- Email: [email protected]
Design and control of a pressure-swing distillation process with vapor recompression
Accepted Manuscript Title: Design and Control of a Pressure-Swing Distillation Process with Vapor Recompression Author: William L. Luyben PII: DOI: Reference:
S0255-2701(17)30869-3 https://doi.org/10.1016/j.cep.2017.09.020 CEP 7084
To appear in:
Chemical Engineering and Processing
Received date: Revised date: Accepted date:
28-8-2017 28-9-2017 29-9-2017
Please cite this article as: William L.Luyben, Design and Control of a PressureSwing Distillation Process with Vapor Recompression, Chemical Engineering and Processing https://doi.org/10.1016/j.cep.2017.09.020 This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.
Submitted to Chemical Engineering and Processing: Process Intensification
Design and Control of a PressureSwing Distillation Process with Vapor Recompression
William L. Luyben
Department of Chemical Engineering Lehigh University Bethlehem, PA 18015 USA
September 1, 2017 Revised September 28, 2017
[email protected]; 610-758-4256; FAX 610-758-5057
High Lights
Pressure-swing distillation using vapor recompression in both columns is studied. An effective control structure is developed for this unique process. Energy costs are 16% lower than those given in a published paper.
Abstract Pressure-swing distillation is widely used to separate minimum-boiling azeotropes when the azeotropic composition has significant pressure dependence. The two columns operate at different pressures with distillate streams having compositions close to their respective azeotropes. Heat integration is typically used since column temperatures are sufficiently different to permit heat transfer. A novel modification to achieve process intensification has recently been proposed that uses vapor recompression on both columns. The study developed a complex heat exchanger network along with the design of the two columns. Dynamic controllability was not studied. The purpose of this paper is to develop a robust control structure for this process. A novel control scheme features the flowrate of the distillate recycle from the highpressure column as the throughput manipulator with the fresh feed brought in on level control.
Key Words Pressure-swing distillation; azeotropic distillation; vapor recompression
1. Introduction Vapor recompression (heat pumping) has many applications in a wide variety of domestic and industrial uses. The basic thermodynamic principle of operation is to use work to produce temperature differences that permit heat transfer. The home refrigerator
2
is the most widely applied and most simple example in which high-value electrical energy is used to achieve transfer of heat from the low temperature inside the refrigerator to the warmer air in the kitchen. The closed cycle of the circulating refrigerant can be visually conveyed on an enthalpy-pressure diagram with energy flowing in from the lowtemperature source (air in refrigerator) and energy flowing out to the high-temperature sink (room air). Work is required to raise the pressure of the vapor refrigerant so it can be condensed by the high-temperature sink. Then expanding the high-pressure liquid refrigerant to a low pressure produces a partially vaporized low-temperature stream that can be completely vaporized by heat from the source. The final step is compressing the refrigerant vapor from the low-pressure “evaporator” to the high pressure in the “condenser.” Early industrial applications that involve vapor recompression of process streams include distillation and evaporation. The separation of components with similar boiling points (low relative volatility) such as propylene/propane1 can be achieved in a distillation column by compressing the overhead vapor and condensing it in the reboiler to provide vapor boilup in the column. Several recent papers have studied vapor recompression in more complex distillation systems. Li et al2 proposed using vapor recompression to provide the heat required in one of the reboilers of a divided-wall column separating a heterogeneous azeotrope. Effective dynamic control of this system has been demonstrated3. Luo et al4 proposed the use of vapor recompression in a single divided-wall column for bioethanol dehydration, and Patrascu et al5 developed a control system for this complex interacting system. An improved process design and a more robust control structure for this process has been presented6. The use of vapor recompression in pressure-swing distillation systems has recently been proposed by Xia et al7. This is a novel switch from conventional pressureswing distillation in which heat integration is typically used between the two columns operating at two pressure and temperature levels. The economics of a conventional pressure-swing distillation process with and without heat integration were compared with the vapor-recompression design. The process with vapor recompression had a much higher capital investment (73 %) compared to the heat integrated case because of the two
3
expensive compressors. However the energy cost is much smaller (48 %), producing a modest reduction in total annual cost of 8.6 %. A complex heat-exchanger network was proposed in the Xia et al7 paper that used both the sensible heat and the latent heat of the compressed vapor streams to achieve “self heating” with no external source of heat required. The only energy cost was the compressor power of the two compressors. Complex process flow patterns were proposed with two process streams split for preheating in parallel heat exchangers with different heat source streams. A total of nine separate heat exchangers were used. A much more simple heat-exchanger system is used in this paper.
