BR0210928B1 - "Process for the Liquidation of a Natural Gas Chain Containing Methane and Heavier Hydrocarbon Components". - Google Patents

Description

Patent Descriptive Report for "PROCESS FOR LIQUILATING A NATURAL GAS CURRENT CONTAINING METHANE AND HIGHER HYDROCARBON COMPONENTS".

SPECIFICATION

BACKGROUND OF THE INVENTION The present invention relates to a process for processing natural gas or other methane rich gas streams to produce a liquefied natural gas (LNG) stream having a high purity of methane and a stream of methane. containing, at first, hydrocarbons heavier than methane. Applicants claim benefits under Title 35, United States Code, Paragraph 119 (e) of former United States provisional application Serial No. 60 / 296,848, which was filed on June 8, 2001. Natural gas It is typically recovered from holes drilled in underground reservoirs. Usually it has a higher proportion of methane, ie methane comprises at least 50 moles per cent of the gas. Depending on the particular underground reservoir, natural gas also contains relatively smaller 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 of conveying natural gas from the cavity to gas processing facilities and then to natural gas consumers is in high pressure gas transmission lines. In a number of circumstances, however, it has been found to be necessary and / or desirable to liquefy natural gas for transportation or use. In remote locations, for example, often there is no infrastructure that allows for the convenient transportation of natural gas to the market. In these cases, the much smaller specific volume of LNG relative to gaseous natural gas can greatly reduce transportation costs by allowing LNG distribution using freighter and haulage trucks.

Another circumstance that favors the liquefaction of natural gas is its use as motor vehicle fuel. In large metropolitan areas there are bus, taxi and truck fleets that could be powered by LNG if an economic source of LNG is available. These LNG-powered 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 burn larger molecular weight hydrocarbons. In addition, if LNG is of high purity (ie methane purity of 95 mol percent or more), the amount of carbon dioxide (a "greenhouse gas") produced is considerably lower due to the proportion Lower carbon: hydrogen to methane compared to all other hydrocarbon fuels. The present invention generally relates to the liquefaction of natural gas while producing as a by-product a stream of liquid consisting mainly of hydrocarbons heavier than methane such as liquid natural gas ( NGL), composed of ethane, propane, butane and heavier hydrocarbon components, liquefied petroleum gas (LPG) composed of propane, butane and heavier hydrocarbon components, or condensed butane compounds and heavier hydrocarbon components heavy. Producing the co-product liquid stream has two important benefits: LNG produced has a high purity of methane and co-product liquid is a valuable product that can be used for many other purposes. A typical analysis of a natural gas stream to be processed according to the present invention would be, in approximate mol percent, 84.2% methane, 7.9% ethane and other C 2 components, 4.9% methane. propane and other C3 components, 1.0% isobutane, 1.1% normal butane, 0.8% pentanes plus, with the remainder composed of nitrogen and carbon dioxide. Sulfur-containing gases are sometimes also present. There are a number of known methods for liquefying natural gas. For example, see Finn, Adrian J., Grant L. Johnson and Terry R.

Tomlinson, "LNG Technology for Offshore and Mid-Scale Plants", Minutes of the Seventy-Ninth Annual Convention of the Gas Processors Association, p. 429-450, Atlanta, Georgia, March 13-15, 2000, and Kikkawa, Yoshitsugi, Masaaki Ohishi and Noriyoshi Nozawa, "Optimize the Power System of Baseload LNG Plant", Minutes of the 80th Annual Processor Association Convention Gas, San Antonio, Texas, March 12-14, 2001, for consideration of a number of these cases. US Patents 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; 6,308,531 B1; 6,324,867 B1; and 6,347,532 B1 also describe relevant processes. These methods generally include steps in which natural gas is purified (by removing water and imported compounds, such as carbon dioxide and sulfur compounds), cooled, condensed and expanded. Cooling and condensing natural gas can be accomplished in many different ways. "Casco cooling" employs natural gas heat exchange with various refrigerants having successively lower boiling points such as propane, ethane and methane. As an alternative, this heat exchange can be performed using a single refrigerant by evaporating the refrigerant at various different pressure levels. "Multicomponent refrigeration" employs natural gas heat exchange with one or more refrigerants composed of several refrigerant components rather than multiple single component refrigerants. Natural gas expansion can be carried out isententally (using Joule-Thomson expansion, for example) and isentropically (using a work expansion turbine, for example).

