Optimization of Toluene-Benzene Reactor Aaron Vigil & Eric Vasko 10/28/2013 1
Introduction The production of benzene is an important process for the petrochemical industry. This chemical process will use toluene and hydrogen at elevated temperatures to react and form benzene and methane. At high temperatures, there is a side reaction that is undesirable which takes benzene and forms diphenyl and hydrogen. These reactions can be observed below. Despite the low rate of formation of diphenyl, the undesired products play a significant role in the economics of the system. The goal of this project is to achieve a 99.97% molar purity of benzene in the product stream and to produce and sell enough of the product to make a reasonable profit. This process will use an adiabatic plug flow reactor (PFR) to create the conditions necessary for a successful reaction. The rate law for the expression is as follows: A list of conversions and selectivities were given in the problem statement ranging from 50% to 85% and 99% to 93% respectively. Using the base-case economics given the optimum conversion and corresponding selectivity were found to be 60% and 98.5% respectively. This what the basis for all further calculations and design work. Summary The final design of this process was much larger than the initial estimates called for. The added process equipment to the original design were a compressor, additional flash drum, and additional distillation column. The desired purity of benzene could not be achieved without the additional separation equipment. The first step in solving this problem was to do mass balances around the different components to get the necessary sizes and capacities. The table below shows the comparison of the hand calculations and ASPEN results. 2
Table 1: Calculations for the capacities of the necessary equipment needed for the chemical process. These are the ASPEN results when the process was fully optimized and a product purity of 99.99% was achieved, exceeding the minim required for the process. Discussion Background The goal of this design was to create a viable and economic process plant for the hydrodealkylation of toluene to benzene. This is a highly valued process in today s modern chemical industry, and optimizing a process such as this can save and earn a company tens or hundreds of thousands of dollars. The base selling price for benzene range was given from $1.70 and $2.20 per gallon. We assumed that the PFR had a feed supply of 2,000 barrels per day of toluene and that there are 42 U.S. gallons in a barrel. Table A-1 located in the appendix section displays the preliminary economics for the initial process design. This table was used to create a base-case scenario so that the equipment could be properly sized and equipment cost determined. Optimization Once the base case had been modeled in Aspen, the process was optimized to obtain a pure product. The first task was to obtain a stronger separation of toluene and benzene. The initial design separated benzene from toluene in a distillation column at 450 psig. However, this leads to an azeotrope at higher temperatures as shown in Figure 1 below. This creates an undesirably high concentration of benzene in the bottoms and decreases product yield. Figure 2 shows that at lower pressures the separation is more favorable and no azeotrope is formed. This allows for a smaller distillation column and for a more complete separation of toluene and benzene. The cost of pressurizing the recycle stream is far outweighed by the savings in the column. Therefore, the distillation columns were operated at one atmosphere. 3
Figure 1: Txy for Benzene/Toluene @ 450 psig Figure 2: Txy for Benzene/Toluene @ 1 atm The next problem faced was the unfavorably high concentration of methane in the benzene product stream. Distillation of methane from benzene results in either an extreme cold temperature distillation (-200 C) or a loss of benzene in the methane waste stream. A second flash tank at one atmosphere was designed to further separate methane from the benzene stream. At this point a distillation column was used to finish the separation by producing a liquid benzene stream and a vapor methane stream. The reduced quantity of methane allowed the benzene vapor concentration to be large enough to raise the temperature without losing an appreciable amount of product. A final molar purity of 99.99% benzene was achieved. PFD & Description For the base case scenario shown in the process flow diagram below, Figure 3, the process equipment consisted of a fired heater, plug flow reactor (PFR), three heat exchangers, a flash drum, a recycle compressor, and two distillation columns. The fired heater was to run at 1150 F to prepare the feed streams for the upcoming reaction that takes place in the PFR. This was the only specification from the problem statement for this stage of the process. The next step in the design was to size the PFR. The PFR had design constraints which were that the maximum outlet of the reactor could not exceed 1300 F, there would be a 6 inch layer of insulation which would reduce the steel to 900 F, and the quench stream would reduce the effluent stream to 1150 F before it would enter the first heat exchanger. The heat exchangers were to be sized to fit the necessary flow and temperature to produce high pressure steam, low pressure steam, and hot water. The flash drum that needed to be installed had operating conditions of 100 F and 450 PSIG and the maximum vapor velocity was calculated using the equation: The installed distillation columns were said to run optimally at an R value of 1.3*R min, have a tray efficiency of 60%, contain at least a purity of 99.97% benzene and 0.03% toluene in the product stream, and a limit of 4% benzene in the recycle stream was allowable. For the initial hand calculations a perfect 4
