Patent Grant
8590340
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, 92.5% methane, 4.2% ethane and other C2 components, 1.3% propane and other C3 components, 0.4% iso-butane, 0.3% normal butane, 0.5% 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. 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; reissue U.S. Pat. No. 33,408; and co-pending application Ser. Nos. 11/430,412 and 11/839,693 describe relevant processes (although the description of the present invention in some cases is based on different processing conditions than those described in the cited U.S. Patents).
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+ or C3+ 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 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 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; and 5,881,569, co-pending application Ser. No. 11/430,412, 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, Tex., Mar. 11-13, 2002.
The present invention also employs an upper rectification section (or a separate rectification column in some embodiments). However, two reflux streams are provided for this rectification section. The upper reflux stream is a recycled stream of residue gas as described above. In addition, however, a supplemental reflux stream is provided at one or more lower feed points by using a side draw of the vapors rising in a lower portion of the tower (which may be combined with a portion of the tower overhead vapor). Because the vapor streams lower in the tower contain a modest concentration of C2 components and heavier components, this side draw stream can be substantially condensed by moderately elevating its pressure and using only the refrigeration available in the cold vapor leaving the upper rectification section. This condensed liquid, which is predominantly liquid methane and ethane, can then be used to absorb C2 components, C3 components, C4 components, and heavier hydrocarbon components from the vapors rising through the lower portion of the upper rectification section and thereby capture these valuable components in the bottom liquid product from the demethanizer. Since this lower reflux stream captures much of the C2 components and essentially all of the C3+ components, only a relatively small flow rate of liquid in the upper reflux stream is needed to absorb the C2 components remaining in the rising vapors and likewise capture these C2 components in the bottom liquid product from the demethanizer.
In accordance with the present invention, it has been found that C2 component recoveries in excess of 97 percent can be obtained. Similarly, in those instances where recovery of C2 components is not desired, C3 recoveries in excess of 98% can be maintained. In addition, the present invention makes possible essentially 100 percent separation of methane (or C2 components) and lighter components from the C2 components (or C3 components) and heavier components at reduced energy requirements compared to the prior art while maintaining the same recovery levels. 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 cooled in heat exchanger 10 by heat exchange with a portion (stream 46) of cool distillation stream 39a at −17° F. [−27° C.], bottom liquid product at 79° F. [26° C.] (stream 42a) from the demethanizer bottoms pump, 19, demethanizer reboiler liquids at 56° F. [14° C.] (stream 41), and demethanizer side reboiler liquids at −19° F. [−28° C.] (stream 40). Note that in all cases exchanger 10 is 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 31a enters separator 11 at 6° F. [−14° C.] and 1025 psia [7,067 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33).
The vapor (stream 32) from separator 11 is divided into two streams, 34 and 36. Stream 34, containing about 30% of the total vapor, is combined with the separator liquid (stream 33). The combined stream 35 then passes through heat exchanger 12 in heat exchange relation with cold distillation stream 39 at −142° F. [−96° C.] where it is cooled to substantial condensation. The resulting substantially condensed stream 35a at −138° F. [−94° C.] is then flash expanded through an appropriate expansion device, such as expansion valve 13, to the operating pressure (approximately 423 psia [2,916 kPa(a)]) of fractionation tower 17. The expanded stream 35b leaving expansion valve 13 reaches a temperature of −140° F. [−96° C.] and is supplied to fractionation tower 17 at a mid-column feed point.
The remaining 70% of the vapor from separator 11 (stream 36) enters a work expansion machine 14 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 14 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 36a to a temperature of approximately −75° F. [−60° C.]. The typical commercially available expanders are capable of recovering on the order of 80-88% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item 15) that can be used to re-compress the heated distillation stream (stream 39b), for example. The partially condensed expanded stream 36a is thereafter supplied to fractionation tower 17 at a second mid-column feed point.
