Control of heat-integrated extractive distillation processes

Control of heat-integrated extractive distillation processes

Computers and Chemical Engineering 111 (2018) 267–277 Contents lists available at ScienceDirect Computers and Chemical Engineering journal homepage:...

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Computers and Chemical Engineering 111 (2018) 267–277

Contents lists available at ScienceDirect

Computers and Chemical Engineering journal homepage: www.elsevier.com/locate/compchemeng

Control of heat-integrated extractive distillation processes William L. Luyben Department of Chemical Engineering, Lehigh University, Bethlehem, PA 18015, United States

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Article history: Received 5 November 2017 Revised 30 November 2017 Accepted 18 December 2017 Available online 29 December 2017 Keywords: Extractive distillation Azeotropes Distillation control

a b s t r a c t Pressure in many distillation columns is set such that cooling water can be used in the condenser. Pressure selection is more involved in some columns such as reactive distillation in which there is a trade-off between temperatures favorable for reaction kinetics and temperatures favorable for vapor-liquid equilibrium. In azeotropic systems, pressure selection is critical in achieving the desired separation by consideration of distillation boundaries and isovolatility curves. A recent paper presented a striking example of pressure selection in extractive distillation. Operation at 1 atm required a solvent-to-feed (S/F) ratio of 3.52 while operating at 10 atm cut the S/F to only 0.717. The purpose of this paper is to explore the dynamic controllability of this low S/F extractive distillation system. Results show that the control structure must be modified from the conventional configuration used in most extractive distillation processes. © 2017 Elsevier Ltd. All rights reserved.

1. Introduction There is a vast literature dealing with the separation of azeotropic mixtures using a variety of distillation methods: pressure-swing, heterogeneous azeotropic and extractive. Kossack et al. (2008) presented a systematic approach to selecting a suitable solvent for extractive distillation. Knapp and Doherty (1990) studied heat-integration of binary homogeneous azeotropic systems using extractive distillation methods with the acetone/methanol system as a specific example. Dynamic controllability was not considered in these papers. Grassi (1992) studied a hypothetical A/B system with a solvent S and developed a control structure that is widely used in industry. The effect of the choice of solvent on dynamic controllability was studied (Luyben, 2008) using the acetone/methanol system with three alternative solvents: water, dimethyl sulfur oxide (DMSO) and chlorobenzene. The best solvent on the basis of both economics and controllability is DMSO. The chlorobenzene solvent has the unusual effect of driving the methanol overhead in the extractive column despite methanol’s higher boiling point compared to acetone. You et al. (2017) recently reexamined the acetone/methanol extractive distillation system using chlorobenzene as solvent by exploring the effect of pressure in the extractive column. They used isovolatility curves to explain the large reduction required in the amount of solvent by raising pressure. No study of the effect of pressure on dynamic controllability was presented. That is the pur-

E-mail address: [email protected] https://doi.org/10.1016/j.compchemeng.2017.12.008 0098-1354/© 2017 Elsevier Ltd. All rights reserved.

pose of this paper. The Uniquac physical property package is used in the Aspen simulations. In the previous work (Luyben, 2008) at 1 atm with a large chlorobenzene solvent flowrate, the process was shown to be able to handle large disturbances in throughput and feed composition. In this paper we want to see if the 10 atm case with a much smaller solvent flowrate can also handle large disturbances. 2. System studied The numerical example is based on the process used by Knapp and Doherty (1990) and You et al. (2017). Feed is 540 kmol/hr of 50/50 mol% acetone/methanol. Product purities are 99.5 mol%. Knapp and Doherty assumed that both columns operated at atmospheric pressure, which is high enough to permit the use of cooling water in both condensers. Fig. 1 gives the xy-curves at 1 and 10 atm for the acetone/methanol system. Clearly pressure-swing distillation is a potential alternative to extractive distillation. Fig. 2 shows the ternary diagrams at these two pressures for the acetone-methanolchlorobenzene system. The isovolatility curve at 1 atm (Fig. 2A) intersects the vertical axis far away from the methanol point, which indicates the need for a large S/F ratio and results in high energy requirements. At 10 atm (Fig. 2B) the isovolatility curve intersects the vertical axis quite close to the methanol corner, so a smaller S/F ratio is expected. The flowsheets shown in Figs. 3 and 4 demonstrate that the pressure effect in this extractive distillation system is quite drastic. The solvent flowrate at 1 atm is 1900 kmol/h, while at 10 atm it drops to only 387 kmol/h. Note that an economizer is used in

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Fig. 1. xy diagram at 1 and 10 atm.

both flowsheets to reduce reboiler duty in the extractive column C1. Preheating the fresh feed using the hot solvent from the base of the solvent-recovery column C2 reduces the heat duty in C1 from 14.0 to 10.9 MW in the 1 atm design. It also reduces the heat load in the solvent cooler. The economizer was not used in the previous work (Luyben, 2008) but it seems like an obvious improvement. Some unexpected results of making this small change are presented in a later section of this paper in which the performance for the 1 atm case is degraded when an economizer is used.

