Shale Gas to Ethylene (G2)
This report is part of an ongoing project to design a process to produce ethylene and NGLs from shale gas. The inlet raw shale gas is taken from the Barnett Shale site in Texas with an inlet flow rate of shale is 171,200 kg/hr. The design would be implemented in Johnson County, TX, which is nearby the shale site. The Barnett shale contains about 17% ethane, which we seek to separate from the inlet gas and convert into ethylene via steam cracking. Additionally, because the shale contains significant amounts of other hydrocarbons, natural gas, propane, and NGLs will be recovered from the shale to be sold.
The proposed design consists of six main sections which were modeled in HYSYS: acid gas removal, dehydration, fractionation train, hydrogenation reactor, ethylene splitter, and steam cracking. The acid gas removal section uses Diethylamine (DEA) to remove the CO2 from the raw shale gas. The DEA is fully recovered, thus no make-up stream is required, and the CO2 content is reduced from 0.66 to 0.04 mol%. The purpose of the dehydration section is to remove all of the water from the shale before it is sent through the fractionation train, as water will freeze due to the low temperatures in this section. Triethylene glycol (TEG) is used to absorb water from the gas, reduce the water content from 2.44 to 0.01 mol%. TEG is then partially recovered and recycled. The fractionation train consists of 4 distillation columns. The first is the demethanizer, which yields a methane-rich tops that is sent to a second distillation column to purify methane for sale, and the bottoms is sent to the third column, the deethanizer. The deethanizer ethane-rich tops is sent on for further processing, while the bottoms is sent to the fourth column, the depropanizer, with yields a propane top stream and an NGL bottom stream. The next section is the hydrogenation reactor, into which the tops of the deethanizer is fed to selectively hydrogenate acetylene, which is an impurity from the cracking reactor, into ethylene and ethane, reducing the content of acetylene in the gas from 100 ppm to 4 ppm. Next, they hydrogenated gas proceeds to an ethylene splitter which separates ethylene for sale, methane to recycle, and ethane which proceeds to the final section, the cracking reactor. The cracking reactor operates at about 820C and converts ethane to ethylene, with the cracked gas containing 48% ethylene. The cracked gas is then quenched and recycled for purification.
This discussed design results in the following product streams: 26,600 kg/hr of >98% purity ethylene , 6,300 kg/hr of 94% pure hydrogen, 68,600 kg/hr of 96% pure methane, 30,600 kg/hr of 97% pure propane, and 38,200 kg/hr of NGLs, which is 72% butane and 26% pentane. The total profits from the sale of these products surmounts to about $328 million annually. The annual cost of this design includes the cost of purchasing feeds, offsite organic waste treatment, utilities (determined from Aspen Energy Analyzer), labor, and other miscellaneous costs. This annual cost is $250 million. Additionally, the total capital cost, determined from estimated equipment sizes as well as added contingency for piping and other fixtures, sums to $120 million. From these costs, the design has a 10 year NPV of $547 million, a payback period of 3 years and an internal rate of return (IRR) of 72%. These calculations indicate a profitable design that should be pursued. However, this design and the resulting calculations should be refined with additional researching and alternatives beyond what is presented in this report.
This report summarizes a project to design a process to produce ethylene and NGLs from shale gas. The production of ethylene is desirable because ethylene is a valuable product in the production of polymers, chemicals and other consumer products. In this final report, the design and simulation in HYSYS are discussed in detail. In addition, the rationale behind the selection of the specific technology chosen for the design is provided. The objectives when choosing specific technology were to maximize efficiency and conversion while minimizing capital and operating costs of the facility. Finally, an economic analysis was conducted, including capital costs, variable costs and fixed costs analyses with an overall NPV calculation. Additionally, a sensitivity analysis was conducted to determine the effect of certain design tradeoffs and changes in market pricing.
The inlet to the process is raw shale gas that is taken from the Barnett Shale site in Texas (specifically, shale gas extracted from well C1 of the Newark East field). The inlet flow rate of shale is 171,200 kg/hr and the composition of this raw shale gas is shown in Table A.1 of the appendix (American Petroleum Institute, 2011). It is assumed that there is little to no sulfur in the raw shale gas. The plant will be located nearby in Johnson County, Texas. This location is near the extraction site of the Barnett shale gas to avoid large transportation distances and costs but far enough away from the Dallas-Ft. Worth area to ensure adequate land availability and privacy.
While ethylene is the only product chemically produced, hydrogen, methane, propane, and NGLs are recovered in the process and can be sold. To be sold, the following purities must be met or exceeded: 98% pure ethylene, 90% pure hydrogen, 95% pure methane and propane, and 95% purity NGLs (Ethylene Gas Specifications, 2014 and Keller, 2012). Thus, the approach taken in the development of this design is to take raw inlet shale, refine it to separate out methane, ethane, propane and NGLs and then use the ethane to produce ethylene via steam cracking. Further processing and refinement also leads to the recovery of hydrogen (an input added to the hydrogenation reactor and produced in the steam cracking process).