The recompressed and cooled distillation stream 39e is divided into two streams. One portion, stream 47, is the volatile residue gas product. The other portion, recycle stream 48, flows to heat exchanger 22 where it is cooled to −6° F. [−21° C.] (stream 48a) by heat exchange with a portion (stream 45) of cool distillation stream 39a. The cooled recycle stream then flows to exchanger 12 where it is cooled to −138° F. [−94° C.] and substantially condensed by heat exchange with cold distillation stream 39 at −142° F. [−96° C.]. The substantially condensed stream 48b is then expanded through an appropriate expansion device, such as expansion valve 23, to the demethanizer operating pressure, resulting in cooling of the total stream to −144° F. [−98° C.]. The expanded stream 48c is then supplied to fractionation tower 17 as the top column feed. The vapor portion (if any) of stream 48c combines with the vapors rising from the top fractionation stage of the column to form distillation stream 39, which is withdrawn from an upper region of the tower.
The demethanizer in tower 17 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 two sections. The upper section 17a is a separator wherein the partially vaporized top feed is divided into its respective vapor and liquid portions, and wherein the vapor rising from the lower distillation or demethanizing section 17b is combined with the vapor portion (if any) of the top feed to form the cold demethanizer overhead vapor (stream 39) which exits the top of the tower at −142° F. [−96° C.]. The lower, demethanizing section 17b contains the trays and/or packing and provides the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section 17b also includes reboilers (such as the reboiler and side reboiler described previously) 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 42, of methane and lighter components.
Liquid product stream 42 exits the bottom of the tower at 75° F. [24° C.], based on a typical specification of a methane to ethane ratio of 0.025:1 on a molar basis in the bottom product. It is pumped to a pressure of approximately 650 psia [4,482 kPa(a)] in demethanizer bottoms pump 19, and the pumped liquid product is then warmed to 116° F. [47° C.] as it provides cooling of stream 31 in exchanger 10 before flowing to storage.
The demethanizer overhead vapor (stream 39) passes countercurrently to the incoming feed gas and recycle stream in heat exchanger 12 where it is heated to −17° F. [−27° C.] (stream 39a), and in heat exchanger 22 and heat exchanger 10 where it is heated to 84° F. [29° C.] (stream 39b). The distillation stream is then re-compressed in two stages. The first stage is compressor 15 driven by expansion machine 14. The second stage is compressor 20 driven by a supplemental power source which compresses stream 39c to sales line pressure (stream 39d). After cooling to 120° F. [49° C.] in discharge cooler 21, stream 39e is split into the residue gas product (stream 47) and the recycle stream 48 as described earlier. Residue gas stream 47 flows to the sales gas pipeline at 1040 psia [7,171 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
The feed stream 31 is cooled in heat exchanger 10 by heat exchange with a portion of the cool distillation column overhead stream (stream 46) at −76° F. [−60° C.], demethanizer bottoms liquid (stream 42a) at 87° F. [31° C.], demethanizer reboiler liquids at 62° F. [17° C.] (stream 41), and demethanizer side reboiler liquids at −42° F. [−41° C.] (stream 40). The cooled stream 31a enters separator 11 at −46° F. [−43° C.] and 1025 psia [7,067 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33).
The separator vapor (stream 32) enters a work expansion machine 14 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 14 expands the vapor substantially isentropically to the tower operating pressure of 461 psia [3,178 kPa(a)], with the work expansion cooling the expanded stream 32a to a temperature of approximately −111° F. [−79° C.]. The partially condensed expanded stream 32a is thereafter supplied to fractionation tower 17 at a mid-column feed point.
The recompressed and cooled distillation stream 39e is divided into two streams. One portion, stream 47, is the volatile residue gas product. The other portion, recycle stream 48, flows to heat exchanger 22 where it is cooled to −70° F. [−57° C.] (stream 48a) by heat exchange with a portion (stream 45) of cool distillation stream 39a at −76° F. [−60° C.]. The cooled recycle stream then flows to exchanger 12 where it is cooled to −133° F. [−92° C.] and substantially condensed by heat exchange with cold distillation column overhead stream 39. The substantially condensed stream 48b is then expanded through an appropriate expansion device, such as expansion valve 23, to the demethanizer operating pressure, resulting in cooling of the total stream to −141° F. [−96° C.]. The expanded stream 48c is then supplied to the fractionation tower as the top column feed. The vapor portion (if any) of stream 48c combines with the vapors rising from the top fractionation stage of the column to form distillation stream 39, which is withdrawn from an upper region of the tower.
