Simulation and Optimization of Distillation Processes for Separating

Article pubs.acs.org/IECR

Simulation and Optimization of Distillation Processes for Separating the Methanol−Chlorobenzene Mixture with Separate Heat-Pump Distillation Xiaoxin Gao,†,‡ Zhengfei Ma,*,† Limin Yang,‡ and Jiangquan Ma‡ †

College of Chemistry and Chemical Engineering, Nanjing University of Technology, Nanjing 210009, P.R. China College of Petrochemical Engineering, Changzhou University, Changzhou 213164, P.R. China

‡

S Supporting Information *

ABSTRACT: The methanol−chlorobenzene mixture has a larger relative volatility in the low composition range than in the high composition range. Based on this characteristic, the mixture can be effectively separated by separate heat-pump distillation (SHPD) with significant energy savings. The binary interaction parameters of the UNIQUAC equation were used to predict the vapor−liquid equilibrium by means of the binary interaction parameters included in the Aspen Plus database. To minimize the overall annual operating costs, simulations for SHPD were carried out using Aspen Plus software, including the RadFrac and Compr blocks, and the optimal operating conditions, such as the split-point concentration, were determined. Simulations for conventional distillation, conventional heat-pump distillation, and multieffect distillation processes were also carried out for comparison. The simulated results showed that the SHPD process has obvious advantages over the other distillation processes in the assessment of energy savings and overall economic efficiency. specific distillation processes, such as extractive distillation,9 azeotropic distillatio,n10 and dividing-wall columns (DWCs),11 have also been the focus of energy-saving research because of their significant energy-saving effects. A large number of experimental studies and industrial practices have shown that heat-pump distillation has the most potential and might be the most effective technique because of its easy implementation, easy control, and high energy-saving efficiency compared with conventional distillation.12−14 Heatpump distillation has unique energy-saving characteristics in energy integration and intensification. Conventional heat-pump distillation is usually applied to a distillation column with a small temperature difference between the bottom and top of the column. For a special distillation column with a large temperature difference between the bottom and top, the direct heating of the bottom by compressing the top stream would cause excessive energy consumption by the compressor, which would result in an increase in the overall energy consumption and operating costs; therefore, this method would not meet the energy-saving principle of heat-pump distillation.15 However, if a distillation column were divided into upper and lower towers, heat-pump distillation could be applied to the upper tower because of the smaller temperature difference between its top and bottom, whereas the lower tower could retain the same function as in conventional distillation. Using this method, a separate heat-pump distillation (SHPD) process is formed,16 and the SHPD process can significantly save energy.

1. INTRODUCTION Global climate change has recently begun to affect human life, as global energy consumption continues to increase, especially in developing countries such as China and India. Many comprehensive energy reduction programs are being pursued to reduce energy usage and promote sustainable development for the near future. In the chemical industry, energy plays a very important role because its consumption can reach as high as 10−12% of the total annual energy consumption. Distillation is widely used in the chemical industry as an important method for the separation of liquid or gas mixtures. Distillation processes represent approximately 3% of the world energy consumption.1 Therefore, research on how to save energy during distillation operation would provide not only efficiency and productivity but also environmental benefits. In conventional distillation, there are usually two sections: a rectifying section with a top condenser in the upper part of the column and a stripping section with a reboiler in the lower part. The heat is added in the reboiler and removed at the top condenser. Because of the lower energy quality of the streams around the condenser, the energy usage is of little value even when these streams exchange heat with other fluids. Therefore, large amounts of energy are wasted, which lead to a low energy efficiency. Over the past few years, energy-saving distillation technologies have attracted increasing attention.2,3 Energysaving measures in distillation processes have been carried out to best use the supplied energy and to promote thermodynamic efficiency according to the first and second laws of thermodynamics and the principle of distillation. Many energy-saving techniques have been developed, including the application of inter-reboiler and intercondenser distillation,4 feedback stream distillation,5 heat-pump distillation,6 multieffect distillation,7 and thermally coupled distillation.8 Other © XXXX American Chemical Society

Received: May 10, 2013 Revised: July 15, 2013 Accepted: July 30, 2013

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dx.doi.org/10.1021/ie401467r | Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Industrial & Engineering Chemistry Research