Regardless of the method used to liquefy the natural gas stream, it is common to require the removal of a significant fraction of 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 by right. Unfortunately, little attention has been focused so far 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 heavy hydrocarbon liquid product, using significantly less energy than the prior art processes. The present invention, although applicable at lower pressures, is particularly advantageous when supplying gas in the range of 400 to 1500 psia [2,758 to 10,342 kPa (a))] or more.

For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the Drawings: Figure 1 is a flow chart of a natural gas liquefaction plant adapted for NGL co-production in accordance with the present invention; Figure 2 is a pressure-enthalpy phase diagram for methane used to illustrate the advantages of the present invention over prior art processes; Figure 3 is a flow chart of an alternative natural gas liquefaction plant adapted for NGL co-production in accordance with the present invention; Figure 4 is a flow chart of an alternative natural gas liquefaction plant adapted for LPG co-production in accordance with the present invention; Fig. 5 is a flow chart of an alternative natural gas liquefaction plant adapted for condensate co-production according to the present invention; Figure 6 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Fig. 7 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Fig. 8 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Figure 9 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Fig. 10 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Figure 11 is a flowchart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Figure 12 is a flowchart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Fig. 13 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Fig. 14 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Figure 15 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Figure 16 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Fig. 17 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Fig. 18 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Figure 19 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; Fig. 20 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention; and Fig. 21 is a flow chart of an alternative natural gas liquefaction plant adapted for co-producing a liquid stream in accordance with the present invention.

In the following explanation of the above figures, tables are provided summarizing the calculated flow rates for representative process conditions. In the tables that appear here, values for flow rates (in moles per hour) have been rounded to get closer to an integer for convenience. The total current rates shown in the tables include all non-hydrocarbon components and therefore are generally greater than the sum of current flow rates for hydrocarbon components. The 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 leakage from (or to) the process vicinity. The quality of commercially available insulation 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 traditional British units and International System of Units (SI) units. The molar flow rates given in the tables may be interpreted as moles per hour or moles per hour. Power consumption reported as horsepower (HP) and / or thousands of English thermal units per hour (MBTU / h) corresponds to the molar flow rates quoted in soft pounds per hour. Energy consumptions reported as kilowatts (kW) correspond to molar flow rates in soft kilograms per hour. Production rates reported as pounds per hour (Lb / h) correspond to the milled runoff rates mentioned in soft kilograms per hour.

DESCRIPTION OF THE INVENTION

Example 1 Referring now to Figure 1, we begin with an illustration of a process according to the present invention wherein it is desired to produce an NGL co-product containing most of the heaviest and ethane components in the stream. natural gas supply. In this embodiment of the present invention, the inlet gas enters the installation at 32 ° C [90 ° F] and 8,860 kPa (a) [1285 psia] as stream 31. If the inlet gas contains a concentration of dioxide carbon and / or sulfur compounds, which would prevent product streams from meeting specifications, these compounds are removed by appropriate pretreatment of the feed gas (not shown). In addition, the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccant has typically been used for this purpose. The feed stream 31 is cooled in heat exchanger 10 by heat exchange with demethanizer side referrer refrigerant and liquid streams at -55 ° C [-68 ° F] (stream 40). Note that in all cases the heat exchanger 10 is representative of a plurality of individual heat exchangers or a single multiphase heat exchanger or any combination thereof. (The decision 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, heat, heat temperatures, etc.). The cooled current 31a enters the separator 11 at -34 ° C [-30 ° F] and 8,812 kPa (a) [1278 psia], where vapor (stream 32) is separated from condensed liquid (stream 33). The steam (stream 32) of separator 11 is divided into two streams, 34 and 36. Stream 34, containing about 20% of total steam, is combined with condensed liquid stream 33 to form stream 35. The combined current 35 passes through the heat exchanger 13 in relation to heat exchange with the refrigerant current 71 and, resulting in substantial cooling and condensation of the current 35a. The substantially condensed current 35a at -85 ° C [-120 ° F] is then instantaneously expanded through an appropriate expansion device such as an expansion valve 14 to operating pressure (approximately 3,206 fractionation tower kPa (a) [465 psia]) 19. During expansion, a portion of the stream is vaporized, resulting in the cooling of the total stream. In the process illustrated in Figure 1, the expanded current 35b leaving the expansion valve 14 reaches a temperature of -86 ° C [-122 ° F] and is supplied at a midpoint feed position in the demethanization section 19b of the tower. of fractionation 19.