A portion of the distillation vapor (stream 49) is withdrawn from fractionation tower 17 at −119° F. [−84° C.] and is compressed to about 727 psia [5,015 kPa(a)] by reflux compressor 24. The separator liquid (stream 33) is expanded to this pressure by expansion valve 16, and the expanded stream 33a at −62° F. [−52° C.] is combined with stream 49a at −66° F. [−54° C.]. The combined stream 35 is then cooled from −68° F. [−56° C.] to −133° F. [−92° C.] and condensed (stream 35a) in heat exchanger 12 by heat exchange with the cold demethanizer overhead stream 39 exiting the top of demethanizer 17 at −137° F. [−94° C.]. The resulting substantially condensed stream 35a is then flash expanded through expansion valve 13 to the operating pressure of fractionation tower 17, cooling stream 35b to a temperature of −135° F. [−93° C.] whereupon it is supplied to fractionation tower 17 at a mid-column feed point.
The liquid product stream 42 exits the bottom of the tower at 82° F. [28° C.], based on a typical specification of a methane to ethane ratio of 0.025:1 on a molar basis in the bottom product. Pump 19 delivers stream 42a to heat exchanger 10 as described previously where it is heated from 87° F. [31° C.] to 116° F. [47° C.] before flowing to storage.
The demethanizer overhead vapor stream 39 is warmed in heat exchanger 12 as it provides cooling to combined stream 35 and recycle stream 48a as described previously, and further heated in heat exchanger 22 and heat exchanger 10. The heated stream 39b at 96° F. [36° C.] is then re-compressed in two stages, compressor 15 driven by expansion machine 14 and compressor 20 driven by a supplemental power source. After stream 39d is cooled to 120° F. [49° C.] in discharge cooler 21 to form stream 39e, recycle stream 48 is withdrawn as described earlier to form residue gas stream 47 which flows to the sales gas pipeline at 1040 psia [7,171 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
Comparison of the recovery levels displayed in Tables I and II shows that the liquids recovery of the
In the simulation of the
The vapor (stream 32) from separator 11 is divided into two streams, 34 and 36. Likewise, the liquid (stream 33) from separator 11 is divided into two streams, 37 and 38. Stream 34, containing about 10% of the total vapor, is combined with stream 37, containing about 50% of the total liquid. The combined stream 35 then passes through heat exchanger 12 in heat exchange relation with cold distillation stream 39 at −137° F. [−94° C.] where it is cooled to substantial condensation. The resulting substantially condensed stream 35a at −133° F. [−92° C.] is then flash expanded through an appropriate expansion device, such as expansion valve 13, to the operating pressure (approximately 460 psia [3,172 kPa(a)]) of fractionation tower 17, cooling stream 35b to −135° F. [−93° C.] before it is supplied to fractionation tower 17 at a mid-column feed point.
The remaining 90% of the vapor from separator 11 (stream 36) enters a work expansion machine 14 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 14 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 36a to a temperature of approximately −103° F. [−75° C.]. The partially condensed expanded stream 36a is thereafter supplied as feed to fractionation tower 17 at a second mid-column feed point.
The remaining 50% of the liquid from separator 11 (stream 38) is flash expanded through an appropriate expansion device, such as expansion valve 16, to the operating pressure of fractionation tower 17. The expansion cools stream 38a to −65° F. [−54° C.] before it is supplied to fractionation tower 17 at a third mid-column feed point.
The recompressed and cooled distillation stream 39e is divided into two streams. One portion, stream 47, is the volatile residue gas product. The other portion, recycle stream 48, flows to heat exchanger 22 where it is cooled to −1° F. [−18° C.] (stream 48a) by heat exchange with a portion (stream 45) of cool distillation stream 39a. The cooled recycle stream then flows to exchanger 12 where it is cooled to −133° F. [−92° C.] and substantially condensed by heat exchange with cold distillation stream 39. The substantially condensed stream 48b is then expanded through an appropriate expansion device, such as expansion valve 23, to the demethanizer operating pressure, resulting in cooling of the total stream to −141° F. [−96° C.]. The expanded stream 48c is then supplied to fractionation tower 17 as the top column feed. The vapor portion (if any) of stream 48c combines with the vapors rising from the top fractionation stage of the column to form distillation stream 39, which is withdrawn from an upper region of the tower.