Article

It is interesting to note that the methanol−chlorobenzene mixture has a large relative volatility in the low composition range and a small relative volatility in the high composition range. Taking these features into consideration, it is possible to design a distillation process to separate the mixture by dividing a conventional distillation column into separate upper and lower parts. The upper part uses heat-pump distillation, and the lower part uses conventional distillation; hence, two distillation columns are used, and an SHPD process is formed. As shown in Figure 2, the SHPD process contains two distillation columns. The upper tower is similar to the conventional direct-type heat-pump distillation but with an extra feed opening at the bottom. Because of the small temperature difference between the top and bottom streams, part of the top vapor can be recompressed and then used as the heat source for heating the bottom. To control the heat balance of the upper tower, an auxiliary condenser is set to condense the remainder of the top vapor. Because the temperature difference is quite small, the required recompression ratio is small. Thus, the required power consumption is also small. The lower tower is similar to the stripper section of a conventional distillation column. The liquid feed from the bottom of the upper tower enters the top of the lower tower, and the vapor from the reboiler of the lower tower enters at the bottom. As can be seen in Figure 1, the methanol−chlorobenzene mixture has a pure methanol-terminal clamp region at x = 0.8−1.0 (mass fraction). In the clamping zone, the composition lines of liquid (x) and vapor (y) are very close, as the temperatures of the liquid and vapor are very close and the relative volatility is near 1. In the x = 0−0.8 range, however, the relative volatility is large, so the reflux ratio in the x = 0−0.8 range is much smaller than that at x = 0.8−1.0. Hence, the reflux ratio of the lower tower is much smaller than that of the upper tower. Under these two energy-saving measures of heat-pump equipment in the upper tower and reduced reflux ratio for the lower tower, the energy-saving effect is quite obvious. The SHPD approach is poised to solve the issue of distillation with a large temperature difference between the top and bottom of a distillation column and broaden its applications.

In the present study, the simulation of an SHPD process for the separation of the methanol−chlorobenzene mixture was carried out using the universal quasichemical (UNIQUAC) equation for predicting the vapor−liquid equilibrium along with Aspen Plus software.17 The results were compared to optimize the distillation process with overall annual operating costs as the main objective function. The optimal operating conditions, such as the split-point concentration and equilibrium stages, for the SHPD process were determined. Other distillation processes, including conventional distillation, conventional heat-pump distillation, and multieffect distillation, were also simulated for the mixture. By comparing the results from the various processes, it was found that the SHPD process is the most suitable process for the separation of the mixture.

2. PROPOSAL OF SHPD PROCESS FOR THE METHANOL−CHLOROBENZENE MIXTURE A temperature versus composition [t−x(y)] phase diagram of the methanol−chlorobenzene mixture at atmospheric pressure is shown in Figure 1.

Figure 1. t−x(y) phase equilibrium diagram for the methanol− chlorobenzene system.

Figure 2. Schematic diagram of the separate heat-pump distillation process. B



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Table 1. Simulation Results for Different Split-Point Concentrationsa upper tower

lower tower

xmethanol

N

D (m)

Qc (kW)

Wm (kW)

Qp (kW)

Tb (°C)

N

D (m)

QR (kW)

0.3 0.35 0.4 0.45 0.5 0.55 0.6 0.65 0.68 0.7

40 40 40 40 40 40 40 40 40 40

0.843 0.843 0.843 0.843 0.843 0.8436 0.8436 0.8436 0.8436 0.8436

269.27 247.84 221.52 203.39 180.42 160.32 147.99 170.71 220.12 300.36

236.61 208.13 172.31 146.76 115.19 83.33 56.30 48.84 42.75 35.09

1343.00 1335.97 1326.47 1319.07 1310.47 1298.73 1284.00 1259.89 1216.44 1108.54

109.12 105.00 100.16 96.46 92.19 87.99 83.69 79.51 77.00 75.53

5 5 5 5 5 5 5 5 5 5

0.15 0.16 0.17 0.17 0.18 0.19 0.20 0.23 0.26 0.35

21.00 28.70 38.00 45.50 55.00 66.00 81.00 107.00 149.00 256.00

a

xmethanol, split-point concentration; N, number of stages; D, diameter of tower; Qc, load of condenser; Wm, input power of compressor; Qp, load of heat pump; Tb, bottom temperature; Qr, load of reboiler.

