Patent Application
20190170435
Ethylene, ethane, propylene, propane, and/or heavier hydrocarbons can be recovered from a variety of gases, such as natural gas, refinery gas, and synthetic gas streams obtained from other hydrocarbon materials such as coal, crude oil, naphtha, oil shale, tar sands, and lignite. Natural gas usually has a major proportion of methane and ethane, i.e., methane and ethane together comprise at least 50 mole percent of the gas. The gas also contains relatively lesser amounts of heavier hydrocarbons such as propane, butanes, pentanes, and the like, as well as hydrogen, nitrogen, carbon dioxide, and other gases.
The present invention is generally concerned with the recovery of ethylene, ethane, propylene, propane and heavier hydrocarbons from such gas streams. A typical analysis of a gas stream to be processed in accordance with this invention would be, in approximate mole percent, 69.4% methane, 11.8% ethane and other C2 components, 5.6% propane and other C3 components, 0.9% iso-butane, 1.8% normal butane, and 1.1% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present.
The historically cyclic fluctuations in the prices of both natural gas and its natural gas liquid (NGL) constituents have at times reduced the incremental value of ethane, ethylene, propane, propylene, and heavier components as liquid products. This has resulted in a demand for processes that can provide more efficient recoveries of these products, for processes that can provide efficient recoveries with lower capital investment, and for processes that can be easily adapted or adjusted to vary the recovery of a specific component over a broad range. Available processes for separating these materials include those based upon cooling and refrigeration of gas, oil absorption, and refrigerated oil absorption. Additionally, cryogenic processes have become popular because of the availability of economical equipment that produces power while simultaneously expanding and extracting heat from the gas being processed. Depending upon the pressure of the gas source, the richness (ethane, ethylene, and heavier hydrocarbons content) of the gas, and the desired end products, each of these processes or a combination thereof may be employed.
The cryogenic expansion process is now generally preferred for natural gas liquids recovery because it provides maximum simplicity with ease of startup, operating flexibility, good efficiency, safety, and good reliability. U.S. Pat. Nos. 3,292,380; 4,061,481; 4,140,504; 4,157,904; 4,171,964; 4,185,978; 4,251,249; 4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063; 4,690,702; 4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,566,554; 5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664; 6,182,469; 6,578,379; 6,712,880; 6,915,662; 7,191,617; 7,219,513; 8,590,340; 8,881,549; 8,919,148; 9,021,831; 9,021,832; 9,052,136; 9,052,137; 9,057,558; 9,068,774; 9,074,814; 9,080,810; and 9,080,811; reissue U.S. Pat. No. 33,408; and co-pending application Ser. Nos. 11/839,693; 12/750,862; 12/772,472; 12/781,259; 12/868,993; 12/869,007; 12/869,139; 14/462,056; and 14/462,083 describe relevant processes (although the description of the present invention in some cases is based on different processing conditions than those described in the listed U.S. Patents and applications).
In a typical cryogenic expansion recovery process, a feed gas stream under pressure is cooled by heat exchange with other streams of the process and/or external sources of refrigeration such as a propane compression-refrigeration system. As the gas is cooled, liquids may be condensed and collected in one or more separators as high-pressure liquids containing some of the desired C2+ components. Depending on the richness of the gas and the amount of liquids formed, the high-pressure liquids may be expanded to a lower pressure and fractionated. The vaporization occurring during expansion of the liquids results in further cooling of the stream. Under some conditions, pre-cooling the high pressure liquids prior to the expansion may be desirable in order to further lower the temperature resulting from the expansion. The expanded stream, comprising a mixture of liquid and vapor, is fractionated in a distillation (demethanizer or deethanizer) column. In the column, the expansion cooled stream(s) is (are) distilled to separate residual methane, nitrogen, and other volatile gases as overhead vapor from the desired C2 components, C3 components, and heavier hydrocarbon components as bottom liquid product, or to separate residual methane, C2 components, nitrogen, and other volatile gases as overhead vapor from the desired C3 components and heavier hydrocarbon components as bottom liquid product.
If the feed gas is not totally condensed (typically it is not), the vapor remaining from the partial condensation can be split into two streams. One portion of the vapor is passed through a work expansion machine or engine, or an expansion valve, to a lower pressure at which additional liquids are condensed as a result of further cooling of the stream. The pressure after expansion is essentially the same as the pressure at which the distillation column is operated. The combined vapor-liquid phases resulting from the expansion are supplied as feed to the column.