The remaining 80% of steam from separator 11 (current 36) enters a work expansion machine 15 where mechanical energy is extracted from that portion of the high pressure feed. Machine 15 expands the steam substantially and isotropically from a pressure of about 8,812 kPa (a) [1278 psia] to the operating pressure of the tower, with work expansion cooling the expanded current 36a to a temperature of at approximately -75 ° C [-103 ° F]. Typical commercially available expanders are capable of recovering about 80 - 85% of the theoretically available work in an ideal isentropic expansion. Recovered work is often used to drive a centrifugal compressor (such as item 16) that can be used to recompress the tower's suspended gas (chain 38), for example. The partially condensed expanded current 36a is supplied as feed to the distillation column 19 at a lower mid-column feed point. The demethanizer in the fractionation tower 19 is a conventional distillation column containing a plurality of vertically spaced trays, one or more filled beds, or a combination of trays and fillers. As is often the case in natural gas processing facilities, the fractionation tower can consist of two sections. Upper section 19a is a separator wherein the upper feed is divided into its respective vapor and liquid portions and wherein the steam rising from the lower distillation or demethanization section 19b is combined with the steam portion (if any) of the top feed to form the cold suspended steam from the demethanizer (stream 37), which exits from the top of the tower at -93 ° C [-135 ° F]. The lower section of demethanizer 19b contains the trays and / or fill and provides contact. between falling liquids and rising vapors. The demethanization section also includes one or more referrers (such as referrer 20), which heat and vaporize a portion of the liquids circulating below the column to provide extraction vapors that circulate above the column. Liquid product stream 41 exits the tower bottom at 46 ° C [115 ° F], based on a typical specification of a methane to ethane ratio of 0.020: 1 on a molar basis in the background product. The demethanizer overhead steam (stream 37) is heated to 32 ° C [90 ° F] in heat exchanger 24 and a portion of the heated demethanizer overhead steam is withdrawn to serve as combustible gas (stream 48) for the installation. (The amount of combustible gas to be removed is largely determined by the fuel required for the engines and / or turbines that drive the gas compressors in the facility, such as refrigerant compressors 64, 66, and 68 in this example) . The remainder of the heated demethanizer overhead steam (stream 38) is compressed by compressor 16 driven by expansion machines 15, 61 and 63. After cooling to 38 ° C [100 ° F] in discharge chiller 25, stream 38b is cooled once again to -86 ° C [-123 ° F] in heat exchanger 24 by cross-exchanging with cold demethanizer suspended steam, stream 37. Current 38c then enters heat exchanger 60 and is heat exchanged. stopped once more by the 71d refrigerant stream. After cooling to an intermediate temperature, current 38c is divided into two portions. The first portion, stream 49, is cooled once again in heat exchanger 60 to -160 ° C [-257 ° F] to condense and subcool as a result of what it enters a machine. expansion work 61 where mechanical energy is extracted from the current. Machine 61 expands liquid stream 49 substantially and isentropically from a pressure of about 3,878 kPa (a) [562 psia] to the LNG storage pressure (107 kPa (a) [15.5 psia]), slightly. above atmospheric pressure. The work expansion cools the expanded current 49a to a temperature of approximately -161 ° C [-258 ° F], as a result of which is then directed to the LNG 62 storage tank, which holds the product of LNG (current 50). Stream 39, the other portion of stream 38c, is withdrawn from heat exchanger 60 at -107 ° C [-160 ° F] and instantaneously expanded through a suitable expansion device such as an expansion valve 17 to fractionation tower operating pressure 19. In the process illustrated in Figure 1, there is no vaporization in the expanded current 39a, so its temperature drops only slightly to -107 ° C [-161 ° F] leaving the expansion valve 17. A The expanded current 39a is then supplied to the separating section 19a in the upper region of the fractionation tower 19. The liquids separated therein make the upper feed to the demethanization section 19b.