3. FORMULATION OF THE PROCESSES 3.1. Rule of Simulation Process. A mixture of methanol and chlorobenzene (35:65, w/w) was fed at a rate of 5000 kg/h at ambient temperature and pressure. The required methanol product purity at the top of the upper tower was 99.5% or above, and the required chlorobenzene product purity at the bottom of the lower tower was 99.5% or above. The operation was under atmospheric pressure, and cooling water was used to cool the top condenser. The reboiler was heated by saturated steam (0.2 MPa). To compare the calculated results from the proposed SHPD approach, the results for the conventional single-tower distillation process under the same conditions were calculated. The equilibrium data and enthalpy were obtained using Aspen Plus software with the UNIQUAC property model. The simulation of the distillation column and compressor was achieved using the RadFrac and Compr blocks, respectively (compressor, etc., isentropic efficiency of 0.75). The simulation of the heat exchangers was performed with the Heatx block. The temperature difference between the compressor outlet and the bottom of the upper tower was limited to within 10 °C. 3.2. Formulation of SHPD Process. SHPD does not simply divide a tower into two towers or set the rectifying section as the upper tower and the stripping section as the lower tower. Instead, SHPD uses a split point to divide the conventional tower into two towers. The split point where the towers are divided will greatly affect the capital investment and operating costs. Therefore, the concentration for the split point is a key parameter for SHPD. The split point can be regulated by changing the reboiler duty of the lower tower. We simulated the effect of the split-point concentration on the equilibrium stages, the loads of the condenser and reboiler, the compressor power, and the overall annual operating costs as reported in Table 1. For a fixed separation requirement, the higher the concentration at the separation point, the smaller the temperature difference in the upper tower, and hence, the more obvious the energy-saving effects and the lower the operating costs for the upper tower. However, the separating duty for the lower tower would increase, resulting in an increase in the operating costs and capital investment for the lower tower. Therefore, the optimal split point is equal to the split point where the overall operating costs is the lowest, as shown in Figure 3.

Figure 3. Simulated relationship between the overall annual operating costs and the split-point concentration.

From an economic point of view, the split point for SHPD can be determined by examining the overall annual operating costs. The overall annual operating costs (Ctot) contain the following two parts: the total operating costs of the two towers [including the cost of reboiler steam (Cα), cooling water (Cβ), and compressor electricity (Cγ)] and the total annual cost of investment (Cδ) (annual depreciation charges for all of the equipment, including towers, heat exchangers, and compressors and other equipment). The cost of equipment and utilities was estimated with formulas obtained from the book by Luyben18 and listed in the Appendix. If the equipment lifetime is assumed to be 10 years and the annual operating time is 8000 h, then the cost calculation formulas are as follows Cα = 8000CsQ r /γ Cβ = 8000CwQ c/(6Cp) Cγ = 8000CmWm

where Cs is the price of steam ($/t), Qr is the load of the reboiler (kW), γ is the latent heat of steam (kJ/kg), Cw is the price of cooling water ($/t), Qc is the load of the condenser (kW), Cp is the specific heat capacity (kJ/kg·°C), Cm is the price of electricity ($/kW), and Wm is the input power of the compressor (kW). The objective function is set to the overall annual operating costs of SHPD and can be expressed as C



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min C tot = min(Cα + Cβ + Cγ + Cδ /10)

Using Aspen Plus simulation software, the relationship between the overall annual operating costs and the split-point concentration was simulated for the mixture system, as shown in Figure 3. It can be seen that, as the split-point concentration (xmethanol) increases, the overall annual operating costs decrease sharply until reaching the lowest value at a split-point concentration of approximately 0.6. As the split-point concentration increases further, the overall annual operating costs first increase gradually and then increase sharply after the concentration is over 0.65. Therefore, the optimal split-point concentration is 0.6, and its corresponding overall annual operating costs are $180,200 (see section S1 of the Supporting Information). 3.3. Temperature Distribution and Composition Distributions of Vapor and Liquid Phases within the Distillation Towers. By simulating the process using a splitpoint concentration of 0.6, the upper tower was found to have 40 equilibrium stages and the lower tower to have 5 equilibrium stages. For the SHPD process, the simulated temperature distribution along the two towers is shown in Figure 4, and the

