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Control of Highly Heat-Integrated EnergyEfficient Extractive Distillation Processes Hui Zheng, Ye Li, and Chunjian Xu Ind. Eng. Chem. Res., Just Accepted Manuscript • Publication Date (Web): 24 Apr 2017 Downloaded from http://pubs.acs.org on April 28, 2017

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Industrial & Engineering Chemistry Research

Control of Highly Heat-Integrated Energy-Efficient Extractive Distillation Processes Hui Zheng, Ye Li, Chunjian Xu*

School of Chemical Engineering and Technology, State Key Laboratory of Chemical Engineering, Chemical Engineering Research Center, Collaborative Innovation Center of Chemical Science and Engineering, Tianjin University, Tianjin 300072, China Tel: +86 022-27892145. Fax: +86 022-27404440. Email: [email protected]

Abstract In our previous papers, two highly heat-integrated energy-efficient extractive distillation processes were presented by combining preconcentration distillation column (PDC) with extractive distillation column (EDC) or entrainer recovery column (ERC) (Ind. Eng. Chem. Res. 2014, 53, 7121; Chem. Eng. Sci. 2015, 135, 166). However, these novel synthesis systems with highly heat-integrated configurations may exhibit complicated issues of dynamic controllability. Thus, this paper extends the previous work to investigate the dynamic controllabilities of these new synthesis systems. The singular value decomposition (SVD) method is applied to analyze the temperature profiles and select the appropriate control trays. As the main column is highly heat-integrated from two columns, several degrees of freedom are lost. Thus, three temperature control loops are installed in the combined column. The dynamic performances of each system are tested by introducing ± 20% disturbances of feed flow rate and composition, respectively. The results show that robust control can be achieved for both systems. The controllabilities of conventional three-column processes also have been explored and compared with that of synthesis systems.

Keywords:

Dynamic

controllability;

Extractive

distillation;

Energy-Efficient; Azeotrope

ACS Paragon Plus Environment

Highly

Heat-Integrated;

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1 Introduction Extractive distillation has been widely applied in industry for the separation of binary mixtures with close boiling point or forming azeotropes, due to its low energy consumption, mass production capacity and flexible selection of possible entrainers.1-6 In recent decades, the design and control strategies of extractive distillation process have been investigated by many researchers. Grassi7 described a general procedure to design the optimal extractive distillation process and the effective control strategy. Luyben8-10 compared extractive distillation and pressure-swing distillation for acetone/methanol and acetone/chloroform separation. The design results revealed that the extractive distillation system brought much lower total annual costs (TAC) and simple control structures could be developed to resist large feed disturbances. Chien et al.11-12 optimally designed a heterogeneous azeotropic distillation system and an extractive distillation system for isopropyl alcohol (IPA) dehydration process, and he concluded the extractive distillation process was more economical. On their basis of the steady-state designs, control structures have also been developed and product purities can be kept quite well. Knapp and Doherty13 used ethylene glycol as the solvent to separate ethanol from water and reduced huge energy consumption by multi-effect and heat-integrated extractive distillation systems. Kiss and Ignat14 developed an innovative distillation to concentrate and dehydrate bioethanol in a single extractive dividing-wall column, with an energy savings of 17%. Abushwireb et al.15 investigated several technologies for the recovery of aromatics from pyrolysis gasoline, the results proved that the thermally-integrated extractive distillation sequence was the best candidate. The dynamic performance of distillation process has been well studied by many researchers. An overall control strategy of three-column IPA dehydration process was investigated by Luyben16. His work is an important reference for the three-column control design in this paper. Luyben17 also found a typical trade-off between the economics and controllability by comparing the control structures of ACS Paragon Plus Environment

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a conventional process and a heat-integrated process for separating a ternary mixture. Wang et al.18 separated ethanol (EtOH) and tetrahydrofuran mixture by extractive distillation and partially heat-integrated pressure-swing distillation, respectively. Results showed the latter one had more advantages in controllability. Errico19 studied the control of a similar process without the intermediate reboiler and concluded that the two-column configuration offered the better dynamic performances. Segovia-Hernández and Hernández20-22 studied the control of thermally coupled distillation schemes and found effective controllabilities by using the integral of the absolute error (IAE) criterion or the integral of the square error (ISE) criterion. Our previous work established two energy-efficient extractive distillation processes23-24. One combines preconcentration distillation column (PDC) with entrainer recovery column (ERC) for IPA/water separation using dimethyl sulfoxide (DMSO) as the entrainer. For the system of ethyl acetate (EtAc)/EtOH using furfural (FURF) as the entrainer, in which the entrainer reverses the relative volatility of the feed components, the PDC was combined with extractive distillation column (EDC). However, as dynamic property is a significant factor in assessing the distillation process, the controllabilities of these energy-efficient processes are still unknown. Furthermore, as novel processes are highly heat-integrated, the loss of several degrees of freedom may pose a challenge to achieve effective control. The purpose of this paper is to investigate the dynamic controllabilities of these complex integrated systems. Conventional temperature control structures are established based on singular value decomposition (SVD) analysis. The final control structures are improved from the base ones for the novel processes and conventional processes. The dynamic performances of the systems are tested by introducing ± 20% disturbances of feed flow rate and composition, respectively. The simulation results are compared by the dynamic performances and IAE criterion.