All cooling for currents 35 and 38c is provided by a closed loop refrigeration circuit. The working fluid for this cycle is a mixture of hydrocarbons and nitrogen, with the mixture composition adjusted as needed to provide the required refrigerant temperature during condensation at a reasonable pressure using the available cooling medium. In this case, condensation with cooling water has thus been assumed that a refrigerant mixture composed of nitrogen, methane, ethane, propane and heavier hydrocarbons is used in the simulation process of Figure 1. The current composition, in approximate mol percent. , is 7.5% nitrogen, 41.0% methane, 41.5% ethane and 10.0% propane, with the remaining heavier hydrocarbon compound. Refrigerant stream 71 leaves the discharge cooler 69 at 38 ° C [100 ° F] and 4,185 kPa (a) [607 psia]. It enters heat exchanger 10 and is cooled to -35 ° C [-31 ° F] is partially condensed by partially heated expanded refrigerant stream 71fe and other refrigerant streams. For the simulation of Figure 1, it was assumed that these other refrigerant streams are commercial grade propane refrigerant at three different temperature and pressure levels. Partially condensed refrigerant stream 71a then enters heat exchanger 13 for further cooling to -81 ° C [-114 ° F] by partially heated expanded refrigerant stream 71 and, partially condensing and subcooling the refrigerant (current 71b). The refrigerant is further subcooled to -160 ° F [-257 ° F] in heat exchanger 60 by means of the expanded coolant stream 71 d. The subcooled liquid stream 71c enters a work expansion machine 63 wherein the machine energy is extracted from the stream as it is substantially expanded and isentropically expanded from a pressure of about 4,040 kPa (a). [586 psia] to about 234 kPa (a) [34 psia]. During expansion, a portion of the current is vaporized, resulting in the total current cooling to -164 ° C [-263 ° F] (current 71 d). The expanded stream 71d then re-enters the heat exchangers 60, 13 and 10, where it provides cooling for stream 38c, stream 35 and refrigerant (streams 71, 71a and 71b) as It is vaporized and overheated. Overheated refrigerant vapor (current 71 g) 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 refrigerators 65, 67, and 69) to remove compression heat. Compressed current 71 from discharge cooler 69 returns to heat exchanger 10 to complete the cycle.

A summary of current flow rates and power consumption for the process illustrated in Figure 1 is presented in the following table: Table 1 (FIGURE 1) Current Flow Summary - Lb. Moles / h [kg mol / h] NGL recoveries * Ethane 95.06% Propane 100.00% Butanes + 100.00% Production rate 308.147 kg / hr [308.147 Lb / hr] LNG product

Production Rate 610,813 kg / hr [610,813 Lb / hr] Purity 99.52% Lowest Heating Value 33,99 MJ / m3 [912,3 BTU / SCF] Power Refrigerant Compression 170,904 kW [1033,957 HP] Propane Compression 55,591 kW [33,815 HP] Total compression 226,495 kW [226,495 HP] Utility heat Demethanizer referrer 18,969 kW [29,364 MBTU / Hr] * (Based on unrounded flow rates) LNG production process efficiency is typically compared to using the required "specific energy consumption", which is the ratio of total cooling compression energy to total liquid production rate. Published energy-specific information for prior art processes for producing LNG indicates a range from 0.168 HP-h / Lb [0.276 kW-h / kg] to 0.182 HP-h / Lb [0.300 kW-h / kg] ], which is believed to be based on a current factor of 340 days per year for the LNG production facility. On the same basis, the specific energy consumption for the embodiment of Figure 1 of the present invention is 0.161 HP-h / Lb [0.265 kW-h / kg], which gives an efficiency improvement of 4-13% over the processes of the invention. prior art. In addition, it should be noted that the specific energy consumption for prior art processes is based on the co-production of only one LPG (C3 and heavier hydrocarbons) or condensate (C4 and heavier hydrocarbons) liquid stream in relatively low recovery levels, not a stream of NGL liquid (C2 and heavier hydrocarbons) as shown for this example of the present invention. Prior art processes require considerably more cooling power to co-produce an NGL stream rather than an LPG stream or a condensate stream. There are two main 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 main constituent of this stream is methane, the thermodynamic properties of methane can be used for comparison purposes of the liquefaction cycle employed in prior art processes versus the cycle used in the present invention. Figure 2 contains a pressure-entrapment phase diagram for methane. In most prior art liquefaction cycles, all cooling of the gas stream is performed while the stream is at high pressure (stroke AB), as a result of which the stream is then expanded (stroke BC) to LNG storage vessel pressure (slightly above atmospheric pressure).

