US20190170435A1

US20190170435A1 - Hydrocarbon Gas Processing

Info

Publication number: US20190170435A1
Application number: US16/269,098
Authority: US
Inventors: J. Ascencion Anguiano; John D. Wilkinson; Joe T. Lynch; Hank M. Hudson
Original assignee: Ortloff Engineers Ltd
Current assignee: Honeywell UOP LLC
Priority date: 2014-09-04
Filing date: 2019-02-06
Publication date: 2019-06-06
Also published as: US20160069610A1; MX2015011073A

Abstract

A process and an apparatus are disclosed for the recovery of ethane, ethylene, propane, propylene, and heavier hydrocarbon components from a hydrocarbon gas stream. The stream is divided into first and second streams. The first stream is cooled to condense substantially all of it, expanded to lower pressure, and supplied to a fractionation tower at an upper mid-column feed position. The second stream is cooled sufficiently to partially condense it and separated into vapor and liquid streams. The vapor stream is divided into first and second portions. The first portion is cooled to condense substantially all of it, expanded to the tower pressure, and supplied to the tower at the top feed position. The second portion is expanded to the tower pressure and supplied to the fractionation tower at an intermediate mid-column feed position. The liquid stream is expanded to the tower pressure and supplied to the column at a lower mid-column feed position. The quantities and temperatures of the feeds to the fractionation tower are effective to maintain the overhead temperature of the fractionation tower at a temperature whereby the major portion of the desired components is recovered.

Description

BACKGROUND OF THE INVENTION

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.
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 C₂components, 5.6% propane and other C₃components, 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 C₂components, C₃components, and heavier hydrocarbon components as bottom liquid product, or to separate residual methane, C₂components, nitrogen, and other volatile gases as overhead vapor from the desired C₃components 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 C₂components, C₃components, C₄components, 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 C₂components, C₃components, C₄components, 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 C₂components, C₃components, C₄components, 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 C₂recovery in excess of 93% and C₃and 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 C₂components 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:

FIG. 1 is a flow diagram of a prior art natural gas processing plant in accordance with U.S. Pat. No. 4,278,457;

FIG. 2 is a flow diagram of a prior art natural gas processing plant in accordance with U.S. Pat. No. 5,568,737;

FIG. 3 is a flow diagram of a prior art natural gas processing plant in accordance with U.S. Pat. No. 7,191,617;

FIG. 4 is a flow diagram of a prior art natural gas processing plant in accordance with assignee's co-pending application Ser. No. 11/839,693;

FIG. 5 is a flow diagram of a natural gas processing plant in accordance with the present invention; and

FIGS. 6 through 9 are flow diagrams illustrating alternative means of application of the present invention to a natural gas stream.

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.

DESCRIPTION OF THE PRIOR ART

FIG. 1 is a process flow diagram showing the design of a processing plant to recover C₂+ components from natural gas using prior art according to U.S. Pat. No. 4,278,457. In this simulation of the process, inlet gas enters the plant at 104° F. [40° C.] and 896 psia [6,181 kPa(a)] as stream 31. If the inlet gas contains a concentration of sulfur compounds which would prevent the product streams from meeting specifications, the sulfur compounds are removed by appropriate pretreatment of the feed gas (not illustrated). In addition, the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccant has typically been used for this purpose.
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 35 a), and propane refrigerant. The resulting substantially condensed stream 32 a 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 FIG. 1, the expanded stream 32 b leaving expansion valve 17 reaches a temperature of −159° F. [−106° C.] before it is supplied to fractionation tower 19 as the top column feed.
The remaining portion of feed stream 31, stream 33, is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 41 a) 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 33 a 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 35 a is heated from −63° F. [−53° C.] to 23° F. [−5° C.] in heat exchanger 13 as described earlier before heated stream 35 b 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 34 a 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 41 b), for example. The partially condensed expanded stream 34 a 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 19 a 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 19 b 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 19 b contains the trays and/or packing to provide the necessary contact between the vapor portions of the expanded streams 34 a and 35 b rising upward and cold liquid falling downward. The lower stripping (demethanizing) section 19 c contains additional trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. Demethanizing section 19 c also includes one or more reboilers (such as the reboiler 21 and side reboiler 20 shown in FIG. 1) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 44, of methane and lighter components. Stream 34 a enters demethanizer 19 at an intermediate feed position located below rectifying section 19 b and above demethanizing section 19 c. The liquid portion of the expanded stream 34 a commingles with liquids falling downward from rectifying section 19 b and the combined liquid continues downward into demethanizing section 19 c of column 19. The vapor portion of the expanded stream 34 a commingles with vapors rising upward from demethanizing section 19 c and the combined vapor rises upward through rectifying section 19 b and is contacted with cold liquid falling downward to condense and absorb the C₂components, C₃components, and heavier components from the vapor.
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 41 a), and countercurrently to the other portion of the feed gas in heat exchanger 10 where it is heated to 88° F. [31° C.] (stream 41 b). 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 41 d) to sales line pressure. After cooling to 110° F. [43° C.] in discharge cooler 23, the residue gas product (stream 41 e) 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 FIG. 1 is set forth in the following table:

