Patent Grant
7565815
This invention relates to a process for processing natural gas or other methane-rich gas streams to produce a liquefied natural gas (LNG) stream that has a high methane purity and a liquid stream containing predominantly hydrocarbons heavier than methane. The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. provisional application Ser. No. 60/296,848 which was filed on Jun. 8, 2001.
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 were 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 while producing as a co-product a liquid stream consisting primarily of hydrocarbons heavier than methane, such as natural gas liquids (NGL) composed of ethane, propane, butanes, and heavier hydrocarbon components, liquefied petroleum gas (LPG) composed of propane, butanes, and heavier hydrocarbon components, or condensate composed of butanes and heavier hydrocarbon components. Producing the co-product liquid stream has two important benefits: the LNG produced has a high methane purity, and the co-product liquid is a valuable product that may be used for many other purposes. A typical analysis of a natural gas stream to be processed in accordance with this invention would be, in approximate mole percent, 84.2% methane, 7.9% ethane and other C2 components, 4.9% propane and other C3 components, 1.0% iso-butane, 1.1% normal butane, 0.8% 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; 6,062,041; 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; and 6,347,532 B1 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).
Regardless of the method used to liquefy the natural gas stream, it is common to require removal of a significant fraction of the hydrocarbons heavier than methane before the methane-rich stream is liquefied. The reasons for this hydrocarbon removal step are numerous, including the need to control the heating value of the LNG stream, and the value of these heavier hydrocarbon components as products in their own right. Unfortunately, little attention has been focused heretofore on the efficiency of the hydrocarbon removal step.
In accordance with the present invention, it has been found that careful integration of the hydrocarbon removal step into the LNG liquefaction process can produce both LNG and a separate heavier hydrocarbon liquid product using significantly less energy than prior art processes. The present invention, although applicable at lower pressures, is particularly advantageous when processing feed gases in the range of 400 to 1500 psia [2,758 to 10,342 kPa(a)] or higher.
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 production rates reported as pounds per hour (Lb/Hr) correspond to the stated molar flow rates in pound moles per hour. The production rates reported as kilograms per hour (kg/Hr) 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 refrigerant streams and demethanizer side reboiler liquids at −68° F. [−55° C.] (stream 40). 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 −30° F. [−34° C.] and 1278 psia [8,812 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 20% of the total vapor, is combined with the condensed liquid, stream 33, to form stream 35. Combined stream 35 passes through heat exchanger 13 in heat exchange relation with refrigerant stream 71e, resulting in cooling and substantial condensation of stream 35a. The substantially condensed stream 35a at −120° F. [−85° C.] is then flash expanded through an appropriate expansion device, such as expansion valve 14, to the operating pressure (approximately 465 psia [3,206 kPa(a)]) of fractionation tower 19. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
The remaining 80% of the vapor from separator 11 (stream 36) enters a work expansion machine 15 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 15 expands the vapor substantially isentropically from a pressure of about 1278 psia [8,812 kPa(a)] to the tower operating pressure, with the work expansion cooling the expanded stream 36a to a temperature of approximately −103° F. [−75° C.]. The typical commercially available expanders are capable of recovering on the order of 80-85% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item 16) that can be used to re-compress the tower overhead gas (stream 38), for example. The expanded and partially condensed stream 36a is supplied as feed to distillation column 19 at a lower mid-column feed point.
The demethanizer in fractionation tower 19 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. As is often the case in natural gas processing plants, the fractionation tower may consist of two sections. The upper section 19a is a separator wherein the top feed is divided into its respective vapor and liquid portions, and wherein the vapor rising from the lower distillation or demethanizing section 19b is combined with the vapor portion (if any) of the top feed to form the cold demethanizer overhead vapor (stream 37) which exits the top of the tower at −135° F. [−93° C.]. The lower, demethanizing section 19b 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 one or more reboilers (such as reboiler 20) 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 115° F. [46° C.], based on a typical specification of a methane to ethane ratio of 0.020:1 on a molar basis in the bottom product.