The remaining portion of the vapor is cooled to substantial condensation by heat exchange with other process streams, e.g., the cold fractionation tower overhead. Some or all of the high-pressure liquid may be combined with this vapor portion prior to cooling. The resulting cooled stream is then expanded through an appropriate expansion device, such as an expansion valve, to the pressure at which the demethanizer is operated. During expansion, a portion of the liquid will vaporize, resulting in cooling of the total stream. The flash expanded stream is then supplied as top feed to the demethanizer. Typically, the vapor portion of the flash expanded stream and the demethanizer overhead vapor combine in an upper separator section in the fractionation tower as residual methane product gas. Alternatively, the cooled and expanded stream may be supplied to a separator to provide vapor and liquid streams. The vapor is combined with the tower overhead and the liquid is supplied to the column as a top column feed.
In the ideal operation of such a separation process, the residue gas leaving the process will contain substantially all of the methane in the feed gas with essentially none of the heavier hydrocarbon components, and the bottoms fraction leaving the demethanizer will contain substantially all of the heavier hydrocarbon components with essentially no methane or more volatile components. In practice, however, this ideal situation is not obtained because the conventional demethanizer is operated largely as a stripping column. The methane product of the process, therefore, typically comprises vapors leaving the top fractionation stage of the column, together with vapors not subjected to any rectification step. Considerable losses of C2, C3, and C4+ components occur because the top liquid feed contains substantial quantities of these components and heavier hydrocarbon components, resulting in corresponding equilibrium quantities of C2 components, C3 components, C4 components, and heavier hydrocarbon components in the vapors leaving the top fractionation stage of the demethanizer. The loss of these desirable components could be significantly reduced if the rising vapors could be brought into contact with a significant quantity of liquid (reflux) capable of absorbing the C2 components, C3 components, C4 components, and heavier hydrocarbon components from the vapors.
In recent years, the preferred processes for hydrocarbon separation use an upper absorber section to provide additional rectification of the rising vapors. The source of the reflux stream for the upper rectification section is typically a recycled stream of residue gas supplied under pressure. The recycled residue gas stream is usually cooled to substantial condensation by heat exchange with other process streams, e.g., the cold fractionation tower overhead. The resulting substantially condensed stream is then expanded through an appropriate expansion device, such as an expansion valve, to the pressure at which the demethanizer is operated. During expansion, a portion of the liquid will usually vaporize, resulting in cooling of the total stream. The flash expanded stream is then supplied as top feed to the demethanizer. Typically, the vapor portion of the expanded stream and the demethanizer overhead vapor combine in an upper separator section in the fractionation tower as residual methane product gas. Alternatively, the cooled and expanded stream may be supplied to a separator to provide vapor and liquid streams, so that thereafter the vapor is combined with the tower overhead and the liquid is supplied to the column as a top column feed. Typical process schemes of this type are disclosed in U.S. Pat. Nos. 4,889,545; 5,568,737; 5,881,569; and 9,052,137, assignee's co-pending application Ser. No. 12/717,394, and in Mowrey, E. Ross, “Efficient, High Recovery of Liquids from Natural Gas Utilizing a High Pressure Absorber”, Proceedings of the Eighty-First Annual Convention of the Gas Processors Association, Dallas, Texas, Mar. 11-13, 2002. Unfortunately, these processes require the use of a large amount of compression power to provide the motive force for recycling the reflux stream to the demethanizer, adding to both the capital cost and the operating cost of facilities using these processes.
The present invention also employs an upper rectification section (or a separate rectification column in some embodiments). However, the reflux for this rectification section is provided by cooling two streams derived from the feed gas to substantial condensation and then expanding both streams to the operating pressure of the fractionation tower. During expansion, a portion of each stream is vaporized, resulting in cooling of each total stream. One cooled, expanded stream is then supplied to the fractionation tower at a top column feed point and the other cooled, expanded stream is supplied to the tower at an upper mid-column feed point. The condensed liquid in the top column feed, which is predominantly liquid methane, can then be used to absorb C2 components, C3 components, C4 components, and heavier hydrocarbon components from the vapors rising through the upper rectification section and thereby capture these valuable components in the bottom liquid product from the demethanizer.