TABLE I

(FIG. 1)
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]

Stream	Methane	Ethane	Propane	Butanes+	Total

31	22,875	3,898	1,843	1,240	32,972
32	6,428	1,095	518	348	9,265
33	16,447	2,803	1,325	892	23,707
34	14,127	1,566	382	94	18,283
35	2,320	1,237	943	798	5,424
41	22,867	504	20	2	26,507
44	8	3,394	1,823	1,238	6,465

Recoveries*

	Ethane	87.06%
	Propane	98.92%
	Butanes+	99.88%

Power

Residue Gas Compression	16,044	HP	[26,376 kW]
Refrigerant Compression	7,492	HP	[12,317 kW]
Total Compression	23,536	HP	[38,693 kW]

*(Based on un-rounded flow rates)

FIG. 2 represents an alternative prior art process according to U.S. Pat. No. 5,568,737. The process of FIG. 2 has been applied to the same feed gas composition and conditions as described above for FIG. 1. In the simulation of this process, as in the simulation for the process of FIG. 1, operating conditions were selected to maximize the recovery level for a given energy consumption.
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 35 a), and propane refrigerant. The resulting substantially condensed stream 32 a 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 32 b 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 45 a) and propane refrigerant, and cooled stream 33 a 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 35 a is heated from −54° F. [−48° C.] to 23° F. [−5° C.] in heat exchanger 13 as described earlier before heated stream 35 b 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 34 a to a temperature of approximately −89° F. [−67° C.]. The partially condensed expanded stream 34 a is thereafter supplied as feed to fractionation tower 19 at an intermediate mid-column feed point.
The recompressed and cooled distillation vapor stream 45 e 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 45 a. The cooled recycle stream 48 a 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 48 b 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 48 c is then supplied to fractionation tower 19 as the top column feed. The vapor portion of stream 48 c 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 48 a) in heat exchanger 13 where it is heated to −27° F. [−33° C.] (stream 45 a) 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 46 a), 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 47 a). Streams 46 a and 47a recombine as stream 45 b 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 45 e 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 FIG. 2 is set forth in the following table:

TABLE II

(FIG. 2)
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]

Stream	Methane	Ethane	Propane	Butanes+	Total

31	22,875	3,898	1,843	1,240	32,972
32	5,513	939	444	299	7,946
33	17,362	2,959	1,399	941	25,026
34	15,187	1,761	450	114	19,761
35	2,175	1,198	949	827	5,265
48	3,035	69	0	0	3,518
45	25,902	592	0	0	30,022
41	22,867	523	0	0	26,504
44	8	3,375	1,843	1,240	6,468

Recoveries*

	Ethane	86.57%
	Propane	100.00%
	Butanes+	100.00%

Power

Residue Gas Compression	17,363	HP	[28,544 kW]
Refrigerant Compression	6,214	HP	[10,216 kW]
Total Compression	23,577	HP	[38,760 kW]

*(Based on un-rounded flow rates)

A comparison of Tables I and II shows that, compared to the FIG. 1 process, the FIG. 2 process has slightly lower ethane recovery (86.57% versus 87.06%), but improves propane recovery from 98.92% to 100.00% and butanes+recovery from 99.88% to 100.00%. Comparison of Tables I and II further shows that the improvement in yields for the FIG. 2 process was achieved using essentially the same power requirements.
FIG. 3 represents an alternative prior art process according to U.S. Pat. No. 7,191,617. The process of FIG. 3 has been applied to the same feed gas composition and conditions as described above for FIGS. 1 and 2. In the simulation of this process, as in the simulation for the processes of FIGS. 1 and 2, operating conditions were selected to maximize the recovery level for a given energy consumption.
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 35 a), and propane refrigerant. The resulting substantially condensed stream 32 a 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 32 b 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 41 a) and propane refrigerant, and cooled stream 33 a 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 35 a is heated from −66° F. [−54° C.] to 21° F. [−6° C.] in heat exchanger 13 as described earlier before heated stream 35 b 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 34 a to a temperature of approximately −100° F. [−74° C.]. The partially condensed expanded stream 34 a 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 34 a. Distillation vapor stream 49 is then cooled from −100° F. [−73° C.] to −146° F. [−99° C.] and partially condensed (stream 49 a) 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 48 a) 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 48 a 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 51 a 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 C₂components, C₃components, 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 41 a), and countercurrently to the other portion of the feed gas in heat exchanger 10 where it is heated to 88° F. [31° C.] (stream 41 b). 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 41 e) 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 FIG. 3 is set forth in the following table:

