EBTAX: The Conversion of Ethane to Aromatics via Catalytic Conversion

Transcription

1 University of Wyoming Wyoming Scholars Repository Honors Theses AY 15/16 Undergraduate Honors Theses 2016 EBTAX: The Conversion of Ethane to Aromatics via Catalytic Conversion Emily Schwichtenberg University of Wyoming, Aaron Cheese University of Wyoming, Bridger Martin University of Wyoming, Saud Alshahri University of Wyoming, Follow this and additional works at: Recommended Citation Schwichtenberg, Emily; Cheese, Aaron; Martin, Bridger; and Alshahri, Saud, "EBTAX: The Conversion of Ethane to Aromatics via Catalytic Conversion" (2016). Honors Theses AY 15/16. Paper 7. This Dissertation/Thesis is brought to you for free and open access by the Undergraduate Honors Theses at Wyoming Scholars Repository. It has been accepted for inclusion in Honors Theses AY 15/16 by an authorized administrator of Wyoming Scholars Repository. For more information, please contact

2 5/6/2016 Ethane to Aromatics EBTAX Group Saud Alshahri, Aaron Cheese, Bridger Martin, Emily Schwichtenberg CHE 4080 PROCESS DESIGN II

3 1 Table of Contents I. Table of Tables... 2 II. Table of Figures... 4 III. Executive Summary (Emily)... 6 IV. Scope of Work (Bridger)... 8 V. Introduction (Saud) VI. Description of Base Case Section 100: Feed processing, reaction, and initial separation (Bridger) Section 200: Lights Separation Section (Emily) Section 300: Separation and Recovery of BTX and Heavy Aromatics Products (Saud) Section 400/500: Propane and Ethylene Refrigeration (Aaron) VII. Design Alternatives (Bridger) Possible Reactor Alterations Product Recovery (Section 300) alternative designs Continuous Catalyst Regeneration Fuel gas reallocation, C2 through C4 repurposing VIII. Permitting and Environmental Concerns (Emily) IX. Safety and Risk Management (Emily) X. Project Economics (Aaron) Equipment and Capital Cost Pricing, Revenue and Production Cost Cash Flow Analysis Sensitivities XI. Global Impacts (Saud) XII. Conclusions and Recommendations (Bridger) XIII. Future Work (Aaron) XIV. Acknowledgements (Saud) XV. References XVI. Appendices... 79

4 2 I. Table of Tables Table 1: List of reactions, their respective conversions, and the heats of reaction/ Table 2: Pressure, temperature and enthalpy data for propane Table 3: Pressure, temperature and enthalpy data for ethylene Table 4: Various operating conditions specified for the different refrigeration cycles involved along with the Aspen unit operation or stream that it corresponds to Table 5: This table shows the emissions of thermal NO x with and without control measures. The emission limit for needing a permit from the EPA is 100 tons/year, which can be obtained with control measures in this process Table 6: Specific information involved in sizing and costing compressors Table 7: Specific information involved in sizing and costing turbines Table 8: Specific information involved in sizing and costing furnaces Table 9: Specific information involved in sizing and costing heat exchangers Table 10: Specific information involved in sizing and costing air coolers Table 11: Specific information involved in sizing and costing vessels Table 12: Specific information for costing the PSA unit Table 13: Specific information for sizing and costing the amount of catalyst used Table 14: Specific information for sizing and costing the amount of catalyst used Table 15: Specific information for distillation column and tray sizing and costing... 61

5 3 Table 16: Fixed Capital Investment for the various equipment involved in the process along with the resulting total Table 17: Income or cost of each of the materials consumed or produced Table 18: Cost of utilities Table 19: Various fixed costs associated with the design Table 20: Results of the cash flow analysis conducted on this design Table 21: Sensitivities run, along with the resulting IRR... 66

6 4 II. Table of Figures Figure 1: Overall Process Flow Diagram. Part A: Section 100, 200, and 300 up to M301. Part B: Remaining portion of section 300. Part C: Section 400 and Figure 2 : Feed and reactor (Section 100). Feed ethane is mixed with hydrocarbon and hydrogen recycles. The presence of hydrogen significantly reduces catalyst coking. The flash tank, D101, sunders gaseous C1-C5 to section 200 and C6-C9 to section Figure 3: Section 200, the Lights Separation Section. This section consists of a mixer to combine the vapor stream from the flash drum and a recycle from the product recovery section, two distillation towers and a pressure swing adsorption (PSA) unit to separate the product and recycle streams, one splitter to allow for the hydrogen sale stream to be separated, and two compressors to pressurize the recycle streams to appropriate pressures to be mixed with the feed stream Figure 4: Section 200 Up To T201. From the flash drum, the vapor stream is mixed in M201 with a vapor recovery stream in the product recovery section before being sent to a distillation tower (T201) to remove hydrogen and methane from the product stream as the vapor distillate (S203), with the remainder exiting the tower in the bottoms stream (S206) Figure 5: Section 200, T201 Distillate Path After T201. The hydrogen and methane stream is sent to a heat exchanger to warm it up to room temperature before being sent to the pressure swing adsorption (PSA) unit, where the methane is removed to a fuel gas stream and the hydrogen stream is sent to a splitter (S201). This splits the hydrogen stream into a sale product stream and a recycle stream, which will be sent compressed and sent back to the reactor to prevent coking of the catalyst Figure 6: Section 200, T201 Bottoms Path. The bottoms of T201 is sent to T202, where C2 and C3 hydrocarbons are distilled off and sent to a compressor before being recycled back to the recycled to increase the overall conversion of the reactor. The bottoms stream is sent to a mixer in the product recovery section to recover any BTX products that could have been lost Figure 7: Heavy Separation (Section 300). From flash tank D101, the heavy stream is separated remaining light hydrocarbons. The remaining heavies are separated into Benzene, Toluene, and Xylene product and TMB byproduct Figure 8: Heavies Separation (Section 300). From flash tank D101, the liquid stream is separated fed into the first distillation column (T301) to remove the remaining light hyrdrocarbons as well as recover TMB as a product. The aromatic rich stream is sent on for further processing by S

7 5 Figure 9: Purge Stream. BTX rich streams are fed into tower T302, one of which comes from the lights separation section, and one of which come from the previous tower, T301. T302 separates out any remaining lights and purges them from the system. The bottoms of the tower is sent on for product recovery as it mainly consists of BTX Figure 10: Benzene Recovery. Benzene is recovered from the BTX rich stream leaving T302. The bottoms of the tower is sent on to recover the remaining Toluene and Xylene Figure 11: Toluene and Xylene Recovery. T304 separates toluene from para-xylene that is fed to the tower from T Figure 12: Propane and Ethylene Refrigeration. Section 400 consists of two propane refrigeration cycles operating at different pressures. Section 500 consists of only one ethylene refrigeration cycle. For each cycle the refrigerant is compressed, condensed, expanded, and evaporated in order to complete the cycle. Propane refrigeration is used in condensing the process fluid in T202 along with condensing the ethylene in section 500. Ethylene refrigeration is only required to condense the process fluid in T Figure 13: Tornado Diagram. This plot chose the change on IRR based on different variation of various uncertain parameters Figure 14: Industry Rivalry. This figure illustrates the possible industry pressure associated with a new competitor... 69

8 6 III. Executive Summary (Emily) Team EBTAX was formed with the goal of taking ethane from natural gas refineries and processing it into various aromatics, including benzene, toluene, and para-xylene (BTX). The market for BTX chemicals is fairly stable, as they can be converted into larger molecules that are critical components in polymer synthesis, producing plastics, textiles, and other consumer goods. The glut of natural gas in the US has caused ethane prices to drop to around half of what they were at the beginning of This makes it an ideal feedstock for our process, which includes a catalyzed reaction and several separation units to separate the reaction products into pure component products. The plant will be located on an existing oil refinery in the gulf coast area. This will provide easy allocation of products, as well as access to the oil refineries and chemical plants that would purchase and further process the products. To catalyze the reactor, a platinum-zeolite catalyst was chosen for its high selectivity toward BTX compared to similar catalysts. US Patent B2 only provides conversion and reaction information for this catalyst from lab scale tests. The unfamiliarity with this catalyst and the lack of information at diverse conditions led to several assumptions when modeling the process. Firstly, the amount of catalyst required for the process scales linearly and ideally with the reactor inlet. Secondly, the conversion of the reactions would not depend strongly on pressure. The patent also provides a functional pressure for the catalyst specified from 20 to 2000 psia without providing correlations for pressure and conversion. Catalyst lifetime use is assumed to outlive the life of the plant, and regeneration is assumed to recover 100% of the catalyst. Regeneration alone will keep the catalyst active for the lifetime of the project. The process is broken down into five sections. To begin, the reactor section (Section 100), consists of an ethane feed stream and two recycles from the separation sections. The feed is mixed with the recycles and then heated before entering a gas-phase, fixed-bed, catalytic reactor at 1150 F. The reactor houses 26 separate reactions. C2 and C3 hydrocarbons (HCs) undergo multiple equilibrium-based reactions to form C1 through C5 linear and C6 through C9 aromatic HCs and hydrogen. After the stream

9 7 has reacted, it is cooled before being sent to the separation sections. An initial flash tank separates the reactor effluent into light and heavy HC streams, which are sent to sections 200 and 300, respectively. In the lights separation, two distillation columns and a hydrogen pressure swing adsorption (PSA) unit are used. The four product streams from the light separation section are high purity hydrogen (99.5 mol%), methane fuel gas, and a C2-C3 HC recycle. The hydrogen stream will be split into a sale stream and a recycle stream, which will prevent coking of the catalyst. Sections 400 and 500 are propane and ethylene refrigeration, respectively, which will be used to cool the condensers in the lights separation to allow the small hydrocarbons to condense. The liquids from the initial flash tank are pumped to the product separation (Section 300). Four distillation towers are used in this section. The first tower (T301) functions as a light HC recovery unit, with the remaining light vapors entering section 200. The liquids enter the next distillation tower (T302) to remove TMB. The last two towers separate the BTX into its components to be sold. The fixed capital investment for this project is $320 million, which is largely due to the many compressors needed for the refrigeration section. With a 20 year project life, this project yields an IRR of 30.3%, with a payback period of 2.6 years.

10 8 IV. Scope of Work (Bridger) The preliminary mission on team EBTAX is to design an industrial plant that takes advantage of the abundant natural gas supplies in order to produce benzene, toluene, and xylene. Research into the natural gas industry, along with common natural gas refinery processes, revealed ethane to be the most probable feedstock. This was due to a number of contributing factors, including the recent practice of ethane rejection, where ethane is allowed to flow with methane into the pipe gas stream. Ethane rejection further lowers the cost of inexpensive ethane feeds. US Patent US A1 was provided to team EBTAX as a starting point for the conversion of ethane into valuable aromatics, including benzene, toluene, and xylene. This catalyst became the basis for the plant design and created many constraints that had to be met by the plant. Constraints The primary design constraints are equilibrium constraints within the reactor. The catalyst conversion is highly specific and creates the high volume reflux of C2 and C3 hydrocarbons. The catalyst operating conditions also set the reactor temperature at 1150 F. These high temperatures can also ignite the HCs if oxygen is present in the system. All process streams are run above atmospheric to prevent oxygen from entering the system in the case of a leak. The reactions taking place are exothermic. This creates the opportunity for a runaway reaction if released heat and built up pressures go beyond controllable conditions. This creates additional safety constraints, such as the inclusion of cooling systems, which will also be considered in more detailed designs. Thermodynamic constraints are present outside the reactor as well. To separate the C1-C9 HC stream, extremely low temperatures are required for the lighter components. Two types of refrigeration were included to reach the low temperatures. These refrigeration systems will be discussed in the Section 400/500 below.

