Patent Application
20210116174
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. Hydrocarbon bearing gas often contains components more volatile than methane as well as unsaturated hydrocarbons (e.g., ethylene, propylene, etc.) in addition to methane, ethane and hydrocarbons of higher molecular weight such as propane, butane, and pentane.
The present invention is generally concerned with improving 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, 42.4% methane, 22.8% ethane and other C2 components, 7.6% propane and other C3 components, 3.1% iso-butane, 2.7% normal butane, and 2.7% pentanes plus, with the balance made up of hydrogen, nitrogen, carbon monoxide, and carbon dioxide. Sulfur-containing gases are also sometimes present.
Recent changes in ethylene demand have created increased markets for ethylene and derivative products. In addition, fluctuations in the prices of both natural gas and its natural gas liquid (NGL) constituents have increased the incremental value of ethane, ethylene, propane, propylene, and heavier components as liquid products. These market conditions have resulted in a demand for processes that can provide high recoveries and more efficient recoveries of these products, and for processes that can provide efficient recoveries with lower capital investment. 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; 9,080,811; 9,476,639; 9,637,428; 9,783,470; 9,927,171; 9,933,207; 9,939,195; 10,227,273; 10,553,794; 10,551,118; 10,551,119; and 10,753,678; reissue U.S. Patent No. 33,408; and published U.S. applications US 2008/0078205 A1; US 2011/0067441 A1; US 2011/0067443 A1; US 2016/0069610 A1; US 2016/0377341 A1; US 2018/0347898 A1; US 2018/0347899 A1; US 2019/0170435 A1; and US 2020/0292230 A1 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 and published U.S. 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) 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.
If the feed gas is not totally condensed (typically it is not), the vapor remaining from the partial condensation can be 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.
In the ideal operation of such a separation process, the residue gas leaving the process will contain substantially all of the methane and more volatile components 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 residue gas from 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. This problem is exacerbated if the gas stream(s) being processed contain relatively large quantities of components more volatile than methane (e.g., hydrogen, nitrogen, etc.), because the volatile vapors rising up the column strip C2+ components from the liquids flowing downward. The loss of these desirable C2+ 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. For many of these processes, a portion of the vapor remaining from the partial condensation of the feed gas 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. Unfortunately, this feed stream is not very effective at capturing the desired C2+ components when the feed gas contains significant quantities of components more volatile than methane because the stream cannot be substantially condensed. This results in large amounts of flash vapor in the stream, which carries away equilibrium quantities of the C2+ components rather than recovering them in the column.
Many processes combine this condensed flash expanded stream with another source of top reflux to the column. The source of this reflux stream is 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. Typical process schemes of this type are disclosed in U.S. Pat. Nos. 4,889,545; 5,568,737; 5,881,569; 9,052,137; and 9,080,811, 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. Unfortunately, this method is not efficient when the residue gas contains significant quantities of components more volatile than methane because the recycle stream cannot be cooled to substantial condensation.
Another means of providing a reflux stream for the upper rectification section is to withdraw a distillation vapor stream from a lower location on the tower (and perhaps combine it with a portion of the tower overhead vapor). This vapor (or combined vapor) stream is compressed to higher pressure, then cooled to substantial condensation, expanded to the tower operating pressure, and supplied as top feed to the tower. Typical process schemes of this type are disclosed in U.S. Pat. Nos. 5,275,005 and 9,476,639 and in published applications US 2008/0078205 A1 and US 2011/0067443 A1. However, this stream is generally not sufficient to provide the desired rectification by itself and combining it with the condensed flash expanded stream described earlier is not effective when there is a significant quantity of components more volatile than methane in the feed gas.
The present invention also employs an upper rectification section (or a separate rectification column in some embodiments) and a withdrawn distillation vapor stream supplied under pressure. However, much of the reflux for this upper rectification section is provided by cooling a liquid stream derived from the feed gas and then expanding the stream to the operating pressure of the fractionation tower. During expansion, a portion of the stream is vaporized, resulting in cooling of the total stream. The cooled, expanded stream is supplied to the tower at an upper mid-column feed point, where along with the condensed liquid in the stream supplied to the top column feed point (which is predominantly liquid methane), it can then be used to absorb C2 components, C3 components, C4 components, and heavier hydrocarbon components from the vapors rising through the upper rectification section and thereby capture these valuable components in the bottom liquid product from the demethanizer.