In accordance with the present invention, it has been found that C2 recovery in excess of 93% and C3 and C4+ recoveries in excess of 99% can be obtained. In addition, the present invention makes possible essentially 100% separation of methane and lighter components from the C2 components and heavier components at higher recovery levels compared to the prior art while maintaining the same energy requirements. The present invention, although applicable at lower pressures and warmer temperatures, is particularly advantageous when processing feed gases in the range of 400 to 1500 psia [2,758 to 10,342 kPa(a)] or higher under conditions requiring NGL recovery column overhead temperatures of −50° F. [−46° C.] or colder.
For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings:
In the following explanation of the above figures, tables are provided summarizing flow rates calculated for representative process conditions. In the tables appearing herein, the values for flow rates (in moles per hour) have been rounded to the nearest whole number for convenience. The total stream rates shown in the tables include all non-hydrocarbon components and hence are generally larger than the sum of the stream flow rates for the hydrocarbon components. Temperatures indicated are approximate values rounded to the nearest degree. It should also be noted that the process design calculations performed for the purpose of comparing the processes depicted in the figures are based on the assumption of no heat leak from (or to) the surroundings to (or from) the process. The quality of commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art.
For convenience, process parameters are reported in both the traditional British units and in the units of the Système International d′Unités (SI). The molar flow rates given in the tables may be interpreted as either pound moles per hour or kilogram moles per hour. The energy consumptions reported as horsepower (HP) and/or thousand British Thermal Units per hour (MBTU/Hr) correspond to the stated molar flow rates in pound moles per hour. The energy consumptions reported as kilowatts (kW) correspond to the stated molar flow rates in kilogram moles per hour.
The feed stream 31 is divided into two portions, streams 32 and 33. Stream 32 is cooled to substantial condensation in heat exchanger 13 by heat exchange with cold residue gas (stream 41), flashed separator liquids (stream 35a), and propane refrigerant. The resulting substantially condensed stream 32a at −132° F. [−91° C.] is then flash expanded through expansion valve 17 to the operating pressure (approximately 316 psia [2,181 kPa(a)]) of fractionation tower 19. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
The remaining portion of feed stream 31, stream 33, is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 41a) and propane refrigerant. Note that in all cases exchangers 10 and 13 are representative of either a multitude of individual heat exchangers or a single multi-pass heat exchanger, or any combination thereof (The decision as to whether to use more than one heat exchanger for the indicated cooling services will depend on a number of factors including, but not limited to, inlet gas flow rate, heat exchanger size, stream temperatures, etc.) The cooled stream 33a enters separator 12 at −30° F. [−35° C.] and 870 psia [6,002 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35). The separator liquid (stream 35) is expanded to slightly above the operating pressure of fractionation tower 19 by expansion valve 18. Expanded stream 35a is heated from −63° F. [−53° C.] to 23° F. [−5° C.] in heat exchanger 13 as described earlier before heated stream 35b is supplied to fractionation tower 19 at a lower mid-column feed point.
The vapor (stream 34) from separator 12 enters a work expansion machine 15 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 15 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −99° F. [−73° C.]. The typical commercially available expanders are capable of recovering on the order of 80-85% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item 16) that can be used to re-compress the residue gas (stream 41b), for example. The partially condensed expanded stream 34a is thereafter supplied as feed to fractionation tower 19 at an upper mid-column feed point.
The demethanizer in tower 19 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. As is often the case in natural gas processing plants, the fractionation tower may consist of three sections. The upper section 19a is a separator wherein the partially vaporized top feed is separated into its respective vapor and liquid portions, and wherein the vapor rising from the rectifying section 19b below is combined with the vapor portion of the top feed to form the cold demethanizer overhead vapor (stream 41) which exits the top of the tower at −137° F. [−94° C.]. The middle absorbing (rectifying) section 19b contains the trays and/or packing to provide the necessary contact between the vapor portions of the expanded streams 34a and 35b rising upward and cold liquid falling downward. The lower stripping (demethanizing) section 19c contains additional trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. Demethanizing section 19c also includes one or more reboilers (such as the reboiler 21 and side reboiler 20 shown in
The liquid product (stream 44) exits the bottom of tower 19 at 79° F. [26° C.], based on a typical specification of a methane concentration of 0.07% on a volume basis in the bottom product. The cold residue gas stream 41 passes countercurrently to one portion of the feed gas in heat exchanger 13 where it is heated to −31° F. [−35° C.] (stream 41a), and countercurrently to the other portion of the feed gas in heat exchanger 10 where it is heated to 88° F. [31° C.] (stream 41b). The residue gas is then re-compressed in two stages. The first stage is compressor 16 driven by expansion machine 15. The second stage is compressor 22 driven by a supplemental power source which compresses the residue gas (stream 41d) to sales line pressure. After cooling to 110° F. [43° C.] in discharge cooler 23, the residue gas product (stream 41e) flows to the sales gas pipeline at 918 psia [6,330 kPa(a)], sufficient to meet line requirements (usually on the order of the inlet pressure).