11 9 Climate restrictions were also taken into account. The Gulf Coast air and water temperatures will be higher than temperatures elsewhere in the country, particularly during the summer. The water temperature was estimated at 105 F and the air temperature was estimated at 95 F. The majority of other operating constraints are the purities of the products we intend to sell. If hydrogen, benzene, toluene, or p-xylene are not at the correct purity, they will not be able to be sold at as high of a price. These could also be considered economic constraints since they directly affect the economic income of the plant. By incorporating the above design requirements, the EBTAX team created a preliminary design for a profitable, industrial scale plant. Team EBTAX verified the plant design using Aspen+ modeling design. Capital costing was primarily performed by hand using cost graphs in Peters and Timmerhaus. Compiled economic analyses were performed via Microsoft Excel to very that the plant design is profitable. A full scale industrial plant was designed to fulfill these requirements from the initial design concept, through the preliminary design analysis.

12 10 V. Introduction (Saud) Recently, there has been a glut of natural gas in the United States. This has driven the price of natural gas, including ethane, to almost half of what it was at the beginning of This remarkably low cost of ethane has led to natural gas refineries no longer separating ethane from the natural gas, and simply rejecting the ethane to pipeline natural gas that gets sent to residential homes. While the natural gas plant does save money from rejecting ethane, there is significant potential loss and waste considering ethane can be converted to valuable BTX products. Due to the large excess of natural gas liquids in the United States, ethane prices are currently very low. From this information, team EBTAX was charged with the task of researching and designing a way to capitalize on this low-priced feedstock. A recent patent, US describes a platinum-zeolite catalyst which converts ethane into valuable aromatics, primarily benzene, toluene, and para-xylenes (BTX). The plant is based on this catalyst, and is still in the design phase, with the goal of producing 700 MMlb/yr of the benzene, toluene, and para-xylene products. Some assumptions were needed to assist in modeling our catalyst and the plant. The catalyst was chosen to achieve the highest selectivity of BTX product. This conversion data has only been tested at the lab scale. US Patent A1 stated that the catalyst could be used with pressures ranging from 20 to 2000 psia. The correlation between pressure and conversion will need to be determined experimentally because to date, no data has been recorded for this catalyst. Although the single regeneration life of the catalyst is unknown, similar catalysts have been shown to need regeneration every one to six months. A continuous regeneration process will eliminate the plant shutting down due to time needed for catalyst regeneration. Economic calculations assume that the overall catalyst life is more than 20 years, the entire length of the project, so it will never need to be replaced.

13 11 VI. Description of Base Case Part A: Sections 100, 200, and 300 Up to the Mixer (M301) S120 C102 S220 FUELGAS H2SALE SP201 PR201 S210 S205 S204 H201 C101 S203 S101 S201 M201 S202 T201 T202 S110 M101 S102 H101C S103 F101 S104 R101 S105 H101H S107 S106 H102 D101 S303 S206 S207 M301 X101 S302 FEED S301 T301 TMBPR OD

14 12 Part B: Section 300 After the Mixer (M301) PUR GE M301 BENZPR OD S304 T303 S305 TOLPR OD T302 V301 S307 S306 T304 XYPROD

15 13 Part C: Refrigeration S410 H204C S409 HPFEED M401 LPFEED C402 S402 H401 V401 SP401 S401 C401 S403 S404 S405 H501C V402 S408 S407 S406 S504 S503 H202C ETHYLEN E S502 S501 H501H V501 C501

16 14 Figure 1: Overall Process Flow Diagram. Part A: Section 100, 200, and 300 up to M301. Part B: Remaining portion of section 300. Part C: Section 400 and 500.

17 15 Section 100: Feed processing, reaction, and initial separation (Bridger) Figure 2 : Feed and reactor (Section 100). Feed ethane is mixed with hydrocarbon and hydrogen recycles. The presence of hydrogen significantly reduces catalyst coking. The flash tank, D101, sunders gaseous C1-C5 to section 200 and C6-C9 to section 300. The EBTAX plant design begins with a stream of feed ethane purchased from a local natural gas refinery. With the EBTAX plant located in the Houston area, purchasing the feedstock from a local refinery will reduce shipping and processing cost associated with plant operation. The cost of feed ethane is further reduced due to the recent practice of ethane rejection. Ethane that would now be attributed to the household natural gas stream for little to no profit can instead be purchased by EBTAX at a very low cost for conversion into valuable products. The purchased ethane has a minimum purity standard of 95% ethane with the primary impurities of propane and carbon dioxide. The maximum allowable range for propane content is 0% to 5%. (Lonestar) These small amounts of propane create no detrimental effects within the system. Propane, being a light hydrocarbon, also reacts via catalysis to form the valuable BTX products. Propane gas actually has a higher conversion to our most valuable product, benzene, than ethane; however, ethane remains the ideal choice of feedstock due its low cost and availability.

18 16 The maximum allowable range for carbon dioxide content is 0% to 0.1%. (Lonestar) Very few changes occur throughout the system due to varying carbon dioxide concentrations. Carbon dioxide gas remains inert within the reactor, so no additional reactions take place. No new products are formed and the product selectivities remain constant. This is an assumption, but ethane is not available without trace carbon dioxide. Additional problems should have been included in the literature, but this may need to be verified. The carbon dioxide present in the reactor effluent continues through the separations process. The flash separator, D101, diverts 100% of contaminant carbon dioxide to the lights separation process (Section 200). During lights separation, 100% of the carbon dioxide is separated via the pressure swing adsorption unit, PR201. The carbon dioxide remains in the methane gas stream, which is used to fuel the heating processes for the EBTAX plant. The small carbon dioxide contents will only slightly contribute to the overall emissions from the reactor furnace, F101, and reboiler furnace, F301. Feed processing The feed ethane begins at high pressure, 835 psia, at a rate of pounds per hour. The ethane flow was determined to produce 700 million pounds of desired aromatics; benzene, toluene, and xylene, annually at 8250 operating hours per year. This raw feed is passed through an expander, X101, to reduce the pressure to 320 psia. This produces 1460hp of energy that can be rerouted to other energy intensive unit operations in the process. Lowering the pressure allows the raw feed to be mixed with the recycle streams near operating pressure. The low pressure feed stream, S101, mixes with two recycle streams, S110 and S120, in mixer M101 before being heated to operating temperature. Stream S110 consists of a 99.5 mol% hydrogen with the remainder being methane. Hydrogen is recycled at a 1 to 10 mole ratio with the hydrocarbons (HC) entering the reactor, including the HC recycle S120. This amounts to 2100 pounds per hour at standard operating conditions. The recycled hydrogen reduces catalyst coking within the reactor. This decreases the catalyst coking rate, allowing for more time to pass between catalyst regeneration cycles.

19 17 Recycle stream S120 contains the HC recycle stream from the light separation section (Section 200). The primary component in S120 is unreacted ethane at 91.2%. This stream also contains 4.5% propane, and trace C4 and C5 HCs. The propane will react to desired aromatics using similar mechanisms and selectivities as ethane, with a trend towards more benzene, our most valuable aromatic product. Conversions for each reaction can be seen in table 1. The trace C4 and C5 HCs do not react further and have no detriment to the reaction system. To prevent a build-up of inert components, a purge stream in the lights separation (Section 200) is used to vent excess C4 and C5. These mid-range HCs are used as fuel gas to heat the furnaces. This purge is used as fuel gas before use of excess produced methane in order to minimize waste streams and associated processing and handling costs. After mixing with the recycle streams, the complete feed is heated to the operating pressure of 1150 F through two units, a counter-current, cross-reactor heat exchanger, H101, and furnace F101. The H101 heat exchanger utilizes our hot, 1150 F, reactor effluent as the heating fluid to simultaneously raise the reactor inlet fluid temperature and cool the reactor effluent to prepare for separation. The hot effluent is located on the tube side and the cool reactor inlet is on the shell side. This exchanger heats the inlet fluid to a temperature of700 F, and cools the reactor effluent to 703 F. The reactor feed flow is further heated through the furnace F101. The operating temperature for the reactor is set to 1150 F. To achieve this temperature, the furnace operates using the C4 and C5 purge stream as fuel gas. The purge stream contains other light HCs as well, which burn normally within the system. Furnace F101 also uses the excess methane produced in the plant as a fuel source. The PSA separation unit, PS201, diverts enough methane from the process that no additional methane or fuel is required for purchase. Reactor mechanics At this point, the reactor inlet, S104, is at the proper inlet conditions. The temperature is 1150 F and the pressure is at 300psia. The stream then enters two identical parallel reactors. The operation

20 18 utilizes three reactors. Two reactors are operated at any given time while the third is being used for catalyst regeneration. Each reactor is composed of a glass-lined stainless steel vessel with a packed bed of catalyst. Stainless steel is used in all components in contact with a hydrogen stream, particularly the reactor, to protect against hydrogen embrittlement and possible explosion risk. The glass lining within the reactors is used to protect against the chlorine gas used in the regeneration process. The catalyst is composed of a zeolite base with germanium inserted into the structure. This germanium-enhanced zeolite is used as an anchor for platinum, which provides the catalyst its reactivity. A platinum content of.0441wt% is used according to the conversion and selectivity data available in US patent US The single pass conversion for this catalyst at these operating conditions is 46%. The overall selectivity towards desired aromatic products is 61%. A full list of reaction conversions can be seen in Table 1.