This expansion step may employ a job expansion machine, which is typically capable of recovering about 75 - 80% of the theoretically available work in an ideal isentropic expansion. For the sake of simplicity, the completely isentropic expansion is shown in figure 2 for course B-C. Even so, the reduction in enthalpy provided by this expansion per work is quite small, because the constant entropy lines are almost vertical in the phase dia- minal fluid region.

Contrast this now with the liquefaction cycle of the present invention. After partial cooling at high pressure (stroke A-A '), the gas stream is expanded by working stroke A'-A ") to an intermediate pressure. (Again, completely isentropic expansion is shown for simplicity's sake). .) The remainder of the cooling is performed at the intermediate pressure (stroke A "-B ') and the current is then expanded (stroke B'-C) to the LNG storage vessel pressure. Since the constant entropy lines slope less abruptly in the vapor region of the phase diagram, a significantly greater enthalpy reduction is provided by the first expansion per work step (stroke A'-A "). Thus, the total amount of cooling required for the present invention (the sum of strokes AA 'and A "-B') is less than the cooling required for prior art processes (stroke AB ), reducing the cooling (and thus the cooling compression) required to liquefy the gas stream. The second factor to be taken into account 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 prior art processes is performed at high pressure typically using a scrubber column which employs a cold hydrocarbon liquid as the absorbent stream to remove the heavier hydrocarbons from the stream. of inlet gas. Operating the scrubber column at high pressure is not very efficient as it results in the co-absorption of a significant fraction of methane and ethane from the gas stream, which must subsequently be extracted from the absorbent liquid and cooled to cool. become part of LNG product. In the present invention, the hydrocarbon removal step is conducted at 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.

Example 2 If the specifications for the LNG product 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. Figure 3 illustrates this alternative embodiment. The composition and conditions of the inlet gas considered in the process shown in Figure 3 are the same as those in Figure 1. Consequently, the process of Figure 3 can be compared with the embodiment shown in Figure 1.

In the process simulation of Figure 3, the inlet gas cooling, separation and expansion scheme for the NGL recovery section is essentially the same as that used in Figure 1. The inlet gas enters the installation at 32 ° C [90 ° F] and 8,860 kPa (a) [1285 psia] as stream 31 and is cooled in heat exchanger 10 by heat exchange with demethanizer side cooler and liquid streams at -37 ° C [-35 ° F] (current 40). The cooled current 31a enters separator 11 at -34 ° C [-30 ° F] and 8,812 kPa (a) [1278 psia], where vapor (current 32) is separated from condensed liquid (current 33). . The steam (stream 32) of separator 11 is divided into two streams, 34 and 36. Stream 34, containing about 20% of total steam, is combined with condensed liquid stream 33 to form stream 35. Combined current 35 passes through heat exchanger 13 in relation to heat exchange with refrigerant current 71 and, resulting in substantial cooling and condensation of current 35a. The substantially condensed current 35a at -85 ° C [-120 ° F] is then instantaneously expanded through a suitable expansion device, such as expansion valve 14, to operating pressure (approx. 3.206 kPa (a) [465 psia] Fractionation Tower 19. During expansion, a portion of the stream is vaporized, resulting in the cooling of the total stream.In the process illustrated in Figure 3, the expanded stream 35b leaving the expansion valve 14 reaches a temperature of -86 ° C [-122 ° F] and is supplied to the separator section in the fractionation tower region 19. The liquids separated thereon become the upper feed for the demethanization section in the lower fractionation tower region 19