2. Process Studied The numerical example studied by Xia et al7 is the separation of di-isopropyl ether (DIPE) and isopropyl alcohol (IPA) whose normal boiling points are 63.3 oC and 82.26 oC, respectively. These components form minimum-boiling azeotropes as shown in Figure 1. At 1.01 bar, the composition of the azeotrope is 77.43 mol% DIPE at 65.97 oC. At 4.05 bar, the composition of the azeotrope is 61.36 mol% DIPE at 110.8 oC. The reasonable pressure dependence of the azeotropic composition permits the economic use of pressure-swing distillation to achieve high-purity products. Aspen steady-state and dynamic simulations are used in this study with NRTL physical properties. Figure 2 gives the xy-diagram for the two columns with the feed and products streams of the two columns. With a feed of 25 mol% DIPE, the bottoms from the lowpressure column C1 is high-purity IPA with composition xB1. The distillate D1 has a composition xD1 slightly smaller than the azeotropic composition at 1.01 bar. The distillate is fed to the high-pressure column C2 in which the bottoms is high-purity DIPE with composition xB2. The distillate has a composition xD2 slightly larger than the azeotropic composition at 4.04 bar. The approach composition difference (yaz – xD) is a key design optimization variable that is discussed in the next section. The distillate D2 from the high-pressure column is recycled back to the low-pressure column. Figure 3 gives the flowsheet of the pressure-swing distillation process with vapor recompression studied in this paper. The column configurations are those used in the Xia
4
et al7 paper (number of trays, feed locations and column pressures), but a more simple heat exchanger network is used to avoid dynamic controllability problems. The overhead vapor from each column is compressed to a pressure that is high enough to provide a differential temperature driving force ∆T in the two reboilers of 5 oC. This is the design ∆T used by Xia et al7 and is reasonable because expensive compression is involved, which pushes the design trade-off between heat exchanger area and compressor power to large area and small ∆T to reduce compression costs. The low-pressure column operates at 1.01 bar and has 20 stages with 0.01 bar pressure drop per stage. The bottoms product is high-purity IPA, giving a base temperature of 86 oC. Therefore the condensing temperature on the hot side of the LPC reboiler must be 91 oC. The compressor discharge pressure that gives this temperature of the LPC overhead vapor is 2.43 bar. The high-pressure column operates at 4.05 bar and has 28 stages with 0.01 bar pressure drop per stage. The bottoms product is high-purity DIPE, giving a base temperature of 121 oC. Therefore the condensing temperature on the hot side of the HPC reboiler must be 126 oC. The compressor discharge pressure that gives this temperature of the HPC overhead vapor is 6.08 bar. The energy requirement to achieve the specified separation in the LPC is 1101 kW, but the heat required to desuperheat and condense the stream coming from the compressor K1 exceeds this amount by 57.0 kW. Therefore, an auxiliary condenser is installed using cooling water. The flowrate of the cooling water will be used to control a temperature in the LPC. The energy requirement to achieve the specified separation in the HPC is 767 kW, but the heat released in desuperheating and condensing the stream coming from the compressor K2 is only 677 kW. A steam-heated auxiliary reboiler is installed to provide the required additional vapor boilup (90 kW). The flowrate of the steam will be used to control a temperature in the HPC. The total energy requirement of the process shown in Figure 3 is 171 kW of expensive electric power to drive the compressors and 90 kW of inexpensive lowpressure steam. It is important to note that the compressor power requirements of the Xia et al7 paper were 248.1 kW. So the proposed process shown in Figure 3 achieves the
5
separation with 31% less high-level compressor power requirements. However, the proposed process requires 90 kW of energy in the auxiliary reboiler. Assuming energy costs of $16.8 per GJ for compressor power and $7.78 per GJ for low-pressure steam, the Xia et al process energy cost is 16% higher than the energy cost of the proposed process.
3. Optimum Distillate Composition There are two important design optimization variables in heat-integrated pressureswing distillation systems. The first is the selection of the two operating pressures, which affects the differences in the two azeotropic compositions and the temperature differentials for heat transfer. The second critical variable is the closeness of the distillate composition to the azeotropic composition. This ∆x = yaz – xD specification has a very significant effect on the flowrates of the distillate streams and on the difficulty of the separation (the required reflux ratio). The design trade-off must balance these two competing effects. A small ∆x means the separation requirement are larger, which results in higher reflux ratios. Vapor boilup and compressor power requirements tend to increase. On the other hand, a small ∆x results in smaller distillate flowrates, as discussed below. A logical criterion for selecting the optimum ∆x is to determine the value that minimizes compressor power requirements. First let us analyze the overall system, which is binary so only two material balances can be applied. The overall balances given below apply for any selection of ∆x. F B1 B2 Fz B1 xB1 B2 xB 2
If the feed flowrate (F), feed composition (z) and the two bottoms product compositions (xB1 and xB2) are given, the two product flowrates (B1 and B2) can be calculated. These values apply for any pressure selection, any distillation configuration and any choice of distillate compositions.