2 Control for IPA/Water Separation System ACS Paragon Plus Environment


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2.1 Process integration-combining PDC with ERC The conventional three-column extractive distillation sequence is commonly applied in industry application when the feed is relatively diluted. One typical case is the EtOH or IPA dehydration process in which the distillation sequence consisting of PDC, EDC and ERC is applied, as shown in Figure 1. The diluted feed is preconcentrated in PDC first, and then the light component IPA is obtained from the top of EDC with the heavy component water from the top of ERC. Make-up 0.00184 kmol/h S 1.00 DMSO 1.048 bar 1.013 bar F 250.0 kmol/h 0.200 IPA 0.800 H2O 360.0 K 1.206 bar

1.013 bar

D1

D2

76.896 kmol/h 0.65 IPA 0.35 H2O 353.7 K 1.013 bar 173.104 kmol/h 0.0001 IPA 0.9999 H2O 378.5 K 1.227 bar B1

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94.416 kmol/h 1E-5 IPA 0.285 H2O 0.715 DMSO 431.3 K 1.344 bar B2

Extractive distillation column

1.013 bar

49.987 kmol/h 0.9999 IPA 9.8E-5 H2O 2.0E-6 DMSO 355.2 K 1.013 bar

D3 26.912 kmol/h 3.5E-5 IPA 0.9999 H2O 6.5E-5 DMSO 373.1 K 1.013 bar

B3

67.505 kmol/h 0.0001 H2O 0.9999 DMSO 469.3 K 1.150 bar

Entrainer recovery column

S

Figure 1. Conventional three-column extractive distillation process for IPA dehydration system. In order to conserve energy, a novel distillation sequence was proposed in our previous paper in which PDC and ERC are heat-integrated into one column as shown in Figure 2.23 The new column functions as both a preconcentration section and an entrainer recovery section, in which the heavy component is acquired as its sidestream. Two cases were clearly developed based on global economic optimization with the goal of minimizing TAC and the results showed that the integrated processes bring more than 13% savings in energy consumption compared with the conventional process.


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Industrial & Engineering Chemistry Research Make-up S 2.2E-3 kmol/h 1.00 DMSO 345.1 K 1.048 bar 250.0 kmol/h 0.200 IPA 0.800 H2O

F

77.013 kmol/h 0.65 IPA D1 0.35 H2O 353.3 K 1.013 bar

358.1 K 1.210 bar

D2 49.986 kmol/h 0.9999 IPA 9.7E-5 H2O 3.0E-6 DMSO 355.2 K 1.013 atm

200.015 kmol/h 9.3E-5 IPA 0.9999 H2O 7E-6 DMSO 378.7 K 1.234 bar

B1

67.500 kmol/h 0.0001 H2O 0.9999 DMSO 476.2 K 1.378 bar

94.526 kmol/h 1E-6 IPA 0.286 H2O 0.714 DMSOB 2 413.4 K 1.344 bar

Combined S preconcentration/recovery column

Extractive distillation column

Figure 2. Combined PDC/ERC process for IPA dehydration system. The IPA dehydration case using DMSO as the heavy entrainer in that paper is the basis of dynamic analysis. The columns are operated at atmospheric pressure. In this system, the light component IPA boils at 355.20 K and the heavy component water boils at 373.15 K. The IPA/water mixture form a minimum azeotrope with a composition of about 67.28 mol % IPA at 353.33 K. The fresh feed is 250 kmol/h saturated liquid with 20 mol % IPA and 80 mol % water. The purity of water product and IPA product is 99.99 mol %. The numerical results are directly applied in the dynamic control study. The final control strategy will be tested by two typical disturbances: ± 20% changes in the fresh feed flow rate (F) (300 kmol/h, 200 kmol/h) and ± 20% changes in the mole composition of IPA (24 mol % IPA + 76 mol % water, 16 mol % IPA + 84 mol % water). 2.2 Basic control loops Before exporting the Aspen Plus files into Aspen Dynamics, some sizing parameters should be determined. The following rules are generally applicable throughout this paper. Column diameters are specified by Aspen Plus tray sizing function. The pump heads and control valve pressure drop are set as 3 atm. The column bases and reflux drums are sized to provide 5 min of holdup when at ACS Paragon Plus Environment


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the 50% liquid level. After checking the pressure, the Aspen Plus files are exported to Aspen Dynamics as pressure driven simulations. Since product purities are maintained by temperature controllers in most industrial cases, the control structures in this paper are limited to the use of temperature control loops. Some simple control loops are the same part of different control schemes. These regulatory control loops are developed first and the details are described as below: (1) Feed flow rate is flow-controlled (reverse acting). (2) The pressure in each column is controlled by manipulating the heat removal rate in its condenser (reverse acting). (3) Reflux drum level in each column is held by manipulating the flow rate of the distillate (direct acting). (4) Entrainer stream temperature is held by manipulating the cooler HX heat duty (reverse acting). (5) Reflux ratio in each column is fixed. Flow rate controllers are set at Kc = 0.5 and τI = 0.3 min. All pressure controllers are set at Kc = 20 and τI = 12 min. All liquid level controllers are set at Κc = 2. For all temperature control loops, the temperature control loop contains 1 min deadtime element. The relay-feedback test is run and Tyreus-Luyben tuning is used to determine the controller parameters. 2.3 Control structure of combined PDC/ERC process 2.3.1 Base control structure For the novel process, the base control scheme which only contains single loops is evaluated first. (1) The base level in C1 is held by manipulating the makeup flow rate (reverse acting). Grassi7 suggested that the makeup flow rate is manipulated for an effective control of the base level in C1. (2) The base level in C2 is held by manipulating bottoms flow rate (direct acting). There are still 3 manipulated variables left for temperature control loops in C1: reboiler duty (Qr1), ACS Paragon Plus Environment

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intermediate reboiler duty, and sidestream flow rate (Fss). The remained manipulated variable in C2 is the reboiler duty (Qr2). In order for these manipulated variables to be used rationally in temperature control loops, it is necessary to analyze the temperature profiles. 480

C1 C2

Temperature (K)

460 440 420 400 380 360 340 0

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Figure 3. Temperature profiles of IPA dehydration system in combined PDC/ERC process. The temperature profiles are presented in Figure 3. In general, it is analyzed by slope criterion method when the temperature information is rich enough. Considering its inapplicability with complex temperature profiles, the SVD method by Moore25 and Luyben26 is employed to choose control trays in both columns. Figure 4 displays the gain matrices vectors plotted against stage number. (a) 30 0.4