The remaining 80% of steam from separator 11 (current 36) enters a work expansion machine 15 where mechanical energy is extracted from that portion of the high pressure feed. Machine 15 expands the steam substantially and isotropically from a pressure of about 8,812 kPa (a) [1278 psia] to the operating pressure of the tower, with work expansion cooling the expanded current 36a to a temperature of at approximately -75 ° C [-103 ° F]. The partially condensed expanded current 36a is supplied as feed to the distillation column 19 at a mid-column feed point. Cold suspended steam from the demethanizer (stream 37) exits from the top of the fractionation tower 19 at -86 ° C [-123 ° F]. The liquid product stream 41 exits the bottom of the tower at 48 ° C [118 ° F], based on a typical specification of a methane to ethane ratio of 0.020: 1 on a molar basis in the background product. The suspended demethanizer steam (stream 37) is heated to 32 ° C [90 ° F] in heat exchanger 24 and a portion (stream 48) is then withdrawn to serve as combustible gas for the installation. The remainder of the heated demethanizer suspended steam (stream 49) is compressed by compressor 16. After cooling to 38 ° C [100 ° F] in discharge chiller 25, stream 49b is cooled once again. to -80 ° C [-112 ° F] on heat exchanger 24 by cross-exchanging with cold demethanizer suspended steam, stream 37. Stream 49c then enters heat exchanger 60 and is cooled further. once through the refrigerant stream 71 d to -160 ° C [-257 ° F] to condense and subcool, as a result of which it enters a work expansion machine 61 where mechanical energy is extracted of the current. Machine 61 expands the 49d liquid stream substantially and isentropically from a pressure of about 4,021 kPa (a) [583 psia] to the LNG storage pressure (107 kPa (a) [15.5 psia]), slightly above of atmospheric pressure. The work expansion cools the expanded current 49e to a temperature of approximately -161 ° C [-258 ° F], as a result of which is then directed to the LNG 62 storage tank, which holds the product of LNG (current 50).

Similar to the process of Fig. 1, all cooling for streams 35 and 49c is provided by a closed loop refrigeration circuit. The composition of the current used as the working fluid in the cycle for the process of Figure 3, in approximate mol percent, is 7.5% nitrogen, 40.0% methane, 42.5% ethane and 10%. 0.0% propane, with the remaining heavy hydrocarbon compound. Refrigerant stream 71 leaves the discharge cooler 69 at 38 ° C [100 ° F] and 4,185 kPa (a) [607 psia]. It enters heat exchanger 10 and is cooled to -35 ° C [-31 ° F] and partially condensed by partially heated expanded refrigerant stream 71 f and other refrigerant streams. For the simulation of figure 3, these other refrigerant streams were assumed to be commercial grade propane refrigerants at three different levels of temperature and pressure. Partially condensed coolant stream 71a then enters heat exchanger 13 for further cooling to -85 ° C [-121 ° F] by partially heated expanded coolant stream 71 and partially condensing and subcooling the refrigerant (current 71b). The refrigerant is further cooled to -160 ° F [-257 ° F] in heat exchanger 60 by the expanded refrigerant stream 71 d. The subcooled liquid stream 71c enters a work expansion machine 63 wherein mechanical energy is extracted from the stream as it is substantially expanded and isentropially of a pressure of about 4,040 kPa (a) [586 psia] to about 234 kPa (a) [34 psia]. During expansion, a portion of the current is vaporized, resulting in the total current cooling to -164 ° C [-263 ° F] (current 71 d). Expanded current 71d then again penetrates heat exchangers 60, 13 and 10, where it provides cooling to current 49c, current 35 and refrigerant (currents 71, 71a and 71b) as It is vaporized and overheated. Overheated refrigerant vapor (current 71 g) leaves heat exchanger 10 at 34 ° C [93 ° F] and is compressed in three stages to 4,254 kPa (a) [617 psia]. 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 refrigerators 65, 67, and 69) to remove compression heat. Compressed current 71 from discharge cooler 69 returns to heat exchanger 10 to complete the cycle.

