Patent Application
20110232327
The invention concerns the efficient processing of off gas to recover ethane, ethylene, and higher hydrocarbons.
Off gas, for example that produced by a refinery (refinery off gas) or an olefins plant, is generally composed of methane, hydrogen, ethane, ethylene, propane, propene, and heavier hydrocarbons. If recovered, the hydrocarbons are valuable product which otherwise would be lost with the off gas in the plant's fuel gas system.
Refinery off-gas usually contains H2, CO, CO2, O2, CH4, C2H4, C2H6, C3H8, C3H6 together with some trace impurities such as such as oxygen, ammonia, nitriles, acetylenes, sulfur compounds, butadiene, chlorides, arsenic, mercury, and water in addition to acid gases H2S, CO2, and COS. These off-gases are produced from refinery units that manufacture conversion products such as hydrotreaters, alkylation units, fluid catalytic cracking units, platformers, etc. Valuable products including hydrogen, olefins, natural gas liquids (NGL) and higher Btu fuel gas can be recovered from the off-gas if an off gas processing unit is installed.
Similar to refinery off gas, the off gas from olefins plants can also be processed to recover valuable products. The off gas from olefins plants typically is richer in ethylene or propylene and the off gas has different species of trace impurities from those in the refinery off gas.
Other plants, as well, may produce off gas with C2 and higher hydrocarbons, for which the method of the present invention may be useful in providing cost effective recovery of valuable C2 and heavier hydrocarbons.
Currently, these valuable hydrocarbons may be recovered from the off gas by at least two different methods. A circulating lean oil process may be used to absorb propylene and heavier components from refinery off gases. Although the absorption process provides a reasonable recovery of propylene and heavier components, it is energy intensive and requires several pieces of operating equipment. The amount of equipment needed generally leads to an increased quantity of control loops and the need for expensive plot space.
Cryogenic expander based technologies are increasingly used in preference to the lean oil absorption methods, because these technologies provide higher ethylene and ethane recoveries. A typical cryogenic expander based process involves a series of progressive cool-down steps in plate fin heat exchangers and vapor-liquid separation steps, followed by demethanization.
Currently, turbo expanders are used in combination with external refrigeration to increase the thermodynamic efficiency of the process, thus achieving higher percentages of natural gas liquids (“NGL”) recovery. The requirement of external direct refrigeration requires more equipment, controls, and instrumentation, as well as storage and handling of the refrigerant that is used. The storage of refrigerant also raises additional safety considerations due to these extra hydrocarbons being stored at the plant site.
Off gas is usually available at a relatively low pressure of about eighty psia. To achieve higher NGL recoveries, the cryogenic expander based units require feed gas compression. The compression of dirty off gas is troublesome during operations. The off gas composition is a mix of waste gas coming out of various units. These units may operate at different capacities, and any one or more of them may not be operating at any particular time. Thus, an off gas stream will vary appreciably in composition and flow rate depending on the source and the types of units operating at a particular time.
Generally, the compressors can be designed for a range of composition for the feed gas. However, it is difficult to predict the range of composition and flow fluctuation for the off gas. Any change in composition outside the design range will result in reduced capacity or loss of recovery of NGL. Similar problems are faced in turbo expander operations. Moreover, if the content of heavier hydrocarbons increases in the off gases then condensation of these hydrocarbons takes place at higher pressure in the upstream section, resulting in loss of valuable NGL.
Various contaminants that appear in off gas also cause mechanical problems for rotating machinery, resulting in sometimes frequent maintenance downtime and a resulting significant loss of revenue. The variations in off gas feed stream mol weights and flow characteristics also cause problems for turbo expanders used in off gas processing, again often resulting in significant maintenance downtime. Similarly, unsteady operating conditions can result in leakages in heat exchanger cores. The fluctuations in composition of the off gas also affects refrigeration requirements, thereby affecting the external direct refrigeration system.
In an attempt to circumvent at least some of these problems, less efficient reciprocating compressors are often used to compress the off gas feed stream. However, it would be more desirable to process the feed gas without compression.
Thus, it is desirable to provide an efficient process for off gas processing that has good adaptability to the feed composition variation.