R Q

25

R Q

20 0.2

15 10

0.0

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Gain

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


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Industrial & Engineering Chemistry Research 30

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Figure 4. Sensitivity and SVD analysis for IPA dehydration system: (a) C1 (b) C2. The gain matrix G is decomposed into three matrices by Matlab function-SVD (U, 0) according to the equation G=U S VT. Matrix S is a 2 by 2 diagonal matrix and its condition number is used to evaluate the feasibility of applying two or more temperature controllers in the system. In general, small condition number indicates that the column can be possibly controlled by more than one temperature control loops.25 In this case, the condition numbers of matrices for C1 and C2 are only 1.27 and 3.23, and both are relatively small. The stage number corresponding to the larger U values reveal the better choices to be effectively controlled. Figure 4 illustrates the U matrices results of the columns. It is clearly observed that the stage 46 and stage 36 are good choices for column C1. There is also a peak at stage 30. According to SVD analysis results, the temperature of stage 46 in the bottom part of C1 should be controlled by Qr1 (reverse acting). The intermediate reboiler plays its role as a reboiler in the upper section so the temperature on stage 30 will be held constant by adjusting its duty (reverse acting). Based on the SVD analysis results, Fss controller is used to control the stage temperature in the upper part of the lower section, and the temperature on stage 36 is selected. For example, when the temperature on stage 36 decreases, Fss will be increased, and this is a disguised decrease in the “reflux” of the solvent recovery section, leading to a rise in temperature. Thus, Fss is manipulated to hold the temperature on stage 36 constant (direct acting). Figure 4 shows that the peak for column C2 is around stage 43 in the lower part. The temperature for stage 43 of C2 is controlled by ACS Paragon Plus Environment

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manipulating the reboiler input into C2 (reverse acting) for its large response value in U matrix and its location not too close to the bottom. The temperature controllers are tuned one by one. When the relay-feedback of temperature controller on the 36th tray is used, the relay amplitude should be reduced from the default 5% to 1% for its high degrees of coupling. Figure 5 gives the final base control scheme (CS1) for IPA dehydration system. Makeup TC PC

PC FC LC

C1

FC

1 29 30

F

LC

D1

×

1

REC

RR1

TC30

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×

C2

RR2

32 33

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40 36 39 46

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Figure 5. CS1: base control for the combined PDC/ERC process. Feed flow rate and IPA composition disturbances are introduced into the system at 0.5 h and the dynamic responses are shown in Figure 6. (a) 1.0000

1.000 0.9998

+20% feed -20% feed

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XD2 (%IPA)

Xss (%water)

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67.52 67.48 67.44 67.40 67.36 0

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Figure 6. Dynamic responses for CS1 of the combined PDC/ERC process: (a) ± 20% fresh feed flow rate disturbances, (b) ± 20% fresh feed IPA mole composition disturbances. It is shown that the product purities show a relatively large difference to their desired values after introducing the disturbances. When the flow rate or IPA composition of the fresh feed increases, the IPA purity from the top of C2 is finally kept at only 98.12%. Too little entrainer fails to facilitate the azeotrope separation since the entrainer flow rate (S) hardly changes. Clearly, enough entrainer is indispensable under these disturbances. 2.3.2 Modified control structure A modified control structure is proposed to hold the IPA purity at the desired value under all circumstances. Results of CS1 show that inadequate entrainer always leads to poor separation of the ACS Paragon Plus Environment

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azeotropes in C2. Too much entrainer may also lead to an increase in Qr1. Thus, S is supposed to be ratioed to F (S/F) so that S would vary to a right point when F or feed composition changes. The S/F feed forward ratio controller were commonly applied in many papers and became one of the indispensable controllers in the control in extractive distillation processes.16-18, 27-31 In this novel process, S can be either ratioed to F or feed stream of C2 (D1). Both of them can be used to resist feed flow rate disturbances, but S will still be kept constant as feed composition changes when S is proportional to F. Clearly, an increase in the S is necessary to facilitate the mixture separation in C2 when the composition of IPA increases in fresh feed. Thus, the latter option is adopted. This S/F feed forward ratio controller can bring the IPA purity into the acceptable range. The modified control structure (CS2) is shown in Figure 7 and the dynamic responses are shown in Figure 8. Makeup Ratio

TC PC

PC × FC

LC

LC

C1

FC

1

× 29 30

F

TC30

RR1

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×

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40 36 39 46

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Figure 7. CS2: modified control for the combined PDC/ERC process. (a)


Industrial & Engineering Chemistry Research 1.0000

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Figure 8. Dynamic responses for CS2 of the combined PDC/ERC process: (a) ± 20% fresh feed flow rate disturbances, (b) ± 20% fresh feed IPA mole composition disturbances. It is seen from Figure 8 (a) that S changes accordingly and is eventually kept constant at the new values by the feed forward ratio controller. The water product purity experiences larger and faster transient departures than the IPA product stream. The reason is that the introduction of perturbation ACS Paragon Plus Environment

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to C1 is earlier than C2 and it damped large disturbances with leading to smaller deviations in D1 stream composition. D1 is the feed stream to C2 and the small disturbances can be easily handled by C2 control loops. It takes about 4 hours for all manipulated variables to reach the steady state. In Figure 8 (b), the products purities are again held at their desired values. The control loops in this system perform well and the final product purities are higher than 99.985% under all disturbances. The improvement in the performance proves that the S/F ratio controller is critical in the control of extractive distillation processes. 2.4 Control structure of conventional three-column process 2.4.1 Base control structure The control structure of the original three-column process is investigated and to be compared with the control of the novel process. (1) The base level in C1 and C2 is held by manipulating bottoms flow rate (direct acting). (2) The base level in C3 is held by manipulating the makeup flow rate (reverse acting). (3) The total entrainer flow is in proportional to the feed flow of C2. 480