A summary of current flow rates and power consumption for the process illustrated in Figure 3 is presented in the following table: Table II (FIGURE 3) Current Flow Summary - Lb. Moles / h [kg moles / h] NGL recoveries * Ethane 87.57% Propane 97.41% Butanes + 99.47% Production rate 296.175 kg / hr [296.175 Lb / hr] LNG product

Production Rate 625,152 kg / hr [625,152 Lb / hr] Purity * 98.66% Lowest Heating Value 34.27 MJ / m3 [919.7 BTU / SCF] Power Refrigerant Compression 158.743 kW [96.560 HP] propane 57,086 kW [34,724 HP] Total compression 215,829 kW [131,284 HP] Utility heat Demethanizer referrer 14,326 kW [22,177 MNTU / Hr] * (Based on uncirculated flow rates) Assuming a current factor of 340 days a year for LNG production facility, the specific energy consumption for the embodiment of Figure 3 of the present invention is 0.153 HP-h / Lb [0.251 kW-h / kg]. of efficiency is 10-20% for the embodiment of figure 3. As noted earlier for the embodiment of figure 1, such efficiency improvement is possible with the present invention, although an NGL co-product is produced in place of the LPG. or condensed by-product produced by the prior art problems.

Compared to the embodiment of Figure 1, the embodiment of Figure 3 of the present invention requires about 5% less energy per unit of liquid produced. Thus, for a given amount of available compression energy, the embodiment of FIG. 3 could liquefy about 5% more natural gas than the embodiment of FIG. 1, due to the recovery of less C2 and heavier hydrocarbons in the byproduct. NGL The choice between the embodiments of FIG. 1 and FIG. 3 of the present invention for a particular application will generally be dictated by the monetary value of the heavier hydrocarbons in the NGL product versus their corresponding value in the LNG product or the specification of the. heating value for the LNG product (since the heating value of the LNG produced by the embodiment of figure 1 is lower than that produced by the embodiment of figure 3).

Example 3 If the specifications for the LNG product allow more of the ethane contained in the feed gas to be recovered in the LNG product, or, if there is no market for a ketone-containing liquid co-product, an alternative embodiment of the present invention. , as shown in Figure 4, may be employed to produce an LPG co-product stream. The composition and conditions of the inlet gas considered in the process shown in figure 4 are the same as those in figures 1 and 3. Consequently, the process of figure 4 can be compared with the embodiments shown in figures 1 and 3. .