The demethanizer overhead vapor (stream 37) is warmed to 90° F. [32° C.] in heat exchanger 24, and a portion of the warmed demethanizer overhead vapor is withdrawn to serve as fuel gas (stream 48) for the plant. (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 refrigerant compressors 64, 66, and 68 in this example.) The remainder of the warmed demethanizer overhead vapor (stream 38) is compressed by compressor 16 driven by expansion machines 15, 61, and 63. After cooling to 100° F. [38° C.] in discharge cooler 25, stream 38b is further cooled to −123° F. [−86° C.] in heat exchanger 24 by cross exchange with the cold demethanizer overhead vapor, stream 37.
Stream 38c then enters heat exchanger 60 and is further cooled by refrigerant stream 71d. After cooling to an intermediate temperature, stream 38c is divided into two portions. The first portion, stream 49, is further cooled in heat exchanger 60 to −257° F. [−160° C.] to condense and subcool it, whereupon it enters a work expansion machine 61 in which mechanical energy is extracted from the stream. The machine 61 expands liquid stream 49 substantially isentropically from a pressure of about 562 psia [3,878 kPa(a)] to the LNG storage pressure (15.5 psia [107 kPa(a)]), slightly above atmospheric pressure. The work expansion cools the expanded stream 49a to a temperature of approximately −258° F. [−161° C.], whereupon it is then directed to the LNG storage tank 62 which holds the LNG product (stream 50).
Stream 39, the other portion of stream 38c, is withdrawn from heat exchanger 60 at −160° F. [−107° C.] and flash expanded through an appropriate expansion device, such as expansion valve 17, to the operating pressure of fractionation tower 19. In the process illustrated in
All of the cooling for streams 35 and 38c 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 cooling water 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 71 leaves discharge cooler 69 at 100° F. [38° C.] and 607 psia [4,185 kPa(a)]. It enters heat exchanger 10 and is cooled to −31° F. [−35° C.] and partially condensed by the partially warmed expanded refrigerant stream 71f and by other refrigerant streams. For the
The superheated refrigerant vapor (stream 71g) leaves heat exchanger 10 at 93° F. [34° C.] and is compressed in three stages to 617 psia [4,254 kPa(a)]. Each of the three compression stages (refrigerant compressors 64, 66, and 68) is driven by a supplemental power source and is followed by a cooler (discharge coolers 65, 67, and 69) to remove the heat of compression. The compressed stream 71 from discharge cooler 69 returns to heat exchanger 10 to complete the cycle.
A summary of stream flow rates and energy consumption for the process illustrated in
The efficiency of LNG production processes is typically compared using the “specific power consumption” required, which is the ratio of the total refrigeration compression power to the total liquid production rate. Published information on the specific power consumption for prior art processes for producing LNG indicates a range of 0.168 HP-Hr/Lb [0.276 kW-Hr/kg] to 0.182 HP-Hr/Lb [0.300 kW-Hr/kg], which is believed to be based on an on-stream factor of 340 days per year for the LNG production plant. On this same basis, the specific power consumption for the
There are two primary factors that account for the improved efficiency of the present invention. The first factor can be understood by examining the thermodynamics of the liquefaction process when applied to a high pressure gas stream such as that considered in this example. Since the primary constituent of this stream is methane, the thermodynamic properties of methane can be used for the purposes of comparing the liquefaction cycle employed in the prior art processes versus the cycle used in the present invention.
Contrast this now with the liquefaction cycle of the present invention. After partial cooling at high pressure (path A-A′), the gas stream is work expanded (path A′-A″) to an intermediate pressure. (Again, fully isentropic expansion is displayed in the interest of simplicity.) The remainder of the cooling is accomplished at the intermediate pressure (path A″-B′), and the stream is then expanded (path B′-C) to the pressure of the LNG storage vessel. Since the lines of constant entropy slope less steeply in the vapor region of the phase diagram, a significantly larger enthalpy reduction is provided by the first work expansion step (path A′-A″) of the present invention. Thus, the total amount of cooling required for the present invention (the sum of paths A-A′ and A″-B′) is less than the cooling required for the prior art processes (path A-B), reducing the refrigeration (and hence the refrigeration compression) required to liquefy the gas stream.