TABLE III

(FIG. 3)
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]

Stream	Methane	Ethane	Propane	Butanes+	Total

31	22,875	3,898	1,843	1,240	32,972
32	6,062	1,033	488	329	8,737
33	16,813	2,865	1,355	911	24,235
34	14,234	1,525	360	86	18,356
35	2,579	1,340	995	825	5,879
49	4,853	364	18	1	5,443
50	3,384	49	0	0	3,622
51	1,469	315	18	1	1,821
48	19,483	288	2	0	22,698
41	22,867	337	2	0	26,320
44	8	3,561	1,841	1,240	6,652

Recoveries*

	Ethane	91.35%
	Propane	99.88%
	Butanes+	100.00%

Power

Residue Gas Compression	15,932	HP	[26,192 kW]
Refrigerant Compression	7,640	HP	[12,560 kW]
Total Compression	23,572	HP	[38,752 kW]

*(Based on un-rounded flow rates)

A comparison of Tables I, II, and III shows that the FIG. 3 process improves the ethane recovery from 87.06% (for FIGS. 1) and 86.57% (for FIGS. 2) to 91.35%. The propane recovery for the FIG. 3 process (99.88%) is significantly higher than that of the FIG. 1 process (98.92%) but slightly lower than that of the FIG. 2 process (100.00%). The butanes+recovery for the FIG. 3 process (100.00%) is slightly higher than that of the FIG. 1 process (99.88%) and the same as that of the FIG. 2 process (100.00%). Comparison of Tables I, II, and III further shows that the improvement in yields for the FIG. 3 process was achieved using essentially the same power requirements.
FIG. 4 represents an alternative prior art process according to co-pending application Ser. No. 11/839,693. The process of FIG. 4 has been applied to the same feed gas composition and conditions as described above for FIGS. 1 through 3. In the simulation of this process, as in the simulation for the process of FIGS. 1 through 3, operating conditions were selected to maximize the recovery level for a given energy consumption.
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 35 a), and propane refrigerant. The resulting substantially condensed stream 32 a 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 32 b 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 41 a) and propane refrigerant, and cooled stream 33 a 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 35 a is heated from −73° F. [−59° C.] to 4° F. [−16° C.] in heat exchanger 13 as described earlier before heated stream 35 b 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 34 a to a temperature of approximately −107° F. [−77° C.]. The partially condensed expanded stream 34 a 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 34 a at −102° F. [−75° C.] and is compressed to approximately 486 psia [3,353 kPa(a)] by vapor compressor 27. The compressed stream 49 a is then cooled from −52° F. [−47° C.] to −151° F. [−101° C.] and substantially condensed (stream 49 b) in heat exchanger 13 as described earlier.
Since substantially condensed stream 49 b 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 49 c is then supplied as cold top column feed (reflux) to demethanizer 19. The vapor portion of stream 49 c 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 C₂components, C₃components, 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 41 a), and countercurrently to the other portion of the feed gas in heat exchanger 10 where it is heated to 92° F. [33° C.] (stream 41 b). 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 41 e) 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 FIG. 4 is set forth in the following table:

TABLE IV

(FIG. 4)
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]

Stream	Methane	Ethane	Propane	Butanes+	Total

31	22,875	3,898	1,843	1,240	32,972
32	4,529	772	365	245	6,529
33	18,346	3,126	1,478	995	26,443
34	15,105	1,522	341	79	19,366
35	3,241	1,604	1,137	916	7,077
49	2,429	179	8	0	2,722
41	22,867	263	1	0	26,244
44	8	3,635	1,842	1,240	6,728

Recoveries*

	Ethane	93.26%
	Propane	99.95%
	Butanes+	100.00%

Power

Residue Gas Compression	15,859	HP	[26,072	kW]
Vapor Compression	351	HP	[577	kW]
Refrigerant Compression	7,366	HP	[12,110	kW]
Total Compression	23,576	HP	[38,759	kW]

*(Based on un-rounded flow rates)

A comparison of Tables I, II, III, and IV shows that the FIG. 4 process improves the ethane recovery from 87.06% (for FIG. 1), 86.57% (for FIGS. 2), and 91.35% (for FIGS. 3) to 93.26%. The propane recovery for the FIG. 4 process (99.95%) is significantly higher than that of the FIG. 1 process (98.92%) and higher than that of the FIG. 3 process (99.88%), but slightly lower than that of the FIG. 2 process (100.00%). The butanes+recovery for the FIG. 4 process (100.00%) is slightly higher than that of the FIG. 1 process (99.88%) and the same as that of the FIG. 2 process and the FIG. 3 process (100.00% for both). Comparison of Tables I, II, III, and IV further shows that the improvement in yields for the FIG. 4 process was achieved using essentially the same power requirements.