Another challenge for this recovery process is to keep the operating temperatures above certain levels to reduce the risk of blue oil formation.
It is also an object of the invention to recover the valuable hydrocarbons (C2+) from off gas without, or with minimal, compression of the feed gas.
It is a further object of this invention to extract the valuable hydrocarbons from off gas by using as part of the apparatus a turbo expander for which the refrigerant is product, feed gas, reflux formed during an intermediate part of the process, or a mixture of two or more of these. Using these sources for the refrigerant eliminates the need for storage of a specific refrigerant type. Further, use of a turbo expander in the refrigerant loop also helps to startup the plant at reduced capacity, allowing the plant to generate the required refrigerant needed to attain the full capacity of the plant.
It is yet another object of the invention to efficiently recover ethane and ethylene from the off gas in a cost effective process.
The method of the present invention alleviates many of the concerns discussed above. Utilizing this method, no feed gas compression is required at the inlet for most cases. In some cases, feed gas may be available at a lower pressure than usual, for example, approximately fifty psia. In such cases, it may be desirable to compress the feed gas to about eighty psia, but even in such circumstances, the amount of compression needed is minimal compared to the prior art, the expected ranges more predictable and easier to design for, and the expected stress on (and resultant maintenance needs of) compression equipment will be significantly lowered.
Feed gas is chilled and the heavier hydrocarbons are separated in the low pressure separator. This low pressure operating point maintains the feed gas far away from the phase envelope, almost eliminating the possibility of hydrocarbon condensation upstream of the dehydrators.
Additionally, the method of this invention minimizes the effects of changes in operating temperature due to process operating conditions changes upstream, or due to ambient heat loss. No temperature controls are required for feed gas, eliminating expensive control systems.
Further, there is no turbo expander in the feed gas stream, eliminating the exposure of turbo expanders to dirty feed gas. A turbo expander is utilized on the refrigerant side, and is thus exposed only to clean refrigerant, thus reducing expected turbo expander maintenance downtime. Because the refrigerant used is a mixture of partially processed feed gas and product streams, no storage of refrigerant is required. In fact, the plant may be started up on feed gas without the refrigerant present, allowing production of refrigerant “on the fly.” This method also allows higher recovery of ethane and ethylene without lowering the operating temperature below certain cryogenic temperatures, thereby avoiding formation of blue oil.
In one embodiment of the invention, dehydrated off gas arrives as feed gas at a temperature of approximately 100° F. and a pressure of approximately 80 psia. This feed gas is cooled to approximately −80° F. in a first heat exchanger (preferably a brazed aluminum plate fin exchanger), yielding condensed hydrocarbon as part of the feed. The condensed hydrocarbon is separated in a low pressure separator and pumped back through the heat exchanger, where it aids in cooling the feed gas, then to a distillation column. The condensed hydrocarbon is warmed to approximately 90° F., and preferably arrives at the distillation column at a pressure of approximately 355 psia.
The vapor separated from the low pressure separator is routed through the first heat exchanger and also aids in cooling the feed gas, and is then compressed to approximately 580 psia in a two stage centrifugal compressor. The inlet feed to the compressor is preferably at approximately 68° F. This compressed gas is then cooled in steps, first in a second heat exchanger (preferably an air cooler or cooling water heat exchanger), and next in the first heat exchanger to about −108° F. The hydrocarbon liquid formed as a result of cooling to a very low temperature is separated in a high pressure separator and is fed to the distillation column, preferably on the top most tray section. The separated vapor from the high pressure separator is heated in the first heat exchanger and is sent out as lean gas at about 85° F.
The distillation column preferably operates at approximately 350 psia at the bottom and approximately 340 psia at the top. The distillation column overhead is cooled in the first heat exchanger to create reflux. Condensed liquid from the reflux is separated out in a reflux drum and is then is fed to the column top tray section. The vapor from the reflux drum is combined with the vapor from the high pressure separator and leaves as lean gas after being heated in the first heat exchanger.
C2+ product is recovered from the distillation column bottom. The column bottom temperature is preferably maintained at about 88° F. This temperature makes it possible to utilize a reboiler at the distillation column bottom that exchanges heat with, and cools, the refrigerant after the final stage of refrigerant compression.