The second factor accounting for the improved efficiency of the present invention is the superior performance of hydrocarbon distillation systems at lower operating pressures. The hydrocarbon removal step in most of the prior art processes is performed at high pressure, typically using a scrub column that employs a cold hydrocarbon liquid as the absorbent stream to remove the heavier hydrocarbons from the incoming gas stream. Operating the scrub column at high pressure is not very efficient, as it results in the co-absorption of a significant fraction of the methane and ethane from the gas stream, which must subsequently be stripped from the absorbent liquid and cooled to become part of the LNG product. In the present invention, the hydrocarbon removal step is conducted at the intermediate pressure where the vapor-liquid equilibrium is much more favorable, resulting in very efficient recovery of the desired heavier hydrocarbons in the co-product liquid stream.
If the specifications for the LNG product will allow more of the ethane contained in the feed gas to be recovered in the LNG product, a simpler embodiment of the present invention may be employed.
In the simulation of the
The vapor (stream 32) from separator 11 is divided into two streams, 34 and 36. Stream 34, containing about 20% of the total vapor, is combined with the condensed liquid, stream 33, to form stream 35. Combined stream 35 passes through heat exchanger 13 in heat exchange relation with refrigerant stream 71e, resulting in cooling and substantial condensation of stream 35a. The substantially condensed stream 35a at −120° F. [−85° C.] is then flash expanded through an appropriate expansion device, such as expansion valve 14, to the operating pressure (approximately 465 psia [3,206 kPa(a)]) of fractionation tower 19. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
The remaining 80% of the vapor from separator 11 (stream 36) enters a work expansion machine 15 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 15 expands the vapor substantially isentropically from a pressure of about 1278 psia [8,812 kPa(a)] to the tower operating pressure, with the work expansion cooling the expanded stream 36a to a temperature of approximately −103° F. [−75° C.]. The expanded and partially condensed stream 36a is supplied as feed to distillation column 19 at a mid-column feed point.
The cold demethanizer overhead vapor (stream 37) exits the top of fractionation tower 19 at −123° F. [−86° C.]. The liquid product stream 41 exits the bottom of the tower at 118° F. [48° C.], based on a typical specification of a methane to ethane ratio of 0.020:1 on a molar basis in the bottom product.
The demethanizer overhead vapor (stream 37) is warmed to 90° F. [32° C.] in heat exchanger 24, and a portion (stream 48) is then withdrawn to serve as fuel gas for the plant. The remainder of the warmed demethanizer overhead vapor (stream 49) is compressed by compressor 16. After cooling to 100° F. [38° C.] in discharge cooler 25, stream 49b is further cooled to −112° F. [−80° C.] in heat exchanger 24 by cross exchange with the cold demethanizer overhead vapor, stream 37.
Stream 49c then enters heat exchanger 60 and is further cooled by refrigerant stream 71d to −257° F. [−160° C.] to condense and subcool it, whereupon it enters a work expansion machine 61 in which mechanical energy is extracted from the stream. The machine 61 expands liquid stream 49d substantially isentropically from a pressure of about 583 psia [4,021 kPa(a)] to the LNG storage pressure (15.5 psia [107 kPa(a)]), slightly above atmospheric pressure. The work expansion cools the expanded stream 49e to a temperature of approximately −258° F. [−161° C.], whereupon it is then directed to the LNG storage tank 62 which holds the LNG product (stream 50).
Similar to the
The superheated refrigerant vapor (stream 71g) leaves heat exchanger 10 at 93° F. [34° C.] and is compressed in three stages to 617 psia [4,254 kPa(a)]. Each of the three compression stages (refrigerant compressors 64, 66, and 68) is driven by a supplemental power source and is followed by a cooler (discharge coolers 65, 67, and 69) to remove the heat of compression. The compressed stream 71 from discharge cooler 69 returns to heat exchanger 10 to complete the cycle.
A summary of stream flow rates and energy consumption for the process illustrated in
Assuming an on-stream factor of 340 days per year for the LNG production plant, the specific power consumption for the
Compared to the
If the specifications for the LNG product will allow all of the ethane contained in the feed gas to be recovered in the LNG product, or if there is no market for a liquid co-product containing ethane, an alternative embodiment of the present invention such as that shown in
In the simulation of the
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 from a pressure of about 1278 psia [8,812 kPa(a)] to a pressure of about 440 psia [3,034 kPa(a)] (the operating pressure of separator/absorber tower 18), with the work expansion cooling the expanded stream 32a to a temperature of approximately −81° F. [−63° C.]. The expanded and partially condensed stream 32a is supplied to absorbing section 18b in a lower region of separator/absorber tower 18. The liquid portion of the expanded stream comingles with liquids falling downward from the absorbing section and the combined liquid stream 40 exits the bottom of separator/absorber tower 18 at −86° F. [−66° C.]. The vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier components.