6
x z B1 F B 2 xB 2 xB1 B2 F B1 Now the total and component balances for just the HPC column C2 can be combined to find the important relationship that shows how the recycle flowrate D2 is affected by the selection of the two distillate compositions. D1 D2 B2 D1 xD1 D2 xD 2 B2 xB 2
Combining these two equations gives: D1
D2 xD 2 B2 xB 2 xD1
D2 B2
D2 xD 2 B2 xB 2 xD1
Solving for the recycle flowrate D2 gives:
x x D2 B2 B 2 D1 xD1 xD 2 The recycle flowrate is strongly affected by the difference between the selected distillate compositions. The smaller this difference is, the larger the recycle flowrate. It should be emphasized that this material balance relationship clearly shows that small changes in the distillate compositions will strongly affected the recycle flowrate. Now the distillate compositions must be related to the azeotropic compositions at the two pressures. The LPC distillate composition must be somewhat smaller than the azeotropic composition at 1 atm. The HPC distillate composition must be somewhat larger than the azeotropic composition at 4 atm. We define this difference as ∆x, which for simplicity is assumed to be the same for both columns.
xD1 y1az x 0.7743 x xD 2 y2az x 0.6136 x Figure 4 shows how a number of important variables are affected by the selection of ∆x. Product purities (xB1 and xB2) are kept constant for all cases, so the product flowrates B1 and B2 are also constant. The upper left graph demonstrates that the distillate flowrates increase as ∆x is increased. However, since the separation is easier, the reflux
7
flowrates R1 and R2 decrease. The net effect is non-monotonic changes in the overhead vapor flowrates through the compressors. A ∆x of 1.5 mol% gives the minimum total compressor power.
4. Simulation Issues Any process with recycle streams can present simulation convergence problems. The procedure used in this work is to “tear” three streams in the steady-state simulation (D2, R1 and R2) as shown in Figure 5. For a given ∆x case, the flowrate and composition of D2 are known, so these are fixed. Two Aspen flowsheet design specs are used to achieve the two specified distillate compositions by varying the two reflux flowrates R1 and R2. These tear streams are closed in Aspen Dynamics. The heat-transfer relationships are handled in Aspen Dynamics by using Flowsheet Equations, as given in Figure 6. In the LPC, the net reboiler duty is calculated by subtracting from the heat transferred in the “REB1” block the heat removed in the auxiliary condenser (the output signal from the temperature controller in C1). In the HPC, the total reboiler duty is calculated by adding to the heat transferred in the “REB2” block the heat added in the auxiliary reboiler (the output signal from the temperature controller in C2).
5. Plantwide Control Many process systems with recycles can be sensitive to small changes in throughput that can produce large changes recycle flowrates. This “snowballing” phenomenon8 has been reported in a variety of systems and led to the suggestion made two decades ago of using a control structure that places a flow controller somewhere in the liquid recycle loop. A number of papers have discussed the control of various types of pressure-swing distillation systems with and without heat integration. Minimum-boiling azeotropes have distillate recycles. Maximum-boiling azeotropes have bottoms recycles. A review of many aspects of pressure-swing distillation, including control structures, has been
8
recently presented9. A pioneering paper10 proposed a control structure for the minimumboiling THF/water azeotropic separation with partial heat integration. The snowball effect was not observed despite having no flow controller in the distillate recycle loop. Successful control of the dimethyl carbonate/methanol separation was reported11 using a control structure without a flow controller in the distillate recycle loop, but heat integration was not considered. Ye et al12 studied a partially heat-integrated pressureswing distillation system separating the maximum-boiling ethylene diamine/water azeotropic mixture. The proposed control structure fixed the ratio of the bottoms recycle from the LPC to the fresh feed, so a flow controller was in the recycle loop. However, their control scheme appears to not control the level in the base of the LPC, so the effectiveness of this plantwide control structure is questionable. The process studied in this paper does not use heat integration between the two columns. Vapor recompression of the overhead vapor from each column is used to provide heat to each reboiler. It is not unexpected to find that a different control structure will be required with this new process configuration. A number of alternative control structures were studied. The only structure that provided robust and effective control is described below and shown in Figure 7. All of the other schemes resulted in snowballing with the recycle D2 stream undergoing very large changes when small disturbances were introduced into the system, which resulted in valve and reboiler duty saturation. The key feature of the proposed control structure is to flow control the recycle flowrate D2 and use it as the throughput manipulator. Then the fresh feed is brought into the system to control the level in the reflux drum of the first column. The location of the temperature control trays are selected by looking at the temperature profiles given in Figure 8. Level controllers are proportional with KC = 2. Pressure controller tuning used the Aspen default values. Temperature controllers had 1-minute deadtimes and were tuned using relay-feedback tests and Tyreus-Luyben tuning rules. The loops are enumerated below. 1. Recycle D2 is flow controlled. 2. Reflux-drum level in C2 is controlled by manipulating the flowrate D1, which is the feed to C2. 3. Reflux-drum level in C1 is controlled by manipulating the flowrate F.
9
4. Base levels are controlled by manipulating bottoms flowrates. 5. Column pressures are controlled by manipulating compressor power. 6.
Reflux flowrates are ratioed to distillate flowrates.
7. The temperature on Stage 16 in C1 is controlled by manipulating heat removal in the auxiliary condenser. 8. The temperature on Stage 25 in C2 is controlled by manipulating heat input in the auxiliary reboiler.