DESCRIPTION OF THE INVENTION

FIG. 5 illustrates a flow diagram of a process in accordance with the present invention. The feed gas composition and conditions considered in the process presented in FIG. 5 are the same as those in FIGS. 1 through 4. Accordingly, the FIG. 5 process can be compared with that of the FIGS. 1 through 4 processes to illustrate the advantages of the present invention.
In the simulation of the FIG. 5 process, inlet gas enters the plant at 104° F. [40° C.] and 896 psia [6,181 kPa(a)] as stream 31. 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 35 a), and propane refrigerant. The resulting substantially condensed stream 32 a at −148° F. [−100° 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 32 b to −164° 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 41 a) and propane refrigerant, and cooled stream 33 a 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 35 a is heated from −68° F. [−56° C.] to 10° F. [−12° C.] in heat exchanger 13 as described earlier before heated stream 35 b 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 36 a) in heat exchanger 13 as described earlier. Substantially condensed stream 36 a 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 FIG. 5, the expanded stream 36 b leaving expansion valve 14 reaches a temperature of −169° F. [−112° C.] before it is supplied as cold top column feed (reflux) to demethanizer 19. This cold liquid reflux absorbs and condenses the C₂components, C₃components, and heavier components rising in the upper region of the rectifying section of demethanizer 19.
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 39 a to a temperature of approximately −103° F. [−75° C.]. The partially condensed expanded stream 39 a 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 19 a is a separator wherein the partially vaporized top feed (stream 36 b) is separated into its respective vapor and liquid portions, and wherein the vapor rising from the rectifying section 19 b 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 19 b contains the trays and/or packing to provide the necessary contact between the vapor portions of the expanded streams 32 b, 39 a, and 35 b rising upward and cold liquid falling downward. The lower stripping (demethanizing) section 19 c contains additional trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. Demethanizing section 19 c also includes one or more reboilers (such as the reboiler 21 and side reboiler 20 shown in FIG. 5) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 44, of methane and lighter components. Stream 39 a enters demethanizer 19 at an intermediate feed position located below rectifying section 19 b and above demethanizing section 19 c. The liquid portion of the expanded stream 39 a commingles with liquids falling downward from rectifying section 19 b and the combined liquid continues downward into demethanizing section 19 c of column 19. The vapor portion of the expanded stream 39 a commingles with vapors rising upward from demethanizing section 19 c and the combined vapor rises upward through rectifying section 19 b and is contacted with cold liquid falling downward to condense and absorb the C₂components, C₃components, and heavier components from the vapor.
In demethanizing section 19 c 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 41 a) and countercurrently to the other portion of the feed gas in heat exchanger 10 where it is heated to 91° F. [33° C.] (stream 41 b) 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 41 e) 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 FIG. 5 is set forth in the following table:

TABLE V

(FIG. 5)
Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]

Stream	Methane	Ethane	Propane	Butanes+	Total

31	22,875	3,898	1,843	1,240	32,972
32	4,804	819	387	260	6,924
33	18,071	3,079	1,456	980	26,048
34	15,151	1,588	368	87	19,496
35	2,920	1,491	1,088	893	6,552
36	2,379	249	58	14	3,061
39	12,772	1,339	310	73	16,435
41	22,867	265	6	0	26,251
44	8	3,633	1,837	1,240	6,721

Recoveries*

	Ethane	93.22%
	Propane	99.70%
	Butanes+	99.99%

Power

Residue Gas Compression	15,881	HP	[26,108 kW]
Refrigerant Compression	7,678	HP	[12,623 kW]
Total Compression	23,559	HP	[38,731 kW]

*(Based on un-rounded flow rates)

