Patent Grant
6889523
This invention relates to a process for processing natural gas to produce liquefied natural gas (LNG) that has a high methane purity. In particular, this invention is well suited to co-production of LNG by integration into natural gas processing plants that recover natural gas liquids (NGL) and/or liquefied petroleum gas (LPG) using a cryogenic process.
Natural gas is typically recovered from wells drilled into underground reservoirs. It usually has a major proportion of methane, i.e., methane comprises at least 50 mole percent of the gas. Depending on the particular underground reservoir, the natural gas also contains relatively lesser amounts of heavier hydrocarbons such as ethane, propane, butanes, pentanes and the like, as well as water, hydrogen, nitrogen, carbon dioxide, and other gases.
Most natural gas is handled in gaseous form. The most common means for transporting natural gas from the wellhead to gas processing plants and thence to the natural gas consumers is in high pressure gas transmission pipelines. In a number of circumstances, however, it has been found necessary and/or desirable to liquefy the natural gas either for transport or for use. In remote locations, for instance, there is often no pipeline infrastructure that would allow for convenient transportation of the natural gas to market. In such cases, the much lower specific volume of LNG relative to natural gas in the gaseous state can greatly reduce transportation costs by allowing delivery of the LNG using cargo ships and transport trucks.
Another circumstance that favors the liquefaction of natural gas is for its use as a motor vehicle fuel. In large metropolitan areas, there are fleets of buses, taxi cabs, and trucks that could be powered by LNG if there was an economic source of LNG available. Such LNG-fueled vehicles produce considerably less air pollution due to the clean-burning nature of natural gas when compared to similar vehicles powered by gasoline and diesel engines which combust higher molecular weight hydrocarbons. In addition, if the LNG is of high purity (i.e., with a methane purity of 95 mole percent or higher), the amount of carbon dioxide (a “greenhouse gas”) produced is considerably less due to the lower carbon:hydrogen ratio for methane compared to all other hydrocarbon fuels.
The present invention is generally concerned with the liquefaction of natural gas as a co-product in a cryogenic gas processing plant that also produces natural gas liquids (NGL) such as ethane, propane, butanes, and heavier hydrocarbon components. A typical analysis of a natural gas stream to be processed in accordance with this invention would be, in approximate mole percent, 92.3% methane, 4.4% ethane and other C2 components, 1.5% propane and other C3 components, 0.3% iso-butane, 0.3% normal butane, 0.3% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present.
There are a number of methods known for liquefying natural gas. For instance, see Finn, Adrian J., Grant L. Johnson, and Terry R. Tomlinson, “LNG Technology for Offshore and Mid-Scale Plants”, Proceedings of the Seventy-Ninth Annual Convention of the Gas Processors Association, pp. 429-450, Atlanta, Ga., Mar. 13-15, 2000 and Kikkawa, Yoshitsugi, Masaaki Ohishi, and Noriyoshi Nozawa, “Optimize the Power System of Baseload LNG Plant”, Proceedings of the Eightieth Annual Convention of the Gas Processors Association, San Antonio, Tex., Mar. 12-14, 2001 for surveys of a number of such processes. U.S. Pat. Nos. 4,445,917; 4,525,185; 4,545,795; 4,755,200; 5,291,736; 5,363,655; 5,365,740; 5,600,969; 5,615,561; 5,651,269; 5,755,114; 5,893,274; 6,014,869; 6,053,007; 6,062,041; 6,119,479; 6,125,653; 6,250,105 B1; 6,269,655 B1; 6,272,882 B1; 6,308,531 B1; 6,324,867 B1; 6,347,532 B1; International Publication Number WO 01/88447 A1 published Nov. 22, 2001; our co-pending U.S. patent application Ser. No. 09/839,907 filed Apr. 20, 2001; our co-pending U.S. patent application Ser. No. 10/161,780 filed Jun. 4, 2002; and our co-pending U.S. patent application Ser. No. 10/278,610 filed Oct. 23, 2002 also describe relevant processes. These methods generally include steps in which the natural gas is purified (by removing water and troublesome compounds such as carbon dioxide and sulfur compounds), cooled, condensed, and expanded. Cooling and condensation of the natural gas can be accomplished in many different manners. “Cascade refrigeration” employs heat exchange of the natural gas with several refrigerants having successively lower boiling points, such as propane, ethane, and methane. As an alternative, this heat exchange can be accomplished using a single refrigerant by evaporating the refrigerant at several different pressure levels. “Multi-component refrigeration” employs heat exchange of the natural gas with one or more refrigerant fluids composed of several refrigerant components in lieu of multiple single-component refrigerants. Expansion of the natural gas can be accomplished both isenthalpically (using Joule-Thomson expansion, for instance) and isentropically (using a work-expansion turbine, for instance).