Fig. 2. (A.) ternary diagram at 1 atm. (B.) Ternary diagram at 10 atm.

An alternative control structure is required to handle large disturbances. The optimum design presented by You et al. (2017) is used to set the number of stages in each column and feed locations. Since there is much less solvent used in the 10 atm case, the concentration of methanol in the extractive column is larger as shown in

Fig. 3. 1 atm case with economizer.

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Fig. 4. 10 atm case with economizer.

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Fig. 5A. At a given pressure, the higher methanol composition would mean lower temperature. However, the pressure is also higher in the 10 atm case, which would mean higher temperature for the same composition. The net effect of these competing effects in the extractive column is temperatures that are somewhat higher in the 10 atm case (Fig. 5B). The temperature difference makes heat integration possible as shown in Fig. 4. A portion of the overhead vapor from C1 at 410 K is condensed in the reboiler of C2 at 400 K (with pressure set at 0.8 atm). The remainder of the C1 vapor is condensed in an auxiliary water cooled condenser. 3. Control dystem design

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The Aspen Plus file is exported to Aspen Dynamics as a pressure-driven simulation after dynamic parameters are specified. Aspen Plus tray sizing is used to fix column diameters. Pump heads and valve pressure drops are selected to give reasonable rangeability so that 20% increases in throughput can be handled without valve saturation. All reflux drums and column bases are sized to provide 5 min of holdup when 50% full at steady-state conditions. 3.1. Basic control structure for 10 atm case The well-established control structure proposed by Grassi (1992) is the basis for the design, which was used in the previous work for the high solvent flowrate case at 1 atm. Fig. 6 shows the structure in which the solvent is ratioed to the feed and both reflux ratios are controlled. Conventional PI controller are used except for level loops, which are proportional with Kc = 2. The basic control structure consists of the following loops.

Fig. 5. (A.) Methanol composition profiles in C1. (B.) Temperature profiles in C1.

1. Feed is flow controlled. This is the throughput manipulator. 2. Solvent flow is ratioed to feed flow. 3. Reflux drum levels are controlled by distillate flows.

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Fig. 6. Base-case control structure using S/F, RR1 and RR2.

4. Base level in the extractive column is controlled by bottoms flow. 5. The reflux ratios are controlled in each column by manipulating reflux flowrate. 6. Base level in the solvent recovery column is controlled by the makeup flow of solvent. 7. Column pressure in C1 is controlled by heat removal in the auxiliary condenser. 8. Column pressure in C2 is controlled by condenser heat removal. 9. The temperature of the solvent entering the extractive column is controlled by the heat removal in the solvent cooler. 10. Stage 39 temperature is controlled in C1 by manipulating the steam-to-feed ratio (not shown). 11. Because of the sharp temperature profile in C2, an average temperature in C2 (Stages 4, 5 and 6) is controlled by manipulating vapor fed to the C2 reboiler. Temperature controllers are tuned by running relay-feedback tests and applying Tyreus–Luyben tuning rules. A 1 min deadtime is included in the temperature loops. The responses to 20% feed flowrate disturbances are shown in Fig. 7A. Solid lines are 20% increases and dashed lines are 20% decreases. Just as was found in the previous work at 1 atm with high solvent flowrates, product purities xD1(M) and xD2(A) are held quite close to their specifications for these large feed flowrate changes. In the previous work good product quality control was also obtained when the acetone feed composition was changed 5 mol% from its design value of 50 mol%. In the present 10 atm case, control performance was good when the feed composition

was changed from 50 to 55 mol% acetone (solid lines). However, a reduction in the acetone feed composition from 50 to 45 mol% resulted in oscillatory behavior of the process and a large drop in the purity of the methanol product xD1(M) produced as distillate in the extractive column (bottom left graph in Fig. 7B). 3.2. Modified control structure for 10 atm case A number of alternative control structures were explored to see if performance could be improved. The three control degrees that can be set are the flowrates of solvent, reflux in C1 and reflux in C2. These can be implemented as a number of ratios such as solvent-to-feed, reflux-to-feed, reflux-to-distillate, solventto-distillate, etc. Fig. 8 shows the control structure that produced the best results. The only change from the conventional Grassi structure (S/F, RR1 and RR2) is changing to a reflux-to-feed in the solvent recovery column C2. The feed to C2 is the bottoms B1 from the extractive column, so a R2/B1 ratio is used. Fig. 9A shows that the 20% feed flowrate disturbances are well handled with essentially the same responses observed using the R2/D2 ratio. In addition, Fig. 9B shows that the 5% acetone feed composition disturbances are handled with no oscillations for both increases and decreases. The purity of the methanol product xD1(M) is maintained close to specification, dropping to 99.36 mol% methanol. An increase in acetone feed composition means less methanol must go overhead in the extractive column, which requires less reboiler duty QR1 (second graph from the top in the right column of Fig. 9B).

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Fig. 7. (A.) Base-case; 20% feed flowrate disturbances. (B.) Base-case; 5% feed composition disturbances.

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Fig. 8. Modified control structure using S/F, RR1 and R1/B1.