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 FIG. 1), 86.57% (for FIGS. 2), and 91.35% (for FIGS. 3) to 93.22%, with essentially the same recovery as the FIG. 4 process (93.26%). The propane recovery for the present invention (99.70%) is significantly higher than that of the FIG. 1 process (98.92%) and somewhat lower than that of the FIG. 2 process (100.00%), the FIG. 3 process (99.88%), and the FIG. 4 process (99.95%). The butanes+recovery for the present invention (99.99%) is slightly higher than that of the FIG. 1 process (99.88%) and essentially the same as that of the FIG. 2 process, the FIG. 3 process, and the FIG. 4 process (100.00% for all three). Comparison of Tables I through V further shows that the improvement in yields for the present invention shown in FIG. 5 was achieved using essentially the same power requirements as the FIG. 1 through 4 processes.
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 19 b. Compared to the prior art of the FIG. 1 process, the present invention produces a better top reflux stream containing more methane and less C₂+ components. Comparing reflux stream 32 in Table I for the FIG. 1 prior art process with first reflux stream 36 in Table V for the present invention, it can be seen that the present invention provides a top reflux stream with a significantly lower concentration of C₂+ components (10.5% for the present invention versus 21.2% for the FIG. 1 prior art process). Further, the present invention uses a second reflux stream (stream 32) at an intermediate feed point of rectifying section 19 b to provide bulk recovery of the C₂components, C₃components, and heavier hydrocarbon components contained in expanded feed 39 a and the vapors rising from demethanizing section 19 c. This means that less rectification is required in the upper region of rectifying section 19 b, minimizing the required flow rate for top reflux stream 36 so that more of the separator vapor (stream 34) flows to expansion machine 15 in stream 39, which produces more power to drive compressor 16 and reduces the power required by compressor 22. Note that the total reflux provided to rectification section 19 b in streams 36 and 32 for the present invention is nearly 8% higher than that provided in the single reflux stream 32 of the FIG. 1 process.
Compared to the prior art of the FIG. 2 process, the present invention supplies its feed streams to column 19 at significantly lower temperatures, reducing the quantity of vapor entering rectification section 19 b and reducing the quantity of reflux required. Recycle stream 48 used in the FIG. 2 process to produce top reflux stream 48 c adds to the cooling load imposed on cold demethanizer overhead stream 45, reducing the cooling available in heat exchanger 10 such that the portion of the feed gas in stream 33 a entering separator 12 for the FIG. 2 process is much warmer than that of the present invention. As can be seen from Table II for the FIG. 2 process, despite top reflux stream 48 having almost 15% more flow than the present invention and a much lower concentration of C₂+ components (2.0% for the FIG. 2 prior art process versus 10.5% for the present invention), it cannot rectify the vapors rising in rectification section 19 b as well as top reflux stream 36 of the present invention.
Compared to the prior art of the FIG. 3 process, the present invention produces a much greater quantity of top reflux containing more methane and less C₂+ components. Comparing reflux stream 51 in Table III for the FIG. 3 prior art process with top reflux stream 36 in Table V for the present invention, it can be seen that the present invention provides a top reflux stream with 68% more flow and with a significantly lower concentration of C₂+ components (10.5% for the present invention versus 18.3% for the FIG. 3 prior art process). Note that although the total reflux provided to rectification section 19 b in streams 51 and 32 of the FIG. 3 process is nearly 6% higher than that of the present invention, the much higher concentration of C₂components in top reflux stream 51 of the FIG. 3 process (nearly twice that in top reflux stream 36 of the present invention) prevents it from providing efficient rectification.
Compared to the prior art of the FIG. 4 process, the present invention achieves essentially the same rectification with its two reflux streams. Comparing reflux stream 49 in Table IV for the FIG. 4 prior art process with top reflux stream 36 in Table V for the present invention, it can be seen that the present invention provides a top reflux stream with 12% more flow but with a higher concentration of C₂+ components (10.5% for the present invention versus 6.9% for the FIG. 4 prior art process). Further, the total reflux provided to rectification section 19 b for the present invention is nearly 8% higher than that in streams 49 and 32 of the FIG. 4 process. This higher reflux flow allows the present invention to match the rectification of the FIG. 4 prior art process despite the higher concentration of C₂+ components in its top reflux stream.
However, the present invention is able to match the recovery of the FIG. 4 process without vapor compressor 27 required by the prior art process. Vapor compressors in this service are expensive to install and to operate, adding to both the capital cost and the maintenance cost of the plant, reducing revenue and reducing the return on investment. Rotating equipment like this also adds to the complexity of operating the plant, making it more difficult to optimize the process for maximum recovery with minimum energy consumption. Thus, the present invention is less expensive to build, less expensive to maintain, and easier to operate than the prior art of the FIG. 4 process.