While any of these methods could be employed to produce vehicular grade LNG, the capital and operating costs associated with these methods have generally made the installation of such facilities uneconomical. For instance, the purification steps required to remove water, carbon dioxide, sulfur compounds, etc. from the natural gas prior to liquefaction represent considerable capital and operating costs in such facilities, as do the drivers for the refrigeration cycles employed. This has led the inventors to investigate the feasibility of integrating LNG production into cryogenic gas processing plants used to recover NGL from natural gas. Such an integrated LNG production method would eliminate the need for separate gas purification facilities and gas compression drivers. Further, the potential for integrating the cooling/condensation for the LNG liquefaction with the process cooling required for NGL recovery could lead to significant efficiency improvements in the LNG liquefaction method.
In accordance with the present invention, it has been found that LNG with a methane purity in excess of 99 percent can be co-produced from a cryogenic NGL recovery plant without reducing the NGL recovery level using less energy than prior art processes. 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 International System of Units (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 LNG production rates reported as gallons per day (gallons/D) and/or pounds per hour (Lbs/hour) correspond to the stated molar flow rates in pound moles per hour. The LNG production rates reported as cubic meters per day (m3/D) and/or kilograms per hour (kg/H) correspond to the stated molar flow rates in kilogram moles per hour.
Referring now to
The feed stream 31 is cooled in heat exchanger 10 by heat exchange with cool demethanizer overhead vapor at −66° F. [−55° C.] (stream 36a), bottom liquid product at 56° F. [13° C.] (stream 41a) from demethanizer bottoms pump 18, demethanizer reboiler liquids at 36° F. [2° C.] (stream 40), and demethanizer side reboiler liquids at −35° F. [−37° C.] (stream 39). Note that in all cases heat 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 −43° F. [−42° C.] and 725 psia [4,999 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 35).
The vapor (stream 32) from separator 11 is divided into two streams, 33 and 34. Stream 33, containing about 27% of the total vapor, passes through heat exchanger 12 in heat exchange relation with the demethanizer overhead vapor stream 36, resulting in cooling and substantial condensation of stream 33a. The substantially condensed stream 33a at −142° F. [−97° C.] is then flash expanded through an appropriate expansion device, such as expansion valve 13, to the operating pressure (approximately 320 psia [2,206 kPa(a)]) of fractionation tower 17. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
The remaining 73% of the vapor from separator 11 (stream 34) 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 from a pressure of about 725 psia [4,999 kPa(a)] to the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −107° F. [−77° 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 15) that can be used to re-compress the residue gas (stream 38), for example. The expanded and partially condensed stream 34a is supplied as a feed to the distillation column at an intermediate point. The separator liquid (stream 35) is likewise expanded to the tower operating pressure by expansion valve 16, cooling stream 35a to −72° F. [−58° C.] before it is supplied to the demethanizer in fractionation tower 17 at a lower mid-column feed point.
The demethanizer in fractionation 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 of the top feed to form the cold demethanizer overhead vapor (stream 36) which exits the top of the tower at −150° F. [−101° 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 also includes reboilers which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column.
The liquid product stream 41 exits the bottom of the tower at 51° F. [10° C.], based on a typical specification of a methane to ethane ratio of 0.028:1 on a molar basis in the bottom product. The stream is pumped to approximately 650 psia [4,482 kPa(a)] (stream 41a) in pump 18. Stream 41a, now at about 56° F. [13° C.], is warmed to 85° F. [29° C.] (stream 41b) in heat exchanger 10 as it provides cooling to stream 31. (The discharge pressure of the pump is usually set by the ultimate destination of the liquid product. Generally the liquid product flows to storage and the pump discharge pressure is set so as to prevent any vaporization of stream 41b as it is warmed in heat exchanger 10.)
The demethanizer overhead vapor (stream 36) passes countercurrently to the incoming feed gas in heat exchanger 12 where it is heated to −66° F. [−55° C.] (stream 36a) and heat exchanger 10 where it is heated to 68° F. [20° C.] (stream 36b). A portion of the warmed demethanizer overhead vapor is withdrawn to serve as fuel gas (stream 37) for the plant, with the remainder becoming the residue gas (stream 38). (The amount of fuel gas that must be withdrawn is largely determined by the fuel required for the engines and/or turbines driving the gas compressors in the plant, such as compressor 19 in this example.) The residue gas is re-compressed in two stages. The first stage is compressor 15 driven by expansion machine 14. The second stage is compressor 19 driven by a supplemental power source which compresses the residue gas (stream 38b) to sales line pressure. After cooling to 120° F. [49° C.] in discharge cooler 20, the residue gas product (stream 38c) flows to the sales gas pipeline at 740 psia [5,102 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
In the simulation of the
The inlet gas to the NGL recovery plant (stream 31) was not treated for carbon dioxide removal prior to processing. Although the carbon dioxide concentration in the inlet gas (about 0.5 mole percent) will not create any operating problems for the NGL recovery plant, a significant fraction of this carbon dioxide will leave the plant in the demethanizer overhead vapor (stream 36) and will subsequently contaminate the feed stream for the LNG production section (stream 71). The carbon dioxide concentration in this stream is about 0.4 mole percent, well in excess of the concentration that can be tolerated by this prior art process (about 0.005 mole percent). Accordingly, the feed stream 71 must be processed in carbon dioxide removal section 50 before entering the LNG production section to avoid operating problems from carbon dioxide freezing. Although there are many different processes that can be used for carbon dioxide removal, many of them will cause the treated gas stream to become partially or completely saturated with water. Since water in the feed stream would also lead to freezing problems in the LNG production section, it is very likely that the carbon dioxide removal section 50 must also include dehydration of the gas stream after treating.