A portion of the distillation vapor (stream 49) is withdrawn from the lower region of absorbing section 17b of fractionation tower 17 at −129° F. [−90° C.] and is compressed to an intermediate pressure of about 697 psia [4,804 kPa(a)] by reflux compressor 24. The compressed stream 49a flows to exchanger 12 where it is cooled to −133° F. [−92° C.] and substantially condensed by heat exchange with cold distillation column overhead stream 39. The substantially condensed stream 49b is then expanded through an appropriate expansion device, such as expansion valve 25, to the demethanizer operating pressure, resulting in cooling of stream 49c to a temperature of −137° F. [−94° C.], whereupon it is supplied to fractionation tower 17 at a fourth mid-column feed point.
The demethanizer in tower 17 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: an upper separator section 17a wherein the top feed is divided into its respective vapor and liquid portions, and wherein the vapor rising from the intermediate absorbing section 17b is combined with the vapor portion (if any) of the top feed to form the cold demethanizer overhead vapor (stream 39); an intermediate absorbing (rectification) section 17b that contains the trays and/or packing to provide the necessary contact between the vapor portion of the expanded stream 36a rising upward and cold liquid falling downward to condense and absorb the C2 components, C3 components, and heavier components; and a lower, stripping (demethanizing) section 17c that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section 17c also includes reboilers (such as the reboiler and side reboiler described previously) 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 42, of methane and lighter components.
Stream 36a enters demethanizer 17 at a feed position located in the lower region of absorbing section 17b of demethanizer 17. The liquid portion of expanded stream 36a commingles with liquids falling downward from the absorbing section 17b and the combined liquid continues downward into the stripping section 17c of demethanizer 17. The vapor portion of expanded stream 36a rises upward through absorbing section 17b and is contacted with cold liquid falling downward to condense and absorb the C2 components, C3 components, and heavier components.
The expanded substantially condensed stream 49c is supplied as cold liquid reflux to an intermediate region in absorbing section 17b of demethanizer 17, as is expanded substantially condensed stream 35b. These secondary reflux streams absorb and condense most of the C3 components and heavier components (as well as much of the C2 components) from the vapors rising in the lower rectification region of absorbing section 17b so that only a small amount of recycle (stream 48) must be cooled, condensed, subcooled, and flash expanded to produce the top reflux stream 48c that provides the final rectification in the upper region of absorbing section 17b. As the cold top reflux stream 48c contacts the rising vapors in the upper region of absorbing section 17b, it condenses and absorbs the C2 components and any remaining C3 components and heavier components from the vapors so that they can be captured in the bottom product (stream 42) from demethanizer 17.
In stripping section 17c of demethanizer 17, the feed streams are stripped of their methane and lighter components. The resulting liquid product (stream 42) exits the bottom of tower 17 at 86° F. [30° C.], based on a typical specification of a methane to ethane ratio of 0.025:1 on a molar basis in the bottom product. Pump 19 delivers stream 42a to heat exchanger 10 as described previously where it is heated to 116° F. [47° C.] (stream 42b) before flowing to storage.