21 19 Table 1: List of reactions, their respective conversions, and the heats of reaction Reactions Conversion ΔHrxn (BTU/lb) C2H6 + H2 --> 2 CH C2H6 --> C2H4 + H C2H6 --> 2C3H6 + 3H C2H6 --> 2C3H8 + H C2H6 --> C6H6 + 6H C2H6 --> C7H8 +CH4 + 6H C2H6 --> C8H10 +7H C2H6 --> C9H12 + CH4 + 7H C2H4 + 2H2 --> 2CH C2H4 --> 2C3H C2H4 + 2H2 --> 2C3H C2H4 --> C6H6 + 3H C2H4 --> C7H8 + 2H2 + CH C2H4 --> C8H10 + 3H C2H4 --> C9H12 + 2H2 + CH C3H8--> C6H6 + 5H C3H8 --> C7H8 + C2H6 + 5H C3H8 --> C8H10 + CH4 + 5H C3H6 + H2 --> CH4 + C2H C3H6 --> C6H6 + 3H C3H6 --> C7H8 + C2H6 + 2H C3H6 --> C8H10 + CH4 + 2H C3H8 --> C9H12+ 6H C3H6 --> C9H12 + 3H C2H6-->C4H10+H C2H6-->C5H12+CH4+H

22 20 An operating temperature of 1150 F is the essential variable which controls the extent of the reaction and gives the desired selectivities. Adjusting this temperature lowers the single pass conversion and adjusts selectivity away from the desired aromatic products. The operating pressure is a more flexible variable. Conversion and selectivity data in US patent US is provided for lab conditions at atmospheric pressure, and the patent lists functional pressure from psia. The chosen pressure, 300psia, reduces the size of the reactor significantly. According to La Chatlier s, principle, this increased pressure should reduce conversion due to a higher number of moles in the products of many reactions taking place. However, when lowering pressure to the patent tested values, the reactor becomes too large, and catalyst based capital costs become too high. The large functional pressure range, combined with a lack of tests for pressure correlation in US patent US , creates the largest assumption that must be verified at the pilot scale and allows for multiple design options; refer to Design Alternatives VII and Future Work XIII. At the high temperatures occurring within the reactor, the catalyst undergoes coking as the reactions occur. After approximately several months the catalyst in the reactor will become coked enough to affect reaction conversions and selectivities. When coking becomes significant, a catalyst regeneration procedure can be performed to return the catalyst to its previous, fully active state. The regeneration procedure follows a 4 step procedure outlined by patent US First, coke is removed via high temperature oxidation. Temperatures in the reactor and during coke removal are sufficient so that sintering of the catalyst occurs. The sintering causes catalyst particles to group up, reducing catalyst efficiency and selectivities. The second step of regeneration is to redisperse the platinum over the catalyst surface using a gas stream containing chlorine gas, oxygen, and steam. The glass lining protects the stainless steel reactor from the chlorine gas. The chlorine gas is then removed from the stream and steam continues to flow. The steam without chlorine allows the platinum to rebind to the surface of the zeolite. The final step in catalyst regeneration is the reduction of the catalyst using hydrogen. This counteracts the original oxidation and returns the catalyst to its original, active state.

23 21 Once the catalyst regeneration process is completed, the catalyst returns to its original activity. The estimated lifetime for a single load of catalyst is estimated to be 20 years, the life of the project. Sensitivities were performed to account for the possible replacement of catalyst one time during the 20 year plant life. These showed very little impact in the economic analysis; refer to Economics. Catalyst regeneration procedures are the reason that three parallel reactors are used. During one reactor s regeneration cycle, the other reactor remains operational under normal operating conditions. This allows the plant to maintain a continuous product stream, even while a reactor undergoes maintenance. Once regeneration procedures are complete, the offline reactor can be started up to continue normal operations. Inside the reactor, the ethane, along with recycled ethylene, propane, and propylene, react while interacting with the catalyst surface. These reactants react in 26 parallel reactions, converting light C2 and C3 HCs into a mixed stream of C1 through C9 HCs and hydrogen. Each of these reactions is exothermic, producing heat and creating the possibility of a runaway reaction. Precautions regarding the reactor, including cooling and other measures, will be installed to minimize the associated risks; refer to Safety IX. A comprehensive reactions list and reaction enthalpies can be found in table 1. Initial separation The desired products consist of C6 through C8, benzene, toluene, and para-xylene. The catalyst in use is unique in its ability to create p-xylene exclusively, without alternative conformations. The heavy C9 by-product is 1, 3, 5 trimethylbenzene (TMB). These heavy components are later processed in Section 300, product recovery. The light C1 through C5 products consist of straight chain paraffins and olefins. The C2 and C3 HCs are recycled into the reactor via stream S120, and can react further to form the desired products. The C1 methane, C4, and C5 HCs cannot react again and are isolated for purge in the lights separation (Section 200). These gasses are used to entirely fuel the furnaces F101, and F301.

24 22 Hydrogen gas is purified from the mixed stream and can be sold. The necessary hydrogen for recycle via stream S110 is removed from the 99.5mol% pure hydrogen stream via splitter SP201. The streams from the parallel reactors reconverge in reactor effluent stream S105. Stream S105, containing the full mixture of gasses, is cooled in the cross-reactor heat exchanger H101 to 703 F, as mentioned above. In H101, reactor effluent is used as the heating fluid and is located on the tube side of the exchanger. The still gaseous stream is then cooled dramatically in heat exchanger H102 to facilitate the condensation of heavy aromatic products. The product stream is lowered to a temperature of 70 F. The pressure is dropped to 290psia. The process stream is located on the tube side of heat exchanger H102, which is actually broken down into two different heat exchangers. Normal cooling water modeled to enter at 95 F cools the process stream down to 105 F. After this, cooling water cooled in section 200 of the process is used to cool the process stream down the remaining 35 F. This cooling water is set to enter at 60 F and exit at 80 F. Once the reactor effluent has been cooled, the separations processes can begin, allowing for the separation of all valuable components. The first step in isolating the products is the initial splitting of the stream to form a light HC rich stream and a heavy HC rich stream. The light HC stream, which still contains some heavy components due to only having a single stage flash, is routed to the light separation section (Section 200). The heavy HC stream, along with some light components, is routed to the product separation section (section 300). These initial split streams have their trace contaminants separated and diverted back to their respective purification sections. The extra separation processes in these sections maximizes the product and feed recovery, as well as allow for maximum product purity. The mixed effluent enters a flash drum, D101, to initialize this splitting. From the flash vessel, gaseous light HCs are expelled from the top in stream S201 to the light separation (Section 200). This stream contains 98% light hydrocarbons, C1 through C5 and hydrogen, and is contaminated with 2%

25 23 benzene, and trace heavier components. The liquid stream from flash drum D101, stream S301, begins the heavy separation section (Section 300). This stream is composed of 82% desired aromatics; 49% benzene, 26.4% toluene, and 7% p-xylene. Some light hydrocarbons remain in this liquid at 7% C2 through C5. The remaining impurity consists of TMB at 9%

26 24 Section 200: Lights Separation Section (Emily) S120 C102 S220 FUELG AS H2SALE SP201 PR201 S204 S210 S205 C101 S110 S203 T201 H201 M201 S202 S201 T202 S206 S303 S207 Figure 3: Section 200, the Lights Separation Section. This section consists of a mixer to combine the vapor stream from the flash drum and a recycle from the product recovery section, two distillation towers and a pressure swing adsorption (PSA) unit to separate the product and recycle streams, one splitter to allow for the hydrogen sale stream to be separated, and two compressors to pressurize the recycle streams to appropriate pressures to be mixed with the feed stream. Section 200, the lights separation section, starts with the vapor stream exiting the flash tank from section 100 (D101) and mixing with a vapor recovery stream from the 300 section, or the product recovery section. After mixing, these streams enter a distillation column that separates methane and hydrogen from the mixture. The hydrogen and methane mixture is heated and sent to a pressure swing adsorption (PSA) unit where they are separated. The hydrogen product stream is split between a stream that will be sold and a stream that is recycled back to the reactor to prevent coking in the catalyst. The methane that is recovered can be used as fuel for the furnace before the reactor, but it can also be sold based on its heating value. The components left after the hydrogen and methane have been taken out continue on to another distillation column, where the components of benzene, toluene, para-xylene and TMB are sent to the 300 section for further product recovery. The hydrocarbons heavier than methane but lighter than benzene either continue through a recycle back to the reactor or get sent to the furnace to be burned to aid in the pre-heating of the process stream before the reactor.

27 25 S203 T201 S201 M201 S202 S206 D101 S107 S303 Figure 4: Section 200 Up To T201. From the flash drum, the vapor stream is mixed in M201 with a vapor recovery stream in the product recovery section before being sent to a distillation tower (T201) to remove hydrogen and methane from the product stream as the vapor distillate (S203), with the remainder exiting the tower in the bottoms stream (S206). The distillate from D101 exits at 290 psia and 70 F. This vapor stream is then mixed with the vapor distillate of the first product recovery distillation tower (T301) in order to collect all of the light HCs, ideally those lighter than benzene, which is our lightest product from this process. This separation is not perfect, so two recycle streams between Sections 200 and 300 have been incorporated in the design to maximize product recovery and purity. After both the stream from the flash tank and the vapor distillate from T301 have been mixed, they are sent to the first distillation tower (T201) that separates the stream such that methane and hydrogen exit as the distillate at 270 psia and -109 F, and the remaining HCs exit in the bottoms at 276 psia and 19 F. The reason behind separating methane and hydrogen first is that they are the most abundant compounds in the beginning of this section. Removing them before any other separations are made significantly lowers the amount of material flowing through the rest of the 200 section, which can help the rest of the section achieve more successful separations. Because very low temperatures are required to condense ethane, the outlet temperature of the distillate is -109 F.

28 26 FUELGAS H2SALE SP201 PR201 S204 S210 S205 C101 S110 S203 H201 Figure 5: Section 200, T201 Distillate Path After T201. The hydrogen and methane stream is sent to a heat exchanger to warm it up to room temperature before being sent to the pressure swing adsorption (PSA) unit, where the methane is removed to a fuel gas stream and the hydrogen stream is sent to a splitter (S201). This splits the hydrogen stream into a sale product stream and a recycle stream, which will be sent compressed and sent back to the reactor to prevent coking of the catalyst. The methane and hydrogen are then heated to 65 F, which is just under room temperature, and sent to a pressure swing adsorption (PSA) unit where they are separated into a hydrogen stream and a fuel gas stream using the concepts of adsorption and desorption and how they are related to changes in pressure. A PSA unit was chosen for this part of the process mainly because operating temperatures for this technology are generally near or at room temperature. Even though the amount of gas adsorbed to the adsorbent depends on both pressure and temperature, a PSA unit swings the pressure from high to low and back again theoretically with the temperature change of the unit being negligible. This makes the whole unit inherently safer, as there will not be extreme temperatures in the PSA unit like those required in the reactor, and there should be no reason why the temperature would rise suddenly, to pose other safety risks. Any noticeable temperature change could be seen as the PSA unit not operating correctly, and actions could be taken to correct the problem immediately. Under normal operating conditions, PSA units go through a two-phase cycle, which can be repeated without maintenance required at the end of each cycle. Because of this cyclical nature, the PSA unit will consist of several adsorbent vessels with staggered cycle timing so as to provide a constant and continuous flow of the product stream to continue through the process. Each vessel simply contains an adsorbent, much like our reactor will contain a catalyst. The difference between a reactor and a PSA vessel is that there are no chemical reactions that take place in a PSA vessel. The adsorbent being used