Figure 5. Composition distributions of vapor and liquid in the two towers.

calculation (DSTWU) and a rigorous method (RadFrac) to establish the reflux ratio, number of stages, feed stage, and distillate mass flow rate (kg/h). The results are reported in Table S1 (Supporting Information). The specifications of the main streams of the conventional distillation process are detailed in Table S2 (Supporting Information). 3.5. Conventional Heat-Pump Distillation Process. The flowsheet for the conventional heat-pump distillation process simulated in Aspen Plus is shown in Figure 6. A method similar to that described in section 3.4 was used. The optimized parameters are reported in Table S3 (Supporting Information), and the specifications of the main streams of the conventional heat-pump distillation process are detailed in Table S4 (Supporting Information). 3.6. Multieffect Distillation Process. The methanol− chlorobenzene mixture can be separated using a multieffect distillation process because the boiling points of the two components vary obviously with pressure. Multieffect distillation is derived from a particular variant of heat-integrated processes. For binary distillation, it comprises two (or more) distillation columns, one running at high pressure and the other at lower pressure. The system is arranged in such way that the high-pressure condenser acts as the low-pressure reboiler. In this way, the reboiler for a tower under lower operating pressure can act as the condenser of another tower under higher operating pressure, and the latent heat of the top vapor released at the condenser can act as the heat supply for partially vaporizing the bottom liquid in the reboiler. Therefore, the multieffect distillation approach can greatly save energy. According to the relationship between the feed direction and the direction of the operating pressure gradient, the doubleeffect distillation process can be divided into three categories: parallel flow, concurrent flow, and countercurrent flow. For the separation of a methanol−chlorobenzene mixture, concurrent and countercurrent-flow double-effect distillation processes and concurrent and countercurrent-flow triple-effect distillation processes have been proposed. Quadruple-effect distillation processes have not been investigated because of its excessive capital investment and much higher bottom temperature. Having simulated all of these processes, we found that the countercurrent-flow triple-effect distillation process is the most

Figure 4. Temperature distributions of stages in the two towers.

composition distributions within the two towers are shown in Figure 5. It can be seen from Figures 4 and 5 that the temperature difference between the top and bottom of the upper tower is 19.14 °C, and the compositions of the vapor and liquid phases within the upper tower approach the azeotropic composition. Under these conditions, separation of the mixture requires a higher reflux ratio, higher energy consumption, and more equilibrium stages. Because of its small temperature difference, the heat-pump principle can be applied in this case. Figure 5 shows that the compositions of the vapor and liquid phases within the lower tower vary widely, and under these conditions, it is easy to separate the mixture, so both the required reflux ratio and the energy consumption are small. 3.4. Conventional Distillation Process. The flowsheet for the conventional distillation process simulated in Aspen Plus is shown in Figure S1 (Supporting Information). To establish the operating conditions for the conventional distillation process, a sensitivity analysis was carried out. The analyzed parameters were as follows: number of stages, reflux ratio, and feed stage. In addition, the conventional distillation was designed initially by means of a short-cut column D



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Figure 6. Aspen process diagram for the conventional heat-pump distillation column.

Figure 7. Aspen process diagram for the countercurrent-flow triple-effect distillation column.

suitable for energy savings, as seen in the Aspen process diagrams of Figure 7. The optimized parameters are reported in Table S5 (Supporting Information), and the specifications of the main streams of the countercurrent-flow triple-effect distillation process are detailed in Table S6 (Supporting Information).