Figures 9 and 10 give the responses of the system to large disturbances in throughput and feed composition. In Figure 9 the setpoint of the recycle D2 flow controller is changed by 20% at 0.5 hours. Solid lines are 20% increases, and dashed lines are 20% decreases. Both product compositions are held close to their specifications. Note that a 20% change in recycle D2 produces about a 20% change in fresh feed F. Both temperatures are well controlled. An increase in throughput results in more compressor power and larger duties in the two reboilers, the auxiliary reboiler and the auxiliary condenser. The response of the fresh feed to changes in the throughput manipulator recycle D2 is quite interesting. When the recycle D2 is increased, the initial response of the fresh feed is to decrease before eventually increasing to the new steady-state flowrate. This inverse response behavior can be explained by remembering that the D2 distillate from the HPC is at a higher temperature and pressure than the conditions in the LPC. So an increase in D2 increases the vapor going to the LPC condenser and results in a temporary increase in the reflux-drum level. The level controller reduces the fresh feed (top right graph in Figure 9). However, the increase in D2 (and R2 since it is ratioed to D2) causes the reflux-drum level in C2 to drop, which eventually increases D1 (and R1). This starts to drop the level in the C1 reflux drum and pulls in more fresh feed. Figure 10 gives responses to changes in feed composition. Solid lines are when DIPE composition is increased from 25 to 30 mol% with a corresponding reduction in IPA. Dashed lines are when DIPE composition is decreased from 25 to 20 mol% with a corresponding increase in IPA. Keep in mind that the throughput manipulator D2 is constant for these feed composition disturbances. The control structure adjusts the
10
flowrate of fresh feed to handle the changes in feed composition. An increase in DIPE feed composition results in less fresh feed (top right graph in Figure 10) fed to the LPC . Putting more DIPE into the system increases the separation load in the HPC where this component is removed as a bottoms product. With fixed distillate and reflux flowrates in this column, its feed must be reduced to achieve the required bottoms purity. Therefore the distillate and reflux streams in the LPC are reduced, which results in lower fresh feed flowrates. Note that there is more heat removed in the auxiliary cooler in the LPC (Qcool in the 4th right graph in Figure 10). The heat duties in the two auxiliary heat exchangers are relatively small percentages of the total duties (4.9% in the condenser and 12% in the reboiler) and could saturate during severe transients. However, process is demonstrated to handle quite large disturbances in both throughput and feed composition.
Conclusion A pressure-swing distillation process that uses vapor recompression on both columns has been studied in terms of both steady-state design and dynamic controllability. The proposed process has lower energy cost than the process in the published paper. A robust plantwide control structure is demonstrated to effectively handle large disturbances in throughput and feed composition. The control scheme features the distillate recycle flowrate as the throughput variable.
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References 1. Muhrer, C. A., Collura, M. A., Luyben, W. L. “Control of Vapor-Recompression Distillation Columns'', Ind. Eng. Chem. Research 29 (1990) 59. 2. Li, R., Ye, Q., Suo, X., Dai, X., Yu, H., Feng, S., Xia, H. “Improving the performance of heat pump-assisted azeotropic dividing wall distillation” Ind. Eng. Chem. Res. 55 (2016) 6454-6464. 3. Luyben, W. L. “Control of Azeotropic DWC with Vapor Recompression” Chemical Engineering and Processing: Process Intensification 109 (2016) 114124. 4. Luo, H., Bildea, C. S., Kiss, A. A. “Novel heat-pump-assisted extractive distillation for bioethanol purification” Ind. Eng. Chem. Res. 54 (2015) 22082213. 5. Patrascu, I., Bildea, C. S., Kiss, A. A. “Dynamics and control of a heat-pumpassisted extractive dividing-wall column for bioethanol dehydration” Chem. Eng. Res. Des. 119 (2017) 66-74. 6. Luyben, W. L. “Improved plantwide control structure for extractive divided-wall columns with vapor recompression” Chem. Eng. Res. Des. 123 (2017) 152-164. 7. Xia,H., Ye, Q., Feng, S., L, R., Suo, X. “A novel energy-saving pressure seing distillaion process based on self-heat recuperation technology” Energy (2017), submitted 8. Luyben, W. L. “Snowball Effect in Reactor/Separator Processes with Recycle” Ind. Eng. Chem. Res. 33 (1994) 299-305. 9. Liang, S., Cao, Y., Liu, X., Li, X, Zhoa, Y., Wang, Y., Wang, Y. “Insight into pressure-swing distillation from azeotropic phenomenon to dynamic control” Wang Chem. Eng. Res. Design 117 (2017) 318-335. 10. Abu-Eishah, S., Luyben, W. L. “Design and control of two-coulmn azeotropic distillation system” IEC Proc. Des. Dev. 24 (1985) 132-140. 11. Wei, H., Wang, F, Zhang, J., Liao, N., Xiao, F., Wei, W., Sun, Y. “Design and control of dimethyl carbonate-methnaol separation via pressure-sweing distillation” Ind. Eng. Chem. Res. 52 (2013) 11463-11478.
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12. Li, R., Ye, Q., Suo, X., Dai, X., Yu, H. “Heat-integrated pressure-swing distillation process for separation of a maximum-boiling azeotrope ethylene diamine/water” Ye Chem. Eng. Res. Des. 105 (2016) 1- 15. 13.