C1 C2 C3

460 440

Temperature (K)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


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Figure 9. Temperature profiles of IPA dehydration system in conventional three-column process. The temperature profiles are presented in Figure 9. Slope criterion method is used in selecting temperature control trays in each column: stage 29 at 360 K in C1, stage 46 at 400 K in C2, and stage 15 at 427 K in C3. (4) The temperature of stage 29 in C1 is controlled by Qr1 (reverse acting). ACS Paragon Plus Environment


(5) The temperature of stage 46 in C2 is controlled by Qr2 (reverse acting). (6) The temperature of stage 15 in C3 is controlled by Qr3 (reverse acting). The control structure (CS3) is shown in the supporting information. The dynamic responses are shown in Figure 10. (a) 1.000

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Figure 10. Dynamic responses for CS3 of the conventional three-column process: (a) ± 20% fresh feed flow rate disturbances, (b) ± 20% fresh feed IPA mole composition disturbances. All the product purities can be kept constant within 10 hours. When F increases by 20%, the water purity from the bottom of C1 is eventually kept at 99.8578%, which is a bit different from the primary purity (99.99%). The DMSO impurity in D3 also increases (from 0.0001 to 0.000318, but not clear in the figure since the vertical scale is too large) because it is a large temperature difference between its top and bottom parts. In such columns, it is hard for only one control tray to perfectly drive the light components up and the heavy components down. These D2 and D3 products purities also experience large transient departures when F or feed IPA composition decreases. The output signal of the temperature controller lags behind the temperature change on the control tray. Thus, the decrease in the reboiler duty lags behind the drop of tray temperature and the excess energy drives much heavy component to the top. Make the reboiler duty be proportional to the feed (Q/F) can reduce the lag time and eliminate this effect. 2.4.2 Modified control structure The modified control for the conventional process is explored based on the discussion in last section. (1) The Qr1 is proportional to the fresh feed flow rate with the ratio manipulated to control the temperature of stage 29 (reverse acting). The ratio controller is effective in reducing the lag time. An increase in the feed flow rate will cause a rapid increase in the Qr1; thus, more vapor rises from the bottom and drive the IPA away from the bottom. Luyben16 also applied the Q/F ratio controller to improve the performance of the IPA dehydration system with the same flowsheet. (2) The Qr2 is proportional to the D1 with the ratio manipulated to control the temperature of stage 46 (reverse acting). When F of feed IPA composition decreases, the Qr2 will decrease immediately and thus prevents the heavy components (water and DMSO) from being driven to the top. ACS Paragon Plus Environment


(3) The temperature of stage 6 in C3 is controlled by manipulating the reflux ratio (direct acting). This controller is added considering the large temperature range in C3. It prevents too much DMSO driven to the top by the vapor because more condensed reflux will flow back to the column when the temperature on the 6th stage rises. The improved control scheme is shown in Figure 11 and the dynamic results are shown in Figure 12. Makeup Ratio

TC PC

PC

PC

× FC LC

C1

1

×

LC

D1

1 6

REC

RR1

FC TC29

D2

×

1

RR2

6

LC

43

TC43

×

F

C2

LC

D3

×

C3

RR3

9

TC6

15

TC15

40 29

×

32

LC

49

B1

23

B2

B3

LC

Figure 11. CS4: modified control for the conventional process. (a) 0.99995

+20% feed -20% feed

0.9995

XD2 (%IPA)

XB1 (%water)

1.0000

99.8629%

0.9990

+20% feed -20% feed

0.99990

0.99985

0.99980

0.9985 0.99975

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S (kmol/h)

XD3 (%water)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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0.99984 0.99980

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70 65 60

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(b) ACS Paragon Plus Environment

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1.000

0.99990

+20% IPA -20% IPA

XD2 (%IPA)

XB1 (%water)

+20% IPA -20% IPA

0.99995

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0.990

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75

S (kmol/h)

XD3 (%water)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


0.9998

0.9997

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0.9996

55 0.9995 0

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Time (h)

Figure 12. Dynamic responses for CS4 of the conventional process: (a) ± 20% fresh feed flow rate disturbances, (b) ± 20% fresh feed IPA mole composition disturbances. Compared with the results by CS3, all three purities show much smaller transient deviations because these feed forward ratio controllers effectively reduce the lad time and quickly manipulate the variables to hold the temperatures constant. However, the final water purity (99.8629%) in C1 still remains relative large offset to the desired value (99.99%). The control loops in C2 and C3 perform well as we designed and bring expected results. 2.5 Comparison For the combined PDC/ERC process, CS1 failed to resist the disturbances of increase in feed flow rate and feed IPA composition. The proposed CS2 with the S/F control loop handled the disturbances well and kept the product purities at the desired values. For the conventional three-column process, the large transient departures in CS3 indicated poor controllabilities. But the Q/F and S/F feed forward ratio controllers in CS4 greatly improved the dynamic performances. The low water purity (99.8629%) in B1 may not be acceptable. The IAE criterion is applied to make a direct comparison between four control schemes, and the ACS Paragon Plus Environment


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results are shown in Table 1. The IAE calculation is conducted by Aspen dynamics software and each simulation is run by 10 hours. Table 1. IAE values of different control schemes in IPA dehydration process under disturbances.

Disturbances

Base control of the novel process (CS1) Modified control of the novel process (CS2) Base control of three-column process (CS3) Modified control of three-column process (CS4)

H2O-Xss IPA-D2 H2O-Xss1 IPA-D2 H2O-B1 IPA-D2 H2O-D3 H2O-B1 IPA-D2 H2O-D3

Feed rate increased by 20%

Feed rate decreased by 20%

2.90E-04 1.60E-01 3.31E-04 9.91E-04 5.01E-02 9.10E-05 1.34E-03 1.18E-02 1.18E-04 8.60E-05

9.51E-04 8.67E-04 3.95E-04 1.22E-04 3.17E-04 2.31E-01 1.13E-01 3.08E-04 6.70E-05 1.37E-04

Feed IPA composition increased by 20% 2.18E-04 1.60E-01 3.89E-04 9.80E-04 3.63E-02 9.68E-05 1.48E-03 2.69E-02 1.00E-04 8.23E-05

Feed IPA composition decreased by 20% 1.39E-03 8.67E-04 5.75E-05 1.22E-04 3.02E-04 2.32E-01 1.14E-01 2.98E-04 7.43E-05 2.17E-04

By comparing these data, it is easy to conclude that the modified control structure performs much better than the base one. The results also show that CS2 performs better than CS4 when F or feed IPA composition increases by 20%, since the 1.18E-02 and 2.69E-02 by CS4 are much larger than 3.31E-04 and 3.89E-04 by CS2. Thus, the combined PDC/ERC process shows advantages both in steady-state design and dynamic control compared with the conventional three-column process.