Refrigeration is provided by means of a closed loop turbo expander cycle. The refrigerant is made by mixing part of the vapor exiting the reflux drum with part of the bottom product from the distillation column. The refrigerant is compressed in a refrigerant compressor to a pressure of about 700-750 psia and cooled in steps, first in a third heat exchanger (preferably an aircooler or a cooling water heat exchanger), then in the distillation column reboiler, and finally in the first heat exchanger. After passing through the first heat exchanger, the refrigerant is at approximately −52° F.
The refrigerant is flashed in a first refrigerant separator at about 500 psia. The flashed gas is further expanded in a turbo expander to a pressure of approximately 170 psia. The pressure of the separated liquid from the refrigerant separator is let down by a control valve to the same pressure (approximately 170 psia). This liquid is then mixed together with the output gas from the turbo expander, and then enters the first heat exchanger to provide additional refrigeration. The refrigerant exits the first heat exchanger at about 70° F. and passes to a second refrigerant separator.
Gas output from the second refrigerant separator is fed to a turbo compressor associated with the turbo expander. The partially compressed gas from the turbo compressor is cooled in a fourth heat exchanger, then is fed to a third refrigerant separator. Gas output from the third refrigerant separator returns to the refrigerant compressor to complete the closed loop refrigerant system. Liquid remaining in the second and third refrigerant separators may be removed as needed via first and second control valves, respectively. If continuous condensation is observed, pumps may be added to the system to relieve this condition.
In an alternative embodiment of the invention, dehydrated off gas arrives as feed gas at a temperature of approximately 100° F. and a pressure of approximately 85 psia. This feed gas is cooled to approximately −82° F. in a first heat exchanger (preferably a brazed aluminum plate fin exchanger), yielding condensed hydrocarbon as part of the feed. The condensed hydrocarbon is separated in a low pressure separator and pumped back through the heat exchanger, where it aids in cooling the feed gas, then to a distillation column. The condensed hydrocarbon is warmed to approximately 90° F., and preferably arrives at the distillation column at a pressure of approximately 355 psia.
The vapor separated from the low pressure separator is routed through the first heat exchanger and also aids in cooling the feed gas, and is then compressed to approximately 475 psia in a two stage centrifugal compressor. The inlet feed to the compressor is preferably at approximately 68° F. This compressed gas is then cooled in steps, first in a second heat exchanger (preferably an air cooler or cooling water heat exchanger), and next in the first heat exchanger to about −118° F. The hydrocarbon liquid formed as a result of cooling to a very low temperature is separated in a high pressure separator and is fed to the distillation column on the top most tray section. The separated vapor from the high pressure separator is heated in the first heat exchanger and is sent out as lean gas at about 95° F.
The distillation column preferably operates at approximately 330 psia at the bottom and approximately 320 psia at the top. The distillation column overhead is cooled in the first heat exchanger to create reflux. Condensed liquid from the reflux is separated out in a reflux drum and is then is fed preferably to the column top tray section. The vapor from the reflux drum is combined with the vapor from the high pressure separator and leaves as lean gas after being heated in the first heat exchanger.
C2+ product is recovered from the distillation column bottom. The column bottom temperature is preferably maintained at about 82° F. This temperature makes it possible to utilize a reboiler at the distillation column bottom that exchanges heat with, and cools, the refrigerant after the final stage of refrigerant compression. Refrigeration is provided by means of a closed loop turbo expander cycle. The refrigerant is made by mixing part of the vapor exiting the reflux drum with part of the bottom product from the distillation column. The refrigerant is compressed in a refrigerant compressor to a pressure of about 310-330 psia and cooled in steps, first in a third heat exchanger (preferably an aircooler or a cooling water heat exchanger) and then in the distillation column reboiler.
At this stage, the refrigerant is partially condensed. The partially condensed refrigerant is separated in a first separator (preferably an expander suction drum separator). The vapor exiting the first separator is fed to an expander, reducing the pressure to about 135 psia. This expanded refrigerant is then further cooled in the first heat exchanger to about −110° F., then is flashed in a second separator. The vapor feed and the liquid feed from the second separator are further flashed, respectively, by first and second control valves to about 50 psia, then the vapor and liquid feeds are remixed to form a mixed stream.