Figure Captions Figure 1 – Txy diagrams at 1.01 and 4.05 bar Figure 2 – xy diagram Figure 3 – Flowsheet Figure 4 – Effect of ∆x Figure 5 – Aspen Plus PFD with torn streams Figure 6 – Aspen Dynamics flowsheet equations Figure 7 – Control structure Figure 8 – Temperature profiles Figure 9 – 20% changes in D2 Figure 10 - Feed composition disturbances
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S0255-2701(17)30869-3 https://doi.org/10.1016/j.cep.2017.09.020 CEP 7084
To appear in:
Chemical Engineering and Processing
Received date: Revised date: Accepted date:
28-8-2017 28-9-2017 29-9-2017
Please cite this article as: William L.Luyben, Design and Control of a PressureSwing Distillation Process with Vapor Recompression, Chemical Engineering and Processing https://doi.org/10.1016/j.cep.2017.09.020 This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.
Submitted to Chemical Engineering and Processing: Process Intensification
Design and Control of a PressureSwing Distillation Process with Vapor Recompression
William L. Luyben
Department of Chemical Engineering Lehigh University Bethlehem, PA 18015 USA
September 1, 2017 Revised September 28, 2017
[email protected]; 610-758-4256; FAX 610-758-5057
High Lights
Pressure-swing distillation using vapor recompression in both columns is studied. An effective control structure is developed for this unique process. Energy costs are 16% lower than those given in a published paper.
Abstract Pressure-swing distillation is widely used to separate minimum-boiling azeotropes when the azeotropic composition has significant pressure dependence. The two columns operate at different pressures with distillate streams having compositions close to their respective azeotropes. Heat integration is typically used since column temperatures are sufficiently different to permit heat transfer. A novel modification to achieve process intensification has recently been proposed that uses vapor recompression on both columns. The study developed a complex heat exchanger network along with the design of the two columns. Dynamic controllability was not studied. The purpose of this paper is to develop a robust control structure for this process. A novel control scheme features the flowrate of the distillate recycle from the highpressure column as the throughput manipulator with the fresh feed brought in on level control.
Key Words Pressure-swing distillation; azeotropic distillation; vapor recompression
1. Introduction Vapor recompression (heat pumping) has many applications in a wide variety of domestic and industrial uses. The basic thermodynamic principle of operation is to use work to produce temperature differences that permit heat transfer. The home refrigerator
2
is the most widely applied and most simple example in which high-value electrical energy is used to achieve transfer of heat from the low temperature inside the refrigerator to the warmer air in the kitchen. The closed cycle of the circulating refrigerant can be visually conveyed on an enthalpy-pressure diagram with energy flowing in from the lowtemperature source (air in refrigerator) and energy flowing out to the high-temperature sink (room air). Work is required to raise the pressure of the vapor refrigerant so it can be condensed by the high-temperature sink. Then expanding the high-pressure liquid refrigerant to a low pressure produces a partially vaporized low-temperature stream that can be completely vaporized by heat from the source. The final step is compressing the refrigerant vapor from the low-pressure “evaporator” to the high pressure in the “condenser.” Early industrial applications that involve vapor recompression of process streams include distillation and evaporation. The separation of components with similar boiling points (low relative volatility) such as propylene/propane1 can be achieved in a distillation column by compressing the overhead vapor and condensing it in the reboiler to provide vapor boilup in the column. Several recent papers have studied vapor recompression in more complex distillation systems. Li et al2 proposed using vapor recompression to provide the heat required in one of the reboilers of a divided-wall column separating a heterogeneous azeotrope. Effective dynamic control of this system has been demonstrated3. Luo et al4 proposed the use of vapor recompression in a single divided-wall column for bioethanol dehydration, and Patrascu et al5 developed a control system for this complex interacting system. An improved process design and a more robust control structure for this process has been presented6. The use of vapor recompression in pressure-swing distillation systems has recently been proposed by Xia et al7. This is a novel switch from conventional pressureswing distillation in which heat integration is typically used between the two columns operating at two pressure and temperature levels. The economics of a conventional pressure-swing distillation process with and without heat integration were compared with the vapor-recompression design. The process with vapor recompression had a much higher capital investment (73 %) compared to the heat integrated case because of the two
3
expensive compressors. However the energy cost is much smaller (48 %), producing a modest reduction in total annual cost of 8.6 %. A complex heat-exchanger network was proposed in the Xia et al7 paper that used both the sensible heat and the latent heat of the compressed vapor streams to achieve “self heating” with no external source of heat required. The only energy cost was the compressor power of the two compressors. Complex process flow patterns were proposed with two process streams split for preheating in parallel heat exchangers with different heat source streams. A total of nine separate heat exchangers were used. A much more simple heat-exchanger system is used in this paper.