3 Control for EtAc/EtOH Separation System 3.1 Process integration-combining PDC with EDC Different from the distillation sequence in Figure 1, for the systems with the entrainer reversing the relative volatility of the azeotrope, actually, the heavy component B instead of light component A is withdrawn from the top of EDC, while A is withdrawn from the top of ERC, as shown in Figure 13.


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Industrial & Engineering Chemistry Research Make-up 0.43 kmol/h 1.00 FURF 343.4 K 1.048 bar D1 165.17 kmol/h 0.550 EtOH 0.450 EtAc 345.0 K 1.013 bar

F 300.0 kmol/h 0.750 EtOH 0.250 EtAc 347.9 K 1.069 bar

90.93 kmol/h 0.995 EtOH 0.005 EtAc 2.51E-4 FURF 351.5 K 1.013 bar

134.83 kmol/h 0.995 EtOH 0.005 EtAc EtOH 355.2 K 1.179 bar PDC

D3

D2

B1

EDC

613.92 kmol/h 6.06E-4 EtOH 0.121 EtAc 0.878 FURF 412.3 K 1.323 bar B2

74.68 kmol/h 0.005 EtOH 0.995 EtAc 1.2E-5 FURF 349.8 K 1.013 bar 539.24 kmol/h 0.0001 EtAc 0.9999 FURF 439.6 K 1.165 bar ERC B3

S

Figure 13. Conventional three-column extractive distillation process for the EtAc/EtOH system with the relative volatility of feed components reversed by the entrainer. 300.0 kmol/h 0.750 EtOH 0.250 EtAc 345.4 K 1.082 bar

D2

F 225.00 kmol/h 0.995 EtOH 355.7 K 0.004 EtAc 1.199 bar 1.00 FURF

165.2 kmol/h 0.550 EtOH D 1 0.450 EtAc 345.0 K 1.390 bar

Make-up 0.60 kmol/h 1.00 FURF 343.7 K 1.241 bar B1

613.68 kmol/h 7.31E-7 EtOH 0.121 EtAc 0.879 FURF 417.9 K 1.509 bar

Combined PDC/EDC

74.47 kmol/h 0.995 EtAc 0.005 FURF 350.5 K 1.013 bar

539.21 kmol/h 0.0001 EtAc B2 0.9999 FURF 440.0 K ERC 1.241 bar

Figure 14. Combined PDC/EDC process for the EtAc/EtOH system with the relative volatility of feed components reversed by the entrainer. In our previous study24, a novel flowsheet which included a preconcentration section in the upper part and an extractive section in the bottom part was configured as presented in Figure 14. Two cases were investigated in that paper and the optimization results revealed that the new distillation sequence brought more than 10% energy savings. The dynamic study in this section is based on the numerical results of the EtAc/EtOH separation case. In the conventional process, EtAc is the light product in PDC and driven to the bottom in EDC by the entrainer. In the new process, ACS Paragon Plus Environment


EtOH is obtained as the sidestream of the combined column (C1) with a purity of 99.5 mol % and EtAc is obtained from the top of ERC (C2) with a purity of 99.5 mol %. In this system, EtAc boils at 350.35 K and EtOH boils at 351.46 K at atmospheric pressure. The EtAc/EtOH mixture form a minimum azeotrope with a composition of about 53.45 mol % EtAc at 345.29 K. The feed flow rate is 300 mol/h with composition 75 mol % EtOH and 25 mol % EtAc. Two typical disturbances will be used to test the control structures: ± 20% changes in the fresh feed flow rate (360 kmol/h, 240 kmol/h) and ± 20% changes in the mole composition of EtAc (30 mol % EtAc + 70 mol % EtOH, 20 mol % EtAc + 80 mol % EtOH). 3.2 Control structure of the combined PDC/EDC process 3.2.1 Base control structure The Aspen Plus files were exported into Aspen Dynamics after calculating sizing parameters. The Kc and τI of the controllers including flow rate controllers, level controllers, and pressure controllers are the same as mentioned in section 2.2. The temperature control loops contain 1 min deadtime element. Relay-feedback test is run and Tyreus-Luyben tuning is used to determine the parameters of temperature controllers. (1) The base level in C1 is held by manipulating bottoms flow rate (direct acting). (2) The base level in C2 is held by manipulating the makeup of entrainer (reverse acting). 440

C1 C2

420

Temperature (K)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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400

380

360

340 0

5

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20

25

30

35

40

45

50

55

60

65

70

75

Stage

Figure 15. Temperature profiles of EtAc/EtOH system in combined PDC/EDC system. Based on our analysis of control strategy in IPA/water system, it is concluded that tray ACS Paragon Plus Environment

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temperatures in column C1 can be controlled by manipulating Qr1, intermediate reboiler duty and Fss. The Qr2 can be employed to control one tray temperature in C2. Figure 15 gives the temperature profiles for EtAc/EtOH system. The SVD method is used to decide which tray is the best control point in each column and the results are plotted in Figure 16. The condition numbers of the 2 by 2 diagonal matrices are 2.67 and 2.99, respectively. (a) 0.3