29 27 has not been specifically designed, but will probably be either zeolite or activated carbon, as both of these have very high surface area to volume ratios, which is an important factor in the amount of gas that can be adsorbed per mass of adsorbent. During the adsorption phase, the pressure is raised so that the methane will adsorb and the hydrogen will pass through, thereby creating a product stream that is almost pure hydrogen. Once the adsorbent is saturated with methane, the feed to the PSA unit will be drastically lowered to allow the pressure to decrease almost to atmospheric pressure. The swing from a high pressure to atmospheric pressure allows the methane to desorb and exit the system as a fuel gas stream with little hydrogen lost overall. This desorption phase also regenerates the adsorbent and allows it to be used in the next cycle with little adsorbance capacity lost between cycles. The desorption phase also allows the PSA unit to operate without additional maintenance after every cycle. The hydrogen product stream from the PSA unit will contain hydrogen at 99.5 mol% purity in a 6,000 lb/hr stream. Small amounts of methane will also be present due to adsorption not being a perfect separation method. After being purified in the PSA unit, the hydrogen product stream is divided between a recycle stream and a sale stream, with 15 mol%, or 1,050 lb/hr, of the total hydrogen stream being recycled back to the reactor to prevent coking in the reactor catalyst. Before reaching the reactor, this hydrogen will go through a compressor to raise the pressure of the stream to match that of the other streams being mixed to enter the reactor. It will then enter the heat exchanger and furnace that the feed stream enters to attain the 1150 F operating temperature of the reactor. The remaining hydrogen will be sold back to the oil refinery that this chemical process is located on. A global market for hydrogen is basically nonexistent, because hydrogen is difficult to store without it escaping easily due to its small size. Despite the lack of a global market, refineries often use large quantities of hydrogen, since it is a valuable feedstock for many of the refinery s processes. These processes often include hydrogenating large hydrocarbons to break them into smaller pieces that will be used in other processes or simply sold as fuels. Selling the hydrogen in this stream back to the refinery is a natural economic decision, as it increases the revenue of this process

30 28 beyond that from merely the main separated benzene, toluene, and para-xylene products. The separated methane will be released from the PSA unit into a fuel gas stream in the quantity of 74,000 lb/hr. The fuel gas stream will contain roughly 50mol% methane, with impurities of hydrogen, ethane and ethylene. Propane, propylene and C4 hydrocarbons will also be present, but in negligible amounts. This fuel gas stream will be burned in one of the furnaces in the system, either the furnace before the reactor to heat the reactants to the temperature required for the reactor or the furnace used in the reboiler for a distillation column (T301) in the 300 section. S120 C102 T201 S220 T202 S206 S207 M301 S304 Figure 6: Section 200, T201 Bottoms Path. The bottoms of T201 is sent to T202, where C2 and C3 hydrocarbons are distilled off and sent to a compressor before being recycled back to the recycled to increase the overall conversion of the reactor. The bottoms stream is sent to a mixer in the product recovery section to recover any BTX products that could have been lost. The bottoms stream from T201 is pumped to the next distillation tower (T202) to separate the ethane, ethylene, propane, and propylene from the C4 and C5 HCs. The C2 and C3 HCs will be recycled back to the reactor at a rate of 151.7Mlb/hr to achieve a higher overall conversion of the system. This outlet stream will be at 320 psia and 57.9 F straight out of the tower, so it will have to be recycled to the heat exchanger and furnace before the reactor to be heated to reactor conditions once again. Before it

31 29 reaches the heat exchangers, a compressor (C102) will be used to pressurize the stream to match the pressure of the feed stream. The bottoms from T202, which is at 206 psia and 229 F and is mostly C4 and C5 HCs, will be sent to Section 300, the product separation section. Just over 72 mol% of this particular stream is composed of benzene, toluene, and para-xylene products that can and should be recovered to maximize our total product stream, which in turn helps to maximize the economics of this process. As mentioned previously, the low temperatures being used to condense these light HCs requires heat removal beyond the capabilities of cooling water in the condensers of both distillation towers. To solve this problem, two refrigeration systems were introduced to the process to allow for extremely low temperatures in the process. Ethylene and propane were chosen as refrigerants because of their capability of working together to achieve the very cold temperatures that the separation processes requires in the condensers. Both refrigeration systems work in tandem similarly to how they work in LNG plants and are described in much more detail later in the 400 and 500 sections of this report. The -109 F temperature in the condenser of T201 has been addressed with ethylene refrigeration. The propane refrigeration system will primarily be used in the condenser of T202 to drop the temperature to 1.7 F. Propane refrigeration is also required to condense the ethylene used in T201. The reboilers in T201 and T202 use cooling water as their heating fluid. This allows for the cooling water to reach lower temperatures than normal and is integrating into cooling in H102 in the 100 section and in the condenser in T302 in the 300 section.

32 30 Section 300: Separation and Recovery of BTX and Heavy Aromatics Products (Saud) Figure 7: Heavy Separation (Section 300). From flash tank D101, the heavy stream is separated remaining light hydrocarbons. The remaining heavies are separated into Benzene, Toluene, and Xylene product and TMB byproduct. The process of heavy separation of the BTX product starts at the flash drum. The liquid effluent mainly contains aromatics and a small amount of light hydrocarbons. In order to separate this heavy product, the stream is fed into a distillation column (T301). This distillation column will separate the effluent into a BTX mixture (S302), a 1,3,5-trimethylbenzene (TMBPROD) product and will recover the light hydrocarbons. In this step, the 1,3,5-trimethylbenzene (TMBPROD) stream is separated and will be sold without further purification. The BTX mixture stream (S302) will be mixed in a mixer (M301) with the light hydrocarbon stream from Section 200 that contains some of the escaped BTX (S207). The new stream (S304) will then enter another distillation column (T302) for further separation. The separated light hydrocarbons will be directed to the purge stream which will be used for utilities. As for the BTX mixture (S305), it will go into a valve (V301) to drop the operating pressure. Lowering the pressure will help to separate the BTX mixture into its components. The low pressure stream (S306) will then go into another distillation column (T303) for further separation. In this distillation column, benzene is separated from the stream. As for the toluene and xylene mixture (S307), it will go into another distillation column (T304) for further separation. Toluene is then separated from the xylene, and all products will be sold without further purification.

33 31 T301: TMB recovery and light separation: Figure 8: Heavies Separation (Section 300). From flash tank D101, the liquid stream is separated fed into the first distillation column (T301) to remove the remaining light hydrocarbons as well as recover TMB as a product. The aromatic rich stream is sent on for further processing by S302. Starting from the flash drum (D101), the liquid effluent stream will be fed to a distillation column (T301) at a pressure of 290 psia and temperature of 70 F. This stream contains mainly aromatics, but it still contains a fair amount of light hydrocarbons. In order to recover these light hydrocarbons, the stream is fed into a distillation column (T301). This distillation is designed to divert all remaining light hydrocarbons back up to Section 200 (to T201) in stream (S303) while still recovering aromatics and separating the BTX from the by-product, 1,3,5-trimethylbenzene. The liquid distillate stream (S302) containing the BTX mixture will be carried to another distillation column (T302). The by-product 1, 3, 5- trimethylbenzene (TMBPROD) stream will be sold afterwards without further purification. The variations of operating pressure and temperature inside the distillation column is between 280 to 289 psia, and the temperature is 280 to 619 F with a flow rate of lb/hr for the BTX liquid distillate stream (S302). The distillation column (T301) is made of stainless steel and has 51 actual stages.

34 32 T302: Purges and Remaining Light Recovery: Figure 9: Purge Stream. BTX rich streams are fed into tower T302, one of which comes from the lights separation section, and one of which come from the previous tower, T301. T302 separates out any remaining lights and purges them from the system. The bottoms of the tower is sent on for product recovery as it mainly consists of BTX. The liquid distillate of T301 (S302) contains a BTX mixture and small fractions of light hydrocarbons. Stream (S207) contains some of the BTX mixtures that escaped during the separation in distillation column (T301). The stream is at a pressure of 206 psia and a temperature of 229 F. To ensure that the escaped BTX is accounted for, both streams will then be mixed in the mixer (M301). The combined stream (S304) will then go into the distillation tower (T302) for further separation. During this process, this hydrocarbon stream is separated at a pressure of 280 psia and a temperature of 280 F. The distillation tower has 65 stages and made out of stainless steel. The reason for using stainless steel instead of carbon steel is the presence of hydrogen in this separation. The tower is designed to divert 99.9mol% of all the remaining light HCs into the Purge stream. The purge stream will then be used for utilities, specifically the furnaces in sections 100 and 300 (F101 and F301). The tower will also recover 99.9mol% of the aromatics (S305) which is then directed to another distillation column (T303). The variations of

35 33 operating pressure and temperature inside the distillation column (T302) is between 200 to 210 psia, and the temperature is 72 to 418 F. A temperature of 72 F is only obtained through the use of the cooled cooling water created in the 200 section. The bottoms BTX stream undergoes a pressure of 210 psia and a temperature of 418 F. T303: BTX recovery: Figure 10: Benzene Recovery. Benzene is recovered from the BTX rich stream leaving T302. The bottoms of the tower is sent on to recover the remaining Toluene and Xylene. The liquid effluent stream (S306) then enters the distillation column (T303). After a valve (V301) to lower the pressure, the stream will be at a pressure of 50 psia and a temperature of 285 F. This distillation column is designed to separate benzene from the BTX mixture. The liquid distillate stream (BENZPROD) contains 99.6 wt% pure benzene at a pressure of 35 psia and temperature of 233 F with a flow rate of 52,080.2 lb/hr. The liquid bottoms stream (S307) contains toluene and para-xylene at a pressure of 43 psia, a temperature of 318 F, and a flow rate of 32,499 lb/hr. This stream (S307) will then be carried into another distillation column (T304) for further separation. This column (T303) has 37 stages and is made out of carbon steel. The absence of hydrogen in this separation makes carbon steel a viable building material for this tower.

36 34 T304: Toluene, and Xylene recovery Figure 11: Toluene and Xylene Recovery. T304 separates toluene from para-xylene that is fed to the tower from T303. The liquid bottoms stream (S307) from T303 then enters the last distillation column (T304). This inlet stream is at a pressure of 43 psia and a temperature of 318 F. The column is designed to separate the TX mixture into toluene and xylene products. The liquid distillate stream (TOLPROD) contains 99.6 wt% pure toluene at a pressure of 20 psia, a temperature of 254 F, and a flow rate of 25,937.5 lb/hr. The liquid bottoms stream (XYPROD) contains para-xylene with a flow rate of lb/hr, a pressure of 29 psia and a temperature of 331 F. This distillation column (T304) has 57 stages and is made out of carbon steel. Carbon steel can also be used for this tower because of the absence of hydrogen in this separation. Section 400/500: Propane and Ethylene Refrigeration (Aaron) As mentioned in the discussion about the lights recovery section (Section 200), the condensers in both distillation towers require refrigeration. The first tower in section 200 (T201) cools the vapor distillate down to -109 F while the second tower (T202) cools the vapor distillate down to 2 F. A

37 35 temperature of -109 F requires the use of ethylene refrigeration (section 500) and a temperature of 2 F requires the use of propane refrigeration (section 400). Propane refrigeration is also needed in order to condense ethylene and so it is modeled as multi-stage refrigeration in order to obtain the proper temperatures. The Aspen flow diagram for these sections is shown in Figure 12. The process is also modeled using the Redlich-Kwong Wilson property method. Optimal operating conditions were determined through the use of pressure-enthalpy data of both the propane and ethylene. T202 Condenser Section 400 T201 Condenser Section 500 Figure 12: Propane and Ethylene Refrigeration. Section 400 consists of two propane refrigeration cycles operating at different pressures. Section 500 consists of only one ethylene refrigeration cycle. For each cycle the refrigerant is compressed, condensed, expanded, and evaporated in order to complete the cycle. Propane refrigeration is used in condensing the process fluid in T202 along with condensing the ethylene in section 500. Ethylene refrigeration is only required to condense the process fluid in T201.