2. Process Studied The numerical example studied by Xia et al7 is the separation of di-isopropyl ether (DIPE) and isopropyl alcohol (IPA) whose normal boiling points are 63.3 oC and 82.26 oC, respectively. These components form minimum-boiling azeotropes as shown in Figure 1. At 1.01 bar, the composition of the azeotrope is 77.43 mol% DIPE at 65.97 oC. At 4.05 bar, the composition of the azeotrope is 61.36 mol% DIPE at 110.8 oC. The reasonable pressure dependence of the azeotropic composition permits the economic use of pressure-swing distillation to achieve high-purity products. Aspen steady-state and dynamic simulations are used in this study with NRTL physical properties. Figure 2 gives the xy-diagram for the two columns with the feed and products streams of the two columns. With a feed of 25 mol% DIPE, the bottoms from the lowpressure column C1 is high-purity IPA with composition xB1. The distillate D1 has a composition xD1 slightly smaller than the azeotropic composition at 1.01 bar. The distillate is fed to the high-pressure column C2 in which the bottoms is high-purity DIPE with composition xB2. The distillate has a composition xD2 slightly larger than the azeotropic composition at 4.04 bar. The approach composition difference (yaz – xD) is a key design optimization variable that is discussed in the next section. The distillate D2 from the high-pressure column is recycled back to the low-pressure column. Figure 3 gives the flowsheet of the pressure-swing distillation process with vapor recompression studied in this paper. The column configurations are those used in the Xia
4
et al7 paper (number of trays, feed locations and column pressures), but a more simple heat exchanger network is used to avoid dynamic controllability problems. The overhead vapor from each column is compressed to a pressure that is high enough to provide a differential temperature driving force ∆T in the two reboilers of 5 oC. This is the design ∆T used by Xia et al7 and is reasonable because expensive compression is involved, which pushes the design trade-off between heat exchanger area and compressor power to large area and small ∆T to reduce compression costs. The low-pressure column operates at 1.01 bar and has 20 stages with 0.01 bar pressure drop per stage. The bottoms product is high-purity IPA, giving a base temperature of 86 oC. Therefore the condensing temperature on the hot side of the LPC reboiler must be 91 oC. The compressor discharge pressure that gives this temperature of the LPC overhead vapor is 2.43 bar. The high-pressure column operates at 4.05 bar and has 28 stages with 0.01 bar pressure drop per stage. The bottoms product is high-purity DIPE, giving a base temperature of 121 oC. Therefore the condensing temperature on the hot side of the HPC reboiler must be 126 oC. The compressor discharge pressure that gives this temperature of the HPC overhead vapor is 6.08 bar. The energy requirement to achieve the specified separation in the LPC is 1101 kW, but the heat required to desuperheat and condense the stream coming from the compressor K1 exceeds this amount by 57.0 kW. Therefore, an auxiliary condenser is installed using cooling water. The flowrate of the cooling water will be used to control a temperature in the LPC. The energy requirement to achieve the specified separation in the HPC is 767 kW, but the heat released in desuperheating and condensing the stream coming from the compressor K2 is only 677 kW. A steam-heated auxiliary reboiler is installed to provide the required additional vapor boilup (90 kW). The flowrate of the steam will be used to control a temperature in the HPC. The total energy requirement of the process shown in Figure 3 is 171 kW of expensive electric power to drive the compressors and 90 kW of inexpensive lowpressure steam. It is important to note that the compressor power requirements of the Xia et al7 paper were 248.1 kW. So the proposed process shown in Figure 3 achieves the
5
separation with 31% less high-level compressor power requirements. However, the proposed process requires 90 kW of energy in the auxiliary reboiler. Assuming energy costs of $16.8 per GJ for compressor power and $7.78 per GJ for low-pressure steam, the Xia et al process energy cost is 16% higher than the energy cost of the proposed process.
3. Optimum Distillate Composition There are two important design optimization variables in heat-integrated pressureswing distillation systems. The first is the selection of the two operating pressures, which affects the differences in the two azeotropic compositions and the temperature differentials for heat transfer. The second critical variable is the closeness of the distillate composition to the azeotropic composition. This ∆x = yaz – xD specification has a very significant effect on the flowrates of the distillate streams and on the difficulty of the separation (the required reflux ratio). The design trade-off must balance these two competing effects. A small ∆x means the separation requirement are larger, which results in higher reflux ratios. Vapor boilup and compressor power requirements tend to increase. On the other hand, a small ∆x results in smaller distillate flowrates, as discussed below. A logical criterion for selecting the optimum ∆x is to determine the value that minimizes compressor power requirements. First let us analyze the overall system, which is binary so only two material balances can be applied. The overall balances given below apply for any selection of ∆x. F B1 B2 Fz B1 xB1 B2 xB 2
If the feed flowrate (F), feed composition (z) and the two bottoms product compositions (xB1 and xB2) are given, the two product flowrates (B1 and B2) can be calculated. These values apply for any pressure selection, any distillation configuration and any choice of distillate compositions.
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x z B1 F B 2 xB 2 xB1 B2 F B1 Now the total and component balances for just the HPC column C2 can be combined to find the important relationship that shows how the recycle flowrate D2 is affected by the selection of the two distillate compositions. D1 D2 B2 D1 xD1 D2 xD 2 B2 xB 2
Combining these two equations gives: D1
D2 xD 2 B2 xB 2 xD1
D2 B2
D2 xD 2 B2 xB 2 xD1
Solving for the recycle flowrate D2 gives:
x x D2 B2 B 2 D1 xD1 xD 2 The recycle flowrate is strongly affected by the difference between the selected distillate compositions. The smaller this difference is, the larger the recycle flowrate. It should be emphasized that this material balance relationship clearly shows that small changes in the distillate compositions will strongly affected the recycle flowrate. Now the distillate compositions must be related to the azeotropic compositions at the two pressures. The LPC distillate composition must be somewhat smaller than the azeotropic composition at 1 atm. The HPC distillate composition must be somewhat larger than the azeotropic composition at 4 atm. We define this difference as ∆x, which for simplicity is assumed to be the same for both columns.