The treated feed gas enters the LNG production section at 120° F. [49° C.] and 730 psia [5,033 kPa(a)] as stream 72 and is cooled in heat exchanger 51 by heat exchange with a refrigerant mixture at −261° F. [−163° C.] (stream 74b). The purpose of heat exchanger 51 is to cool the feed stream to substantial condensation and, preferably, to subcool the stream so as to eliminate any flash vapor being generated in the subsequent expansion step. For the conditions stated, however, the feed stream pressure is above the cricondenbar, so no liquid will condense as the stream is cooled. Instead, the cooled stream 72a leaves heat exchanger 51 at −256° F. [−160° C.] as a dense-phase fluid. (The cricondenbar is the maximum pressure at which a vapor phase can exist in a multi-phase fluid. At pressures below the cricondenbar, stream 72a would typically exit heat exchanger 51 as a subcooled liquid stream.)
Stream 72a enters a work expansion machine 52 in which mechanical energy is extracted from this high pressure stream. The machine 52 expands the dense-phase fluid substantially isentropically from a pressure of about 728 psia [5,019 kPa(a)] to the LNG storage pressure (18 psia [124 kPa(a)]), slightly above atmospheric pressure. The work expansion cools the expanded stream 72b to a temperature of approximately −257° F. [−160° C.], whereupon it is then directed to the LNG storage tank 53 which holds the LNG product (stream 73).
All of the cooling for stream 72 is provided by a closed cycle refrigeration loop. The working fluid for this cycle is a mixture of hydrocarbons and nitrogen, with the composition of the mixture adjusted as needed to provide the required refrigerant temperature while condensing at a reasonable pressure using the available cooling medium. In this case, condensing with ambient air has been assumed, so a refrigerant mixture composed of nitrogen, methane, ethane, propane, and heavier hydrocarbons is used in the simulation of the
The refrigerant stream 74 leaves partial condenser 56 at 120° F. [49° C.] and 140 psia [965 kPa(a)]. It enters heat exchanger 51 and is condensed and then subcooled to −256° F. [−160° C.] by the flashed refrigerant stream 74b. The subcooled liquid stream 74a is flash expanded substantially isenthalpically in expansion valve 54 from about 138 psia [951 kPa(a)] to about 26 psia [179 kPa(a)]. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream to −261° F. [−163° C.] (stream 74b). The flash expanded stream 74b then reenters heat exchanger 51 where it provides cooling to the feed gas (stream 72) and the refrigerant (stream 74) as it is vaporized and superheated.
The superheated refrigerant vapor (stream 74c) leaves heat exchanger 51 at 110° F. [43° C.] and flows to refrigerant compressor 55, driven by a supplemental power source. Compressor 55 compresses the refrigerant to 145 psia [1,000 kPa(a)], whereupon the compressed stream 74d returns to partial condenser 56 to complete the cycle.
A summary of stream flow rates and energy consumption for the process illustrated in
As stated earlier, the NGL recovery plant operates exactly the same in the
The net increase in compression power for the
For this adaptation of the prior art LNG production process where the NGL recovery plant residue gas is used as the source of feed gas for LNG production, no provisions have been included for removing heavier hydrocarbons from the LNG feed gas. Consequently, all of the heavier hydrocarbons present in the feed gas become part of the LNG product, reducing the purity (i.e., methane concentration) of the LNG product. If higher LNG purity is desired, or if the source of feed gas contains higher concentrations of heavier hydrocarbons (inlet gas stream 31, for instance), the feed stream 72 would need to be withdrawn from heat exchanger 51 after cooling to an intermediate temperature so that condensed liquid could be separated, with the uncondensed vapor thereafter returned to heat exchanger 51 for cooling to the final outlet temperature. These condensed liquids would preferentially contain the majority of the heavier hydrocarbons, along with a considerable fraction of liquid methane, which could then be re-vaporized and used to supply a part of the plant fuel gas requirements. Unfortunately, this means that the C2 components, C3 components, and heavier hydrocarbon components removed from the LNG feed stream would not be recovered in the NGL product from the NGL recovery plant, and their value as liquid products would be lost to the plant operator. Further, for feed streams such as the one considered in this example, condensation of liquid from the feed stream may not be possible due to the process operating conditions (i.e., operating at pressures above the cricondenbar of the stream), meaning that removal of heavier hydrocarbons could not be accomplished in such instances.