38 36 Process Description Both refrigeration loops are modeled using a generic refrigeration process. In both sections, a refrigerant is compressed to a desired pressure, one that brings the refrigerant into a temperature range that allows it to be condensed through the use of the cooling method available. For the propane refrigeration, the pressure was chosen so that air cooling would be capable of condensing the propane. The ethylene refrigeration was compressed such that it could be condensed from propane refrigeration (H501). After condensation, the refrigerant goes through adiabatic expansion in a Joule-Thompson valve in order to decrease the pressure while simultaneously decreasing the temperature to the value necessary to be used in the evaporators, or in other words the process heat exchangers in need of refrigeration. These temperature values are found by giving a 10 F difference between the refrigerant temperature and the process condenser temperature. The refrigerant is then sent to the heat exchanger that required the refrigeration and it is evaporated during the process. Since evaporation is an endothermic process it requires the intake of energy and it is this process that works as the actual refrigeration. The vapor then is recompressed and the cycle is repeated. Since the propane refrigeration needs to be multi-stage refrigeration, there are a couple of differences. In this case, there are two cycles for two different operating conditions that are integrated in order to save on overall compression and energy costs. Both cycles still however follow the same basic refrigeration concept outlined above where the refrigerant is compressed, condensed, expanded, and evaporated. The lower loop that can be traced corresponds to the lower pressure cycle, which is also the refrigeration for the evaporator (H501C) used to condense the ethylene in section 500 (H501H). Starting out the low pressure cycle, the propane is compressed (C402) up to the pressure of the high pressure propane and the streams are combined. The combined streams are then compressed to a determined pressure and condensed in an air cooler as described previously. The Joule-Thompson valve then that follows (V401) drops the pressure of the combined streams down to that of the higher pressure loop and the amount needed for the condenser of T202 (H204C) is split and diverted to that evaporator. The

39 37 remaining refrigerant is sent to another Joule-Thompson valve (V402) to drop the pressure further to what is required for the evaporator used in section 500 (H501C). Operating Conditions Determining optimal operating conditions for refrigeration is actually very important. When initially designing this section, values were not optimally chosen and resulted in very expensive compression. In fact, it required approximately $75,000, ,000,000 more in FCI just due to an excess of required compression. In order to choose optimal operating conditions for refrigeration, pressure, temperature, and enthalpy data needs to be consulted for the refrigerant being used. The first step in optimizing operating conditions started with finding the temperature that the refrigerant should be at in order to evaporate at an acceptable temperature. A basis of at least 10 F temperature difference between the process fluid being condensed and the refrigerant was chosen, because less than this can cause control issues. The high pressure refrigeration cycle in section 400 is designed for the condenser in T202 (H204C), which operates at a temperature of 2 F. A temperature of the propane was chosen to be -10 F. This means that the corresponding pressure in the evaporator becomes approximately 30 psia, which is the pressure for the high pressure cycle. The ethylene refrigeration cycle is designed for the condenser in T201 (H202C), which operates at a temperature of -109 F. A temperature of -125 F was chosen meaning that the pressure of the ethylene in the evaporator then becomes approximately 35 psia. The low pressure refrigeration cycle in section 400 is designed for condensing the ethylene in the ethylene refrigeration section (H501C). This temperature is decided by what the ethylene is compressed to. The colder propane is, the lower the required compression is in the ethane refrigeration. In order to hopefully minimize the compression, the temperature of the propane in this section was chosen to be -44 F as this corresponds to saturated propane at 15 psia, or just above atmospheric pressure. It is important to keep the pressure in the process above atmospheric so that if there is a break anywhere in the process line, propane will flow

40 38 out of the system instead of oxygen rushing in. See Table 2 and Table 3 for pressure, temperature, and enthalpy data [Bühner]. Table 2: Pressure, temperature and enthalpy data for propane Propane Pressure Temperature Enthalpy P [bar] P [psi] T [ C] T [ F] h L [J/g] h V [J/g] Δh [J/g]

41 39 Table 3: Pressure, temperature and enthalpy data for ethylene Ethylene Pressure Temperature Enthalpy P [bar] P [psi] T [ C] T [ F] h L [J/g] h V [J/g] Δh [J/g] After finding the different evaporators temperatures and pressures, compressor discharge pressures were determined. These values were determined by what is being used to condense the specific refrigerant. For the propane refrigeration cycle, air cooling was chosen. It is possible that using cooling water could be more economic, but for now air cooling (H401) will be considered. A base temperature of the air being used was set to be around that of air in the gulf coast of 100 F, which means the temperature

42 40 of the condensed propane was determined to be around 115 F. The pressure-temperature data in Table 2 shows that this temperature corresponds to a pressure of approximately 230 psia. A pressure drop of 5 psi across the air cooler was assumed and so the main compressor (C401) was determined to need a discharge pressure of 235 psia. For the ethylene refrigeration cycle, as mentioned previously, propane refrigeration is used. The temperature for the propane in the evaporator (H501C) has already been determined to be - 44 F. This means that the ethylene is condensed (H501H) at approximately -34 F, which corresponds to a saturation pressure of 235 psia. Taking a pressure drop of 5 psi through the condenser, the discharge pressure of the ethylene compressor (C501) was determined to be 240 psia. Considering now that every heat exchanger has a pressure drop of 5 psi, a lot of the needed operating conditions are known. Most of the remaining conditions are set by the refrigeration cycle itself. In all three evaporators, H204C, H501C, and H202C, the discharge was set to be at the dew point. Similarly, in all of the condensers, H401 and H501H, the discharge was set to be at the bubble point. The only thing remaining to be specified is the flow rate required for all of the evaporators in order to meet the process needs. Starting with ethylene refrigeration, the duty of the condenser in T201 is calculated in Aspen and found to be 72 MMBtu/hr. Using the difference in the enthalpy between the input and output streams of the ethylene in the evaporator, the mass flow rate of ethylene in the refrigeration cycle was determined. Taking the condenser duty and dividing it by this difference, the flow rate required was determined to be 530,000 lb/hr. This process was then repeated for the propane refrigeration section using the condenser duty of T202 and the condenser in section 500 (H501H) along with the corresponding enthalpy of the inlet and outlet. This calculation resulted in a propane flow of 27,000 lb/hr in the high pressure loop and 1,100,000 lb/hr in the low pressure loop. A summary of this data including flowrates and duties is shown in Table 4.

43 41 Table 4: Various operating conditions specified for the different refrigeration cycles involved along with the Aspen unit operation or stream that it corresponds to. Operating Condition Refrigerant Flowrate (Mlb/hr) Evaporator Duty (MMBtu/hr) Evaporator Pressure (psia) J-T Valve Discharge Pressure (psia) Compressor Discharge Pressure (psia) High Pressure Propane Cycle (Section Upper Loop) Unit Operation or Stream Spec'd Value Spec'd Low Pressure Propane Cycle (Section Upper Loop) Unit Operation or Stream Spec'd Value Spec'd Ethylene Cycle Unit Operation or Stream Spec'd Value Spec'd HPFEED 33 LPFEED 1455 ETHYLENE 563 H204C 3 H501C 119 H202C 76 H204C 30 H501C 15 H202C 35 V V V C C C As far as modeling this process in Aspen, the only needed piece of information that hasn t been cover is the split fraction for the total flow as to how much propane is diverted to the low pressure and high pressure cycles. This value was calculated by taking the flow rate in the low pressure cycle and dividing it by the combined flow rate of both cycles. This value gives the split fraction that is diverted to the lower pressure cycle. This calculation along with that of each of the mass flows is implemented into different calculators making the simulation more robust to changes made to the process. Overall the process being used is pretty sound. Aspen results for pressure, temperature, and enthalpy data matched up with data that was found in literature (Bühner). It also is quite similar to refrigeration used in natural gas plants. Some things that could be done that might further reduce capital include finding ways to decrease the amount of needed refrigeration in the process, analyzing different refrigerants, and considering water cooling as an alternative to air cooling. A location to search for the best reduction in capital cost and utility cost is in the compression. When optimizing temperatures and

44 42 pressures throughout the process the overall compression was dropped from a total of 70,000 hp down to 44,000 hp. This, while resulting in a fair drop in utility costs, dropped the FCI by $75,000, ,000,000. This process as is, is likely not perfect and can be improved but is likely quite good.

45 43 VII. Design Alternatives (Bridger) Throughout the initial design phase, the production plant for conversion ethane to aromatics underwent much iteration. Several of these design plans were still able to yield successful, economic plants under many different circumstances. These changes vary in the number of units, schematics, operating conditions, and level of product separation. The primary design was created with the intention of maximizing IRR while making the plant as applicable as possible. Possible Reactor Alterations The largest design flexibility lies in the reactor section (Section 100). Some of these available reactor alternatives may likely be used, as pilot testing may reveal unexpected catalyst properties. US patent, US A1, lists operating condition ranges for catalytic functionality of the zeolitegermanium-platinum catalyst. While the functional range for pressure is listed from 20psia to 2000psia, the stoichiometry of the reactions, seen in table 1, reveals that more moles of gas are present in the products than in the reactants. According to La Chatlier s principle, increasing the pressure of a reaction with more moles in the product will slow the reaction, and therefore reduce conversion of the feedstock. This was not reported in the patent, as lab conversion tests were performed solely at atmospheric pressure, but is suspected to occur, even slightly, once a larger scale plant is constructed. Ideally, the reactor pressure of 300psia was chosen to reduce the reactor s size since the catalyst shows no indication of significant conversion losses with increasing pressure. At atmospheric pressure, the reactor becomes so large it becomes wholly infeasible both physically and economically. The cost of filling the high reactor volume exceeded all plant costs. However, the reactor/s can still operate at any pressure between these two points. This allows the pressure to be dropped significantly without reaching gigantic proportions. These changes are primarily dependent on further lab tests, or optimization at the pilot scale.