xD1 y1az x 0.7743 x xD 2 y2az x 0.6136 x Figure 4 shows how a number of important variables are affected by the selection of ∆x. Product purities (xB1 and xB2) are kept constant for all cases, so the product flowrates B1 and B2 are also constant. The upper left graph demonstrates that the distillate flowrates increase as ∆x is increased. However, since the separation is easier, the reflux
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flowrates R1 and R2 decrease. The net effect is non-monotonic changes in the overhead vapor flowrates through the compressors. A ∆x of 1.5 mol% gives the minimum total compressor power.
4. Simulation Issues Any process with recycle streams can present simulation convergence problems. The procedure used in this work is to “tear” three streams in the steady-state simulation (D2, R1 and R2) as shown in Figure 5. For a given ∆x case, the flowrate and composition of D2 are known, so these are fixed. Two Aspen flowsheet design specs are used to achieve the two specified distillate compositions by varying the two reflux flowrates R1 and R2. These tear streams are closed in Aspen Dynamics. The heat-transfer relationships are handled in Aspen Dynamics by using Flowsheet Equations, as given in Figure 6. In the LPC, the net reboiler duty is calculated by subtracting from the heat transferred in the “REB1” block the heat removed in the auxiliary condenser (the output signal from the temperature controller in C1). In the HPC, the total reboiler duty is calculated by adding to the heat transferred in the “REB2” block the heat added in the auxiliary reboiler (the output signal from the temperature controller in C2).
5. Plantwide Control Many process systems with recycles can be sensitive to small changes in throughput that can produce large changes recycle flowrates. This “snowballing” phenomenon8 has been reported in a variety of systems and led to the suggestion made two decades ago of using a control structure that places a flow controller somewhere in the liquid recycle loop. A number of papers have discussed the control of various types of pressure-swing distillation systems with and without heat integration. Minimum-boiling azeotropes have distillate recycles. Maximum-boiling azeotropes have bottoms recycles. A review of many aspects of pressure-swing distillation, including control structures, has been
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recently presented9. A pioneering paper10 proposed a control structure for the minimumboiling THF/water azeotropic separation with partial heat integration. The snowball effect was not observed despite having no flow controller in the distillate recycle loop. Successful control of the dimethyl carbonate/methanol separation was reported11 using a control structure without a flow controller in the distillate recycle loop, but heat integration was not considered. Ye et al12 studied a partially heat-integrated pressureswing distillation system separating the maximum-boiling ethylene diamine/water azeotropic mixture. The proposed control structure fixed the ratio of the bottoms recycle from the LPC to the fresh feed, so a flow controller was in the recycle loop. However, their control scheme appears to not control the level in the base of the LPC, so the effectiveness of this plantwide control structure is questionable. The process studied in this paper does not use heat integration between the two columns. Vapor recompression of the overhead vapor from each column is used to provide heat to each reboiler. It is not unexpected to find that a different control structure will be required with this new process configuration. A number of alternative control structures were studied. The only structure that provided robust and effective control is described below and shown in Figure 7. All of the other schemes resulted in snowballing with the recycle D2 stream undergoing very large changes when small disturbances were introduced into the system, which resulted in valve and reboiler duty saturation. The key feature of the proposed control structure is to flow control the recycle flowrate D2 and use it as the throughput manipulator. Then the fresh feed is brought into the system to control the level in the reflux drum of the first column. The location of the temperature control trays are selected by looking at the temperature profiles given in Figure 8. Level controllers are proportional with KC = 2. Pressure controller tuning used the Aspen default values. Temperature controllers had 1-minute deadtimes and were tuned using relay-feedback tests and Tyreus-Luyben tuning rules. The loops are enumerated below. 1. Recycle D2 is flow controlled. 2. Reflux-drum level in C2 is controlled by manipulating the flowrate D1, which is the feed to C2. 3. Reflux-drum level in C1 is controlled by manipulating the flowrate F.
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4. Base levels are controlled by manipulating bottoms flowrates. 5. Column pressures are controlled by manipulating compressor power. 6.
Reflux flowrates are ratioed to distillate flowrates.
7. The temperature on Stage 16 in C1 is controlled by manipulating heat removal in the auxiliary condenser. 8. The temperature on Stage 25 in C2 is controlled by manipulating heat input in the auxiliary reboiler.