3.0

R Q

2.5

R Q

2.0 1.5

0.2 0.1

1.0 0.5

0.0

-0.5

U

Gain

0.0

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0

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-0.2 -6 -8

-0.4

-10 0

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Stage

Stage

Figure 16. Sensitivity and SVD analysis for EtAc/EtOH separation system: (a) C1 (b) C2. There are peaks at stage 21, 33, 47-53 in C1 and stage 4 and 16 in C2. In C1, the temperature for stage 51 in the lower part is controlled by Qr1 (reverse acting). The temperature for stage 21 is controlled by manipulating the intermediate reboiler input (reverse acting). Fss is manipulated to hold temperature stage 33 constant (reverse acting). The temperature for stage 16 of C2 is controlled by manipulating the reboiler input into C2 (reverse acting). Luyben32 studied the system in which the entrainer reverses the relative volatility between the feed components and concluded that it is difficult to hold product purity well since the ACS Paragon Plus Environment


original heavy component preferentially wants to go out the bottom. It is also a vast variation between the top and bottom in C2. Thus, we add one more temperature control loop to adjust the reflux ratio for holding temperature stage 4 constant (reverse acting). The final base control structure (CS5) for this combined PDC/EDC process is given in the supporting information. (a) 0.998

1.00

XD2 (%Ethyl Acetate)

0.996

Xss (%Ethanol)

0.98

+20% feed -20% feed

0.96

0.94

0.92

0.994 0.992 0.990

+20% feed -20% feed

0.988 0.986 0.984 0.982 0.980 0.978

0.90 0

2

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6

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Time (h)

540.8

+20% feed -20% feed

S (kmol/h)

540.4 540.0 539.6 539.2 538.8 0

2

4

6

8

10

Time (h)

(b) 0.998 0.996

0.992


Xss (%Ethanol)

0.996

+20% Ethyl Acetate -20% Ethyl Acetate

0.988

0.984

0.994 0.992


0.990 0.988 0.986 0.984

0.980 0

2

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6

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0

Time (h)

540.4


540.2 540.0 539.8 539.6 539.4 539.2 539.0 0

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Time (h)

540.6

S (kmol/h)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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6

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Time (h)


8

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Figure 17. Dynamic responses for CS5 of the combined PDC/EDC process: (a) ± 20% fresh feed flow rate disturbances, (b) ± 20% fresh feed EtAc mole composition disturbances. The dynamic responses are given in Figure 17. It is seen that the purity of EtAc decreased a lot when feed flow rate increased. This is due to insufficient entrainer which is critical in the extractive distillation process. Other purities are kept at the acceptable range. 3.2.2 Modified control structure In the modified control scheme, S/F feed forward ratio controller is newly added. Unlike the IPA dehydration system, S can only be ratioed to the F in this flowsheet which means S will still be kept constant when the feed composition changes. According to the results of the CS5, the disturbances of 20% change in the EtAc composition would not be a problem even without the S/F controller. The modified control scheme (CS6) is shown in Figure 18, and the dynamic responses are shown in Figure 19.

PC

PC LC

LC

D1 1

FC

11 21

F

FC

C1

×

TC21

D2 1

RR1

4

27 28

×

TC4

11

33 34

TC33

REC

51 56 16

Ratio×

72

FC

C2

TC51

LC

TC16

24

B1

LC

Makeup B2 TC

Figure 18. CS6: modified control for the combined PDC/EDC process. (a)


RR2

Industrial & Engineering Chemistry Research 1.01

1.00

1.00

0.99

Xss (%Ethanol)

0.98


0.99

+20% feed -20% feed

0.97 0.96 0.95 0.94 0.93 0.92 0.91

+20% feed -20% feed

0.98 0.97 0.96 0.95 0.94 0.93

0.90 0

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640

S (kmol/h)

600

+20% feed -20% feed

560 520 480 440 0

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(b) 0.998

0.998 0.996

0.996

0.992


Xss (%Ethanol)

0.994


0.990 0.988 0.986 0.984 0.982 0.980

0.994 0.992


0.990 0.988 0.986 0.984

0

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1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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540 538 536 534 0

2

4

6

8

10

Time (h)

Figure 19. Dynamic responses for CS6 of the combined PDC/EDC process: (a) ± 20% fresh feed flow rate disturbances, (b) ± 20% fresh feed EtAc mole composition disturbances. The disturbances were introduced into the system at 0.5 h. All desired product purities can be achieved. In Figure 19 (a), the product purities experience about 8% deviations in two hours. In Figure 19 (b), the EtOH product purity has slightly decreased but is still acceptable. If higher purity ACS Paragon Plus Environment

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is needed, more stages and possibly more energy consumption is required in steady-state design or advanced composition controller should be used. 3.3 Control structure of conventional three-column process In this system, the control property of the conventional three-column process is explored for a direct comparison with the combined PDC/EDC process. (1) The base level in C1 and C2 is held by manipulating bottoms flow rate (direct acting). (2) The base level in C3 is held by manipulating the makeup flow rate (reverse acting). (3) The total entrainer flow is in proportional to the feed flow of C2. 440

C1 C2 C3

420

Temperature (K)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


400

380

360

340 0

5

10

15

20

25

30

35

40

45

Stage

Figure 20. Temperature profiles of EtAc/EtOH system in conventional three-column process. Figure 20 gives the temperature profiles. Slope criterion method is applied in selecting temperature control trays in each column: stage 20 at 351 K in C1, stage 42 at 379 K in C2, stage 4 at 351 K and stage 16 at 423 K in C3. Remember that we also selected the same two trays in the ERC in CS5 and CS6. (4) The temperature of stage 20 in C1 is controlled by Qr1 (reverse acting). (5) The temperature of stage 42 in C2 is controlled by Qr2 (reverse acting). (6) The temperature of stage 16 in C3 is controlled by Qr3 (reverse acting). (7) The temperature of stage 4 in C3 is controlled by manipulating the reflux ratio (direct acting). The control scheme of the novel process is shown in the supporting information and the dynamic responses are shown in Figure 21. ACS Paragon Plus Environment