46 44 Economically, lowering the pressure creates larger, or a larger number of, reactors, thereby increasing the overall capital costs. These reactors must be filled with a larger amount of catalyst, further increasing the cost. At a constant conversion, the highest-pressure reactor is the most economically viable since it require the least materials and smallest equipment to produce the same amount of product. If further research tests reveal that increased pressure reduces conversion, an optimization will have to be performed that take into account the costs of increasing the reactor size as well as the conversion losses. It is important to note that due to the C2 and C3 recycle, stream S120, the feed is still utilized to 100%. The majority of losses associated with the loss of single pass conversion are associated with larger recycle stream and increased equipment size, along with additional heating and cooling. An optimized reactor sizing will likely remain within the process specifications and retain viability. Product Recovery (Section 300) Alternative Designs Another key area with the possibility of alternative design is the product recovery section (Section 300). In the proposed plant, benzene, toluene, and xylene are each distilled to their pure chemical standards, typically above 99.5% purity. In the current design, this allows for the maximum profit yield. Previous design schematics forewent the final separations and sold a simple, mixed benzene, toluene, and xylene (BTX) stream. Various pros and cons are present within the mixed stream sale design, but it was eventually foregone for a more robust process which can applied to many more opportunities and remain stand alone as a producer in the market. In previous design iterations, a mixed BTX stream was sold. Selling a mixed BTX stream reduces the costs associated with separation. This includes the capital cost of multiple towers (T303 and T304) as well as associated labor, maintenance, and utilities. The original idea was to sell this mixed stream at a reduced chemical price to a BTX distillery. Many operations that involve BTX are equipped with the proper units to process and distill the BTX to its pure components. By selling the produced BTX

47 45 at a slightly reduced price, this allows the partner distillation company to make a small profit while processing the products. This approach was foregone primarily due to the idea that finding a company to purchase mixed BTX at a defined price was much more niche than creating a stand-alone plant. To provide a more secure basis for the company, assumed associations, complications, and potential falling points were removed from the design. The mixed stream sale is also difficult to estimate using economic analyses. The prices of individual components are set based on the market, while selling a mixed stream to a distillery is based on the individual contract, as well as market conditions. If the BTX stream is treated as a fuel component, its individual chemical value is lost. The fuel market is the primary consumer of BTX products, in which case the sale price falls significantly. If a company is found to purchase mixed BTX near the chemical price, reducing the equipment capital cost is one potential method to improve economically. Using the mixed BTX strategy still proved economically viable, even before the economic improvements made in other sections. This shows a high potential for a mixed BTX selling plant if the proper conditions can be met in terms of sale price, etc. The previous plant IRR was roughly 25% before extensive optimization of the refrigeration section. The IRR for the designed plant is 32.7%. It comes to reason that an optimized plant with a mixed BTX stream should have an IRR between these values. Another design alternative concerning the product stream is to include the first of the two product separation distillation columns, T303 and T304. The first column (T303) separates chemical grade benzene, which can be sold for 44 cents per pound. Benzene is the most valuable aromatic product, which allows for high value sales without the separation of toluene and xylene. The toluene and xylene stream can be sold similarly to the mixed BTX stream above, either to be distilled, or as a fuel additive. The toluene and xylene separation tower is fairly small, but still contributes significant capital cost. This design scheme could be best utilized if a reliable company cannot be found for high price mixed aromatics sales, and there are capital cost constraints that must be met.

48 46 Continuous Catalyst Regeneration One alternative design that could be applied to improve the functionality of the catalyst regeneration system is to include continuous regeneration. Currently, three reactor vessels are used, and alternated to run two reactors at any given time while the third undergoes regeneration. This method is fairly efficient, as the catalyst will need to be regenerated often and takes a short time to go through the four regeneration steps. By running two reactors, the plant remains at full capacity at all times. The primary drawback to having three reactors is the capital cost associated with both the extra reactor vessel and catalyst, particularly in a unit that is not used for one third of operating hours. By implementing a continuous catalyst regeneration cycle, the extra reactors and catalyst can be removed. This dramatically reduces capital costs associated with the reactor section, but will also add a complex continuous unit. More research would need to be performed regarding the catalyst regeneration to create a continuous process. In a continuous regeneration operation, the catalyst is continuously removed from the reactor and reinserted in its regenerated state. This allows fresh catalyst to continuously be present inside the reactor. This is a benefit compared to the batch regeneration because in batch, the catalyst slowly decreases in activity before reaching the regeneration threshold. With the continuous presence of fresh catalyst, the entire system will be able to reach steady state. This will provide constant conversions and selectivities compared to a deactivating batch process. A continuous catalyst regeneration process will simplify the management of the reactor system. The parallel reactors approach requires the switching of reactors every several months, along with the operations of the regeneration cycle. Switching and manipulating a series of gas streams repeatedly creates increased opportunity for operator error and hazards. A continuous process, although more complicated to design, is an inherently safer process.

49 47 Fuel gas reallocation, C2 through C4 repurposing Another opportunity for an alternative design is the reallocation of the C4 and C5 purge stream. Although this stream is primarily used to vent the low levels in inert C4 and C5 in the system, this stream also contains a number of other light hydrocarbons. The C2 and C3 components in the purge stream make up only a small percentage of the feed, but these HCs still have the potential to react into higher value products. Similarly to mixed BTX, mixed light HCs could be sold to a company prepared to process the stream into its components for either reaction or sale. This would provide more revenue for the stream as opposed to using it as a fuel gas. This small percentage of waste HCs were not processed in the current design due to excessively low temperatures and therefore high refrigeration costs. This HC stream is very small relative to the other product streams. The economic impact of changing this stream design appears negligible, but is once again difficult to estimate due to the unreliability of selling a pre-product to the further refined. The prices are not readily available and a specific price with the purchasing company would have to be negotiated.

50 48 VIII. Permitting and Environmental Concerns (Emily) Environmental issues associated with the design of this chemical plant include those of accidental as well as operational releases of any of the process chemicals. Accidental releases include any incidences of the process fluid escaping the process in a place where and/or when it is not designed to leave the system. These events could be as small as a leak in a joint of two pipes or as large as an explosion, and include everything in between such as a pressure relief valve opening due to a buildup of pressure. These events could correlate to a range of releases from almost negligible to extremely large and everything in between. Accidental releases usually can t be predicted, and often occur in an emergency situation when a major part of the process is failing and safety is the most pressing issue due to the concerns of keeping the workers safe and trying to save the process from being destroyed. Environmental impacts of such releases are hard to predict, as many hydrocarbons could be released and they may respond differently to being released into the environment. Very small releases may simply dilute very quickly and react or combust on a very small scale as to almost be not noticeable. Very large releases may explode or start a fire if contained within a building, or they may disperse in the ambient air and the larger molecules could potentially settle out of the atmosphere before reacting and affect soils or bodies of water near the release. Operational releases for this particular process will simply be the products of combustion associated with burning the fuel gas stream in the furnaces of the process. Two separate furnaces will be used to heat the reactor inlet stream before it is reacted and to act as the reboiler of T301. The furnace before the reactor is required because of the very high temperature of 1150 F required in the reactor that can t be feasibly reached with superheated steam in a heat exchanger. The reboiler of T301, which requires a temperature of 618 F, is also hot enough to make superheated steam in a heat exchanger become unfeasible. All combustion reactions in these furnaces are expected to produce CO 2, but are also assumed to produce thermal NO x due to the high temperatures of the flames and the fact that air, which is mostly nitrogen, will be used to provide oxygen to the flame to continue combustion. CO 2 emissions can be estimated using a simple mass balance: any carbon that goes in to be burned must come out, assumedly

51 49 in CO 2. NO x emissions are harder to predict since nitrogen is not purposefully being burned, and there is a large quantity of it available in the air being used to supply oxygen to the flames. The United States Environmental Protection Agency (EPA) is in charge of keeping track of how much and what kinds of pollutants are emitted by industrial processes, and to do this, they devised an equation that could predict emissions based on the what is being combusted, how the combustion unit is configured, and what control methods are being used to lower the emissions. Equation 1 is the equation given by the EPA in AP 42, Compilation of Air Pollutant Emission Factors, to estimate emissions ( Emissions Factors ). Equation 1 is commonly used in estimation calculations, because it simplifies the interactions between the specifications of the design so that the general population can use the equation. Many charts make up the rest of AP 42, and they contain values for some of the variables, such as emission factor and overall emission reduction efficiency, based on measurements of actual emissions from different combustion units. E = A x EF x (1-ER/100) Equation 1 Where: E = emissions, tons/year; A = activity rate, MMBtu/year; EF = emission factor, lb/mmbtu; and ER =overall emission reduction efficiency, % Initial estimates show that the furnace fuel streams will contain no more than MMBtu/hr of thermal energy combined. These streams are made up of mainly methane, which is very similar to natural gas, so the EPA natural gas combustion estimates and factors will be used for the emissions calculations for reasons of simplicity. Because of the large capacity of both furnaces, they will both be considered Large Wall-Fired Boilers to accurately decide which factors should be used. Without control measures,

52 50 these estimates provide a calculation resulting in 205 tons/year of thermal NO x being emitted, which is over the limit of 100 tons/year, since NO x is a criteria pollutant, for needing a permit from the EPA to operate the furnace. Possible control measures include using low NO x burners either alone or with flue gas recirculation (FGR). According to AP 42, low NOx burners achieve this result by breaking up the combustion process into stages. The smaller stages draw out the process of combustion over a longer period of time, which results in a cooler flame than normal combustion and suppression of the formation of thermal NOx. AP 42 also says that low NOx burners generally reduce emissions by 40 to 85%. FGR works with recycled flue gas to dilute the combustion air. This mainly reduces combustion temperatures, but also reduces the oxygen concentration in the primary flame zone. Both of these factors contribute to the formation of thermal NO x, so the more they are reduced, the more the NO x emissions are reduced. The combination of low NO x burners with FGR can be estimated to reduce emissions by 60 to 90%. From the tables in AP 42, the Emission Factor for uncontrolled boilers is 190 lb/10 6 scf. In the same units, the Emission Factors for low NO x burners alone and low NO x burners combined with FGR are 140 and 90, respectively ( Emissions Factors ). Using these numbers in the equation above, the emissions can be lowered to less than 100 tons/year of thermal NO x, which is within the acceptable limits, as shown in Table 5. Table 5: This table shows the emissions of thermal NO x with and without control measures. The emission limit for needing a permit from the EPA is 100 tons/year, which can be obtained with control measures in this process. Emissions (tons/year) Pre-Reactor Furnace Reboiler Furnace Total Uncontrolled Low NOx burners Low NOx burners and Fuel Gas Recirculation

53 51 An air permit from the EPA may be required because the uncontrolled emissions are higher than those allowed without a permit, but control measures will be implemented to lower the emissions to acceptable rates. Analyzing the Best Available Control Technology (BACT) is slightly difficult, as it takes into account the cost of the control technology and compares it to how well it works at controlling emissions. Many options will need to be considered for this analysis, as certain controls may work in one furnace but not the other. Both of the control methods explained above rely on lowering the flame temperature, which may be quite suitable for the reboiler furnace, since it only has to reach a temperature greater than 618 F to transfer heat to the process fluid at that point. The furnace before the reactor needs to have a flame temperature greater than 1150 F in order to heat the feed and recycle streams to the reactor s operating temperature. If the control methods lower the flame temperature below 1150 F, thermodynamic laws will have to be violated to raise the temperature of the process stream to the operating temperature of the reactor. To avoid violating thermodynamic laws, other control methods will have to be considered that operate on principles other than lowering the flame temperature.