These results demonstrate that the low S/F extractive process at 10 atm can be effectively control by using a slightly modified control structure. 4. Effect of economizer for 1 atm case The extractive column studied in the previous paper (Luyben, 2008) operated at 1 atm and handled large disturbances. No economizer was used in the previous flowsheet. It should be noted that the response of the 1 atm case without an economizer (given in Figure 18 of the previous paper (Luyben, 2008)) for the negative feed composition disturbance is essentially the same with that for the 10 atm case shown in Fig. 9B. There is a small offset of xD1(M) in both cases. An unexpected result was observed when the dynamic controllability of the 1 atm case was explored with an economizer. Fig. 3 shows the 1 atm flowsheet in which an economizer preheats the feed to 337 K. At 1 atm, the feed enters the column partially vaporized (V/F = 0.24). The reboiler duty in the extractive column is 14.0 MW and the reflux ratio is 1.70. Fig. 4 shows that at 10 atm, the feed does not flash and enters the column at 385 K. Fig. 10 shows the original flowsheet without the economizer in which the feed enters at 322 K as liquid. The reboiler duty is 14.94 MW and the reflux ratio is 1.55. So the use of an economizer affects the condition of the feed in terms of fraction vapor. This relatively small difference would not be expected to affect dynamic performance. However, an unexpected result was observed.

Small decreases in feed flowrate were made in the 1 atm process with an economizer. Fig. 11 gives results for 2, 5 and 7% decreases in feed flowrate. Methanol product purity drops drastically for these small decreases in feed flowrate. Remember that the process at 1 atm without an economizer handles 20% decreases in feed flowrate. It appears that the partially vaporized feed affects the separation. An explanation may be changes in the V/F ratio. When the feed flowrate is decreased, the temperature of the feed leaving the economizer increases, which increases the V/F ratio. More vapor in the feed requires an increase in the reflux ratio to maintain the same separation in the column. But in an extractive distillation column, there is a conflict between reflux and solvent. More reflux can dilute the solvent in the column and result in poorer separation. A number of approaches were studied to solve this problem. A composition controller was used to maintain the methanol product purity by adjusting the solvent-to-feed ratio over a range of feed flowrates. The stars in Fig. 12 show the steadystate values of the solvent flowrate that keep the methanol product xD1(M ) = 0.995 for five feed flowrates varying from 20% less than design up to 20% greater than design. Then the coefficients of a quadratic function were calculated using the design point and the two 20% change points.

 Solvent (kmol/h ) = 617.4 + 334.9

Feed 100



 − 18.69

Feed 100

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Fig. 9. (A.) Modified; 20% feed flowrate disturbances. (B.) Modified; 5% feed composition disturbances.

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Fig. 10. 1 atm case without economizer.

Fig. 11. 1 atm case with economizer; decreases in feed flowrate.

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Fig. 12. Nonlinear S/F for 1 atm process.

Fig. 13. Aspen Dynamics implementation of nonlinear S/F.

The revised control structure uses a nonlinear functional relationship between solvent flowrate and feed flowrate. Fig. 13 shows the Aspen Dynamics implementation of generating the setpoint of the solvent flow controller from the measured feed flowrate.

The solvent flow controller is on “cascade” and gets its setpoint from the MultiSum block labeled “sol/feed.” Note that the proposed control scheme does not require an online composition measurement.

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Fig. 14. (A.) Nonlinear S/F; 20% feed flowrate disturbances. (B.) Nonlinear S/F; 50% feed composition disturbances.

The effectiveness of this control structure is demonstrated in Fig. 14 for large disturbances in both feed flowrate and feed composition. The 20% decrease in feed flowrate exhibits a large dynamic transient methanol product purity xD1(M) , but the control structure eventually returns the purity to very close to its specification.

5. Conclusion A control structure is developed that provides effective control of the 10 atm extractive distillation column with a small solventto-feed ratio. A reflux-to-feed ratio is used in the solvent recovery column instead of the conventional reflux-to-distillate.

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Using an economizer in the 1 atm process is shown to produce unexpected problems in handling decreases in feed flowrate. A nonlinear solvent-to-feed control structure is developed to handle the problem, which appears to be caused by the partially vaporized feed when an economizer is used at low pressure. References Grassi, V.G., 1992. Process design and control of extractive distillation. Practical Distillation Control Chapter 18 in. van Nostrand Reinhold, New York.

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Knapp, J.P., Doherty, M.F., 1990. Thermal integration of homogeneous azeotropic distillation sequences. AIChE J. 36, 969–983. Kossack, S., Kraemer, K., Gani, R., Marquardt, W, 2008. A systematic synthesis framework for extractive distillation processes. Chem. Eng. Res. Des. 86, 781–792. Luyben, W.L., 2008. Effect of solvent on controllability in extractive distillaiton. Ind. Eng. Chem. Res. 47, 4425–4439. You, X., Gu, J., Peng, C., Shen, W., Liu, H., 2017. Improved design and optimization for separating azeotropes with heavy component as distillate through energy-saving extractive distillation by varying pressure. Ind. Eng. Chem. Res. 56, 9156–9166.