The process of
In the simulation of the
The vapor (stream 32) from separator 11 is divided into two streams, 33 and 34. Stream 33, containing about 25% of the total vapor, passes through heat exchanger 12 in heat exchange relation with the cold demethanizer overhead vapor stream 36a where it is cooled to −142° F. [−97° C.]. The resulting substantially condensed stream 33a is then flash expanded through expansion valve 13 to the operating pressure (approximately 291 psia [2,006 kPa(a)]) of fractionation tower 17. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
The remaining 75% of the vapor from separator 11 (stream 34) 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 from a pressure of about 725 psia [4,999 kPa(a)] to the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −116° F. [−82° C.]. The expanded and partially condensed stream 34a is thereafter supplied as a feed to fractionation tower 17 at an intermediate point. The separator liquid (stream 35) is likewise expanded to the tower operating pressure by expansion valve 16, cooling stream 35a to −80° F. [−62° C.] before it is supplied to fractionation tower 17 at a lower mid-column feed point.
The liquid product (stream 41) exits the bottom of tower 17 at 42° F. [6° C.]. This stream is pumped to approximately 650 psia [4,482 kPa(a)] (stream 41a) in pump 18 and warmed to 83° F. [28° C.] (stream 41b) in heat exchanger 10 as it provides cooling to stream 31. The distillation vapor stream forming the tower overhead (stream 36) leaves demethanizer 17 at −154° F. [−103° C.] and is divided into two portions. One portion (stream 43) is directed to heat exchanger 51 in the LNG production section to provide most of the cooling duty in this exchanger as it is warmed to −42° F. [−41° C.] (stream 43a). The remaining portion (stream 42) bypasses heat exchanger 51, with control valve 21 adjusting the quantity of this bypass in order to regulate the cooling accomplished in heat exchanger 51. The two portions recombine at −146° F. [−99° C.] to form stream 36a, which passes countercurrently to the incoming feed gas in heat exchanger 12 where it is heated to −69° F. [−56° C.] (stream 36b) and heat exchanger 10 where it is heated to 72° F. [22° C.] (stream 36c). Stream 36c combines with warm HP flash vapor (stream 73a) from the LNG production section, forming stream 44 at 72° F. [22° C.]. A portion of this stream is withdrawn (stream 37) to serve as part of the fuel gas for the plant. The remainder (stream 45) is re-compressed in two stages, compressor 15 driven by expansion machine 14 and compressor 19 driven by a supplemental power source, and cooled to 120° F. [49° C.] in discharge cooler 20. The cooled compressed stream (stream 45c) is then divided into two portions. One portion is the residue gas product (stream 38), which flows to the sales gas pipeline at 740 psia [5,102 kPa(a)]. The other portion (stream 71) is the feed stream for the LNG production section.
The inlet gas to the NGL recovery plant (stream 31) was not treated for carbon dioxide removal prior to processing. Although the carbon dioxide concentration in the inlet gas (about 0.5 mole percent) will not create any operating problems for the NGL recovery plant, a significant fraction of this carbon dioxide will leave the plant in the demethanizer overhead vapor (stream 36) and will subsequently contaminate the feed stream for the LNG production section (stream 71). The carbon dioxide concentration in this stream is about 0.4 mole percent, well in excess of the concentration that can be tolerated by this prior art process (0.005 mole percent). As for the
The treated feed gas enters the LNG production section at 120° F. [49° C.] and 730 psia [5,033 kPa(a)] as stream 72 and is cooled in heat exchanger 51 by heat exchange with LP flash vapor at −200° F. [−129° C.] (stream 75), HP flash vapor at −164° F. [−109° C.] (stream 73), and a portion of the demethanizer overhead vapor (stream 43) at −154° F. [−103° C.] from the NGL recovery plant. The purpose of heat exchanger 51 is to cool the LNG feed stream 72 to substantial condensation, and preferably to subcool the stream so as to reduce the quantity of flash vapor generated in subsequent expansion steps in the LNG cool-down section. For the conditions stated, however, the feed stream pressure is above the cricondenbar, so no liquid will condense as the stream is cooled. Instead, the cooled stream 72a leaves heat exchanger 51 at −148° F. [−100° C.] as a dense-phase fluid. At pressures below the cricondenbar, stream 72a would typically exit heat exchanger 51 as a condensed (and preferably subcooled) liquid stream.
Stream 72a is flash expanded substantially isenthalpically in expansion valve 52 from about 727 psia [5,012 kPa(a)] to the operating pressure of HP flash drum 53, about 279 psia [1,924 kPa(a)]. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream to −164° F. [−109° C.] (stream 72b). The flash expanded stream 72b then enters HP flash drum 53 where the HP flash vapor (stream 73) is separated and directed to heat exchanger 51 as described previously. The operating pressure of the HP flash drum is set so that the heated HP flash vapor (stream 73a) leaving heat exchanger 51 is at sufficient pressure to allow it to join the heated demethanizer overhead vapor (stream 36c) leaving the NGL recovery plant and subsequently be compressed by compressors 15 and 19 after withdrawal of a portion (stream 37) to serve as part of the fuel gas for the plant.