A summary of stream flow rates and energy consumption for the process illustrated in
After feed stream 31 is divided into two portions, streams 32 and 33, stream 32 is cooled to substantial condensation in heat exchanger 13 by heat exchange with cold residue gas (stream 41), flashed separator liquids (stream 35a), and propane refrigerant. The resulting substantially condensed stream 32a at −140° F. [−96° C.] is then flash expanded through expansion valve 17 to the operating pressure (approximately 340 psia [2,346 kPa(a)]) of fractionation tower 19. During expansion a portion of the stream is vaporized, resulting in cooling of expanded stream 32b to −159° F. [−106° C.] before it is supplied to fractionation tower 19 at an upper mid-column feed point.
The remaining portion of feed stream 31, stream 33, is cooled in heat exchanger 10 by heat exchange with a portion (stream 47) of the cool demethanizer overhead vapor (stream 45a) and propane refrigerant, and cooled stream 33a enters separator 12 at −25° F. [−32° C.] and 870 psia [6,002 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35). The separator liquid (stream 35) is expanded to slightly above the operating pressure of fractionation tower 19 by expansion valve 18. Expanded stream 35a is heated from −54° F. [−48° C.] to 23° F. [−5° C.] in heat exchanger 13 as described earlier before heated stream 35b is supplied to fractionation tower 19 at a lower mid-column feed point.
The vapor (stream 34) from separator 12 enters a work expansion machine 15 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 15 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −89° F. [−67° C.]. The partially condensed expanded stream 34a is thereafter supplied as feed to fractionation tower 19 at an intermediate mid-column feed point.
The recompressed and cooled distillation vapor stream 45e is divided into two streams. One portion, stream 41, is the volatile residue gas product. The other portion, recycle stream 48, flows to heat exchanger 11 where it is cooled to 0° F. [−18° C.] by heat exchange with the remaining portion (stream 46) of cool demethanizer overhead vapor stream 45a. The cooled recycle stream 48a then flows to exchanger 13 where it is further cooled to −140° F. [−96° C.] and substantially condensed by heat exchange with cold distillation vapor stream 45 and propane refrigeration. The substantially condensed stream 48b is then expanded through an appropriate expansion device, such as expansion valve 14, to the demethanizer operating pressure, resulting in cooling of the total stream to −167° F. [−111° C.]. The expanded stream 48c is then supplied to fractionation tower 19 as the top column feed. The vapor portion of stream 48c combines with the vapors rising from the top fractionation stage of the column to form distillation vapor stream 45, which is withdrawn from an upper region of the tower.
The liquid product (stream 44) exits the bottom of tower 19 at 88° F. [31° C.], based on a methane concentration of 0.07% on a volume basis in the bottom product. The demethanizer overhead vapor stream 45 passes countercurrently to one portion of the feed gas (stream 32) and the partially cooled recycle stream (stream 48a) in heat exchanger 13 where it is heated to −27° F. [−33° C.] (stream 45a) and then divided into two portions, stream 46 and stream 47. Stream 46 passes countercurrently to recycle stream 48 in heat exchanger 11 and is heated to 105° F. [41° C.] (stream 46a), while stream 47 passes countercurrently to the other portion of the feed gas in heat exchanger 10 where it is heated to 91° F. [33° C.] (stream 47a). Streams 46a and 47a recombine as stream 45b at 92° F. [33° C.], which is then re-compressed in two stages, compressor 16 driven by expansion machine 15 and compressor 22 driven by a supplemental power source. After cooling to 110° F. [43° C.] in discharge cooler 23, stream 45e is split into the residue gas product (stream 41) and the recycle stream 48 as described earlier. Residue gas stream 41 then flows to the sales gas pipeline 918 psia [6,330 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables I and II shows that, compared to the
After feed stream 31 is divided into two portions, streams 32 and 33, stream 32 is cooled to substantial condensation in heat exchanger 13 by heat exchange with cold residue gas (stream 41), flashed separator liquids (stream 35a), and propane refrigerant. The resulting substantially condensed stream 32a at −120° F. [−84° C.] is then flash expanded through expansion valve 17 to the operating pressure (approximately 324 psia [2,235 kPa(a)]) of fractionation tower 19. During expansion a portion of the stream is vaporized, resulting in cooling of expanded stream 32b to −153° F. [−103° C.] before it is supplied to fractionation tower 19 at an upper mid-column feed point.