In accordance with the present invention, it has been found that C2 component recovery in excess of 96% and C3 component and C4+ component recoveries in excess of 99% can be obtained. In addition, the present invention makes possible essentially 100% separation of methane and lighter components from the C2 components and heavier components at the same energy requirements compared to the prior art while increasing the recovery level. The present invention is particularly advantageous when processing feed gases that contain more than 10 mole % of components more volatile than methane (e.g., hydrogen, nitrogen, etc.).
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 Systeme 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 enters compressor 16 driven by expansion machine 15 and is boosted to higher pressure. After cooling to 142° F. [61° C.] in cooler 20, stream 31b at 1146 psia [7,901 kPa(a)] is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 39a) at −57° F. [−49° C.], pumped liquid product (stream 42a) at 12° F. [−11° C.], demethanizer reboiler liquids (stream 41) at −11° F. [−24° C.], and propane refrigerant. 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 31c enters separator 11 at −14° F. [−26° C.] and 1123 psia [7,743 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 35). The separator liquid (stream 35) is expanded to the operating pressure (approximately 141 psia [972 kPa(a)]) of fractionation tower 19 by expansion valve 18, cooling stream 35a to −64° F. [−53° C.] before it is supplied to fractionation tower 19 at a lower mid-column feed position.
The vapor (stream 32) from separator 11 is divided into two streams, 33 and 34. Stream 33, containing about 5% of the total vapor, passes through heat exchanger 12 in heat exchange relation with cold distillation stream 39 at −143° F. [−97° C.] where it is cooled to substantial condensation. The resulting substantially condensed stream 33a at −119° F. [−84° C.] is then flash expanded through an appropriate expansion device, such as expansion valve 17, to the operating pressure of fractionation tower 19. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
The remaining vapor from separator 11 (stream 34) enters work expansion machine 15 in which mechanical energy is extracted from this portion of the high-pressure feed. The machine 15 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −132° F. [−91° 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 compress the feed gas (stream 31), for example. The partially condensed expanded stream 34a is thereafter supplied as feed to fractionation tower 19 at a mid-column feed position.
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 gas processing plants, the fractionation tower may consist of three sections. The upper section 19a is a separator wherein the partially vaporized top feed is divided into its respective vapor and liquid portions, and wherein the vapor rising from the middle absorbing (rectifying) section 19b is combined with the vapor portion of the top feed to form the cold demethanizer overhead vapor (stream 39) which exits the top of the tower at −143° F. [−97° C.]. The lower stripping (demethanizing) section 19c contains the trays and/or packing and provides the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section 19c also includes reboilers (such as the 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 34a enters demethanizer 19 at an intermediate feed position located in the lower region of rectifying section 19b of demethanizer 19. The liquid portion of the expanded stream commingles with liquids falling downward from rectifying section 19b and the combined liquid continues downward into demethanizing section 19c of demethanizer 19. The vapor portion of the expanded stream commingles with vapors arising from demethanizing section 19c and the combined vapor rises upward through rectifying section 19b and is contacted with cold liquid falling downward to condense and absorb the C2 components, C3 components, and heavier components.
A portion of the distillation vapor (stream 40) is withdrawn from an upper region of demethanizing section 19c in fractionation column 19, below the feed position of expanded stream 34a in the lower region of rectifying section 19b. The distillation vapor stream 40 at −93° F. [−77° C.] is heated to 132° F. [55° C.] in heat exchanger 22 as it cools compressed stream 40e. Heated stream 40a is compressed to 412 psia [2,839 kPa(a)] (stream 40d) in two stages by reflux compressors 23 and 25, with cooling to 142° F. [61° C.] after each stage in coolers 24 and 26. The cooled compressed stream 40e is cooled to −55° F. [−48° C.] (stream 40f) in heat exchanger 22 as described earlier, then further cooled to −119° F. [−84° C.] and substantially condensed (stream 40g) in heat exchanger 12 by heat exchange with cold demethanizer overhead stream 39 as described previously. The cold residue gas stream is warmed (stream 39a) as it provides cooling to compressed distillation vapor stream 40f.
The substantially condensed stream 40g is flash expanded to the operating pressure of demethanizer 19 by expansion valve 14. A portion of the stream is vaporized, further cooling stream 40h to −157° F. [−105° C.] before it is supplied as cold top column feed (reflux) to separator section 19a in the upper region of fractionation tower 19. The liquids separated therein become the top feed to rectifying section 19b and the cold liquid reflux absorbs and condenses the C2 components, C3 components, and heavier components rising in the upper region of rectifying section 19b of demethanizer 19.