The HP flash liquid (stream 74) from HP flash drum 53 is flash expanded substantially isenthalpically in expansion valve 54 from the operating pressure of the HP flash drum to the operating pressure of LP flash drum 55, about 118 psia [814 kPa(a)]. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream to −200° F. [−129° C.] (stream 74a). The flash expanded stream 74a then enters LP flash drum 55 where the LP flash vapor (stream 75) is separated and directed to heat exchanger 51 as described previously. The operating pressure of the LP flash drum is set so that the heated LP flash vapor (stream 75a) leaving heat exchanger 51 is at sufficient pressure to allow its use as plant fuel gas.
The LP flash liquid (stream 76) from LP flash drum 55 is flash expanded substantially isenthalpically in expansion valve 56 from the operating pressure of the LP flash drum to the LNG storage pressure (18 psia [124 kPa(a)]), slightly above atmospheric pressure. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream to −254° F. [−159° C.] (stream 76a), whereupon it is then directed to LNG storage tank 57 where the flash vapor resulting from expansion (stream 77) is separated from the LNG product (stream 78).
The flash vapor (stream 77) from LNG storage tank 57 is at too low a pressure to be used for plant fuel gas, and is too cold to enter directly into a compressor. Accordingly, it is first heated to −30° F. [−34° C.] (stream 77a) in heater 58, then compressors 59 and 60 (driven by supplemental power sources) are used to compress the stream (stream 77c). Following cooling in aftercooler 61, stream 77d at 115 psia [793 kPa(a)] is combined with streams 37 and 75a to become the fuel gas for the plant (stream 79).
A summary of stream flow rates and energy consumption for the process illustrated in
The process of
As can be seen by comparing Tables I and III, the plant fuel gas consumption is higher for the
The
In the simulation of the
The vapor (stream 32) from separator 11 is divided into two streams, 33 and 34. Stream 33, containing about 26% of the total vapor, passes through heat exchanger 12 in heat exchange relation with the cold distillation vapor stream 42 where it is cooled to −146° F. [−99° C.]. The resulting substantially condensed stream 33a is then flash expanded through expansion valve 13 to the operating pressure (approximately 306 psia [2,110 kPa(a)]) of fractionation tower 17. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
The remaining 74% of the vapor from separator 11 (stream 34) 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 from a pressure of about 725 psia [4,999 kPa(a)] to the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −110° F. [−79° C.]. The expanded and partially condensed stream 34a is thereafter supplied as a feed to fractionation tower 17 at an intermediate point. The separator liquid (stream 35) is likewise expanded to the tower operating pressure by expansion valve 16, cooling stream 35a to −75° F. [−59° C.] before it is supplied to fractionation tower 17 at a lower mid-column feed point.
The liquid product (stream 41) exits the bottom of tower 17 at 47° F. [8° C.]. This stream is pumped to approximately 650 psia [4,482 kPa(a)] (stream 41a) in pump 18 and warmed to 83° F. [28° C.] (stream 41b) in heat exchanger 10 as it provides cooling to stream 31. The distillation vapor stream forming the tower overhead at −151° F. [−102° C.] (stream 36) is divided into two portions. One portion (stream 43) is directed to the LNG production section. The remaining portion (stream 42) passes countercurrently to the incoming feed gas in heat exchanger 12 where it is heated to −66° F. [−55° C.] (stream 42a) and heat exchanger 10 where it is heated to 72° F. [22° C.] (stream 42b). A portion of the warmed distillation vapor stream is withdrawn (stream 37) to serve as part of the fuel gas for the plant, with the remainder becoming the first residue gas (stream 44). The first residue gas is then re-compressed in two stages, compressor 15 driven by expansion machine 14 and compressor 19 driven by a supplemental power source to form the compressed first residue gas (stream 44b).