Other Embodiments

FIGS. 6 through 9 display other embodiments of the present invention. FIGS. 5 through 7 depict fractionation towers constructed in a single vessel. FIGS. 8 and 9 depict fractionation towers constructed in two vessels, absorber (rectifying) column 28 (a contacting and separating device) and partial rectification stripper (distillation) column 19. In such cases, the overhead vapor stream 52 from partial rectification stripper column 19 flows to the lower section of absorber column 28 to be contacted and further rectified by substantially condensed stream 36 b. Pump 29 is used to route the liquids (stream 53) from the bottom of absorber column 28 to the top of stripper column 19 so that the two towers effectively function as one distillation system. The decision whether to construct the fractionation tower as a single vessel (such as demethanizer 19 in FIGS. 5 through 7) or multiple vessels will depend on a number of factors such as plant size, the distance to fabrication facilities, etc.
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 36 b leaving expansion valve 14 and all or a part of expanded substantially condensed stream 32 b (FIG. 5) or 38 b (FIGS. 6 through 9) from expansion valve 17 can be combined (such as in the piping joining the expansion valves to the demethanizer) and if thoroughly intermingled, the vapors and liquids will mix together and separate in accordance with the relative volatilities of the various components of the total combined streams. Such commingling of the two streams, combined with contacting at least a portion of expanded stream 39 a, shall be considered for the purposes of this invention as constituting an absorbing section.
In accordance with the present invention, the splitting of the feed gas may be accomplished in several ways. In the processes of FIGS. 5, 6, and 8, the splitting of the feed gas occurs before any cooling of the feed gas. In such cases, cooling and substantial condensation of one portion of the feed gas in multiple heat exchangers may be favored in some circumstances, such as heat exchangers 11 and 13 shown in FIGS. 6 and 8. The feed gas may also be split, however, following cooling (but prior to separation of any liquids which may have been formed) as shown in FIGS. 7 and 9.
The high pressure liquid (stream 35 in FIGS. 5 through 9) need not be expanded, heated, and fed to a mid-column feed point on the distillation column. Instead, all or a portion of it may be combined with the portion of the cooled feed gas (stream 32 a in FIGS. 6 and 8 or stream 32 in FIGS. 7 and 9) flowing to heat exchanger 13. (This is shown by the dashed stream 37 in FIGS. 6 through 9.) Any remaining portion of the liquid (stream 40 in FIGS. 6 through 9) may be expanded through an appropriate expansion device, such as expansion valve 18 or an expansion machine, and fed to a mid-column feed point on the distillation column (stream 40a). Stream 40 may also be used for inlet gas cooling or other heat exchange service before or after the expansion step prior to flowing to the demethanizer.
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 C₂components, C₃components, and heavier components from the vapors rising through rectifying section 19 b of demethanizer 19 (FIGS. 5 through 7), or through absorber column 28 and the upper section of partial rectification stripper column 19 (FIGS. 8 and 9). However, the present invention is not limited to this embodiment. It may be advantageous, for instance, to treat only a portion of these vapors in this manner, or to use only a portion of the condensate as an absorbent, in cases where other design considerations indicate portions of the vapors or the condensate should bypass rectifying section 19 b of demethanizer 19 (FIGS. 5 through 7), or absorber column 28 and/or partial rectification stripper column 19 (FIGS. 8 and 9).
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 36 a in FIGS. 5 through 9), the substantially condensed portion of the feed stream (stream 32 a in FIG. 5), or the substantially condensed combined stream (stream 38 a in FIGS. 6 through 9).
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 C₂components, C₃components, and heavier hydrocarbon components or of C₃components 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.

Claims

We claim:

1. In a process for the separation of a gas stream containing methane, C₂components, C₃components, and heavier hydrocarbon components into a volatile residue gas fraction and a relatively less volatile fraction containing a major portion of said C₂components, C₃components, and heavier hydrocarbon components or said C₃components and heavier hydrocarbon components, in which process

(a) said gas stream is cooled under pressure to provide a cooled stream;

(b) said cooled stream is expanded to a lower pressure whereby it is further cooled; and

(c) said further cooled stream is directed into a distillation column and fractionated at said lower pressure whereby the components of said relatively less volatile fraction are recovered;

the improvement wherein prior to cooling, said gas stream is divided into first and second streams; and

(1) said first stream is cooled to condense substantially all of it;

(2) said substantially condensed first stream is expanded to said lower pressure whereby it is further cooled, and is thereafter supplied to said distillation column at an upper mid-column feed position;

(3) said second stream is cooled under pressure sufficiently to partially condense it;

(4) said partially condensed second stream is separated thereby to provide a vapor stream and at least one liquid stream;

(5) said vapor stream is divided into first and second portions;

(6) said first portion is cooled to condense substantially all of it;

(7) said substantially condensed first portion is expanded to said lower pressure whereby it is further cooled, and is thereafter supplied to said distillation column at a top feed position;

(8) said second portion is expanded to said lower pressure and is supplied to said distillation column at a mid-column feed position below said upper mid-column feed position;

(9) at least a portion of said at least one liquid stream is expanded to said lower pressure and is supplied to said distillation column at a lower mid-column feed position below said mid-column feed position; and

(10) the quantities and temperatures of said feed streams to said distillation column are effective to maintain the overhead temperature of said distillation column at a temperature whereby the major portions of the components in said relatively less volatile fraction are recovered.