Liquid product stream 42 exits the bottom of the tower at 8° F. [−13° C.], based on a typical specification of a methane to ethane ratio of 0.010:1 on a molar basis in the bottom product. It is pumped to a pressure of approximately 455 psia [3,135 kPa(a)] in demethanizer bottoms pump 21, and the pumped liquid product is then warmed to 83° F. [28° C.] as it provides cooling of stream 31b in exchanger 10 before flowing to storage.
The residue gas (demethanizer overhead vapor stream 39) passes countercurrently to the incoming feed gas in heat exchanger 12 where it is heated to −57° F. [−49° C.] (stream 39a) and in heat exchanger 10 where it is heated to 137° F. [58° C.] (stream 39b). The residue gas product (stream 39b) then flows to the fuel gas system at 116 psia [801 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
The feed stream 31 enters compressor 16 driven by expansion machine 15 and is boosted to higher pressure. After cooling to 142° F. [61° C.] in cooler 20, stream 31b at 1166 psia [8,042 kPa(a)] is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 39a) at −53° F. [−47° C.], pumped liquid product (stream 42a) at 11° F. [−11° C.], demethanizer reboiler liquids (stream 41) at −12° F. [−24° C.], and propane refrigerant. The cooled stream 31c enters separator 11 at −14° F. [−25° C.] and 1144 psia [7,884 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 35). The separator liquid (stream 35) is expanded to the operating pressure (approximately 141 psia [972 kPa(a)]) of fractionation tower 19 by expansion valve 18, cooling stream 35a to −64° F. [−53° C.] before it is supplied to fractionation tower 19 at a lower mid-column feed position.
The vapor (stream 32) from separator 11 enters 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 32a to a temperature of approximately −133° F. [−92° C.]. The partially condensed expanded stream 32a is thereafter supplied as feed to fractionation tower 19 at a mid-column feed position.
A portion of the distillation vapor (stream 40) is withdrawn from an intermediate region of fractionation column 19, below the feed position of expanded stream 32a. The distillation vapor stream 40 at −86° F. [−66° C.] is heated to 132° F. [55° C.] in heat exchanger 22 as it cools compressed stream 40e. Heated stream 40a is compressed to 411 psia [2,836 kPa(a)] (stream 40d) in two stages by reflux compressors 23 and 25, with cooling to 142° F. [61° C.] after each stage in coolers 24 and 26. The cooled compressed stream 40e is cooled to −46° F. [−44° C.] (stream 40f) in heat exchanger 22 as described earlier, then further cooled to −141° F. [−96° C.] and substantially condensed (stream 40g) in heat exchanger 12 by heat exchange with cold demethanizer overhead stream 39 as described previously and with distillation liquid stream 43 at −127° F. [−88° C.] which is withdrawn from a region of demethanizer 19 immediately below the feed position of expanded stream 32a. The cold residue gas stream and the distillation liquid stream are warmed as they provide cooling to compressed distillation vapor stream 40f, and the warmed distillation liquid stream 43a returns to demethanizer 19 at −82° F. [−63° C.].
The substantially condensed stream 40g is flash expanded to the operating pressure of demethanizer 19 by expansion valve 14. A portion of the stream is vaporized, further cooling stream 40h to −176° F. [−116° C.] before it is supplied as cold top column feed (reflux) to fractionation tower 19. The cold liquid reflux absorbs and condenses the C2 components, C3 components, and heavier components rising in the upper region of demethanizer 19.
Liquid product stream 42 exits the bottom of the tower at 7° F. [−14° C.], based on a typical specification of a methane to ethane ratio of 0.010:1 on a molar basis in the bottom product. It is pumped to a pressure of approximately 455 psia [3,135 kPa(a)] in demethanizer bottoms pump 21, and the pumped liquid product is then warmed to 83° F. [28° C.] as it provides cooling of stream 31b in exchanger 10 before flowing to storage.
The residue gas (demethanizer overhead vapor stream 39) passes countercurrently to cooled compressed distillation vapor stream 40f in heat exchanger 12 where it is heated to −53° F. [−47° C.] (stream 39a), and to the incoming feed gas in heat exchanger 10 where it is heated to 137° F. [58° C.] (stream 39b). The residue gas product (stream 39b) then flows to the fuel gas system at 116 psia [801 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables I and II shows that, compared to the
The feed stream 31 enters compressor 16 driven by expansion machine 15 and is boosted to higher pressure. After cooling to 142° F. [61° C.] in cooler 20, stream 31b at 1060 psia [7,311 kPa(a)] is cooled in heat exchanger 10 by heat exchange with cold residue gas (stream 39) at −155° F. [−104° C.], pumped liquid product (stream 42a) at 9° F. [−13° C.], demethanizer reboiler liquids (stream 41) at −13° F. [−25° C.], and propane refrigerant. The cooled stream 31c enters separator 11 at −51° F. [−46° C.] and 1037 psia [7,153 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 35).