54 52 IX. Safety and Risk Management (Emily) When operating as designed, this design should pose no safety issues. However, issues will arise from time to time simply because life is unpredictable. This process involves hydrogen as well as many different hydrocarbons, which means that the risk of fire and explosion can be extremely high if all chemicals are not handled properly. Any leaks, even if they are small, can produce hazardous working conditions for the operators and maintenance crews, as well as anyone else who happens to be working near this process. Benzene leaks are especially hazardous, as benzene is a known carcinogen. Safety data sheets for each chemical present in this process are available in Appendix 4 for further information on their safety concerns. Besides being hazardous to the environment and the workers near it, leaks are also hazardous to the process itself, as a leak was not designed for, and the system may not be able to react to such an unexpected situation. As seen in the HAZOP analysis in Appendix 3, there are many plausible situations that could result in unknown problems to the rest of the process. These situations include temperatures and pressures outside of normal operating conditions, as well as improper mixing in the several mixers throughout the design. The many unit operations included in the process design present many opportunities for malfunctions or failures to occur, as they all include pipes that go in and out, as well as valves that control the flow in these pipes, in addition to the configuration of the equipment itself and the control technology that it includes. Some unit operations have inherently safe design considerations, such as the PSA, which operates at ambient temperatures, but even these more robust pieces of equipment cannot save the process from other pieces of equipment not working properly. The problems that are predicted to happen on a regular basis, such as the catalyst deactivating and heat exchangers scaling, can be addressed with regular maintenance, and should not cause catastrophic consequences as long as these problems are taken care of before they create emergency situations. The reactor is the least inherently safe unit operation in the design simply because it operates at very high temperatures and contains extremely flammable materials. Because many problems can be predicted from this unit operation, it will be under

55 53 careful supervision to ensure than none of these problems arise. For example, normal operating conditions are such that the flammable materials in the reactor are at concentrations larger than the upper explosive limit. Large deviations in pressure or temperature in the reactor could lead to runaway reactions, use of a pressure relief valve, and even an explosion if oxygen is allowed to enter the system through a leak or if the process gas is allowed to leak out into the ambient air. Although reactor failure may have the most severe consequences of all the unit operations, almost any malfunction or failure could cause the whole process to be shut down simply because of the recycle streams and fuel gas streams that connect each section to one another, and also because these unit operations will have to occupy a small total geographical space to accommodate the requirement that this process be built on or near an existing oil refinery.

56 54 X. Project Economics (Aaron) Equipment and Capital Cost The costs for all of the individual unit operations, except for the PSA unit and pumps, were determined through the use of Peters and Timmerhaus [Peters]. These costs were determined as FOB costs from The prices of the equipment were scaled through time using average chemical engineering cost indices from 2002 and 2015 to obtain the cost in 2015 dollars [Economic Indicators 1 and 2]. This FOB cost was then converted into fixed capital investment. The FCI for a PSA unit was found separately in 2011 dollars and was also scaled into 2015 dollars using cost indexes [Economic Indicators 1 and 2]]. Taking an estimate of OSBL being 10% of ISBL, and knowing that FCI is equal to ISBL plus OSBL, ISBL and OSBL were back calculated. Pumps Pumps are generally rather small pieces of equipment. Since this process mostly exists in the gaseous phase, the only pumps needed in the process are those needed for reflux in all distillation towers along with pumps to move cooling water around. This makes pumps somewhat difficult to model but a general rule of thumb exists that says pumps account for around 5% of the FOB cost of all equipment so this assumption will be used for this case [Myers]. Compressors Since this process mainly exists in the gas phase and requires refrigeration, compressors account for a large chunk of the overall capital cost. All compressors were modeled in Aspen as polytropic compressors with an efficiency of 0.75 using the ASME method. Sizing compressors depends on the inlet actual volume flow and the power used. Compressors were considered to be centrifugal and driven by a motor. Carbon steel was used for all compressors that didn t contain significant hydrogen. If significant hydrogen was present, they were sized as stainless steel. Some compressors required more

57 55 power than what are commonly available. For these cases, the compressor modeled in Aspen is broken up into multiple units where in each unit falls within the range of commonly available compressors. For specifics on compressor sizing and costing, see Table 6. Aspen ID Table 6: Specific information involved in sizing and costing compressors. Description Inlet Actual Volume Flow [Mcf/hr] KW Used Number of Units KW/Unit FOB Cost MM$ C101 Hydrogen Rec $ 0.14 C102 Light Rec $ 0.40 C401 Prop Ref HP $ C402 Prop Ref LP $ 3.80 C501 Ethylene Ref $ 6.40 Total $ Turbines Only one turbine is used in this process and it was sized using the same information as that of compressors; inlet actual volume flow and power produced. The turbine was modeled in Aspen as an isentropic turbine with an isentropic efficiency of The turbine was designed to be made out of carbon steel since there isn t any hydrogen present. For specifics on turbine sizing and costing, see Table 7. Aspen ID Table 7: Specific information involved in sizing and costing turbines Description Inlet Actual Volume Flow [Mcf/hr] KW Produced Number of Units KW/Unit FOB Cost MM$ (2002) X101 Feed Pres Drop $ 0.2

58 56 Furnaces Due to the excess fuel gas made in this process, furnaces are reasonably available methods to heat streams that need to be heated to fairly hot temperatures. Furnaces also present an opportunity to create steam that is to be used around the plant. Since the heat is made from combustion, the hot gases created during the process can be used as a heat source to generate the steam. This integration has not been taken into account as of yet. These furnaces were designed as box type with horizontal radiant tubes. For specifics on furnace sizing and costing, see Table 8. Table 8: Specific information involved in sizing and costing furnaces Aspen ID Description MMBtu/min Material FOB Cost MM$ (2002) F101 Feed Preparation 2.17 Stainless Steel $ 3.10 F301 T301 Reboiler 2.03 Carbon Steel $ 1.70 Total $ 4.80 Heat Exchangers All heat exchangers were considered to be fixed-tube-sheet heat exchangers with 0.75 inch OD x 1 inch square pitch and ft bundles. Pressure factors were taken into consideration. Stainless-steel was used for all heat exchangers that contained a significant amount of hydrogen. Values for the overall heat transfer coefficient, U, were found by using common values for that between the process fluids involved [Myers]. Steam was considered to be saturated at 525 F. Generally, cooling water was modeled to enter at 95 F and exit at 115 F unless it was used to heat where in the cooling water was modeled to enter at 95 F and exit at 60 F. In a few instances, 95 F was not cold enough and so the cooling water that was used to heat was then integrated to be used to cool as well. In order to justify the ability to do this, look at the energy transferred to the water in the reboilers and the energy needed to be transferred using this special cooling water. Around 40 MMBtu/hr is

59 57 removed from the cooling water in the reboilers of T201 and T202 while only 10 MMBtu/hr is needed from the special cooling water and so there is a sufficient amount of special cooling water. Some heat exchangers that were sized fell outside of the range of the normal size of heat exchanger. If this was the case, it was broken up into separate units until each unit fell within the common size range. For specifics on heat exchanger sizing and costing, see Table 9. Table 9: Specific information involved in sizing and costing heat exchangers ID Description Material Duty, [MMBtu/hr] U [Btu/hr- F-ft2] LMTD [ F] Pres [psia] Area [ft 2 ] Units Area per unit [ft 2 ] FOB [M$] (2002) H101 Reactor Feed-Effluent SS $ Reactor Out Cooler H102-1 (General CW) SS $ Reactor Out Cooler H102-2 (Special CW) SS $ H201 Hydrogen Heater SS $ H202 T201 Condenser SS $ H203 T201 Reboiler CS $ H204 T202 Condenser CS $ 4.97 H205 T202 Reboiler CS $ 7.31 H301 T301 Condenser CS $ H303 T302 Condenser CS $ 8.82 H304 T302 Reboiler CS $ 7.32 H305 T303 Condenser CS $ H306 T303 Reboiler CS $ 7.28 H307 T304 Condenser CS $ 2.02 H308 T304 Reboiler CS $ 4.83 Ethylene Cooled by H501 Propane CS $ Total [MM$] $ 0.83 Air Coolers The only air cooler required in the process is the propane condenser in the refrigeration section. This particular piece of equipment may be replaced by just a regular heat exchanger but for now this type

60 58 of exchanger works just fine. It was designed to be made out of carbon steel. Air was assumed to enter the exchanger at 95 F and exit at 105 F. This approximates conditions in the gulf coast on a fairly warm day. For specifics on air cooler sizing and costing, see Table 10. Aspen ID Table 10: Specific information involved in sizing and costing air coolers Description Material Duty [Btu/hr] U [Btu/hr- F-ft 2 ] LMTD [ F] Pres [psia] Area [ft 2 ] Units Area per unit [ft 2 ] FOB [MM$] (2002) H401 Air Cooler CS $ 0.34 Vessels Vessels involved in this process include the flash tank in section 100, along with all of the reflux drums that are required. The size of the tank was determined to be the amount of volume that would be filled by 10 minutes of the flow to the tank. It was then assumed that the length is three times the diameter and so resulting in the diameter and the length dimensions. The flash tank was considered to be vertical which was corrected for by a 10% increase to the FOB cost. Pressure factors and material factors were taken into account where the material was chosen to be carbon steel if there isn t a significant amount of hydrogen involved. If there is significant hydrogen, the material was chosen to be stainless steel. For specifics on vessel sizing and costing, see Table 11.

61 59 Aspen Id Table 11: Specific information involved in sizing and costing vessels Description Horiz or Vert Total Liquid Vol Flow [Mcf/hr] Drum Total Volum e [ft3] L [ft] D [ft] Pres [psia] Materia l FOB [M$] (2002) D101 Light/Heavy Sep Vert SS $ D201 T201 Reflux Drum Horiz SS $ D202 T202 Reflux Drum Horiz CS $ 6.96 D301 T301 Reflux Drum Horiz CS $ D302 T302 Reflux Drum Horiz CS $ 4.35 D303 T303 Reflux Drum Horiz CS $ D304 T304 Reflux Drum Horiz CS $ 7.00 Total [MM$] $ 0.24 PSA unit Pressure swing adsorption units are not as common of equipment. In order to cost and size this unit a similar unit was found that had been already been costed and sized [Analysis of Natural Gas-to Liquid Transportation Fuels via Fischer-Tropsch]. The particular PSA unit found sized was done in 2011 dollars and had a hydrogen production of 7091 lb/hr. Using the Six-Tenths Factor Rule [Myers], the cost of the PSA unit found was adjusted to the size of PSA for the EBTAX process. For specifics on PSA sizing and costing, see Table 12. Table 12: Specific information for costing the PSA unit Aspen ID Description EBTAX H 2 Production [Mlb/hr] Similar Costed PSA Production [Mlb/hr] Similar Costed PSA FCI [MM$] EBTAX PSA FCI [MM$] (2011) SP201 Hydrogen Recovery $ $ 17.46

62 60 Catalyst For this particular case, catalyst is assumed to last the life of the project and is a one-time cost. A patent for the catalyst chosen gives a gas hourly space velocity for lab tests that were conducted. Using the volumetric flow of gas to the reactor, the volume of catalyst was calculated. In order to allow regeneration of the catalyst to occur without changing production, three reactors are used, containing half of the calculated volume of catalyst each. It is difficult to price the catalyst as it is a rather specific makeup but generally these catalysts cost $/lb and so the price was chosen to be 100 $/lb [Myers]. The density of similar catalysts was found to be lb/ft 3 and so the density was chosen to be 59 lb/ft 3. For specifics on sizing and costing catalyst used, see Table 13. Table 13: Specific information for sizing and costing the amount of catalyst used Description GHSV [1/hr] Total Vol. Flow To Reactor [Mcf/hr] Catalyst Mass Needed [lb] FOB [MM$] (2015) Catalyst $ 6.02 Reactor As described in the catalyst section, three reactors are used in parallel. Each reactor uses half of the needed catalyst for the desired production. The size of each reactor is taken to be twice the volume of catalyst used to account for support, screening, and distribution of feed. The reactor itself was sized the same way as vessels from here. The length was determined to be three times the diameter. Pressure and material factors were taken into account where in the material for the three reactors is glass lined stainless-steel, since there is a significant amount of hydrogen present and chlorine is used in the regeneration process. The vessels were also all taken to be vertical and so the cost was increased by 10%. For specifics on sizing and costing the reactors, see Table 14.