The separator/absorber tower 18 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 separator/absorber tower may consist of two sections. The upper section 18a is a separator wherein any vapor contained in the top feed is separated from its corresponding liquid portion, and wherein the vapor rising from the lower distillation or absorbing section 18b is combined with the vapor portion (if any) of the top feed to form the cold distillation stream 37 which exits the top of the tower. The lower, absorbing section 18b contains the trays and/or packing and provides the necessary contact between the liquids falling downward and the vapors rising upward to condense and absorb the C3 components and heavier components.
The combined liquid stream 40 from the bottom of separator/absorber tower 18 is routed to heat exchanger 13 by pump 26 where it (stream 40a) is heated as it provides cooling of deethanizer overhead (stream 42) and refrigerant (stream 71a). The combined liquid stream is heated to −24° F. [−31° C.], partially vaporizing stream 40b before it is supplied as a mid-column feed to deethanizer 19. The separator liquid (stream 33) is flash expanded to slightly above the operating pressure of deethanizer 19 by expansion valve 12, cooling stream 33 to −46° F. [−43° C.] (stream 33a) before it provides cooling to the incoming feed gas as described earlier. Stream 33b, now at 85° F. [29° C.], then enters deethanizer 19 at a lower mid-column feed point. In the deethanizer, streams 40b and 33b are stripped of their methane and C2 components. The deethanizer in tower 19, operating at about 453 psia [3,123 kPa(a)], is also a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The deethanizer tower may also consist of two sections: an upper separator section 19a wherein any vapor contained in the top feed is separated from its corresponding liquid portion, and wherein the vapor rising from the lower distillation or deethanizing section 19b is combined with the vapor portion (if any) of the top feed to form distillation stream 42 which exits the top of the tower; and a lower, deethanizing section 19b that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The deethanizing section 19b also includes one or more reboilers (such as reboiler 20) which heat and vaporize a portion of the liquid at the bottom of the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 41, of methane and C2 components. A typical specification for the bottom liquid product is to have an ethane to propane ratio of 0.020:1 on a molar basis. The liquid product stream 41 exits the bottom of the deethanizer at 214° F. [101° C.].
The operating pressure in deethanizer 19 is maintained slightly above the operating pressure of separator/absorber tower 18. This allows the deethanizer overhead vapor (stream 42) to pressure flow through heat exchanger 13 and thence into the upper section of separator/absorber tower 18. In heat exchanger 13, the deethanizer overhead at −19° F. [−28° C.] is directed in heat exchange relation with the combined liquid stream (stream 40a) from the bottom of separator/absorber tower 18 and flashed refrigerant stream 71e, cooling the stream to −89° F. [−67° C.] (stream 42a) and partially condensing it. The partially condensed stream enters reflux drum 22 where the condensed liquid (stream 44) is separated from the uncondensed vapor (stream 43). Stream 43 combines with the distillation vapor stream (stream 37) leaving the upper region of separator/absorber tower 18 to form cold residue gas stream 47. The condensed liquid (stream 44) is pumped to higher pressure by pump 23, whereupon stream 44a is divided into two portions. One portion, stream 45, is routed to the upper separator section of separator/absorber tower 18 to serve as the cold liquid that contacts the vapors rising upward through the absorbing section. The other portion is supplied to deethanizer 19 as reflux stream 46, flowing to a top feed point on deethanizer 19 at −89° F. [−67° C.].
The cold residue gas (stream 47) is warmed from −94° F. [−70° C.] to 94° F. [34° C.] in heat exchanger 24, and a portion (stream 48) is then withdrawn to serve as fuel gas for the plant. The remainder of the warmed residue gas (stream 49) is compressed by compressor 16. After cooling to 100° F. [38° C.] in discharge cooler 25, stream 49b is further cooled to −78° F. [−61° C.] in heat exchanger 24 by cross exchange with the cold residue gas, stream 47.