Turning now to the LNG production section, feed stream 71 enters heat exchanger 51 at 120° F. [49° C.] and 740 psia [5,102 kPa(a)]. The feed stream 71 is cooled to −120° F. [−84° C.] in heat exchanger 51 by heat exchange with cool LNG flash vapor (stream 83a), the distillation vapor stream from the NGL recovery plant at −151° F. [−102° C.] (stream 43), flash liquids (stream 80), and distillation column reboiler liquids at −142° F. [−97° C.] (stream 76). (For the conditions stated, the feed stream pressure is above the cricondenbar, so no liquid will condense as the stream is cooled. Instead, the cooled stream 71a leaves heat exchanger 51 as a dense-phase fluid. For other processing conditions, it is possible that the feed gas pressure will be below its cricondenbar pressure, in which case the feed stream will be cooled to substantial condensation.) The resulting cooled stream 71a is then flash expanded through an appropriate expansion device, such as expansion valve 52, to the operating pressure (420 psia [2,896 kPa(a)]) of distillation column 56. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
Distillation column 56 serves as an LNG purification tower, recovering nearly all of the carbon dioxide and the hydrocarbons heavier than methane present in its feed stream (stream 71b) as its bottom product (stream 77) so that the only significant impurity in its overhead (stream 74) is the nitrogen contained in the feed stream. Reflux for distillation column 56 is created by cooling and condensing the tower overhead vapor (stream 74 at −144° F. [−98° C.]) in heat exchanger 51 by heat exchange with cool LNG flash vapor at −155° F. [−104° C.] (stream 83a) and flash liquids at −157° F. [−105° C.] (stream 80). The condensed stream 74a, now at −146° F. [−99° C.], is divided into two portions. One portion (stream 78) becomes the feed to the LNG cool-down section. The other portion (stream 75) enters reflux pump 55. After pumping, stream 75a at −145° F. [−98° C.] is supplied to LNG purification tower 56 at a top feed point to provide the reflux liquid for the tower. This reflux liquid rectifies the vapors rising up the tower so that the tower overhead (stream 74) and consequently feed stream 78 to the LNG cool-down section contain minimal amounts of carbon dioxide and hydrocarbons heavier than methane.
The feed stream for the LNG cool-down section (condensed liquid stream 78) enters heat exchanger 58 at −146° F. [−99° C.] and is subcooled by heat exchange with cold LNG flash vapor at −255° F. [−159° C.] (stream 83) and cold flash liquids (stream 79a). The cold flash liquids are produced by withdrawing a portion of the partially subcooled feed stream (stream 79) from heat exchanger 58 and flash expanding the stream through an appropriate expansion device, such as expansion valve 59, to slightly above the operating pressure of fractionation tower 17. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream from −156° F. [−104° C.] to −160° F. [−106° C.] (stream 79a). The flash expanded stream 79a is then supplied to heat exchanger 58 as previously described.
The remaining portion of the partially subcooled feed stream is further subcooled in heat exchanger 58 to −169° F. [112° C.] (stream 82). It then enters a work expansion machine 60 in which mechanical energy is extracted from this intermediate pressure stream. The machine 60 expands the subcooled liquid substantially isentropically from a pressure of about 414 psia [2,854 kPa(a)] to the LNG storage pressure (18 psia [124 kPa(a)]), slightly above atmospheric pressure. The work expansion cools the expanded stream 82a to a temperature of approximately −255° F. [−159° C.], whereupon it is then directed to LNG storage tank 61 where the flash vapor resulting from expansion (stream 83) is separated from the LNG product (stream 84).
Tower bottoms stream 77 from LNG purification tower 56 is flash expanded to slightly above the operating pressure of fractionation tower 17 by expansion valve 57. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream from −141° F. [−96° C.] to −156° F. [−105° C.] (stream 77a). The flash expanded stream 77a is then combined with warmed flash liquid stream 79b leaving heat exchanger 58 at −155° F. [−104° C.] to form a combined flash liquid stream (stream 80) at −157° F. [−105° C.] which is supplied to heat exchanger 51. It is heated to −90° F. [−68° C.] (stream 80a) as it supplies cooling to LNG feed stream 71 and tower overhead vapor stream 74 as described earlier, and thereafter supplied to fractionation tower 17 at a lower mid-column feed point.
The flash vapor (stream 83) from LNG storage tank 61 passes countercurrently to the incoming liquid in heat exchanger 58 where it is heated to −155° F. [−104° C.] (stream 83a). It then enters heat exchanger 51 where it is heated to 115° F. [46° C.] (stream 83b) as it supplies cooling to LNG feed stream 71 and tower overhead stream 74. Since this stream is at low pressure (15.5 psia [107 kPa(a)]), it must be compressed before it can be used as plant fuel gas. Compressors 63 and 65 (driven by supplemental power sources) with intercooler 64 are used to compress the stream (stream 83e). Following cooling in aftercooler 66, stream 83f at 115 psia [793 kPa(a)] is combined with stream 37 to become the fuel gas for the plant (stream 85).