The distillation vapor stream forming the tower overhead (stream 39) is warmed in heat exchanger 12 as it provides cooling to combined stream 35, compressed distillation vapor stream 49a, and recycle stream 48a as described previously to form cool distillation stream 39a. Distillation stream 39a is divided into two portions (streams 45 and 46), which are heated to 116° F. [47° C.] and 92° F. [33° C.], respectively, in heat exchanger 22 and heat exchanger 10. Note that in all cases exchangers 10, 22, and 12 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 heating services will depend on a number of factors including, but not limited to, inlet gas flow rate, heat exchanger size, stream temperatures, etc.) The heated streams recombine to form stream 39b at 94° F. [34° C.] which is then re-compressed in two stages, compressor 15 driven by expansion machine 14 and compressor 20 driven by a supplemental power source. After stream 39d is cooled to 120° F. [49° C.] in discharge cooler 21 to form stream 39e, recycle stream 48 is withdrawn as described earlier to form residue gas stream 47 which flows to the sales gas pipeline at 1040 psia [7,171 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, compared to the prior art processes, the present invention maintains essentially the same ethane recovery, propane recovery, and butanes+ recovery. However, comparison of Tables I, II, and III further shoes that these yields were achieved with substantially lower horsepower requirements than those of the prior art processes. The total power requirement of the present invention 11% lower than that of the
The key feature of the present invention is the supplemental rectification provided by reflux stream 49c in conjunction with stream 35b, which reduces the amount of C2 components, C3 components, and C4+ components contained in the vapors rising in the upper region of absorbing section 17b. Compare these two supplemental reflux streams in Table III with the single supplemental reflux stream, 35b, in Table I for the
A further advantage provided by supplemental reflux stream 49c is that it allows a reduction in the flow rate of supplemental reflux stream 35b, so that there is a corresponding increase in the flow rate of stream 36 to work expansion machine 14. This in turn provides a two-fold improvement in the process efficiency. First, with more flow to expansion machine 14, the increase in power recovery increases the refrigeration generated by the process. Second, the greater power recovery means more power available to compressor 15, reducing the external power consumption of compressor 20.
Compared to the
Note that in the
In the simulation of the
The vapor (stream 32) from separator 11 is divided into two streams, 34 and 36. Likewise, the liquid (stream 33) from separator 11 is divided into two streams, 37 and 38. Stream 34, containing about 11% of the total vapor, is combined with stream 37, containing about 50% of the total liquid. The combined stream 35 then passes through heat exchanger 12 in heat exchange relation with cold distillation stream 39 at −136° F. [−94° C.] where it is cooled to substantial condensation. The resulting substantially condensed stream 35a at −132° F. [−91° C.] is then flash expanded through an appropriate expansion device, such as expansion valve 13, to the operating pressure (approximately 465 psia [3,206 kPa(a)]) of fractionation tower 17, cooling stream 35b to −134° F. [−92° C.] before it is supplied to fractionation tower 17 at a mid-column feed point.
The remaining 89% of the vapor from separator 11 (stream 36) enters a work expansion machine 14 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 14 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 36a to a temperature of approximately −99° F. [−73° C.]. The partially condensed expanded stream 36a is thereafter supplied as feed to fractionation tower 17 at a second mid-column feed point.
The remaining 50% of the liquid from separator 11 (stream 38) is flash expanded through an appropriate expansion device, such as expansion valve 16, to the operating pressure of fractionation tower 17. The expansion cools stream 38a to −60° F. [−51° C.] before it is supplied to fractionation tower 17 at a third mid-column feed point.
The recompressed and cooled distillation stream 39e is divided into two streams. One portion, stream 47, is the volatile residue gas product. The other portion, recycle stream 48, flows to heat exchanger 22 where it is cooled to −1° F. [−18° C.] (stream 48a) by heat exchange with a portion (stream 45) of cool distillation stream 39a. The cooled recycle stream then flows to exchanger 12 where it is cooled to −132° F. [−91° C.] and substantially condensed by heat exchange with cold distillation stream 39. The substantially condensed stream 48b is then expanded through an appropriate expansion device, such as expansion valve 23, to the demethanizer operating pressure, resulting in cooling of the total stream to −140° F. [−96° C.]. The expanded stream 48c is then supplied to fractionation tower 17 as the top column feed. The vapor portion (if any) of stream 48c combines with the vapors rising from the top fractionation stage of the column to form distillation stream 39, which is withdrawn from an upper region of the tower.