63 61 Aspen ID Table 14: Specific information for sizing and costing the amount of catalyst used Description Rx Total Volume [cf] L [ft] D [ft] Pres [psia] FOB [MM$] (2002) R101 Reactor $ 0.24 Tower and Trays Distillation towers require two main components, a shell and trays. The height of each of the columns was determined by the number of real stages. Assuming 75% efficiency for each stage, the number of theoretical stages was converted into the number of real stages for each tower. The height of the column was then determined by assuming a 2 ft tray spacing and adding an additional 14 ft for various tower needs such as support and distribution. The diameter of the tower and trays was calculated in Aspen by designing the tower at 80% flood. An additional 5% was added to the cost of the shell to account for manways and other various column needs. Pressure factors and material factors were taken into account for the shell where in the towers containing significant amounts of hydrogen were considered to be made out of stainless-steel. For specifics on sizing and costing results of trays and distillation columns, see Table 15. Aspen ID Table 15: Specific information for distillation column and tray sizing and costing Description Number of Actual Stages Tower Height [ft] Aspen Diameter [ft] Total FOB [MM$] (2002) T201 Demethanizer $ 0.62 T202 Purges $ 0.11 T301 TMB Recovery/Lights Sep $ 0.23 T302 Purges Remaining Lights $ 0.89 T303 BTX Recovery $ 0.12 T304 TMB/Xylene Recovery $ 0.17 Total $ 2.14

64 62 After compiling this costing information, the FCI for the overall process in 2015 dollars was determined. First, all costs were scaled to 2015 dollars if they needed to be. After that, all costs that were FOB needed to be converted into FCI so that all values could be summed together to result in the FCI for the design. To do this, they were multiplied by 1.1 to account for delivery and then multiplied by 5.04 to convert it into FCI [Peters]. These results are compiled into Table 16. Table 16: Fixed Capital Investment for the various equipment involved in the process along with the resulting total Equipment Quantity Total FCI (2015 MM$) Pumps N/A $ Compressors 7 $ Turbines 1 $ 1.66 Furnaces 2 $ Heat Exchangers 33 $ 6.89 Air Coolers 1 $ 2.86 Vessels 7 $ 2.03 PSA Unit 1 $ Reactor 3 $ 1.99 Catalyst 1 $ Tower+Trays 4 $ Total 60 $ ISBL N/A $ OSBL N/A $ Pricing, Revenue and Production Cost Finding prices for raw materials and chemicals was rather difficult, but reasonable estimates were found. Ethane was found to be around $3.75/MMBtu [Brown]. Benzene, toluene, and xylene were found to be priced at $3.45/gal, $2.8/gal, and $2.84/gal respectively in 2008 dollars [Chemicals A-Z]. This price was scaled into 2015 dollars through the use of economic indexes [Economic Indexes 1 and 2]. TMB was considered to be sold as a gasoline additive and so its price was chosen to be $2.50/gal to reflect a price slightly higher than gasoline itself. Hydrogen was difficult to price since it doesn t really have a market, and most people that need hydrogen just make it themselves. Since this is the case,

65 63 hydrogen is priced in terms of how much it costs to make it, which turns out to be $0.65/lb [James]. Lastly, the purge stream and fuel gas stream are just burned for energy and so it is priced as natural gas at $2.40/MMBtu [U.S. Natural Gas Wellhead Price]. Densities and heating values were gathered in Aspen in order to convert all prices to a c/lb basis in 2015 dollars. A summary of these values and the resulting cost or income is shown in Table 17. Table 17: Income or cost of each of the materials consumed or produced Product or Raw Material Price [c/lb] Flowrate [Mlb/hr] Income or Cost [MM$/yr] Ethane Benzene Toluene Xylene TMB Hydrogen Fuel Gas Purge After this, the remaining costs of utilities and fixed costs were determined. Steam was assumed to be priced at $8.00/MMBtu while cooling water was assumed to be priced at $0.40/MMBtu [Myers]. Natural gas was already found to be priced at $2.40/MMBtu [U.S. Natural Gas Wellhead Price]. Lastly electricity was assumed to be at a price of 4 c/kwh. Table 18 shows the summary of utility costs. Fixed costs were all based off of assumptions made by John Myers in his economic notes [Myers]. For labor, this design was assumed to have four men per shift, each making $50/hr with a 60% increase in cost to account for overtime. Maintenance was assumed to be 4% of FCI. Laboratory costs were assumed to be 10% of labor. Plant overhead was assumed to be 30% of labor, maintenance, and lab costs combined. Lastly, taxes and insurance were assumed to be 3% of FCI. A summary of fixed costs can be found in Table 19.

66 64 Table 18: Cost of utilities Utility Price [$/MMBtu or c/kwh] Energy Requirement [MMBtu or kwh] Cost [MM$/yr] Steam Cooling Water Natural Gas Electricity Table 19: Various fixed costs associated with the design Fixed Cost Price Basis Yearly Cost (MM$/yr) Labor 4 50$/hr *1.6*8760hr/yr 2.80 Maintenance 4% FCI Laboratory 10% labor 0.28 Plant Overhead 30% of (labor, maintenance, lab) 4.76 Taxes and Insurance 3% FCI 9.59 Cash Flow Analysis The cash flow that resulted for this design used assumptions found in Peters and Timmerhaus [Peters]. Start-up cost was determined to be 10% FCI and working capital was determined to be 89% of installed FOB. A build up period of two years was used where the production rate would start at 75% nameplate, move up to 90% nameplate the following year, and finally reaching 100% in the third year. A construction period of 3 years was used where 25% of the FCI was spent in the first year, 50% of the FCI was spent in the second year, and the rest was spent in the third year. The project life was considered to be 20 years and no scrap value was taken into consideration. A tax rate of 35% was used along with MACRS5 depreciation. The minimum annual rate of return, or MARR, was decided to be 25%. The results of the cash flow are compiled into Table 20. For more detail on any sizing, costing, and general economics such as the cash flow and the production cost estimate, see Appendix 1.

67 65 Table 20: Results of the cash flow analysis conducted on this design FCI (MM$) NPV0 (MM$) NPV10 (MM$) PBP (yrs) 2.6 IRR 30% Sensitivities Various sensitivities were run for uncertainties in the design itself. Prices for products and reactants were varied, prices for utilities were varied, and FCI was varied. Some very specific sensitivities were also ran for things like reactor size, purchasing more catalyst throughout the project life, and feed composition. The results for this is compiled into Table 21 and was plotted to form a tornado diagram, Figure 13. Even though the parameters were varied quite a bit, the only parameters that were capable of dropping the IRR below or close to the MARR were the FCI, ethane price, and the benzene price. Decreasing all prices simultaneously, which could happen since all materials scale with natural gas and oil price, also lowered the profitability quite a bit. This, however, makes sense since these prices and costs play the largest contribution to the cash flow.

68 66 Table 21: Sensitivities run, along with the resulting IRR Parameter Variation Min Max Base Range ± 40% FCI 21.78% 46.41% 30.27% 24.63% Ethane Price 4-16 c/lb 14.80% 35.54% 30.27% 20.75% ± 40% Benzene Price 21.94% 37.55% 30.27% 15.61% All Products and Reactants Price Increase or Decrease 19.20% 32.38% 30.27% 13.18% Reactor Size at 30 psia 22.82% 30.27% 30.27% 7.45% ± 40% Toluene Price 27.05% 33.33% 30.27% 6.28% ± 40% Hydrogen Price 27.53% 32.89% 30.27% 5.35% ± 50% NG price 28.97% 31.55% 30.27% 2.58% Replacing the Catalyst every 5 years 28.63% 30.27% 30.27% 1.64% ± 40% Xylene Price 29.45% 31.08% 30.27% 1.62% ± 50% Electricity Price 29.73% 30.80% 30.27% 1.07% ± 30% Catalyst Price 30.14% 30.96% 30.27% 0.82% ± 50% Steam Price 30.07% 30.47% 30.27% 0.39% ± 50% Cold Water Price 30.19% 30.36% 30.27% 0.17% Feed Composition with minimum propane 30.27% 30.33% 30.27% 0.06%

69 67 Figure 13: Tornado Diagram. This plot chose the change on IRR based on different variation of various uncertain parameters. ± 40% FCI Ethane Price 4-16 c/lb ± 40% Benzene Price All Products and Reactants Price Increase or Decrease Reactor Size at 30 psia ± 40% Toluene Price ± 40% Hydrogen Price ± 50% NG price Replacing the Catalyst every 5 years ± 40% Xylene Price ± 50% Electricity Price ± 30% Catalyst Price ± 50% Steam Price ± 50% Cold Water Price Feed Composition with minimum propane 0.00% 5.00% 10.00% 15.00% 20.00% 25.00% 30.00% 35.00% 40.00% 45.00% 50.00% IRIRR

70 68 XI. Global Impacts (Saud) The recent popularity of fracking among oil companies in the USA has been increasing the availability of ethane, especially in the Gulf Coast area, as a feedstock. The BTX market is estimated to have a total value of $81 billion in sales per year. Although the petrochemicals manufacturing companies are more profitable and are expected to grow with an average rate of 3% over the next five years to $98 billion in At this point the BTX market is considered to be developing due to the production and the demand in the market. The BTX market is still considered a good market for business because of the shortage in supply due to an increase in demand. The production of BTX emits significant amounts of CO 2 and NO x which contributes towards global warming which have direct impact on the environment. That is why this process requires government permits in accordance with the federal and state regulations. Currently the plant does not required permit for emitting CO 2. The reason for that it is not regulated by EPA. For the NO x the EPA required permit. To understand the petrochemicals market economically, we must consider a basis to analyze the level of competition within an industry by making a business strategy. Introducing this five forces analysis, also called the Porter Five Forces, is a structure which used widely in industry to evaluate the competitive forces that must be considered when the investor is willing to enter a prospective market. The founder of this idea, Michael Porter, made this tool to analyze the five market forces that determine the competitive industry and study the target market. The five forces are the competitive rivalry, the power of suppliers, the power of the buyers, the threat of substitution, and the threat of new entry.

71 69 Figure 14: Industry Rivalry. This figure illustrates the possible industry pressure associated with a new competitor 1-The competitive rivalry: There are relatively few sources, other than fossil fuels, from which hydrocarbons can be produced, for this reason there is a low potential in the market for new inventions for producing BTX. Some markets are saturated with the amount of BTX being produced. For example, East Asia has a large excess of benzene, which is being addressed by sometimes cutting the production to avoid significant price drops in the market. The producer of the BTX are dependent on the price of the crude oil to indicate their profit, therefore an alternative method should be considered, such as natural gas, which is more economically beneficial. Because of that, the expected amount of the competitive rivalry is high. 2-The power of suppliers: The natural gas that is required for our process, is produced in abundance in the Gulf coast regions that is why the suppliers ability to affect the industry is less significant. One of the reasons for this is that the price of natural gas is determined by the natural gas market and also because the expense of transportation of natural gas from an offshore site is fairly large. Since there is a large number of refineries in the region, which means more suppliers are in the market, hence it will be more convenient