The cold distillation vapor stream from the NGL recovery plant (stream 43) is heated to 115° F. [46° C.] as it supplies cooling to LNG feed stream 71 in heat exchanger 51, becoming the second residue gas (stream 43a) which is then re-compressed in compressor 62 driven by a supplemental power source. The compressed second residue gas (stream 43b) combines with the compressed first residue gas (stream 44b) to form third residue gas stream 45. After cooling to 120° F. [49° C.] in discharge cooler 20, third residue gas stream 45a is divided into two portions. One portion (stream 71) becomes the feed stream to the LNG production section. The other portion (stream 38) becomes the residue gas product, which flows to the sales gas pipeline at 740 psia [5,102 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
Comparing the recovery levels displayed in Table IV for the
In the simulation of the
The vapor (stream 32) from separator 11 is divided into two streams, 33 and 34. Stream 33, containing about 26% of the total vapor, passes through heat exchanger 12 in heat exchange relation with the cold distillation vapor stream 42 where it is cooled to −146° F. [−99° C.]. The resulting substantially condensed stream 33a is then flash expanded through expansion valve 13 to the operating pressure (approximately 306 psia [2,110 kPa(a)]) of fractionation tower 17. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
The remaining 74% of the vapor from separator 11 (stream 34) 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 from a pressure of about 725 psia [4,999 kPa(a)] to the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −110° F. [−79° C.]. The expanded and partially condensed stream 34a is thereafter supplied as a feed to fractionation tower 17 at an intermediate point. The separator liquid (stream 35) is likewise expanded to the tower operating pressure by expansion valve 16, cooling stream 35a to −75° F. [−59° C.] before it is supplied to fractionation tower 17 at a lower mid-column feed point.
The liquid product (stream 41) exits the bottom of tower 17 at 47° F. [9° C.]. This stream is pumped to approximately 650 psia [4,482 kPa(a)] (stream 41a) in pump 18 and warmed to 83° F. [28° C.] (stream 41b) in heat exchanger 10 as it provides cooling to stream 31. The distillation vapor stream forming the tower overhead at −152° F. [−102° C.] (stream 36) is divided into two portions. One portion (stream 43) is directed to the LNG production section. The remaining portion (stream 42) passes countercurrently to the incoming feed gas in heat exchanger 12 where it is heated to −66° F. [−55° C.] (stream 42a) and heat exchanger 10 where it is heated to 72° F. [22° C.] (stream 42b). A portion of the warmed distillation vapor stream is withdrawn (stream 37) to serve as part of the fuel gas for the plant, with the remainder becoming the first residue gas (stream 44). The first residue gas is then re-compressed in two stages, compressor 15 driven by expansion machine 14 and compressor 19 driven by a supplemental power source to form the compressed first residue gas (stream 44b).
The inlet gas to the NGL recovery plant (stream 31) was not treated for carbon dioxide removal prior to processing. Although the carbon dioxide concentration in the inlet gas (about 0.5 mole percent) will not create any operating problems for the NGL recovery plant, a significant fraction of this carbon dioxide will leave the plant in the demethanizer overhead vapor (stream 36) and will subsequently contaminate the feed stream for the LNG production section (stream 71). The carbon dioxide concentration in this stream is about 0.4 mole percent, in excess of the concentration that can be tolerated by the present invention for the
Treated feed stream 72 enters heat exchanger 51 at 120° F. [49° C.] and 730 psia [5,033 kPa(a)]. Note that in all cases heat exchanger 51 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, feed stream flow rate, heat exchanger size, stream temperatures, etc.) The feed stream 72 is cooled to −120° F. [−84° C.] in heat exchanger 51 by heat exchange with cool LNG flash vapor (stream 83a), the distillation vapor stream from the NGL recovery plant at −152° F. [−102° C.] (stream 43), and flash liquids (stream 79b). (For the conditions stated, the feed stream pressure is above the cricondenbar, so no liquid will condense as the stream is cooled. Instead, the cooled stream 72a leaves heat exchanger 51 as a dense-phase fluid. For other processing conditions, it is possible that the feed gas pressure will be below its cricondenbar pressure, in which case the feed stream will be cooled to substantial condensation.)
The feed stream for the LNG cool-down section (dense-phase stream 72a) enters heat exchanger 58 at −120° F. [−84° C.] and is further cooled by heat exchange with cold LNG flash vapor at −254° F. [−159° C.] (stream 83) and cold flash liquids (stream 79a). The cold flash liquids are produced by withdrawing a portion of the partially subcooled feed stream (stream 79) from heat exchanger 58 and flash expanding the stream through an appropriate expansion device, such as expansion valve 59, to slightly above the operating pressure of fractionation tower 17. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream from −155° F. [−104° C.] to −158° F. [−106° C.] (stream 79a). The flash expanded stream 79a is then supplied to heat exchanger 58 as previously described. Note that in all cases heat exchanger 58 is representative of either a multitude of individual heat exchangers or a single multi-pass heat exchanger, or any combination thereof. In some circumstances, combining the services of heat exchanger 51 and heat exchanger 58 into a single multi-pass heat exchanger may be appropriate.
The remaining portion of the partially cooled feed stream is further cooled in heat exchanger 58 to −169° F. [−112° C.] (stream 82). It then enters a work expansion machine 60 in which mechanical energy is extracted from this high pressure stream. The machine 60 expands the subcooled liquid substantially isentropically from a pressure of about 720 psia [4,964 kPa(a)] to the LNG storage pressure (18 psia [124 kPa(a)]), slightly above atmospheric pressure. The work expansion cools the expanded stream 82a to a temperature of approximately −254° F. [−159° C.], whereupon it is then directed to LNG storage tank 61 where the flash vapor resulting from expansion (stream 83) is separated from the LNG product (stream 84).