Figures 9 and 10 give the responses of the system to large disturbances in throughput and feed composition. In Figure 9 the setpoint of the recycle D2 flow controller is changed by 20% at 0.5 hours. Solid lines are 20% increases, and dashed lines are 20% decreases. Both product compositions are held close to their specifications. Note that a 20% change in recycle D2 produces about a 20% change in fresh feed F. Both temperatures are well controlled. An increase in throughput results in more compressor power and larger duties in the two reboilers, the auxiliary reboiler and the auxiliary condenser. The response of the fresh feed to changes in the throughput manipulator recycle D2 is quite interesting. When the recycle D2 is increased, the initial response of the fresh feed is to decrease before eventually increasing to the new steady-state flowrate. This inverse response behavior can be explained by remembering that the D2 distillate from the HPC is at a higher temperature and pressure than the conditions in the LPC. So an increase in D2 increases the vapor going to the LPC condenser and results in a temporary increase in the reflux-drum level. The level controller reduces the fresh feed (top right graph in Figure 9). However, the increase in D2 (and R2 since it is ratioed to D2) causes the reflux-drum level in C2 to drop, which eventually increases D1 (and R1). This starts to drop the level in the C1 reflux drum and pulls in more fresh feed. Figure 10 gives responses to changes in feed composition. Solid lines are when DIPE composition is increased from 25 to 30 mol% with a corresponding reduction in IPA. Dashed lines are when DIPE composition is decreased from 25 to 20 mol% with a corresponding increase in IPA. Keep in mind that the throughput manipulator D2 is constant for these feed composition disturbances. The control structure adjusts the
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flowrate of fresh feed to handle the changes in feed composition. An increase in DIPE feed composition results in less fresh feed (top right graph in Figure 10) fed to the LPC . Putting more DIPE into the system increases the separation load in the HPC where this component is removed as a bottoms product. With fixed distillate and reflux flowrates in this column, its feed must be reduced to achieve the required bottoms purity. Therefore the distillate and reflux streams in the LPC are reduced, which results in lower fresh feed flowrates. Note that there is more heat removed in the auxiliary cooler in the LPC (Qcool in the 4th right graph in Figure 10). The heat duties in the two auxiliary heat exchangers are relatively small percentages of the total duties (4.9% in the condenser and 12% in the reboiler) and could saturate during severe transients. However, process is demonstrated to handle quite large disturbances in both throughput and feed composition.
Conclusion A pressure-swing distillation process that uses vapor recompression on both columns has been studied in terms of both steady-state design and dynamic controllability. The proposed process has lower energy cost than the process in the published paper. A robust plantwide control structure is demonstrated to effectively handle large disturbances in throughput and feed composition. The control scheme features the distillate recycle flowrate as the throughput variable.
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References 1. Muhrer, C. A., Collura, M. A., Luyben, W. L. “Control of Vapor-Recompression Distillation Columns'', Ind. Eng. Chem. Research 29 (1990) 59. 2. Li, R., Ye, Q., Suo, X., Dai, X., Yu, H., Feng, S., Xia, H. “Improving the performance of heat pump-assisted azeotropic dividing wall distillation” Ind. Eng. Chem. Res. 55 (2016) 6454-6464. 3. Luyben, W. L. “Control of Azeotropic DWC with Vapor Recompression” Chemical Engineering and Processing: Process Intensification 109 (2016) 114124. 4. Luo, H., Bildea, C. S., Kiss, A. A. “Novel heat-pump-assisted extractive distillation for bioethanol purification” Ind. Eng. Chem. Res. 54 (2015) 22082213. 5. Patrascu, I., Bildea, C. S., Kiss, A. A. “Dynamics and control of a heat-pumpassisted extractive dividing-wall column for bioethanol dehydration” Chem. Eng. Res. Des. 119 (2017) 66-74. 6. Luyben, W. L. “Improved plantwide control structure for extractive divided-wall columns with vapor recompression” Chem. Eng. Res. Des. 123 (2017) 152-164. 7. Xia,H., Ye, Q., Feng, S., L, R., Suo, X. “A novel energy-saving pressure seing distillaion process based on self-heat recuperation technology” Energy (2017), submitted 8. Luyben, W. L. “Snowball Effect in Reactor/Separator Processes with Recycle” Ind. Eng. Chem. Res. 33 (1994) 299-305. 9. Liang, S., Cao, Y., Liu, X., Li, X, Zhoa, Y., Wang, Y., Wang, Y. “Insight into pressure-swing distillation from azeotropic phenomenon to dynamic control” Wang Chem. Eng. Res. Design 117 (2017) 318-335. 10. Abu-Eishah, S., Luyben, W. L. “Design and control of two-coulmn azeotropic distillation system” IEC Proc. Des. Dev. 24 (1985) 132-140. 11. Wei, H., Wang, F, Zhang, J., Liao, N., Xiao, F., Wei, W., Sun, Y. “Design and control of dimethyl carbonate-methnaol separation via pressure-sweing distillation” Ind. Eng. Chem. Res. 52 (2013) 11463-11478.
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12. Li, R., Ye, Q., Suo, X., Dai, X., Yu, H. “Heat-integrated pressure-swing distillation process for separation of a maximum-boiling azeotrope ethylene diamine/water” Ye Chem. Eng. Res. Des. 105 (2016) 1- 15. 13.
Figure Captions Figure 1 – Txy diagrams at 1.01 and 4.05 bar Figure 2 – xy diagram Figure 3 – Flowsheet Figure 4 – Effect of ∆x Figure 5 – Aspen Plus PFD with torn streams Figure 6 – Aspen Dynamics flowsheet equations Figure 7 – Control structure Figure 8 – Temperature profiles Figure 9 – 20% changes in D2 Figure 10 - Feed composition disturbances
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