separation was assumed. An isentropic compressor with 75% efficiency was used in any vapor compression. Equipment Design Figure 3: Original process flow diagram for the base-case scenario. Before an ASPEN simulation could be completed, hand calculations were required to provide an initial process flow diagram and estimate of stream values and equipment parameters. The equipment was sized in the order of how the process took place (i.e. starting with the fired heater and ending with the last distillation column). Table 2 shows the costs that were calculated using the sizes from the hand calculations. These calculations can be looked at more closely in the hand calculation portion of the appendix. 5
Table 2: This table shows the calculated costs for the base-case scenario. An issue for the equipment design was the first distillation column. The sheer size of the column was too large to calculate an actual cost. This column would never be feasible in a real life application. The reason the column was so large was the formation of an azeotrope at 450 psig. This problem was corrected by later lowering the pressure of the distillation column. For the modified equipment, the specifications can be seen in Appendix B. The hand calculations proved to be a good start for inputs required for the ASPEN simulation to run properly. The required heat duty for the fired heater came out to be 5.27181E+7 BTU/hr. It sent the mixed stream out at 1150 F and 485.304psig. The PFR s limitation was that the outlet temperature could not exceed 1300 F in which case ASPEN calculated the outlet temperature and pressure to be 1227.91 F and 485.304psig respectively. Since this temperature did not exceed the limit the process was safe to continue forward. The first heat exchanger used high pressure water to cool the quenched stream down to 464.698 F from 589.647 F. The energy used to cool the stream produced high pressure steam to be used elsewhere in the process/plant. The second heat exchanger took the exiting stream from the first heat exchanger and cooled it to 302.792 F. This exchanger was cooled by water and produced low pressure steam which may serve other purposes in the plant. The last heat exchanger cooled the outlet stream of the first heat exchanger down to 100 F which is optimal feed temperature for the flash drum. The first flash drum had two design constraints; the operating temperature and pressure were to be 100 F and 450psig respectively. The heat duty required for the flash separation was 20,942.4 BTU/hr. The size of this drum was calculated using the required heat duty and thickness of the unit. This flash drum produced a separation of a distillate and bottoms stream. The top stream was purged and split to be recycled back into the feed. The top compositions of the materials are as follows; benzene: 0.004, toluene: 8.54E-4, methanol: 0.431, hydrogen: 0.564 and diphenyl: trace amount. The bottoms stream molar compositions are as follows; benzene: 0.576, toluene: 0.384, methanol: 0.033, hydrogen: trace 6
amount and diphenyl:.004. Distillation column 1 (B17 with reference to the modified PFD) was corrected and resized to have a 36 stages instead of the 108 initially calculated. The distillate rate was chosen to be 115.5 kmol/hr, the reflux ratio to be 10, and the operating pressure to be 1 atmosphere. The column produced a distillate stream which had a benzene composition of 0.9432. This stream goes on to another flash drum and distillation column to achieve the final purity. The bottoms had a toluene composition of 0.9845; this stream would also be further distilled and recycled into the initial feed stream. The distillate stream, S1 (modified PFD), was fed into a second flash drum (flash drum 2). Flash drum 2 operated at 1 atmosphere and 86 F to produce a bottom-stream (S9) benzene composition of 0.9978. The distillate stream (S10) was purged off and harvested as fuel gas due to the high methane concentration. The benzene and toluene loss from this stream are minute in comparison with the bottoms stream. The bottoms stream (S9) was fed to a final distillation column where a purity of 99.99% was reached. This column (B4) operated with 3 stages, a distillate rate of 0.3 kmol/hr, and a reflux ratio of 2. This column would further separate the methanol and remaining hydrogen from the benzene. The benzene flow leaving this column was 107.704kmol/hr. The distillate stream was also used as fuel gas due to the high methane concentration. The bottoms stream of the first distillation column (B17) was further purified so that toluene could be recycled back into the toluene feed stream (Tol. Dist.). This column (B18) produced a toluene recycle stream that had a purity of 0.9958 and a flow of 59.7453kmol/hr. This column was fairly small with 5 stages, a distillate rate of 60kmol/hr, and a reflux