Stream 49c then enters heat exchanger 60 and is further cooled by refrigerant stream 71d to −255° F. [−160° C.] to condense and subcool it, whereupon it enters a work expansion machine 61 in which mechanical energy is extracted from the stream. The machine 61 expands liquid stream 49d substantially isentropically from a pressure of about 648 psia [4,465 kPa(a)] to the LNG storage pressure (15.5 psia [107 kPa(a)]), slightly above atmospheric pressure. The work expansion cools the expanded stream 49e to a temperature of approximately −256° F. [−160° C.], whereupon it is then directed to the LNG storage tank 62 which holds the LNG product (stream 50).
Similar to the
The superheated refrigerant vapor (stream 71g) leaves heat exchanger 10 at 90° F. [32° C.] and is compressed in three stages to 617 psia [4,254 kPa(a)]. Each of the three compression stages (refrigerant compressors 64, 66, and 68) is driven by a supplemental power source and is followed by a cooler (discharge coolers 65, 67, and 69) to remove the heat of compression. The compressed stream 71 from discharge cooler 69 returns to heat exchanger 10 to complete the cycle.
A summary of stream flow rates and energy consumption for the process illustrated in
Assuming an on-stream factor of 340 days per year for the LNG production plant, the specific power consumption for the
Compared to the
If the specifications for the LNG product will allow all of the ethane and propane contained in the feed gas to be recovered in the LNG product, or if there is no market for a liquid co-product containing ethane and propane, an alternative embodiment of the present invention such as that shown in
In the simulation of the
The vapor (stream 32) from high pressure 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 from a pressure of about 1278 psia [8,812 kPa(a)] to a pressure of about 635 psia [4,378 kPa(a)], with the work expansion cooling the expanded stream 32a to a temperature of approximately −83° F. [−64° C.]. The expanded and partially condensed stream 32a enters intermediate pressure separator 18 where the vapor (stream 42) is separated from the condensed liquid (stream 39). The intermediate pressure separator liquid (stream 39) is flash expanded to slightly above the operating pressure of depropanizer 19 by expansion valve 17, cooling stream 39 to −108 F [−78° C.] (stream 39a) before it enters heat exchanger 13 and is heated as it provides cooling to residue gas stream 49 and refrigerant stream 71a, and thence to heat exchanger 10 to provide cooling to the incoming feed gas as described earlier. Stream 39c, now at −15° F. [−26° C.], then enters depropanizer 19 at an upper mid-column feed point.
The condensed liquid, stream 33, from high pressure separator 11 is flash expanded to slightly above the operating pressure of depropanizer 19 by expansion valve 12, cooling stream 33 to −93 F [−70° C.] (stream 33a) before it enters heat exchanger 13 and is heated as it provides cooling to residue gas stream 49 and refrigerant stream 71a, and thence to heat exchanger 10 to provide cooling to the incoming feed gas as described earlier. Stream 33c, now at 50° F. [10° C.], then enters depropanizer 19 at a lower mid-column feed point. In the depropanizer, streams 39c and 33c are stripped of their methane, C2 components, and C3 components. The depropanizer in tower 19, operating at about 385 psia [2,654 kPa(a)], 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 depropanizer tower may consist of two sections: an upper separator section 19a wherein any vapor contained in the top feed is separated from its corresponding liquid portion, and wherein the vapor rising from the lower distillation or depropanizing section 19b is combined with the vapor portion (if any) of the top feed to form distillation stream 37 which exits the top of the tower; and a lower, depropanizing section 19b that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The depropanizing section 19b also includes one or more reboilers (such as reboiler 20) which heat and vaporize a portion of the liquid at the bottom of the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 41, of methane, C2 components, and C3 components. A typical specification for the bottom liquid product is to have a propane to butanes ratio of 0.020:1 on a volume basis. The liquid product stream 41 exits the bottom of the deethanizer at 286° F. [141° C.].
The overhead distillation stream 37 leaves depropanizer 19 at 36° F. [2° C.] and is cooled and partially condensed by commercial-quality propane refrigerant in reflux condenser 21. The partially condensed stream 37a enters reflux drum 22 at 2° F. [−17° C.] where the condensed liquid (stream 44) is separated from the uncondensed vapor (stream 43). The condensed liquid (stream 44) is pumped by pump 23 to a top feed point on depropanizer 19 as reflux stream 44a.