The liquid separated from the first separator is also further cooled in the first heat exchanger to about −110° F., and is then flashed by a third control valve to about 50 psia. The flashed liquid stream is mixed with the mixed stream to provide refrigerant to the first heat exchanger. The refrigerant exits the first heat exchanger at about 45° F. and passes to a second refrigerant separator.
Gas output from the second refrigerant separator is fed to a turbo compressor associated with the turbo expander. The partially compressed gas from the turbo compressor is cooled in a fourth heat exchanger, then is fed to a third refrigerant separator. Gas output from the third refrigerant separator returns to the refrigerant compressor to complete the closed loop refrigerant system. Liquid remaining in the second and third refrigerant separators may be removed as needed via first and second control valves, respectively. If continuous condensation is observed, pumps may be added to the system to relieve this condition.
In another alternative embodiment of the invention, dehydrated off gas arrives as feed gas at a temperature of approximately 100° F. and a pressure of approximately 85 psia. This feed gas is cooled to approximately −65° F. in a first heat exchanger (preferably a brazed aluminum plate fin exchanger), yielding condensed hydrocarbon as part of the feed. The condensed hydrocarbon is separated in a low pressure separator and pumped back through the heat exchanger, where it aids in cooling the feed gas, then to a distillation column. The condensed hydrocarbon is warmed to approximately 42° F., and preferably arrives at the distillation column at a pressure of approximately 110 psia.
The vapor separated from the low pressure separator is routed through the first heat exchanger and also aids in cooling the feed gas, and is then compressed to approximately 110 psia in a centrifugal compressor. The inlet feed to the compressor is preferably at approximately −65° F. This compressed gas is fed to a distillation column.
The distillation column preferably operates at approximately 110 psia at the bottom and approximately 100 psia at the top. The distillation column overhead is cooled in the first heat exchanger to create reflux. Condensed liquid from the reflux is separated out in a reflux drum and is then is fed to the column top tray section. The vapor from the reflux drum is combined with the vapor from the high pressure separator and leaves as lean gas after being heated in the first heat exchanger. C3+ product is recovered from the distillation column bottom. The column bottom temperature is preferably maintained at about 77° F. This temperature makes it possible to utilize a reboiler at the distillation column bottom that exchanges heat with, and cools, the refrigerant after the final stage of refrigerant compression.
Refrigeration is provided by means of a closed loop turbo expander cycle. The refrigerant is made by mixing part of the vapor exiting the reflux drum with part of the bottom product from the distillation column. The refrigerant is compressed in a refrigerant compressor to a pressure of about 700-800 psia and cooled in steps, first in a third heat exchanger (preferably an aircooler or a cooling water heat exchanger), then in the distillation column reboiler, and finally in the first heat exchanger. After passing through the first heat exchanger, the refrigerant is at approximately −2° F.
The refrigerant is flashed in a first refrigerant separator at about 760 psia. The flashed gas is further expanded in a turbo expander to a pressure of approximately 210 psia. The pressure of the separated liquid from the refrigerant separator is let down by a control valve to the same pressure (approximately 210 psia). This liquid is then mixed together with the output gas from the turbo expander, and then enters the first heat exchanger to provide additional refrigeration. The refrigerant exits the first heat exchanger at about 70° F. and passes to a second refrigerant separator.
Gas output from the second refrigerant separator is fed to a turbo compressor associated with the turbo expander. The partially compressed gas from the turbo compressor is cooled in a fourth heat exchanger, then is fed to a third refrigerant separator. Gas output from the third refrigerant separator returns to the refrigerant compressor to complete the closed loop refrigerant system. Liquid remaining in the second and third refrigerant separators may be removed as needed via first and second control valves, respectively. If continuous condensation is observed, pumps may be added to the system to relieve this condition.
In one embodiment of the invention, dehydrated off gas arrives as feed gas 10 at a temperature of approximately 100° F. and a pressure of approximately 85 psia. Feed gas 10 is cooled to approximately −80° F. in a first heat exchanger 12 (preferably a brazed aluminum plate fin exchanger), yielding partially condensed hydrocarbon 11 as part of the feed. The condensed hydrocarbon 13 is separated in a low pressure separator 14 and pumped by first pump 16 back through the first heat exchanger 12, where it aids in cooling the feed gas 10, then to a distillation column 18. The condensed hydrocarbon 13 is warmed in the first heat exchanger 12 to approximately 90° F., and preferably arrives at the distillation column 18 at a pressure of approximately 355 psia.