ratio of 0.3. This stream was fed into a compressor to turn the distillate vapor back into a liquid. The compressor (Toluene Comp.) took the recycle stream up to 485.304psig and 551.3746 F. This compressor had horsepower rating of 292.429hp (218.064kW). For a more detailed look of the equipment see Appendix B which has each individual stream table for the specified equipment. ASPEN Summary The reactor was modeled using RSTOIC because the conversion and selectivity were known for the process. Simple heat exchangers were used to model shell and tube heat exchangers with the given overall heat transfer coefficients. The flash tank was designed at the given parameters. The light stream separator was designed to purge 13.6% of the stream as determined by hand calculations. The remaining light recycle stream was split evenly to two compressors. The compressors were modeled as isentropic with an efficiency of 75%. The two compressed streams were combined and then mixed with the fresh hydrogen stream before entering the fired heater. The bottom stream from the flash tank was split to form the quench stream. The initial quench stream value, based on hand calculations, was optimized to produce a combined stream temperature of 1150 F by diverting 41.6 kmol/hr back into the system. The remaining stream was fed to the first distillation column. The initial design was based on a McCabe Thiele diagram for the separation of benzene from toluene. After the pressure was changed to one atmosphere a sensitivity analysis was 7
conducted to determine the ideal number of stages and feed locations. The distillate was fed to a flash tank operating at 86 F to remove a very pure methane stream while minimizing benzene yield loss. The remaining liquid stream was fed to a distillation column with three stages to finalize the separation of benzene and methane. The distillate stream was left as a vapor to reduce the heat duty of condensing the cold methane stream. The toluene stream from the first distillation column was fed to a third distillation column to remove the diphenyl. Only five stages were necessary to achieve the desired separation and the toluene distillate was left as a vapor to reduce heat duty. This toluene stream was compressed to 500 psia in another isentropic condenser at 75% efficiency before being mixed with the fresh toluene feed. This stream was fed to the fired heater and then the reactor. It should be noted that despite both streams entering the fired heater together the streams would be separate in reality to prevent the reaction from occurring outside the reactor. 8
Modified PFD Figure 4: Modified process flow diagram from ASPEN simulation. The modified process flow diagram did not change much with optimization but does have some significant differences when compared to the base-case process flow diagram. The modified PFD adds two compressors, a flash drum, and a distillation column. This equipment was all necessary in order to achieve the desired conversion. Comparison of ASPEN & Hand Calculations When looking at the difference between the hand calculations and ASPEN results, the biggest concern was the first distillation column. In the hand calculations the McCabe-Thiele plot proved that with the desired separation the minimum stages required were 108 (see Appendix D for the plots). This column was optimized in ASPEN and only required 36 stages. This significantly decreased the overall cost of the plant by cutting the column size down so greatly. Another difference between the hand calculation and ASPEN simulation was that the ASPEN simulation required a third distillation column, another flash drum, and 2 compressors. The two compressors were required to allow for regular maintenance without sacrificing production. The second flash drum removed most of the remaining methanol from the product stream. The third distillation column was also needed in order to achieve a high enough purity for the final product. 9
Safety Considerations The nature of this process poses several serious risks. All of the components are extremely flammable and form vapors at relatively low temperatures. Extreme caution would need to be taken to prevent leaks and sparks from occurring. All maintenance, especially hot work, would require extreme precaution. A robust fire suppression system could prevent small sparks or fires from developing into larger fires or even an explosion. The vapor leaving the flash tank is purged and contained as fuel. This process creates another potential point for leaks or explosions and should be designed with safety in mind. The plant operates at a very high pressure, over thirty atmospheres. This creates a risk for explosions and ruptures, especially in the reactor at high temperature. The reactor is insulated with refractory bricks and designed with a thick steel shell to prevent ruptures and explosion from occurring. The heat exchangers are designed with the high pressure stream tube-side. While high pressure steam is produced in one exchanger it poses a very low risk compared to the flammable gases produced by the reactor. The liquid leaving the flash tank is depressurized to one atmosphere. At this point it poses a much lower threat. However, the vapor leaving the flash tank is still at high pressure and is compressed to even higher pressures. The dual compressor design