A portion of the distillation vapor (stream 49) is withdrawn from the lower region of the absorbing section of fractionation tower 17 at −129° F. [−89° C.] and is compressed to an intermediate pressure of about 697 psia [4,804 kPa(a)] by reflux compressor 24. The compressed stream 49a flows to exchanger 12 where it is cooled to −132° F. [−91° C.] and substantially condensed by heat exchange with cold distillation column overhead stream 39. The substantially condensed stream 49b is then divided into two portions, streams 51 and 52. The first portion, stream 51 containing about 90% of stream 49b, is expanded through an appropriate expansion device, such as expansion valve 25, to the demethanizer operating pressure, resulting in cooling of stream 51a to a temperature of −136° F. [−94° C.], whereupon it is supplied to fractionation tower 17 at a fourth mid-column feed point as in the
In the stripping section of demethanizer 17, the feed streams are stripped of their methane and lighter components. The resulting liquid product (stream 42) exits the bottom of tower 17 at 88° F. [31° C.]. Pump 19 delivers stream 42a to heat exchanger 10 as described previously where it is heated to 116° F. [47° C.] (stream 42b) before flowing to storage.
The distillation vapor stream forming the tower overhead (stream 39) is warmed in heat exchanger 12 as it provides cooling to combined stream 35, compressed distillation vapor stream 49a, and recycle stream 48a as described previously to form cool distillation stream 39a. Distillation stream 39a is divided into two portions (streams 45 and 46), which are heated to 116° F. [47° C.] and 92° F. [33° C.], respectively, in heat exchanger 22 and heat exchanger 10. The heated streams recombine to form stream 39b at 94° F. [35° C.] which is then re-compressed in two stages, compressor 15 driven by expansion machine 14 and compressor 20 driven by a supplemental power source. After stream 39d is cooled to 120° F. [49° C.] in discharge cooler 21 to form stream 39e, recycle stream 48 is withdrawn as described earlier to form residue gas stream 47 which flows to the sales gas pipeline at 1040 psia [7,171 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables III and IV shows that, compared to the
An alternative method of generating the supplemental reflux streams for the column is shown in another embodiment of the present invention as illustrated in
In the simulation of the
The vapor (stream 32) from separator 11 is divided into two streams, 34 and 36. Likewise, the liquid (stream 33) from separator 11 is divided into two streams, 37 and 38. Stream 34, containing about 10% of the total vapor, is combined with stream 37, containing about 50% of the total liquid. The combined stream 35 then passes through heat exchanger 12 in heat exchange relation with cold vapor stream 43 at −137° F. [−94° C.] where it is cooled to substantial condensation. The resulting substantially condensed stream 35a at −133° F. [−91° C.] is then flash expanded through an appropriate expansion device, such as expansion valve 13, to the operating pressure (approximately 464 psia [3,199 kPa(a)]) of fractionation tower 17, cooling stream 35b to −134° F. [−92° C.] before it is supplied to fractionation tower 17 at a mid-column feed point.
The remaining 90% of the vapor from separator 11 (stream 36) enters a work expansion machine 14 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 14 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 36a to a temperature of approximately −102° F. [−75° C.]. The partially condensed expanded stream 36a is thereafter supplied as feed to fractionation tower 17 at a second mid-column feed point.
The remaining 50% of the liquid from separator 11 (stream 38) is flash expanded through an appropriate expansion device, such as expansion valve 16, to the operating pressure of fractionation tower 17. The expansion cools stream 38a to =65° F. [−54° C.] before it is supplied to fractionation tower 17 at a third mid-column feed point.
The recompressed and cooled vapor stream 43e is divided into two streams. One portion, stream 47, is the volatile residue gas product. The other portion, recycle stream 48, flows to heat exchanger 22 where it is cooled to −1° F. [−18° C.] (stream 48a) by heat exchange with a portion (stream 45) of cool vapor stream 43a. The cooled recycle stream then flows to exchanger 12 where it is cooled to −133° F. [−91° C.] and substantially condensed by heat exchange with cold vapor stream 43. The substantially condensed stream 48b is then expanded through an appropriate expansion device, such as expansion valve 23, to the demethanizer operating pressure, resulting in cooling of the total stream to −140° F. [−96° C.]. The expanded stream 48c is then supplied to fractionation tower 17 as the top column feed. The vapor portion (if any) of stream 48c combines with the vapors rising from the top fractionation stage of the column to form distillation stream 39, which is withdrawn from an upper region of the tower.