Table 2. Comparison of Results for Different Distillation Processes distillation processes

parameter Qr (103 kW) Qc (103 kW) Wm (kW) stream (t/h) cooling water (t/h) Cα ($104/ year) Cβ ($104/ year) Cγ ($104/ year) total capital cost ($104/ year) annual operating cost ($104/ year) TAC ($104/ year) saved TAC ($104/ year)

4. COMPARISON AMONG DIFFERENT DISTILLATION APPROACHES To analyze and compare the energy-saving effects and comprehensive economic effects of the SHPD approach, we also simulated conventional distillation, conventional heatpump distillation, and triple-effect distillation, the results of which are shown in Table 2. From low to high, the operating costs were found to be in the order SHPD < triple-effect distillation < conventional heatpump distillation < conventional distillation. Compared with conventional distillation in a single tower, the capital investment of the conventional heat-pump distillation process increases by 31.48%, the overall annual operating costs decrease by 41.53%, and the energy saved is up to 76.6%, so the energy savings is obvious. For the concurrent-flow triple-effect distillation approach, the total capital investment increases by 102.96%, the overall annual operating costs decrease by 75.83%, and the energy saved is up to 62.88%. For the SHPD approach, the total capital investment decreases by 2.87%, the overall annual operating costs decrease by 83.77%, and the energy saved is up to 89.93%. In a comparison between the two heat-pump distillation approaches, the SHPD process E

conventional distillation

conventional heat-pump distillation

triple-effect countercurrent distillation

separate heat-pump distillation

1363.5 1376.06 − 2.31 197.9

− 330.13 318.77 − 47.34

506.12 580.57 − 0.86 83.25

81.0 147.99 56.3 0.14 21.2

56.9

−

13.37

3.4

19.5

4.66

5.1

2.1

−

39.23

−

6.9

57.78

75.97

117.27

56.12

76.4

43.89

18.47

12.4

82.18

51.49

30.20

18.01

−

30.69

51.98

64.17



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saves up to 71.75% of the overall annual operating costs relative to those of a conventional heat-pump distillation, and both the energy savings and the comprehensive economic effects are better than those of conventional processes. Compared with the concurrent-flow triple-effect distillation, the total capital investment of the SHPD approach decreases by 52.14%, and the overall annual operating costs decrease by 32.86%, with an energy savings of 72.87%. Furthermore, because the SHPD process has the additional advantages of ease of operation and control and higher energy savings, it is a more promising option for the separation of the methanol−chlorobenzene mixture.

5. CONCLUSIONS A separate heat-pump distillation process has been proposed to separate the methanol−chlorobenzene mixture based on an analysis of the mixture characteristics. This process and the other three distillation processes were simulated to compare their capital investments, energy usages, and overall operating costs, and the following conclusions were found: (1) The separate heat-pump distillation process has the best energy-saving effect and the best comprehensive economic effect among all of the examined distillation processes. (2) The separate heat-pump distillation process requires almost the same capital investment as the conventional distillation process and a lower capital investment than the other two processes. (3) The optimal split-point concentration for the separate heat-pump distillation is 60 wt % methanol in the mixture. (4) The separate heat-pump distillation process is easier to operate and control and has much better energy-saving and comprehensive economic effects than the concurrent-flow triple-effect distillation process.

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Additional information as noted in text. This material is available free of charge via the Internet at http://pubs.acs.org.

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*E-mail: [email protected]. Tel.: +86-519-86330255. Notes

The authors declare no competing financial interest.