The vapor (stream 32) from separator 11 enters 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 operating pressure (approximately 141 psia [972 kPa(a)]) of fractionation tower 19, with the work expansion cooling the expanded stream 32a to a temperature of approximately −162° F. [−108° C.]. The partially condensed expanded stream 32a is thereafter supplied as feed to fractionation tower 19 at a mid-column feed position.
The separator liquid (stream 35) is cooled to −141° F. [−96° C.] in heat exchanger 13, and cooled liquid stream 35a is then divided into two streams, stream 36 and stream 37. Stream 37 is expanded to slightly above the operating pressure of fractionation tower 19 by expansion valve 18, cooling stream 37a to −152° F. [−102° C.] before it is heated as it supplies the cooling in heat exchanger 13. The warmed stream 37b at −57° F. [−49° C.] is then supplied to fractionation tower 19 at a lower mid-column feed position.
The remaining portion of cooled liquid stream 35a, stream 36, is flash expanded to the operating pressure of demethanizer 19 by expansion valve 17. A portion of the stream is vaporized, further cooling stream 36a to −162° F. [−108° C.] before it is supplied as cold top column feed (reflux) to fractionation tower 19. The cold liquid reflux absorbs and condenses the C2 components, C3 components, and heavier components rising in the upper region of demethanizer 19.
Liquid product stream 42 exits the bottom of the tower at 4° F. [−15° C.], based on a typical specification of a methane to ethane ratio of 0.010:1 on a molar basis in the bottom product. It is pumped to a pressure of approximately 455 psia [3,135 kPa(a)] in demethanizer bottoms pump 21, and the pumped liquid product is then warmed to 83° F. [28° C.] as it provides cooling of stream 31b in exchanger 10 before flowing to storage.
The residue gas (demethanizer overhead vapor stream 39) passes countercurrently to the incoming feed gas in heat exchanger 10 where it is heated to 136° F. [58° C.] (stream 39a). The residue gas product (stream 39a) then flows to the fuel gas system at 116 psia [801 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables II and III shows that, compared to the
The feed stream 31 enters compressor 16 driven by expansion machine 15 and is boosted to higher pressure. After cooling to 142° F. [61° C.] in cooler 20, stream 31b at 1194 psia [8,234 kPa(a)] is cooled in heat exchanger 10 by heat exchange with cool residue gas (stream 39a) at −43° F. [−42° C.], pumped liquid product (stream 42a) at 8° F. [−13° C.], demethanizer reboiler liquids (stream 41) at −10° F. [−23° C.], and propane refrigerant. The cooled stream 31c enters separator 11 at −6° F. [−21° C.] and 1171 psia [8,075 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 35).
The vapor (stream 32) from separator 11 enters 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 operating pressure (approximately 141 psia [972 kPa(a)]) of fractionation tower 19, with the work expansion cooling the expanded stream 32a to a temperature of approximately −127° F. [−88° C.]. The partially condensed expanded stream 32a is thereafter supplied as feed to fractionation tower 19 at a mid-column feed position below the rectifying section and above the demethanizing section in fractionation tower 19.
The separator liquid (stream 35) is cooled to −132° F. [−91° C.] in heat exchanger 13, and cooled liquid stream 35a is then divided into two streams, stream 36 and stream 37. Stream 37 is expanded to slightly above the operating pressure of fractionation tower 19 by expansion valve 18, cooling stream 37a to −142° F. [−97° C.] before it is heated as it supplies the cooling in heat exchanger 13. The warmed stream 37b at −17° F. [−27° C.] is then supplied to fractionation tower 19 at a lower mid-column feed position.
The remaining portion of cooled liquid stream 35a, stream 36, is flash expanded to the operating pressure of demethanizer 19 by expansion valve 17. A portion of the stream is vaporized, further cooling stream 36a to −143° F. [−97° C.] before it is supplied to fractionation tower 19 at an upper mid-column feed position above expanded stream 32a. The cold liquid portion of stream 36a absorbs and condenses the C2 components, C3 components, and heavier components rising in the lower region of the rectifying section in demethanizer 19.