The remaining portion of feed stream 31, stream 33, is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 41a) and propane refrigerant, and cooled stream 33a enters separator 12 at −34° F. [−36° C.] and 870 psia [6,002 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35). The separator liquid (stream 35) is expanded to slightly above the operating pressure of fractionation tower 19 by expansion valve 18. Expanded stream 35a is heated from −66° F. [−54° C.] to 21° F. [−6° C.] in heat exchanger 13 as described earlier before heated stream 35b is supplied to fractionation tower 19 at a lower mid-column feed point.
The vapor (stream 34) from separator 12 enters a work expansion machine 15 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 15 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −100° F. [−74° C.]. The partially condensed expanded stream 34a is thereafter supplied as feed to fractionation tower 19 at an intermediate mid-column feed point.
A portion of the distillation vapor (stream 49) is withdrawn from an upper region of the demethanizing section in fractionation column 19, below the feed position of expanded stream 34a. Distillation vapor stream 49 is then cooled from −100° F. [−73° C.] to −146° F. [−99° C.] and partially condensed (stream 49a) in heat exchanger 24 by heat exchange with the cold demethanizer overhead stream 48 exiting the top of demethanizer 19 at −150° F. [−101° C.]. The cold demethanizer overhead stream is warmed to −118° F. [−84° C.] (stream 48a) as it cools and condenses a portion of stream 49.
The operating pressure in reflux separator 25 is maintained slightly below the operating pressure of demethanizer 19. This provides the driving force which causes distillation vapor stream 49 to flow through heat exchanger 24 and thence into the reflux separator 25 where the condensed liquid (stream 51) is separated from the uncondensed vapor (stream 50). Stream 50 then combines with the warmed demethanizer overhead stream 48a from heat exchanger 24 to form cold residue gas stream 41 at −123° F. [−86° C.].
The liquid stream 51 from reflux separator 25 is pumped by pump 26 to a pressure slightly above the operating pressure of demethanizer 19, and stream 51a is then supplied as cold top column feed (reflux) to demethanizer 19 at −145° F. [−98° C.]. This cold liquid reflux absorbs and condenses the C2 components, C3 components, and heavier components rising in the upper region of the rectifying section of demethanizer 19.
The liquid product (stream 44) exits the bottom of tower 19 at 81° F. [27° C.], based on a methane concentration of 0.07% on a volume basis in the bottom product. The cold residue gas stream 41 passes countercurrently to one portion of the feed gas in heat exchanger 13 where it is heated to −35° F. [−37° C.] (stream 41a), and countercurrently to the other portion of the feed gas in heat exchanger 10 where it is heated to 88° F. [31° C.] (stream 41b). The residue gas is then re-compressed in two stages, compressor 16 driven by expansion machine 15 and compressor 22 driven by a supplemental power source. After cooling to 110° F. [43° C.] in discharge cooler 23, the residue gas product (stream 41e) flows to the sales gas pipeline at 918 psia [6,330 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables I, II, and III shows that the
After feed stream 31 is divided into two portions, streams 32 and 33, stream 32 is cooled to substantial condensation in heat exchanger 13 by heat exchange with cold residue gas (stream 41), flashed separator liquids (stream 35a), and propane refrigerant. The resulting substantially condensed stream 32a at −151° F. [−101° C.] is then flash expanded through expansion valve 17 to the operating pressure (approximately 319 psia [2,202 kPa(a)]) of fractionation tower 19. During expansion a portion of the stream is vaporized, resulting in cooling of expanded stream 32b to −165° F. [−109° C.] before it is supplied to fractionation tower 19 at an upper mid-column feed point.
The remaining portion of feed stream 31, stream 33, is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 41a) and propane refrigerant, and cooled stream 33a enters separator 12 at −40° F. [−40° C.] and 870 psia [6,002 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35). The separator liquid (stream 35) is expanded to slightly above the operating pressure of fractionation tower 19 by expansion valve 18. Expanded stream 35a is heated from −73° F. [−59° C.] to 4° F. [−16° C.] in heat exchanger 13 as described earlier before heated stream 35b is supplied to fractionation tower 19 at a lower mid-column feed point.