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REFERENCES

(1) Hilde, K. E.; Sigurd, S. Selecting Appropriate Control Variables for a Heat Integrated Distillation System with Prefractionator. Comput. Chem. Eng. 2004, 28, 683−691. (2) Kansha, Y.; Tsuru, N.; Sato, K.; Fushimi, C.; Tsutsumi, A. SelfHeat Recuperation Technology for Energy Saving in Chemical Processes. Ind. Eng. Chem. Res. 2009, 48, 7682−7686. (3) Errico, M.; Tola, G.; Mascia, M. Energy Saving in a Crude Distillation Unit by a Preflash Implementation. Appl. Therm. Eng. 2009, 29, 1642−1647. (4) Lu, E. N.; Li, X. L.; Wu, Z. Inter-Reboiler and Inter-Condenser in Distillation. Chem. Eng. (China) 2008, 36, 74−78. (5) Luo, Y. Q.; Yuan, C. G.; Liu, C. J.; Yu, G. C. Simulation and Characteristic Analysis of Cryogenic Distillation. J. Chem. Ind. Eng. 2003, 56, 761−765. (6) Rivera-Ortega, P.; Picón-Núñez, M.; Torres-Reyes, E.; GallegosMuñoz, A. Thermal Integration of Heat Pumping Systems in Distillation Columns. Appl. Therm. Eng. 1999, 19, 819−829. (7) Engelien, H. K.; Skogestad, S. Multi-Effect Distillation Applied to an Industrial Case Study. Chem. Eng. Process.: Process Intensif. 2005, 44, 819−826. (8) Calzon-McConville, C. J.; Rosales-Zamora, M. B.; SegoviaHernández, J. G. Design and Optimization of Thermally Coupled Distillation Schemes for the Separation of Multicomponent Mixtures. Ind. Eng. Chem. Res. 2006, 45, 724−732. (9) Ligero, E. L.; Ravagnani, T. M. K. Dehydration of Ethanol with Salt Extractive DistillationA Comparative Analysis between Processes with Salt Recovery. Chem. Eng. Process.: Process Intensif. 2003, 42, 543−552. (10) Kansha, Y.; Tsuru, N.; Fushimi, C.; Tsutsumi, A. New Design Methodology Based on Self-Heat Recuperation for Production by Azeotropic Distillation. Energy Fuels 2010, 24, 6099−6102. (11) Yildirim, O.; Kiss, A. A.; Kenig, E. Y. Dividing Wall Columns in Chemical Process Industry: A Review on Current Activities. Sep. Purif. Technol. 2011, 80, 403−417. (12) Fonyo, Z.; Mizsey, P. Economic Application of Heat Pumps in Integrated Distillation Systems. Heat Recovery Syst CHP 1994, 14, 249−263. (13) Meszaros, I.; Fonyo, Z. Design strategy for heat pump-assisted distillation system. Heat Recovery Syst CHP 1986, 6, 469−476.

The height of a distillation column is calculated from the equation 1.2 3.281

The heat-transfer areas of the condenser and reboiler are calculated using the equations

Sr =

AUTHOR INFORMATION

Corresponding Author

Sizing and Economic Basis of the Various Process Schemes

Sc =

ASSOCIATED CONTENT

S Supporting Information *

APPENDIX

H = 2(N − 2)

Cδ = total annual costs of investment ($104/year) Ctot = overall annual operating costs ($104/year) D = diameter (m) H = height (m) N = number of stages Qc = heat duty of condenser (kW) Qr = heat duty of reboiler (kW) Sc = heat-transfer area of condenser (m2) Sr = heat-transfer area of reboiler (m2) SHPD = separate heat-pump distillation ΔT = temperature difference (°C) TAC = total annual costs ($104/year) U = overall heat-transfer coefficient (kW K−1m−2)

Qc UcΔTc

Qr Ur ΔTr

In terms of the preceding size estimations, the capital and energy costs of a distillation column are estimated using the equations column shell cost = 17640D1.066H 0.802 total heat exchanger cost = 7296Sc 0.65 + 7296Sr 0.65 Nomenclature

Cα = costs of reboiler steam ($104/year) Cβ = costs of cooling water ($104/year) Cγ = costs of compressor electricity ($104/year) F



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(14) Annakou, O.; Mizsey, P. Rigorous Investigation of Heat Pump Assisted Distillation. Heat Recovery Syst. CHP 1995, 15, 241−247. (15) Díez, E.; Langston, P.; Ovejero, G.; Romero, M. D. Economic Feasibility of Heat Pumps in Distillation to Reduce Energy Use. Appl. Therm. Eng. 2009, 29, 1216−1223. (16) Zhu, P.; Feng, X. Optimal Design and Optimal Operation of Separate Heat Pump Distillation. Can. J. Chem. Eng. 2003, 81, 963− 967. (17) Luyben, W. L. Distillation Design and Control Using Aspen Simulation; Wiley: New York, 2006. (18) Luyben, W. L.; Chien, I. L. Design and Control of Distillation for Separating Azeotropes; Wiley: New York, 2010.

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Simulation and Optimization of Distillation Processes for Separating

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