The vapor feed 20 separated from the low pressure separator 14 is routed through the first heat exchanger 12 and also aids in cooling the feed gas 10, and is then compressed to approximately 580 psia in a two stage centrifugal compressor 22. The inlet feed 23 to the two stage centrifugal compressor 22 is preferably at approximately 68° F. The compressed gas 25 is then cooled in steps, first in a second heat exchanger 24 (preferably an air cooler or cooling water heat exchanger), and next in the first heat exchanger 12 to about −108° F. The hydrocarbon liquid feed 26 formed as a result of cooling to a very low temperature is separated in a high pressure separator 28 and is fed to the distillation column 18, preferably on the distillation column top tray section 29. The separated vapor 30 from the high pressure separator 28 is heated in the first heat exchanger 12 and is sent out as lean gas 32 at about 85° F.
The distillation column 18 preferably operates at approximately 350 psia at the distillation column bottom 34 and approximately 340 psia at the distillation column top 36. The distillation column overhead 38 is cooled in the first heat exchanger 12 to create reflux 40. Reflux condensed liquid 42 from the reflux 40 is separated out in a reflux drum 44 and is then pumped by second pump 46 to the distillation column top tray section 29. The reflux vapor 48 from the reflux drum 44 is combined with the separated vapor 30 from the high pressure separator 28 and leaves as lean gas 32 after being heated in the first heat exchanger 12. First pressure control valve 50 regulates the pressure of the lean gas 32.
C2+ bottom product 52 is recovered from the distillation column bottom 34. The distillation column bottom 34 temperature is preferably maintained at about 88° F. This temperature makes it possible to utilize a reboiler 54 at the distillation column bottom 34 that exchanges heat with, and cools, the refrigerant stream 56 after the final stage of refrigerant compression.
Refrigeration is provided by means of a closed loop turbo expander cycle. The refrigerant stream 56 is made by mixing part of the reflux vapor 48 exiting the reflux drum 44 with part of the bottom product 52 from the distillation column 18. (Piping omitted for clarity). The refrigerant stream 56 is compressed in a refrigerant compressor 58 to a pressure of about 700-750 psia and cooled in steps, first in a third heat exchanger 60 (preferably an aircooler or a cooling water heat exchanger), then in the distillation column reboiler 54, and finally in the first heat exchanger 12. After passing through the first heat exchanger 12, the refrigerant stream 56 is at approximately −52° F.
The refrigerant stream 56 is flashed in a first refrigerant separator 62 at about 500 psia. The flashed refrigerant gas 64 is further expanded in a turbo expander 66 to a pressure of approximately 170 psia. The pressure of the separated refrigerant liquid 68 from the first refrigerant separator 62 is let down by second pressure control valve 70 to the same pressure (approximately 170 psia). The separated refrigerant liquid 68 is then mixed together with the first gas output 72 from the turbo expander 66, and then enters the first heat exchanger 12 to provide additional refrigeration. The warmed refrigerant stream 74 exits the first heat exchanger 12 at about 70° F. and passes to a second refrigerant separator 76.
Second gas output 78 from the second refrigerant separator 76 is fed to a turbo compressor 80 associated with the turbo expander 66. The partially compressed gas 82 from the turbo compressor 80 is cooled in a fourth heat exchanger 84, then is fed to a third refrigerant separator 86. Third gas output 88 from the third refrigerant separator 86 returns to the refrigerant compressor 58 to complete the closed loop refrigerant system. Liquid remaining in the second and third refrigerant separators 76, 86 may be removed as needed via first and second control valves 90,92, respectively. If continuous condensation is observed, pumps (not shown) may be added to the system to relieve this condition.