The proposed design consists of six main sections: acid gas removal, dehydration, fractionation train, hydrogenation reactor, an ethylene splitter and steam cracking (refer to Appendix B for the process flow diagram). Carbon dioxide must first be removed from the inlet shale gas in the acid gas removal section. This is accomplished using an absorption column. The resulting shale gas is termed “sweet gas”. The sweet gas is then mixed with cracking gas from another part of the plant before it is sent to another absorption column so that water can be removed. This dehydrated stream is then sent onto the fractionation train, where it is cooled then sent through a series of cryogenic distillation columns to separate out methane, ethane, and propane. The methane stream is further refined to separate out methane from light gases (H2, O2, and N2). The ethane stream continues on to the rest of the process. It contains a small amount of acetylene (a product of the cracking reactor) which is removed in a hydrogenation reactor. After the acetylene is removed, the stream is then sent to another cryogenic distillation column, the ethylene splitter. The splitter recovers methane which is recycled back to the demethanizer. Additionally, a side draw separates out ethylene which is to be sold. The remaining stream, consisting mostly of ethane, is fed into a cracking reactor with steam to produce ethylene. After exiting the reactor, it is quenched to stop the reaction. This stream is then recycled and combined with the sweet gas so that it can be purified.
The design described above was simulated in Aspen HYSYS to determine necessary stream specifications, required utilities and sizing information for the various pieces of equipment for this process. A picture of the setup in HYSYS and details of the HYSYS results, including stream and equipment specifications, can be found in Appendices C, D and E.
The fluid package used for the main case of the simulation was Peng-Robinson (PR) which was chosen because it is most compatible with hydrocarbons (National Technical Institute of Athens). However, PR is not compatible with the absorbent, diethylamine (DEA), that is required in the acid gas removal portion of the process. Thus, it was necessary to simulate the acid gas removal (CO2 absorption column & DEA recovery distillation column) in a sub-flowsheet that utilizes the amine fluid package.4 All other processes were simulated within the main case.
Acid Gas Removal
The first step in the process is the removal of carbon dioxide from the inlet shale gas. This is done with an absorption column with DEA as the absorbent. The inlet shale gas is fed into the absorption column, T-100, at a flow rate of 171,200 kg/hr, at 6900 kPa (1000psi) and around 50C (Transform Software and Services). The composition of the inlet shale is shown in Table 1 of the Appendix A (American Petroleum Institute, 2011). Recycled aqueous diethylamine is fed into the column at similar temperature and pressure (Oli Systems, Inc.). The vapor exiting the top of the column is the shale gas without CO2 and comes out the column at 170,000 kg/hr, with a temperature of around 55C and pressure of 6700 kPa. The absorber reduces the CO2 component fraction in the shale gas from 0.66 to 0.04 mol%.
The bottoms, the spent DEA, exit from the column at 32,400 kg/hr at a temperature around 50C and pressure of 6800 kPa. The spent DEA is sent through a distillation column, T-101, in order to regenerate the DEA and minimize material cost. However, the pressure of the spent DEA must first be brought down to 202 kPa for distillation to be viable.7 The distillate is CO2 vapor and the bottoms are recovered DEA. The bottoms are mixed with 900 kg/hr makeup water to compensate for water loss in the tops, pumped back up to 6895 kPa, and then reenter the absorption column. All of the DEA is recovered. The distillation column reduces the CO2 component fraction in the DEA stream from 4.73 to 0.93 mol%.
After exiting T-100, the sweet gas is mixed with the recycled cracking gas and fed into the dehydration absorption column, T-102, along with triethylene glycol (TEG) to absorb the water. The gas comes in at a rate of 271,000 kg/hr and the TEG at a rate of 39,800 kg/hr, both with a temperature around 53C and pressure of 6700 kPa. T-102 reduces the water component fraction of the sweet gas from 2.44 to 0.01 mol%. The distillate of T-102 is the dehydrated gas and the bottoms is the spent TEG. The dehydrated gas exits at 266,000 kg/hr at a temperature of 57C and pressure of 6550 kPa.
The spent TEG comes out at a rate of 45,100 kg/h, with a temperature of 58C and pressure of 6650 kPa. The pressure of the TEG is then reduced to 400 kPa and the temperature is increased to 150C in order to prepare it for distillation.8 The amount of water in the TEG stream is reduced from 49.16 to 1.03 mol%. The tops of T-103 is mostly water vapor and the bottoms is the recovered TEG that is then cooled, mixed with additional pure TEG to compensate for losses in the overhead, and then pumped back up to pressure before being recycled into absorption column T-102. Although these temperature and pressure changes require excess energy, recovery of TEG using distillation is only viable at the lower pressure and higher temperature specified above in order to prevent the degradation of TEG (Nivargi, Gupta, Shaikh and Shah, 2005).
Acid Gas Removal
Separation of Methane and Hydrogen
Fixed and Variable Costs
Overall Economic Feasibility
Conclusion and Recommendations
Appendix A: Shale Gas Composition
Appendix B: Process Flow Diagram
Appendix C: HYSYS Diagram