2. The improvement according to claim 1 wherein

(1) said gas stream is cooled under pressure sufficiently to partially condense it;

(2) said partially condensed gas stream divided into said first stream and said second stream; and

(3) said second stream is separated thereby to provide said vapor stream and said at least one liquid stream.

3. The improvement according to claim 1 wherein

(1) said first stream is combined with at least a portion of said at least one liquid stream to form a combined stream;

(2) said combined stream is cooled to condense substantially all of it;

(3) said substantially condensed combined stream is expanded to said lower pressure whereby it is further cooled, and is thereafter supplied to said distillation column at said upper mid-column feed position; and

(4) any remaining portion of said at least one liquid stream is expanded to said lower pressure and is supplied to said distillation column at said lower mid-column feed position below said mid-column feed position.

4. The improvement according to claim 2 wherein

(2) said combined stream is cooled to condense substantially all of it;

5. The improvement according to claim 1 wherein

(1) said expanded substantially condensed first portion is supplied at a top feed position to a contacting and separating device that produces said volatile residue gas fraction and a bottom liquid stream, whereupon said bottom liquid stream is supplied to said distillation column;

(2) an overhead vapor stream is withdrawn from an upper region of said distillation column and is supplied to said contacting and separating device at a lower column feed position; and

(3) the quantities and temperatures of said feed streams to said contacting and separating device are effective to maintain the overhead temperature of said contacting and separating device at a temperature whereby the major portions of the components in said relatively less volatile fraction are recovered.

6. The improvement according to claim 5 wherein

7. The improvement according to claim 5 wherein

(2) said combined stream is cooled to condense substantially all of it;

(4) any remaining portion of said at least one liquid stream is expanded to said lower pressure and is supplied to said distillation column at said lower mid-column feed position.

8. The improvement according to claim 6 wherein

(2) said combined stream is cooled to condense substantially all of it;

9. The improvement according to claim 1 or 2 wherein said expanded said at least one liquid stream is heated and thereafter supplied to said distillation column at said lower mid-column feed position below said mid-column feed position.

10. The improvement according to claim 3 or 4 wherein said expanded said any remaining portion of said at least one liquid stream is heated and thereafter supplied to said distillation column at said lower mid-column feed position below said mid-column feed position.

11. The improvement according to claim 5 or 6 wherein said expanded said at least one liquid stream is heated and thereafter supplied to said distillation column at said lower mid-column feed position.

12. The improvement according to claim 7 or 8 wherein said expanded said any remaining portion of said at least one liquid stream is heated and thereafter supplied to said distillation column at said lower mid-column feed position.

13. In an apparatus for the separation of a gas stream containing methane, C₂components, C₃components, and heavier hydrocarbon components into a volatile residue gas fraction and a relatively less volatile fraction containing a major portion of said C₂components, C₃components, and heavier hydrocarbon components or said C₃components and heavier hydrocarbon components, in said apparatus there being

(a) a first cooling means to cool said gas stream under pressure connected to provide a cooled stream under pressure;

(b) a first expansion means connected to receive at least a portion of said cooled stream under pressure and expand it to a lower pressure, whereby said stream is further cooled; and

(c) a distillation column connected to receive said further cooled stream, said distillation column being adapted to separate said further cooled stream into said volatile residue gas fraction and said relatively less volatile fraction;

the improvement wherein said apparatus includes

(1) first dividing means prior to said first cooling means to divide said gas stream into first and second streams;

(2) second cooling means connected to said first dividing means to receive said first stream and cool it sufficiently to substantially condense it;

(3) said first expansion means connected to said second cooling means, said first expansion means being adapted to receive said substantially condensed first stream and expand it to said lower pressure, said first expansion means being further connected to said distillation column to supply said expanded substantially condensed first stream to said distillation column at an upper mid-column feed position;

(4) said first cooling means connected to said first dividing means to receive said second stream and cool it under pressure sufficiently to partially condense it;

(5) separating means connected to said first cooling means to receive said partially condensed second stream and separate it into a vapor stream and at least one liquid stream;

(6) second dividing means connected to said separating means to receive said vapor stream and divide it into first and second portions;