The vapor (stream 34) from separator 12 enters a work expansion machine 15 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 15 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −107° F. [−77° C.]. The partially condensed expanded stream 34a is thereafter supplied as feed to fractionation tower 19 at an intermediate mid-column feed point.
A portion of the distillation vapor (stream 49) is withdrawn from an upper region of the demethanizing section in fractionation column 19 below expanded stream 34a at −102° F. [−75° C.] and is compressed to approximately 486 psia [3,353 kPa(a)] by vapor compressor 27. The compressed stream 49a is then cooled from −52° F. [−47° C.] to −151° F. [−101° C.] and substantially condensed (stream 49b) in heat exchanger 13 as described earlier.
Since substantially condensed stream 49b is at a pressure greater than the operating pressure of demethanizer 19, it is flash expanded through expansion valve 14 to the operating pressure of fractionation tower 19. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream to −159° F. [−106° C.]. The expanded stream 49c is then supplied as cold top column feed (reflux) to demethanizer 19. The vapor portion of stream 49c combines with the distillation vapor rising from the upper fractionation stage to form residue gas stream 41 exiting the top of demethanizer 19 at −154° F. [−103° C.], while the cold liquid reflux portion absorbs and condenses the C2 components, C3 components, and heavier components rising in the upper region of the rectifying section of demethanizer 19.
The liquid product (stream 44) exits the bottom of tower 19 at 79° F. [26° C.], based on a methane concentration of 0.07% on a volume basis in the bottom product. The cold residue gas stream 41 passes countercurrently to one portion of the feed gas in heat exchanger 13 where it is heated to −56° F. [−49° C.] (stream 41a), and countercurrently to the other portion of the feed gas in heat exchanger 10 where it is heated to 92° F. [33° C.] (stream 41b). The residue gas is then re-compressed in two stages, compressor 16 driven by expansion machine 15 and compressor 22 driven by a supplemental power source. After cooling to 110° F. [43° C.] in discharge cooler 23, the residue gas product (stream 41e) flows to the sales gas pipeline at 918 psia [6,330 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables I, II, III, and IV shows that the
In the simulation of the
The remaining portion of feed stream 31, stream 33, is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 41a) and propane refrigerant, and cooled stream 33a enters separator 12 at −36° F. [−38° C.] and 870 psia [6,002 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35). The separator liquid (stream 35) is expanded to slightly above the operating pressure of fractionation tower 19 by expansion valve 18. Expanded stream 35a is heated from −68° F. [−56° C.] to 10° F. [−12° C.] in heat exchanger 13 as described earlier before heated stream 35b is supplied to fractionation tower 19 at a lower mid-column feed point.
The vapor (stream 34) from separator 12 is divided into two streams, 36 and 39. Stream 36, containing about 16% of the total vapor, is cooled to −148° F. [−100° C.] and substantially condensed (stream 36a) in heat exchanger 13 as described earlier. Substantially condensed stream 36a is flash expanded through expansion valve 14 to the operating pressure of fractionation tower 19. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
The remaining 84% of the vapor from separator 12 (stream 39) enters a work expansion machine 15 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 15 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 39a to a temperature of approximately −103° F. [−75° C.]. The partially condensed expanded stream 39a is thereafter supplied as feed to fractionation tower 19 at an intermediate mid-column feed point.