The warmed flash liquid stream 79b leaving heat exchanger 58 at −158° F. [−105° C.] is supplied to heat exchanger 51. It is heated to −85° F. [−65° C.] (stream 79c) as it supplies cooling to LNG feed stream 72 as described earlier, and thereafter supplied to fractionation tower 17 at a lower mid-column feed point.
The flash vapor (stream 83) from LNG storage tank 61 passes countercurrently to the incoming dense-phase stream in heat exchanger 58 where it is heated to −158° F. [−105° C.] (stream 83a). It then enters heat exchanger 51 where it is heated to 115° F. [46° C.] (stream 83b) as it supplies cooling to LNG feed stream 72. Since this stream is at low pressure (15.5 psia [107 kPa(a)]), it must be compressed before it can be used as plant fuel gas. Compressors 63 and 65 (driven by supplemental power sources) with intercooler 64 are used to compress the stream (stream 83e). Following cooling in aftercooler 66, stream 83f at 115 psia [793 kPa(a)] is combined with stream 37 to become the fuel gas for the plant (stream 85).
The cold distillation vapor stream from the NGL recovery plant (stream 43) is heated to 115° F. [46° C.] as it supplies cooling to LNG feed stream 72 in heat exchanger 51, becoming the second residue gas (stream 43a) which is then re-compressed in compressor 62 driven by a supplemental power source. The compressed second residue gas (stream 43b) combines with the compressed first residue gas (stream 44b) to form third residue gas stream 45. After cooling to 120° F. [49° C.] in discharge cooler 20, third residue gas stream 45a is divided into two portions. One portion (stream 71) becomes the feed stream to the LNG production section. The other portion (stream 38) becomes the residue gas product, which flows to the sales gas pipeline at 740 psia [5,102 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
Comparing the recovery levels displayed in Table V for the
The present invention also has a lower specific power consumption than the
One skilled in the art will recognize that the present invention can be adapted for use with all types of NGL recovery plants to allow co-production of LNG. The examples presented earlier have all depicted the use of the present invention with an NGL recovery plant employing the process disclosed in U.S. Pat. No. 4,278,457 in order to facilitate comparisons of the present invention with the prior art. However, the present invention is generally applicable for use with any NGL recovery process that produces a distillation vapor stream that is at temperatures of −50° F. [−46° C.] or colder. Examples of such NGL recovery processes are described and illustrated in U.S. Pat. Nos. 3,292,380; 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; 5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664; 6,182,469; reissue U.S. Pat. No. 33,408; and co-pending application Ser. No. 09/677,220, the full disclosures of which are incorporated by reference herein in their entirety. Further, the present invention is applicable for use with NGL recovery plants that are designed to recover only C3 components and heavier hydrocarbon components in the NGL product (i.e., no significant recovery of C2 components), or with NGL recovery plants that are designed to recover C2 components and heavier hydrocarbon components in the NGL product but are being operated to reject the C2 components to the residue gas so as to recover only C3 components and heavier hydrocarbon components in the NGL product (i.e., ethane rejection mode of operation).
When the pressure of the feed gas to the LNG production section (stream 72) is below its cricondenbar pressure, it may be advantageous to withdraw the feed stream after cooling to an intermediate temperature, separate any condensed liquid that may have formed, and then expand the vapor stream in a work expansion machine prior to cooling the expanded stream to substantial condensation, similar to the embodiment displayed in FIG. 6. The condensed liquid (stream 74) removed in separator 52 will preferentially contain the heavier hydrocarbons found in the feed gas, which can then be flash expanded to the operating pressure of fractionation tower 17 by expansion valve 55 and supplied to fractionation tower 17 at a lower mid-column feed point. This allows these heavier hydrocarbons to be recovered in the NGL product (stream 41), increasing the purity of the LNG (stream 84). As shown in
For applications where the plant inlet gas (stream 31 in
In accordance with this invention, the cooling of the feed stream to the LNG production section may be accomplished in many ways. In the processes of
Depending on the quantity of heavier hydrocarbons in the LNG feed gas and the LNG feed gas pressure, the cooled feed stream 72a leaving heat exchanger 51 may not contain any liquid (because it is above its dewpoint, or because it is above its cricondenbar), so that separator 52 shown in
In accordance with this invention, external refrigeration may be employed to supplement the cooling available to the LNG feed gas from other process streams, particularly in the case of a feed gas richer than that used in the example. The use and distribution of flash vapor and flash liquid from the LNG cool-down section for process heat exchange, and the particular arrangement of heat exchangers for feed 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 the stream 72a (FIG. 5), stream 73b (FIG. 6), or stream 73a (
Subcooling of condensed liquid stream 72a (FIG. 5), condensed liquid stream 73b (FIG. 6), or condensed liquid stream 73a (
Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where appropriate. For example, isenthalpic flash expansion may be used in lieu of work expansion for subcooled liquid stream 82 in
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.
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