(7) third cooling means connected to said second dividing means to receive said first portion and cool it sufficiently to substantially condense it;

(8) second expansion means connected to said third cooling means to receive said substantially condensed first portion and expand it to said lower pressure, said second expansion means being further connected to said distillation column to supply said expanded substantially condensed first portion to said distillation column at a top feed position;

(9) third expansion means being connected to said second dividing means to receive said second portion and expand it to said lower pressure, said third expansion means being further connected to said distillation column to supply said expanded second portion to said distillation column at a mid-column feed position below said upper mid-column feed position;

(10) fourth expansion means connected to said separating means to receive at least a portion of said at least one liquid stream and expand it to said lower pressure, said fourth expansion means being further connected to said distillation column to supply said expanded liquid stream to said distillation column at a lower mid-column feed position below said mid-column feed position; and

(11) control means adapted to regulate the quantities and temperatures of said feed streams to said distillation column to maintain the overhead temperature of said distillation column at a temperature whereby the major portions of the components in said relatively less volatile fraction are recovered.

14. The improvement according to claim 13 wherein

(1) said first cooling means is connected to receive said gas stream and cool it under pressure sufficiently to partially condense it;

(2) said first dividing means is connected to said first cooling means to receive said partially condensed gas stream and divide it into said first stream and said second stream; and

(3) said separating means is connected to said first dividing means to receive said second stream and separate it into said vapor stream and said at least one liquid stream.

15. The improvement according to claim 13 wherein

(1) a combining means is connected to said first dividing means and said separating means to receive said first stream and at least a portion of said at least one liquid stream and form a combined stream;

(2) said second cooling means is connected to said combining means to receive said combined stream and cool it sufficiently to substantially condense it;

(3) said first expansion means is connected to said second cooling means to receive said substantially condensed combined stream and expand it to said lower pressure, said first expansion means being further connected to said distillation column to supply said expanded substantially condensed combined stream to said distillation column at said upper mid-column feed position; and

(4) said fourth expansion means is connected to said separating means to receive any remaining portion of said at least one liquid stream and expand it to said lower pressure, said fourth expansion means being further connected to said distillation column to supply said expanded liquid stream to said distillation column at said lower mid-column feed position below said mid-column feed position.

16. The improvement according to claim 14 wherein

17. The improvement according to claim 13 wherein

(1) said second expansion means is connected to a contacting and separating means to supply said expanded substantially condensed first portion to said contacting and separating means at a top feed position, said contacting and separating means being adapted to produce said relatively less volatile fraction and a bottom liquid stream;

(2) said distillation column is connected to receive said bottom liquid stream, said distillation column being adapted to separate said bottom liquid stream into an overhead vapor stream and said relatively less volatile fraction;

(3) said distillation column is further connected to said contacting and separating means to supply said overhead vapor stream to said contacting and separating means at a lower column feed position; and

(4) said control means is adapted to regulate the quantities and temperatures of said feed streams to said contacting and separating means to maintain the overhead temperature of said contacting and separating means at a temperature whereby the major portions of the components in said relatively less volatile fraction are recovered.

18. The improvement according to claim 17 wherein

19. The improvement according to claim 17 wherein

(4) said fourth expansion means is connected to said separating means to receive any remaining portion of said at least one liquid stream and expand it to said lower pressure, said fourth expansion means being further connected to said distillation column to supply said expanded liquid stream to said distillation column at said lower mid-column feed position.

20. The improvement according to claim 18 wherein

21. The improvement according to claim 13 or 14 wherein a heating means is connected to said fourth expansion means to receive said expanded said at least one liquid stream and heat it, said heating means being further connected to said distillation column to supply said heated expanded said at least one liquid stream to said distillation column at said lower mid-column feed point.

22. The improvement according to claim 15 or 16 wherein a heating means is connected to said fourth expansion means to receive said expanded said at least a portion of said at least one liquid stream and heat it, said heating means being further connected to said distillation column to supply said heated expanded said at least a portion of said at least one liquid stream to said distillation column at said lower mid-column feed point.

23. The improvement according to claim 17 or 18 wherein a heating means is connected to said fourth expansion means to receive said expanded said at least one liquid stream and heat it, said heating means being further connected to said distillation column to supply said heated expanded said at least one liquid stream to said distillation column at said lower mid-column feed point.

24. The improvement according to claim 19 or 20 wherein a heating means is connected to said fourth expansion means to receive said expanded said at least a portion of said at least one liquid stream and heat it, said heating means being further connected to said distillation column to supply said heated expanded said at least a portion of said at least one liquid stream to said distillation column at said lower mid-column feed point.