted for C1 column, and stage 4 and stage 10 were controlled by reflux rate and by reboiler duty in the C2 column, respectively, along with the same other control strategy mentioned above in single temperature control schemes. The test results were shown in Figures 13 and 14, from which the acetonitrile product purities have great improvement under the same fluctuation. The acetonitrile product purities are 0.9963 to the feed flow disturbance and 0.9987 to the feed benzene composition disturbance, respectively. In actual production, which one to be adopted for S2 or S3, depends on the purity of the product requirements, or the trade-off between the added value from improved purity of the two products and the increased costs if S3 was adopted.

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AUTHOR INFORMATION

Corresponding Authors

*Tel: +86 13462354555. E-mail: [email protected]. *Tel: +86 15903023407. E-mail: [email protected]. *Tel: +86 13938729275. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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NOMENCLATURE αi,j = separation factor of component i and j C1 = extractive distillation column C2 = entrainer recovery column IDn = internal diameter for column n (m) Ki = K factor of component i NFE = feeding location for the entrainer NF = feeding location for the fresh feed NF2 = feeding location for the feed to column C2 NTn = number of theoretical plates for column n QCn = condenser heat removal for column n (KW) QHX = heat duty of the heat exchanger HX QRn = reboiler heat input for column n (KW) Rn = reflux ratio for column n T = absolute temperature (K) TAC = total annual costs Tn,m = temperature for tray m in column n (K) REFERENCES

(1) Qin, B. W.; Gao, J. J. Chinese Patent, CN 1869023A, 2006. (2) Shri, K.; Raghunath, P. T.; Bachan, S. R. Isobaric Vapor-Liquid Equilibria of Binary Systems of Acetonitrile with Benzene, Toluene, and Methylcyclohexane. J. Chem. Eng. Data 1980, 25, 11−13. (3) Dyk, B. V.; Nieuwoudt, I. Design of Solvents for Extractive Distillation. Ind. Eng. Chem. Res. 2000, 39, 1423−1429. (4) Wang, Q. Y.; Zeng, H.; Song, H.; Liu, Q. S.; Yao, S. Vapor-Liquid Equilibria for the Ternary System Acetonitrile + 1-Propanol + Dimethyl Sulfoxide and the Corresponding Binary Systems at 101.3 kPa. J. Chem. Eng. Data 2010, 55, 5271−5275. (5) Luyben, W. L. Distillation Design and Control Using Aspen Simulation; John Wiley &Sons, Inc: New York, 2006. (6) Knight, J. R.; Doherty, M. F. Optimal Design and Synthesis of Homogeneous Azeotropic Distillation Sequences. Ind. Eng. Chem. Res. 1989, 28, 564−572. (7) Grassi, V. G. Practical Distillation Control; Van Nostand Reinhold Press: New York, 1992. (8) Luyben, W. L. Plantwide Control of an Isopropyl Alcohol Dehydration Process. AIChE J. 2006, 52, 2290−2296.

4. CONCLUSIONS Design and control of an extractive distillation process for separation of benzene/acetonitrile are investigated in our work. DMSO is chosen as a suitable entrainer by comparing the relative volatility. The Wilson model is used to calculate the thermodynamics properties. Using the total annual cost as the objective function, the optimal design of the extractive distillation process is presented. According to the simulation results, it is found that the optimal mass flow rate of the entrainer (S) is 9000 kg/h, and the optimal total number of stages is 47 for the extractive distillation column and 12 for the entrainer recovery distillation column, respectively. The optimal solvent and raw feed position for extractive distillation column are the 4th stage plate and the 31st stage plate, respectively, and the optimal feed position is the 7th stage plate for entrainer recovery column. Two single temperature control structures, i.e., basic and improved control scheme, and a dual temperature control scheme were presented and tested to the feed flow rate and composition disturbance. The improved control scheme with fixed reflux-to-feed ratio has slightly better tested results than the basic control scheme with fixed reflux ratio, while the dual temperature control structure can maintain higher purity of the K


Design and Control of an Extractive Distillation System for Benzene

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