The demethanizer in tower 19 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The demethanizer tower consists of three sections. The upper section 19a is a separator wherein the partially vaporized top feed (stream 36b) is separated into its respective vapor and liquid portions, and wherein the vapor rising from the rectifying section 19b below is combined with the vapor portion of the top feed to form the cold demethanizer overhead vapor (stream 41) which exits the top of the tower at −151° F. [−102° C.]. The middle absorbing (rectifying) section 19b contains the trays and/or packing to provide the necessary contact between the vapor portions of the expanded streams 32b, 39a, and 35b rising upward and cold liquid falling downward. The lower stripping (demethanizing) section 19c contains additional trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. Demethanizing section 19c also includes one or more reboilers (such as the reboiler 21 and side reboiler 20 shown in
In demethanizing section 19c of demethanizer 19, the feed streams are stripped of their methane and lighter components. The resulting liquid product (stream 44) exits the bottom of tower 19 at 80° F. [27° C.] based on a methane concentration of 0.07% on a volume basis in the bottom product. The cold residue gas stream 41 leaves demethanizer 19 and passes countercurrently to one portion of the feed gas and a portion of the separator vapor in heat exchanger 13 where it is heated to −44° F. [−42° C.] (stream 41a) and countercurrently to the other portion of the feed gas in heat exchanger 10 where it is heated to 91° F. [33° C.] (stream 41b) as it provides cooling as previously described. The residue gas is then re-compressed in two stages, compressor 16 driven by expansion machine 15 and compressor 22 driven by a supplemental power source. After cooling to 110° F. [43° C.] in discharge cooler 23, the residue gas product (stream 41e) flows to the sales gas pipeline at 918 psia [6,330 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables I through V shows that, compared to the prior art, the present invention improves the ethane recovery from 87.06% (for
The improvement in the recovery efficiency of the present invention over that of the prior art processes can be understood by examining the improvement in the rectification that the present invention provides for rectifying section 19b. Compared to the prior art of the
Compared to the prior art of the
Compared to the prior art of the
Compared to the prior art of the
However, the present invention is able to match the recovery of the
In accordance with this invention, it is generally advantageous to design the absorbing (rectifying) section of the demethanizer to contain multiple theoretical separation stages. However, the benefits of the present invention can be achieved with as few as two theoretical stages. For instance, all or a part of expanded substantially condensed stream 36b leaving expansion valve 14 and all or a part of expanded substantially condensed stream 32b (
In accordance with the present invention, the splitting of the feed gas may be accomplished in several ways. In the processes of
The high pressure liquid (stream 35 in
As described earlier, a portion of the feed gas (stream 32) and a portion of the separator vapor (stream 36) are substantially condensed and the resulting condensate used to absorb valuable C2 components, C3 components, and heavier components from the vapors rising through rectifying section 19b of demethanizer 19 (
Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine 15, or replacement with an alternate expansion device (such as an expansion valve), is feasible. Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where appropriate. For example, conditions may warrant work expansion of the substantially condensed portion of the separator vapor (stream 36a in
In accordance with the present invention, the use of external refrigeration to supplement the cooling available to the inlet gas and/or separator vapor from other process streams may be employed, particularly in the case of a rich inlet gas. The use and distribution of separator liquids and demethanizer side draw liquids for process heat exchange, and the particular arrangement of heat exchangers for inlet gas and separator vapor cooling must be evaluated for each particular application, as well as the choice of process streams for specific heat exchange services.
It will also be recognized that the relative amount of feed found in each branch of the split vapor feeds will depend on several factors, including gas pressure, feed gas composition, the amount of heat which can economically be extracted from the feed, and the quantity of horsepower available. More feed to the top of the column may increase recovery while decreasing power recovered from the expander thereby increasing the recompression horsepower requirements. Increasing feed lower in the column reduces the horsepower consumption but may also reduce product recovery. The relative locations of the mid-column feeds may vary depending on inlet composition or other factors such as desired recovery levels and amount of liquid formed during inlet gas cooling. Moreover, two or more of the feed streams, or portions thereof, may be combined depending on the relative temperatures and quantities of individual streams, and the combined stream then fed to a mid-column feed position.
The present invention allows reduced capital expenditures and/or provides improved recovery of C2 components, C3 components, and heavier hydrocarbon components or of C3 components and heavier hydrocarbon components per amount of utility consumption required to operate the process. An improvement in utility consumption required for operating the demethanizer or deethanizer process may appear in the form of reduced power requirements for compression or re-compression, reduced power requirements for external refrigeration, reduced energy requirements for tower reboilers, or a combination thereof.
While there have been described what are believed to be preferred embodiments of the invention, those skilled in the art will recognize that other and further modifications may be made thereto, e.g. to adapt the invention to various conditions, types of feed, or other requirements without departing from the spirit of the present invention as defined by the following claims.
This invention relates to a process and an apparatus for the separation of a gas containing hydrocarbons. This application is a continuation of U.S. patent application Ser. No. 14/828,093 filed Aug. 17, 2015 and claims the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application No. 62/045,908 which was filed on Sep. 4, 2014.
| Number | Date | Country | |
|---|---|---|---|
| 62045908 | Sep 2014 | US |
| Number | Date | Country | |
|---|---|---|---|
| Parent | 14828093 | Aug 2015 | US |
| Child | 16269098 | US |