In an alternative embodiment of the invention, dehydrated off gas arrives as feed gas 210 at a temperature of approximately 100° F. and a pressure of approximately 85 psia. This feed gas 210 is cooled to approximately −82° F. in a first heat exchanger 212 (preferably a brazed aluminum plate fin exchanger), yielding partially condensed hydrocarbon 211 as part of the feed. The condensed hydrocarbon 213 is separated in a low pressure separator 214 and pumped by first pump 216 back through the first heat exchanger 212, where it aids in cooling the feed gas 210, then to a distillation column 218. The condensed hydrocarbon 213 is warmed in the first heat exchanger 212 to approximately 90° F., and preferably arrives at the distillation column 218 at a pressure of approximately 355 psia.
The vapor feed 220 separated from the low pressure separator 214 is routed through the first heat exchanger 212 and also aids in cooling the feed gas 210, and is then compressed to approximately 475 psia in a two stage centrifugal compressor 222. The inlet feed 223 to the compressor is preferably at approximately 68° F. The compressed gas 225 is then cooled in steps, first in a second heat exchanger 224 (preferably an air cooler or cooling water heat exchanger), and next in the first heat exchanger 212 to about −118° F. The hydrocarbon liquid feed 226 formed as a result of cooling to a very low temperature is separated in a high pressure separator 228 and is fed to the distillation column 218, preferably on the distillation column top tray section 229. The separated vapor 230 from the high pressure separator 228 is heated in the first heat exchanger 212 and is sent out as lean gas 232 at about 95° F.
The distillation column 218 preferably operates at approximately 330 psia at the bottom and approximately 320 psia at the top. The distillation column overhead 238 is cooled in the first heat exchanger 212 to create reflux 240. Reflux condensed liquid 242 from the reflux 240 is separated out in a reflux drum 244 and is then is fed to the distillation column top tray section 229. The reflux vapor 248 from the reflux drum 244 is combined with the separated vapor 230 from the high pressure separator 228 and leaves as lean gas 232 after being heated in the first heat exchanger 212. First pressure control valve 250 regulates the pressure of the lean gas 232.
C2+ bottom product 252 is recovered from the distillation column bottom 234. The distillation column bottom 234 temperature is preferably maintained at about 82° F. This temperature makes it possible to utilize a reboiler 254 at the distillation column bottom 234 that exchanges heat with, and cools, the refrigerant stream 256 after the final stage of refrigerant compression.
Refrigeration is provided by means of a closed loop turbo expander cycle. The refrigerant stream 256 is made by mixing part of the reflux vapor 248 exiting the reflux drum 244 with part of the bottom product 252 from the distillation column 218. (Piping omitted for clarity). The refrigerant stream 256 is compressed in a refrigerant compressor 258 to a pressure of about 310-330 psia and cooled in steps, first in a third heat exchanger 260 (preferably an aircooler or a cooling water heat exchanger) and then in the distillation column reboiler 254.
At this stage, the refrigerant stream 256 is partially condensed. The partially condensed refrigerant stream 257 is separated in a first refrigerant separator 262 (preferably an expander suction drum separator). The vapor 264 exiting the first refrigerant separator 262 is fed to a turbo expander 266, reducing the pressure to about 135 psia. The first gas output 272 from the turbo expander 266 is then further cooled in the first heat exchanger 212 to about −110° F., then is flashed in an intermediate refrigerant separator 267. The intermediate vapor feed 269 from the intermediate refrigerant separator 267 is further flashed by first control valve 273 to about 50 psia. The intermediate liquid feed 271 is regulated by second control valve 275, and the intermediate vapor feed 269 and the intermediate liquid feed 271 are remixed to form a mixed stream 277.
The separated refrigerant liquid 268 from the first refrigerant separator 262 is also further cooled in the first heat exchanger 212 to about −110° F., and is then flashed by a third control valve 279 to about 50 psia. The flashed liquid stream 281 is mixed with the mixed stream 277 to provide refrigerant 283 to the first heat exchanger 212. The refrigerant 283 exits the first heat exchanger 212 at about 45° F. and passes to a second refrigerant separator 276.