The distillation vapor stream forming the tower overhead (stream 39) leaves fractionation tower 17 at −137° F. [−94° C.] and is divided into two portions, first and second vapor streams 44 and 43, respectively. First vapor stream 44 is combined with a portion of the distillation vapor (stream 49) withdrawn from the lower region of the absorbing section of fractionation tower 17 at −131° F. [−90° C.], and the combined vapor stream 50 is compressed to an intermediate pressure of about 723 psia [4,985 kPa(a)] by reflux compressor 24. The compressed stream 50a flows to exchanger 12 where it is cooled to −133° F. [−91° C.] and substantially condensed by heat exchange with the remaining portion (stream 43) of cold distillation column overhead stream 39. The substantially condensed stream 50b is then expanded through an appropriate expansion device, such as expansion valve 25, to the demethanizer operating pressure, resulting in cooling of stream 50c to a temperature of −137° F. [−94° C.], whereupon it is supplied to fractionation tower 17 at a fourth mid-column feed point.
In the stripping section of demethanizer 17, the feed streams are stripped of their methane and lighter components. The resulting liquid product (stream 42) exits the bottom of tower 17 at 87° F. [31° C.]. Pump 19 delivers stream 42a to heat exchanger 10 as described previously where it is heated to 116° F. [47° C.] (stream 42b) before flowing to storage.
Second vapor stream 43 (the remaining portion of cold distillation column overhead stream 39) is warmed in heat exchanger 12 as it provides cooling to combined steam 35, compressed combined stream 50a, and recycle stream 48a as described previously to form cool second vapor stream 43a. Second vapor stream 43a is divided into two portions (streams 45 and 46), which are heated to 116° F. [47° C.] and 94° F. [34° C.], respectively, in heat exchanger 22 and heat exchanger 10. The heated streams recombine to form stream 43b at 95° F. [35° C.] which is then re-compressed in two stages, compressor 15 driven by expansion machine 14 and compressor 20 driven by a supplemental power source. After stream 43d is cooled to 120° F. [49° C.] in discharge cooler 21 to form stream 43e, recycle stream 48 is withdrawn as described earlier to form residue gas stream 47 which flows to the sales gas pipeline at 1040 psia [7,171 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables III, IV, and V shows that, compared to the
An alternative method of using the supplemental reflux streams for the column is shown in another embodiment of the present invention as illustrated in
In the simulation of the
The vapor (stream 32) from separator 11 is divided into two streams, 34 and 36. Likewise, the liquid (stream 33) from separator 11 is divided into two streams, 37 and 38. Stream 34, containing about 12% of the total vapor, is combined with stream 37, containing about 50% of the total liquid. The combined stream 35 then passes through heat exchanger 12 in heat exchange relation with cold vapor stream 43 at −136° F. [−93° C.] where it is cooled to substantial condensation. The resulting substantially condensed stream 35a at −132° F. [−91° C.] is then flash expanded through an appropriate expansion device, such as expansion valve 13, to the operating pressure (approximately 469 psia [3,234 kPa(a)]) of fractionation tower 17, cooling stream 35b to −134° F. [92° C.] before it is supplied to fractionation tower 17 at a mid-column feed point.
The remaining 88% of the vapor from separator 11 (stream 36) enters a work expansion machine 14 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 14 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 36a to a temperature of approximately −99° F. [−73° C.]. The partially condensed expanded stream 36a is thereafter supplied as feed to fractionation tower 17 at a second mid-column feed point.
The remaining 50% of the liquid from separator 11 (stream 38) is flash expanded through an appropriate expansion device, such as expansion valve 16, to the operating pressure of fractionation tower 17. The expansion cools stream 38a to −59° F. [−51° C.] before it is supplied to fractionation tower 17 at a third mid-column feed point.
The recompressed and cooled vapor stream 43e is divided into two streams. One portion, stream 47, is the volatile residue gas product. The other portion, recycle stream 48, flows to heat exchanger 22 where it is cooled to −1° F. [−18° C.] (stream 48a) by heat exchange with a portion (stream 45) of cool vapor stream 43a. The cooled recycle stream then flows to exchanger 12 where it is cooled to −132° F. [−91° C.] and substantially condensed by heat exchange with cold vapor stream 43. The substantially condensed stream 48b is then expanded through an appropriate expansion device, such as expansion valve 23, to the demethanizer operating pressure, resulting in cooling of the total stream to −140° F. [−95° C.]. The expanded stream 48c is then supplied to fractionation tower 17 as the top column feed. The vapor portion (if any) of stream 48c combines with the vapors rising from the top fractionation stage of the column to form distillation stream 39, which is withdrawn from an upper region of the tower.