The uncondensed vapor (stream 43) from reflux drum 22 is warmed to 94° F. [34° C.] in heat exchanger 24, and a portion (stream 48) is then withdrawn to serve as fuel gas for the plant. The remainder of the warmed vapor (stream 38) is compressed by compressor 16. After cooling to 100° F. [38° C.] in discharge cooler 25, stream 38b is further cooled to 15° F. [−9° C.] in heat exchanger 24 by cross exchange with the cool vapor, stream 43.
Stream 38c then combines with the intermediate pressure separator vapor (stream 42) to form cool residue gas stream 49. Stream 49 enters heat exchanger 13 and is cooled from 38° F. [−39° C.] to −102° F. [−74° C.] by separator liquids (streams 39a and 33a) as described earlier and by refrigerant stream 71e. Partially condensed stream 49a then enters heat exchanger 60 and is further cooled by refrigerant stream 71d to −254° F. [−159° C.] to condense and subcool it, whereupon it enters a work expansion machine 61 in which mechanical energy is extracted from the stream. The machine 61 expands liquid stream 49b substantially isentropically from a pressure of about 621 psia [4,282 kPa(a)] to the LNG storage pressure (15.5 psia [107 kPa(a)]), slightly above atmospheric pressure. The work expansion cools the expanded stream 49c to a temperature of approximately −255° F. [−159° C.], whereupon it is then directed to the LNG storage tank 62 which holds the LNG product (stream 50).
Similar to the
The superheated refrigerant vapor (stream 71g) leaves heat exchanger 10 at 93° F. [34° C.] and is compressed in three stages to 617 psia [4,254 kPa(a)]. Each of the three compression stages (refrigerant compressors 64, 66, and 68) is driven by a supplemental power source and is followed by a cooler (discharge coolers 65, 67, and 69) to remove the heat of compression. The compressed stream 71 from discharge cooler 69 returns to heat exchanger 10 to complete the cycle.
A summary of stream flow rates and energy consumption for the process illustrated in
Assuming an on-stream factor of 340 days per year for the LNG production plant, the specific power consumption for the
Compared to the
One skilled in the art will recognize that the present invention can be adapted for use with all types of LNG liquefaction plants to allow co-production of an NGL stream, an LPG stream, or a condensate stream, as best suits the needs at a given plant location. Further, it will be recognized that a variety of process configurations may be employed for recovering the liquid co-product stream. For instance, the
The disposition of the gas stream remaining after recovery of the liquid co-product stream (stream 37 in
In accordance with the present invention, the cooling of the inlet gas stream and the feed stream to the LNG production section may be accomplished in many ways. In the processes of
Further, the supplemental external refrigeration that is supplied to the inlet gas stream and the feed stream to the LNG production section may also be accomplished in many different ways. In
Subcooling of the condensed liquid stream leaving heat exchanger 60 (stream 49 in
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 feed stream (stream 35a 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.
This is a divisional of U.S. patent application Ser. No. 11/188,297, filed Jul. 22, 2005. This is a divisional of U.S. patent application Ser. No. 10/823,248, filed on Apr. 13, 2004 which is a divisional of U.S. patent application Ser. No. 10/161,780, filed on Jun. 4, 2002, which claims priority under 35 U.S.C. § 199(e) to U.S. Provisional Patent Application No. 60/296,848, filed on Jun. 8, 2001.
| Number | Name | Date | Kind |
|---|---|---|---|
| 4895584 | Buck et al. | Jan 1990 | A |
| 5685170 | Sorensen | Nov 1997 | A |
| 6526777 | Campbell et al. | Mar 2003 | B1 |
| Number | Date | Country | |
|---|---|---|---|
| 20080028790 A1 | Feb 2008 | US |
| Number | Date | Country | |
|---|---|---|---|
| 60296848 | Jun 2001 | US |
| Number | Date | Country | |
|---|---|---|---|
| Parent | 11188297 | Jul 2005 | US |
| Child | 11707787 | US | |
| Parent | 10823248 | Apr 2004 | US |
| Child | 11188297 | US | |
| Parent | 10161780 | Jun 2002 | US |
| Child | 10823248 | US |