Second gas output 278 from the second refrigerant separator 276 is fed to a turbo compressor 280 associated with the turbo expander 266. The partially compressed gas 282 from the turbo compressor 280 is cooled in a fourth heat exchanger 284, then is fed to a third refrigerant separator 286. Gas output from the third refrigerant separator 286 returns to the refrigerant compressor 258 to complete the closed loop refrigerant system. Liquid remaining in the second and third refrigerant separators 276, 286 may be removed as needed via first and second control valves 290, 292, respectively. If continuous condensation is observed, pumps (not shown) may be added to the system to relieve this condition.
In another alternative embodiment of the invention, dehydrated off gas arrives as feed gas 310 at a temperature of approximately 100° F. and a pressure of approximately 85 psia. Feed gas 310 is cooled to approximately −65° F. in a first heat exchanger 312 (preferably a brazed aluminum plate fin exchanger), yielding partially condensed hydrocarbon 311 as part of the feed. The condensed hydrocarbon 313 is separated in a low pressure separator 314 and pumped by first pump 316 back through the first heat exchanger 312, where it aids in cooling the feed gas 310, then to a distillation column 318. The condensed hydrocarbon 313 is warmed in the first heat exchanger 312 to approximately 42° F., and preferably arrives at the distillation column 318 at a pressure of approximately 110 psia.
The vapor feed 320 separated from the low pressure separator 314 is compressed to approximately 110 psia in a centrifugal compressor 322. The inlet feed 323 to the centrifugal compressor 322 is preferably at approximately −65° F. The compressed gas 325 is fed to distillation column 318.
The distillation column 318 preferably operates at approximately 110 psia at the distillation column bottom 334 and approximately 100 psia at the distillation column top 336. The distillation column overhead 338 is cooled in the first heat exchanger 312 to create reflux 340. Reflux condensed liquid 342 from the reflux 340 is separated out in a reflux drum 344 and is then pumped by second pump 346 to the distillation column top tray section 329. The reflux vapor 348 from the reflux drum 344 leaves as lean gas 332 after being heated in the first heat exchanger 312. First pressure control valve 350 regulates the pressure of the lean gas 332.
C3+ bottom product 353 is recovered from the distillation column bottom 334. The distillation column bottom 334 temperature is preferably maintained at about 77° F. This temperature makes it possible to utilize a reboiler 354 at the distillation column bottom 334 that exchanges heat with, and cools, the refrigerant stream 356 after the final stage of refrigerant compression.
Refrigeration is provided by means of a closed loop turbo expander cycle. The refrigerant stream 356 is made by mixing part of the reflux vapor 348 exiting the reflux drum 344 with part of the bottom product 353 from the distillation column 318. (Piping omitted for clarity). The refrigerant stream 356 is compressed in a refrigerant compressor 358 to a pressure of about 700-800 psia and cooled in steps, first in a third heat exchanger 360 (preferably an aircooler or a cooling water heat exchanger), then in the distillation column reboiler 354, and finally in the first heat exchanger 312. After passing through the first heat exchanger 312, the refrigerant stream 356 is at approximately −2° F.
The refrigerant stream 356 is flashed in a first refrigerant separator 362 at about 760 psia. The flashed refrigerant gas 364 is further expanded in a turbo expander 366 to a pressure of approximately 210 psia. The pressure of the separated refrigerant liquid 368 from the first refrigerant separator 362 is let down by second pressure control valve 370 to the same pressure (approximately 210 psia). The separated refrigerant liquid 368 is then mixed together with the first gas output 372 from the turbo expander 366, and then enters the first heat exchanger 312 to provide additional refrigeration. The warmed refrigerant stream 374 exits the first heat exchanger 312 at about 70° F. and passes to a second refrigerant separator 376.
Second gas output 378 from the second refrigerant separator 376 is fed to a turbo compressor 380 associated with the turbo expander 366. The partially compressed gas 382 from the turbo compressor 380 is cooled in a fourth heat exchanger 384, then is fed to a third refrigerant separator 386. Third gas output 388 from the third refrigerant separator 386 returns to the refrigerant compressor 358 to complete the closed loop refrigerant system. Liquid remaining in the second and third refrigerant separators 376, 386 may be removed as needed via first and second control valves 390, 392, respectively. If continuous condensation is observed, pumps (not shown) may be added to the system to relieve this condition.
Those of skill in the art will understand that the above descriptions and operating parameters are provided by way of example only, and do not limit the scope of the invention as described in the following claims.