A portion of the distillation vapor (stream 40) is withdrawn from an upper region of the demethanizing section in fractionation column 19, below the feed position of expanded stream 32a. The distillation vapor stream 40 at −75° F. [−60° C.] is heated to 132° F. [55° C.] in heat exchanger 22 as it cools compressed stream 40e. Heated stream 40a is compressed to 412 psia [2,839 kPa(a)] (stream 40d) in two stages by reflux compressors 23 and 25, with cooling to 142° F. [61° C.] after each stage in coolers 24 and 26. The cooled compressed stream 40e is cooled to −38° F. [−39° C.] (stream 40f) in heat exchanger 22 as described earlier, then further cooled to −175° F. [−115° C.] and substantially condensed (stream 40g) in heat exchanger 12 by heat exchange with cold demethanizer overhead stream 39 as described previously and with distillation liquid stream 43 at −125° F. [−87° C.] which is withdrawn from a region of demethanizer 19 immediately below the feed position of expanded stream 32a. The cold residue gas stream and the distillation liquid stream are warmed as they provide cooling to compressed distillation vapor stream 40f, and the warmed distillation liquid stream 43a returns to demethanizer 19 at −100° F. [−73° C.]
The substantially condensed stream 40g is flash expanded to the operating pressure of demethanizer 19 by expansion valve 14. A portion of the stream is vaporized, further cooling stream 40h to −191° F. [−124° C.] before it is supplied as cold top column feed (reflux) to the separator section in fractionation tower 19. The liquids separated therein become the top feed to the rectifying section in fractionation tower 19 and the cold liquid reflux absorbs and condenses the C2 components, C3 components, and heavier components rising in the upper region of the rectifying section in demethanizer 19.
Liquid product stream 42 exits the bottom of the tower at 3° F. [−16° C.], based on a typical specification of a methane to ethane ratio of 0.010:1 on a molar basis in the bottom product as regulated by control means 29. It is pumped to a pressure of approximately 455 psia [3,135 kPa(a)] in demethanizer bottoms pump 21, and the pumped liquid product is then warmed to 83° F. [28° C.] as it provides cooling of stream 31b in exchanger 10 before flowing to storage.
The residue gas (demethanizer overhead vapor stream 39) passes countercurrently to cooled compressed distillation vapor stream 40f in heat exchanger 12 where it is heated to −43° F. [−42° C.] (stream 39a), and to the incoming feed gas in heat exchanger 10 where it is heated to 136° F. [58° C.] (stream 39b). The residue gas product (stream 39b) then flows to the fuel gas system at 116 psia [801 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables I, II, III, and IV shows that, compared to the prior art processes, the present invention significantly improves C2 component recovery (from 81.68%, 84.81%, and 93.68%, respectively, to 96.37%), improves C3 component recovery (from 99.95%, 99.97%, and 99.56%, respectively, to 99.99%), and maintains the same C4+ component recovery. Comparison of Tables I, II, III, and IV further shows that these yields were achieved using slightly less power than the prior art processes. In terms of the recovery efficiency (defined by the quantity of ethane recovered per unit of power), the present invention represents a 19%, 15%, and 4% improvement, respectively, over the prior art of the
The superior C2 component recovery performance of the present invention compared to that of the prior art processes is most easily understood by examining the feed streams supplied to the rectifying section in demethanizer 19. As explained in the BACKGROUND OF THE INVENTION, the goal is to provide reflux stream(s) capable of capturing the desired C2+ components rising from below (most of which originate in the expanded stream supplied to the mid-column feed position of demethanizer 19, stream 34a in
The
Both the
In the present invention in
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 the expanded substantially condensed stream (stream 40h) leaving expansion valve 14 and all or a part of the expanded cooled liquid stream 36a 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 32a, shall be considered for the purposes of this invention as constituting an absorbing section.
The present invention provides improved recovery of C2 components, C3 components, and heavier hydrocarbon components per amount of utility consumption required to operate the process. An improvement may also be effected in lower utility consumption required for operating the process, which may appear in the form of reduced power requirements for compression or re-compression, reduced power requirements for external refrigeration, reduced energy requirements for supplemental heating, or a combination thereof.
While there have been described what are believed to be preferred embodiments of the invention, those skilled in the art will recognize that other and further modifications may be made thereto, e.g. to adapt the invention to various conditions, types of feed, or other requirements without departing from the spirit of the present invention as defined by the following claims.
This invention relates to a process and apparatus for the separation of a gas containing hydrocarbons and significant quantities of components more volatile than methane (e.g., hydrogen, nitrogen, etc.). The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application No. 62/923,067 which was filed on Oct. 18, 2019.
| Number | Date | Country | |
|---|---|---|---|
| 62923067 | Oct 2019 | US |