(a) 0.998

0.9956

XD2 (%Ethanol)

XB1 (%Ethanol)

+20% feed -20% feed

0.9954

0.996

+20% feed -20% feed

0.994

0.992

0.9952 0.9950 0.9948 0.9946 0.9944

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Time (h) 1.000

640 0.998

+20% feed -20% feed

S (kmol/h)


600 0.996 0.994 0.992 0.990

+20% feed -20% feed

560 520 480

0.988

440 0.986 0

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Time (h)

(b) 0.9965

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XD2 (%Ethanol)

XB1 (%Ethanol)

0.9955 0.9950 0.9945


0.9940 0.9935 0.9930

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0.994

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560


540 520 500 480

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0

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Time (h)

Figure 21. Dynamic responses for CS7 of the conventional process: (a) ± 20% fresh feed flow rate disturbances, (b) ± 20% fresh feed EtAc mole composition disturbances. In this system, the control structure performs quite well and all of the final product purities are higher than 99.2%. The entrainer flow changes smoothly after the disturbances are introduced. ACS Paragon Plus Environment

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3.4 Comparison The IAE results of different control structures are shown in Table 2. Table 2. IAE values of different control schemes in EtAc/EtOH separation process under disturbances.

Disturbances

Base control of novel process (CS5) Modified control of the novel process (CS6) Control of three-column process (CS7)

EtOH-Xss EtAc-D2 EtOH-Xss EtAc-D2 EtOH-B1 EtOH-D2 EtAc-D3

Feed rate increased by 20%

Feed rate decreased by 20%

6.36E-01 1.90E-02 8.89E-02 6.97E-02 2.76E-02 1.23E-03 2.77E-02

1.40E-02 1.10E-02 9.77E-02 1.03E-02 1.40E-02 8.24E-04 1.87E-02

Feed EtAc composition increased by 20% 1.34E-02 8.75E-03 1.43E-02 8.81E-03 1.29E-02 1.08E-03 2.19E-02

Feed EtAc composition decreased by 20% 1.19E-02 1.31E-02 1.20E-02 1.31E-02 2.35E-02 1.67E-03 1.63E-02

Both CS5 and CS6 of the novel process show some advantages and some disadvantages. However, as we discussed in section 3.2, the EtOH purity dropped to an undesired value (92%) in CS5. Some IAE values of the CS6 are larger, which is mainly due to the transient deviations. But the final product purities of the modified control are still acceptable. The IAE values of the CS7 are generally smaller, and this control performs the best. The investigation shows the trade-off between steady-state economics and dynamic controllability. The heat-integrated process is energy-efficient but its dynamic responses to disturbances exhibit slight transient deviations.

4 Conclusion In this paper, the overall control strategies of two heat-integrated distillation processes and the conventional processes are proposed. The proposed strategies are illustrated based on the IPA dehydration process and the separation of EtAc/EtOH system, respectively. The SVD method is employed to analyze the temperature profiles. For the novel processes, the final strategy contains three temperature control loops in the integrated column and the three temperatures are held by manipulating reboiler duty, intermediate reboiler duty, and sidestream flow rate, respectively. The ACS Paragon Plus Environment


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S/F feed forward ratio controller greatly improves the dynamic responses under the disturbances. The results show that the products purities can be held close to the setpoint under the ± 20% disturbances of feed flow rate and composition, respectively. These robust controls indicate both heat-integrated processes can be applicable to extractive distillation. For the conventional processes, effective controls have been presented. For the IPA/water separation system, the novel process shows advantages both in steady-state design and dynamic control. For the EtAc/EtOH separation system, the products purities of the heat-integrated process exhibit slight transient deviations, but the final purities are still acceptable.

Funding sources This research did not receive any specific grant from funding agencies in the public, commercial, or not-for-profit sectors.

Acknowledgment Conflict of interest The authors declare no competing financial interest.

Abbreviations PDC

preconcentration distillation column

EDC

extractive distillation column

ERC

entrainer recovery column

SVD

singular value decomposition

TAC

total annual costs

IPA

isopropyl alcohol

EtOH

ethanol

IAE

integral of the absolute error ACS Paragon Plus Environment

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ISE


integral of the square error

DMSO dimethyl sulfoxide EtAc

ethyl acetate

Kc

gain of controller

τI

integral time of controller

F

fresh feed flow rate

S

entrainer flow rate

S/F

entrainer flow rate to fresh feed flow rate ratio

RR

reflux ratio

Q/F

reboiler duty to fresh feed flow rate ratio

Qr1

reboiler duty in the first column

Qr2

reboiler duty in the second column

Qr3

reboiler duty in the third column

Fss

sidestream flow rate

F

fresh feed flow rate

Xss

sidestream product purity

CS1

base control for the combined PDC/ERC process

CS2

modified control for the combined PDC/ERC process

CS3

base control for the conventional three-column extractive distillation process

CS4

modified control for the conventional three-column extractive distillation process

CS5

base control for the combined PDC/EDC process

CS6

modified control for the combined PDC/EDC process

CS7

base control for the conventional three-column extractive distillation process with the

relative volatility of feed components reversed by the entrainer ACS Paragon Plus Environment