allows for regular maintenance on both compressors to prevent an equipment failure that could create hazardous conditions. Toluene and diphenyl are mildly toxic. Toluene vapors can cause serious health risks if inhaled in large quantities but do not pose a threat at minor levels. Benzene however, is a known carcinogen and long term exposure has been linked to leukemia. Therefore it is extremely important that the product packaging and shipping process eliminates any operator exposure. Conclusions The preliminary process design showed that the process operated under a small profit margin assuming complete separations. Rigorous Aspen analysis shows that the equipment cost required to achieve separation is greater than initial estimates, but the benzene yield was increased. With the new process the cash flow payback period is 2.32 years, as shown in appendix E. The additional separation of methane from benzene was the primary source of the additional costs. However, with the additional equipment additions the required separation was achieved. The plant shows strong profitability and is considered a strong investment. References [1] Peters, Max Stone., and Klaus D. Timmerhaus. Plant Design and Economics for Chemical Engineers. New York: McGraw-Hill, 1991. Print. [2] Seader, J. D., Ernest J. Henley, and D. K. Roper. Seperation Process Principles Chemical and Biochemical Operations. 3rd ed. N.p.: RDC, 2011. Print. [3] Green, Don W. Perry's Chemical Engineers' Handbook. New York: McGraw-Hill, 2008. Print. List of Assumptions 10
A perfect separation was assumed in the hand calculations for the distillation columns Cost of piping, valves, and pumps is negligible relative to major equipment Heat capacity of quench stream is negligible (hand calculations) Isothermal reactor for reactor sizing (hand calculations) Toluene and Hydrogen feed streams come in at the same conditions as the recycle streams. No frictional losses or pressure drop in the system except where specified. 11
Appendix Appendix A: Preliminary economics Table 1: Preliminary Economics 12
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Stream Tables: Initial conditions Appendix B: PFD s with ASPEN calculations 14
Figure B1: Fired Heater 15
Figure B2: PFR 16
Figure B3: Heat Exchanger 1 17
Figure B4: Heat Exchanger 2 18
Figure B5: Heat Exchanger 3 19
Figure B6: Flash Drum 1 20
Figure B7: Flash Drum 2 21
Figure B8: Distillation Column 1(B17) 22
Figure B9: Distillation Column 2 23
Figure B10: Distillation Column 3 24
Figure B11: Centrifugal Pump 1 25
Figure B12: Compressor 1 26
Figure B13: Compressor 2 27
Figure B14: Toluene Recycle Compressor 28
Appendix C: Hand Calculations Flash Drum Tank Thickness - Distillation Column (Sieve Trays) - (Peters and Timmerhaus, Table 12-10) - - - (Peters and Timmerhaus, pg 780) - (Seader, Henley, and Roper pg 227) Heat Exchanger Reactor - 29
Appendix D: McCabe-Thiele Plots 30
Appendix E: Economics of Final and Initial Design Table 2: Toluene Demethylation (Preliminary Economics) Given Calculated Item Unit $/unit /gal of fresh toluene charge Fixed Capital Investment $11,563,278 Manufacturing Cost N/A N/A N/A N/A Chemicals N/A N/A 0.18 0.18 Utilies N/A N/A 9.9 9.9 Total Variable Costs N/A N/A 10.08 10.08 Toluene gal 1.3 130 130 Hydrogen 1000std.cu.f t 2.7 16.5 23.0 Total Charge Cost N/A N/A 146.5 153.0 Diphenyl gal 0.36 0.24 0.38 Fuel Gas MM Btu 2.4 9.18 9.57 Total by-product credit N/A N/A 9.42 9.95 Net Charge Cost N/A N/A 137.08 143.09 Fixed Costs N/A N/A N/A N/A Labor, two men/shift N/A N/A 0.68 Supervision N/A N/A 0.18 Maintenance N/A N/A 1.13 31
Depreciation N/A N/A 3.77 Local taxes and insurance N/A N/A 0.38 Main+Depr+Tax+ins N/A N/A 5.28 General Overhead N/A N/A 0.32 Total fixed Costs N/A N/A 6.46 Total Manufacturing Cost N/A N/A 153.62 159.63 Net Receipts from Benzene Sales gal 2 165.54 165.76 Net Operating Income Before Taxes N/A N/A 11.92 6.13 Federal Income Tax @ 40% N/A N/A 4.77 2.45 Net Income After Taxes N/A N/A 7.15 3.68 Depreciation N/A N/A 3.77 3.77 Cash Flow N/A N/A 10.92 7.45 Cash Flow Payout, year 3.45 5.06 32
Table 3: Toluene Demethylation (Final Economics) Given Calculated Item Unit $/unit /gal of fresh toluene charge Fixed Capital Investment $10,757,311 Manufacturing Cost N/A N/A N/A N/A Chemicals N/A N/A 0.18 0.18 Utilies N/A N/A 9.9 9.9 Total Variable Costs N/A N/A 10.08 10.08 Toluene gal 1.3 130 130 Hydrogen 1000std.cu.ft 2.7 16.5 23.0 Total Charge Cost N/A N/A 146.5 153.0 Diphenyl gal 0.36 0.24 0.36 Fuel Gas MM Btu 2.4 9.18 9.04 Total by-product credit N/A N/A 9.42 9.40 Net Charge Cost N/A N/A 137.08 143.64 Fixed Costs N/A N/A N/A N/A Labor, two men/shift N/A N/A 0.68 Supervision N/A N/A 0.18 Maintenance N/A N/A 1.05 Depreciation N/A N/A 3.51 Local taxes and insurance N/A N/A 0.35 33
Main+Depr+Tax+ins N/A N/A 4.91 General Overhead N/A N/A 0.32 Total fixed Costs N/A N/A 6.09 Total Manufacturing Cost N/A N/A 153.25 159.81 Net Receipts from Benzene Sales gal 2 165.54 179.15 Net Operating Income Before Taxes N/A N/A 12.29 19.34 Federal Income Tax @ 40% N/A N/A 4.92 7.74 Net Income After Taxes N/A N/A 7.37 11.61 Depreciation N/A N/A 3.51 3.51 Cash Flow N/A N/A 10.88 15.11 Cash Flow Payout, year 3.22 2.32 Figure E3: Cash flow pay back versus conversion for the process Table E1: Final Stream table from Aspen 34
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