The distillation vapor stream forming the tower overhead (stream 39) leaves fractionation tower 17 at −136° F. [−93° C.] and is divided into two portions, first and second vapor streams 44 and 43, respectively. First vapor stream 44 is combined with a portion of the distillation vapor (stream 49) withdrawn from the lower region of the absorbing section of fractionation tower 17 at −128° F. [−89° C.], and the combined vapor stream 50 is compressed to an intermediate pressure of about 732 psia [5,047 kPa(a)] by reflux compressor 24. The compressed stream 50a flows to exchanger 12 where it is cooled to −132° F. [−91° C.] and substantially condensed by heat exchange with the remaining portion (stream 43) of cold distillation column overhead stream 39. The substantially condensed stream 50b is then divided into two portions, streams 51 and 52. The first portion, stream 51 containing about 90% of stream 50b, is expanded through an appropriate expansion device, such as expansion valve 25, to the demethanizer operating pressure, resulting in cooling of stream 51a to a temperature of −136° F. [−94° C.], whereupon it is supplied to fractionation tower 17 at a fourth mid-column feed point as in the
In the stripping section of demethanizer 17, the feed streams are stripped of their methane and lighter components. The resulting liquid product (stream 42) exits the bottom of tower 17 at 89° F. [31° C.]. Pump 19 delivers stream 42a to heat exchanger 10 as described previously where it is heated to 116° F. [47° C.] (stream 42b) before flowing to storage.
Second vapor stream 43 (the remaining portion of cold distillation column overhead stream 39) is warmed in heat exchanger 12 as it provides cooling to combined stream 35, compressed combined stream 50a, and recycle stream 48a as described previously to form cool second vapor stream 43a. Second vapor stream 43a is divided into two portions (streams 45 and 46), which are heated to 116° F. [47° C.] and 94° F. [34° C.], respectively, in heat exchanger 22 and heat exchanger 10. The heated streams recombine to form stream 43b at 96° F. [35° C.] which is then re-compressed in two stages, compressor 15 driven by expansion machine 14 and compressor 20 driven by a supplemental power source. After stream 43d is cooled to 120° F. [49° C.] in discharge cooler 21 to form stream 43e, recycle stream 48 is withdrawn as described earlier to form residue gas stream 47 which flows to the sales gas pipeline at 1040 psia [7,171 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables III, IV, V, and VI shows that, compared to the
In accordance with this invention, it is generally advantageous to design the absorbing (rectification) section of the demethanizer to contain multiple theoretical separation stages. However, the benefits of the present invention can be achieved with as few as one theoretical stage, and it is believed that even the equivalent of a fractional theoretical stage may allow achieving these benefits. For instance, all or a part of the expanded substantially condensed recycle stream 48c, all or a part of the supplemental reflux (stream 49c in
As described in the earlier examples, the supplemental reflux (stream 49b in
Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine 14, 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 recycle stream (stream 48b), the supplemental reflux (stream 49b, stream 50b, or streams 51 and/or 52), or the substantially condensed stream (stream 35a).
When the inlet gas is leaner, separator 11 in
In accordance with this invention, the use of external refrigeration to supplement the cooling available to the inlet gas and/or the recycle gas 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 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 feed and the split liquid feed 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. The relative locations of the mid-column feeds and the withdrawal point of distillation vapor stream 49 may vary depending on inlet composition or other factors such as desired recovery levels and amount of liquid formed during inlet gas cooling. In some circumstances, withdrawal of distillation vapor stream 49 below the feed location of expanded stream 36a is favored. 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 intermediate pressure to which distillation stream 49 or combined vapor stream 50 is compressed must be determined for each application, as it is a function of inlet composition, the desired recovery level, the withdrawal point of distillation vapor stream 49, and other factors.
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 for the separation of a gas containing hydrocarbons. The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application No. 60/900,400 which was filed on Feb. 9, 2007.
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