Page 30 of 33

References (1) Kovach, J. W.; Seider, W. D. Heterogeneous azeotropic distillation: Experimental and simulation results. AIChE J. 1987, 33, 1300. (2) Doherty, M. F.; Malone, M. F. Conceptual Design of Distillation Systems. McGraw Hill: New York, 2001. (3) Widagdo, S.; Seider, W. D. Journal review. Azeotropic distillation. AIChE J. 1996, 42, 96. (4) Seader, J.; Henley, E.; Roper, D. Separation Process Principles; Wiley: 2011. (5) Lei, Z. G; Li, C. Y.; Chen, B. H. Extractive Distillation: A Review. Sep. Purif. Rev. 2007, 32, 121. (6) Sucksmith, I. Extractive Distillation Saves Energy. Chem. Eng. 1982, 28, 91. (7) Grassi, V. G. Process design and control of extractive distillation. In Practical distillation control; Van Nostand Reinhold: New York, 1992. (8) Luyben, W. L. Comparison of extractive distillation and pressure-swing distillation for acetone-methanol separation. Ind. Eng. Chem. Res. 2008, 47, 2696. (9) Luyben, W. L. Control of the Maximum-Boiling Acetone/Chloroform Azeotropic Distillation System. Ind. Eng. Chem. Res. 2008, 47, 6140. (10) Luyben, W. L. Comparison of extractive distillation and pressure-swing distillation for acetone/chloroform separation. Comput. Chem. Eng. 2013, 50, 1-7. (11) Chien, I. L.; Zeng, K. L.; Chao, H. Y. Design and Control of a Complete Heterogeneous Azeotropic Distillation Column System. Ind. Eng. Chem. Res. 2004, 43, 2160. (12) Arifin S., Chien, I. L. Design and Control of an Isopropyl Alcohol Dehydration Process via Extractive Distillation Using Dimethyl Sulfoxide as an Entrainer. Ind. Eng. Chem. Res. 2008, 47, 790. (13) Knapp, J. P.; Doherty, M. F. Thermal integration of homogeneous azeotropic distillation sequences. AIChE J. 1990, 36, 969. ACS Paragon Plus Environment

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(14) Kiss, A. A.; Ignat, R. M. Innovative single step bioethanol dehydration in an extractive dividing-wall column. Sep. Purif. Technol. 2012, 98, 290. (15) Abushwireb, F.; Elakrami, H.; Emtir, M. Recovery of aromatics from pyrolysis gasoline by conventional and energy-integrated extractive distillation. Comput. Aided. Chem. Eng. 2007, 24, 1071. (16) Luyben, W. L. Plantwide Control of an Isopropyl Alcohol Dehydration Process. AIChE J. 2006, 52, 2290. (17) Luyben, W. L. Control Comparison of Conventional and Thermally Coupled Ternary Extractive Distillation Processes. Chem. Eng. Res. Des. 2016, 106, 253. (18) Wang, Y. L.; Zhang, Z.; Zhao, Y. J; Liang, S. S.; Bu, G. L. Control of Extractive Distillation and Partially Heat-Integrated Pressure-Swing Distillation for Separating Azeotropic Mixture of Ethanol and Tetrahydrofuran. Ind. Eng. Chem. Res. 2015, 54, 8533. (19) Errico, M.; Ramírez-márquez, C.; Ortega, C. E. T.; Rong, B. G.; Segovia-Hernández, J.G. Design and control of an alternative distillation sequence for bioethanol purification. J. Chem. Technol. Biotechnol. 2015, 90, 2180. (20) Segovia-Hernández, J. G.; Hernández, S. Rico-Ramı́Rez, V.; Jiménez, A. A comparison of the feedback control behavior between thermally coupled and conventional distillation schemes. Comput. Chem. Eng. 2004, 28, 811. (21) Segovia-Hernández, J. G.; Hernández, S.; Jiménez, A. Control Behaviour of Thermally Coupled Distillation Sequences. Chem. Eng. Res. Des. 2002, 80:783. (22) Segovia-Hernández, J. G.; Hernández, S.; Femet, R.; Jiménez, A. Control of Thermally Coupled Distillation Arrangements with Dynamic Estimation of Load Disturbances. Ind. Eng. Chem. Res. 2007, 46, 546. (23) Liang, K.; Li, W. S.; Luo, H. T.; Xia, M.; Xu, C. J. Energy-efficient extractive distillation ACS Paragon Plus Environment


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process by combining preconcentration column and entrainer recovery column. Ind. Eng. Chem. Res. 2014, 53, 7121. (24) An, Y.; Li, W. S.; Li, Y.; Huang, S. Y.; Ma, J.; Shen, C. L.; Xu, C. J. Design/optimization of energy-saving extractive distillation process by combining preconcentration column and extractive distillation column. Chem. Eng. Sci. 2015, 135, 166. (25) Moore, C. F. Selection of controlled and manipulated variables. In Practical Distillation Control; Van Nostrand Reinhold: New York, 1992. (26) Luyben, W. L. Distillation Design and Control Using Aspen Simulation; John Wiley & Sons: New York, 2006. (27) Wang, H. H.; Li, Y.; Su, W. Y.; Zhang, Y. Z.; Guo, J. J.; Li, C. L. Design and Control of Extractive Distillation Based on an Effective Relative Gain Array. Chem. Eng. Technol. 2016, 39, 2339. (28) Ramos, W. B.; Figueirêdo, M. F.; Brito, K. D.; Ciannella, B. S.; Vasconcelos, L. G. S.; Brito, R, B. Effect of Solvent Content and Heat Integration on the Controllability of Extractive Distillation Process for Anhydrous Ethanol Production. Ind. Eng. Chem. Res. 2016, 55, 11315. (29) Wang, H. H.; Xie, L.; Tian, P.; Tian, G. Z. Design and control of extractive dividing wall column and pressure-swing distillation for separating azeotropic mixture of acetonitrile/ N -propanol[J]. Chem. Eng. Process. 2016, 110, 172. (30) Cao, Y. J.; Hu, J. J.; Jia, H.; Bu, G. L.; Zhu, Z. Y.; Wang, Y. L. Comparison of pressure-swing distillation and extractive distillation with varied-diameter column in economics and dynamic control. J. Process Control, 2017, 49, 9. (31) Wang, Q. Y.; Yu, B. R.; Xu, C. J. Design and Control of Distillation System for Methylal/Methanol Separation. Part 1: Extractive Distillation Using DMF as an Entrainer. Ind. Eng. Chem. Res. 2016, 51, 1281. ACS Paragon Plus Environment

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(32) Luyben, W. L. Effect of Solvent on Controllability in Extractive Distillation. Ind. Eng. Chem. Res. 2008, 47, 4425. For Table of Contents Only


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