In the process simulation of Figure 4, the inlet gas enters the installation at 32 ° C [90 ° F] and 8,860 kPa (a) [1285 psia] as current 31 and is cooled in heat exchanger 10 by heat exchange with refrigerant and vaporizer separator streams at -43 ° C [-46 ° F] (current 33a). The cooled stream 31a enters separator 11 at -18 ° C [-1 ° F] and 8,812 kPa (a) [1278 psia], where vapor (stream 32) is separated from condensed liquid (stream 33). Steam (current 32) from separator 11 enters working expansion machine 15 where mechanical energy is extracted from that portion of the high pressure feed. Machine 15 expands steam substantially and isentropically from a pressure of 8,812 kPa (a) [1278 psia] to a pressure of about 3,034 kPa (a) [440 psia] (the operating pressure of the separator / absorber tower 18) , with work expansion cooling the expanded current 32a to a temperature of approximately -63 ° C [-81 ° F]. The partially condensed expanded stream 32a is supplied to the absorption section 18b in a lower region of the separator / absorber tower 18. The liquid portion of the expanded stream mixes with the liquids coming down from the absorption section and the combined liquid stream 40 exits the bottom of the separator / absorber tower 18 at -66 ° C [-86 ° F]. The vapor portion of the expanded stream rises through the absorption section and is contacted with the cold liquid that descends to condense and absorb 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 filled beds, or a combination of trays and fillers. As is often the case in natural gas processing installations, the separator / absorber tower can consist of two sections. Upper section 18a is a separator in which any vapor contained in the upper feed is separated from its corresponding liquid portion and wherein the vapor rising from the distillation or absorption lower section 18b is combined with the steam portion (if any). ) from the upper feed to form the cold distillation vapor 37 exiting the top of the tower. Lower absorption section 18b contains trays and / or fillers and provides the necessary contact between falling liquids and rising vapors to condense and absorb C3 components and heavier components. The combined liquid stream 40 from the bottom of the separator / absorber tower 18 is routed to heat exchanger 13 by pump 26 where it (current 40a) is heated as it provides cooling of the upper desethaniser (current 42 ) and refrigerant (current 71a). The combined liquid stream is heated to -31 ° C [-24 ° F], partially vaporizing stream 40b before it is supplied as a mid-column feed to the deethaniser 19. Separator liquid (current 33) is rapidly expanded slightly above the operating pressure of the deethaniser 19 by the expansion valve 12, cooling current 33 to -43 ° C [-46 ° F] (current 33a) before it provides cooling to the incoming feed gas as described above. Stream 33b, now at 29 ° C [85 ° F], then enters the deethanizer 19 at a lower mid-column feed point. In the desethanizer, streams 40b and 33b are extracted from their methane and C2 components. The tower deethaniser, operating at about 3,123 kPa (a) [453 psia], is also a conventional distillation column containing a plurality of vertically spaced trays, one or more filled beds, or a combination of trays. and filling. The desethanizer tower may also consist of two sections: an upper section 19a in which any vapor contained in the upper feed is separated from its corresponding liquid portion and in which steam rising from the lower distillation section or desethanization 19b is combined with the steam portion (if any) of the upper feed to form distillation stream 42 exiting the top of the tower; and a lower desethanization section 19b contains trays and / or fillers to provide the necessary contact between falling liquids and rising vapors. The deethanization section 19b also includes one or more referrers (such as referrer 20) which heat and vaporize a portion of the liquid at the bottom of the column to provide the extraction vapor that circulates above the column and extracts the liquid product. stream 41 of methane and C2 components. A typical specification for the lower liquid product is to have an ethane to propane ratio of 0.020: 1 on a molar basis. The liquid product stream 41 exits at the bottom of the deethaniser at 101 ° C [214 ° F]. The operating pressure on the deethanizer 19 is kept slightly above the operating pressure of the separator / absorber tower 18.

This allows the suspended steam from the deethaniser (stream 42) to circulate under pressure through the heat exchanger 13 and thence to the upper section of the separator / absorber tower 18. At the heat exchanger 13, the -28 deethanizer suspension ° C [-19 ° F] is driven in relation to heat exchange with the combined lower link current (current 40a) of the separator / absorber tower bottom 18 and vaporized refrigerant current 71 and, cooling the current to -67 ° C [-89 ° F] (current 42a) and partially condensing it. The partially condensed stream enters the flow drum 22, where the condensed liquid (stream 44) is separated from uncondensed steam (stream 43). Stream 43 combines with the distillation steam stream (stream 37) leaving the upper region of the separator / absorber tower 18 to form cold waste gas stream 47. Condensate liquid (stream 44) is pumped for higher pressure by pump 23, as a result of which current 44a is divided into two parts. A portion, stream 45, is prorated to the separate upper section of the separator / absorber tower 18 to serve as the cold liquid that contacts the vapors rising through the absorption section. The other portion is supplied to the desethanizer 19 as reflux stream 46, circulating to a higher feed point on the deethanizer 19 at -67 ° C [-89 ° F]. The cold residual gas (stream 47) is heated from -70 ° C [-94 ° F] to 34 ° C [94 ° F] in heat exchanger 24 and a portion (stream 48) is then withdrawn to serve as fuel gas for the installation. The remainder of the heated residual gas (stream 49) is compressed by compressor 16. After cooling to 38 ° C [100 ° F] in discharge chiller 25, stream 49b is cooled once again to -61 ° C [-78 ° F] in heat exchanger 24 by cross-exchanging with cold residual gas, stream 47. Stream 49c then enters heat exchanger 60 and is cooled once again by refrigerant stream 71 d to -160 ° C [-255 ° F] to condense and subcool, as a result of which it enters a work expansion machine 61 in which mechanical energy is extracted from the current. Machine 61 expands liquid stream 49d substantially and isentropically from a pressure of about 4,465 kPa (a) [648 psia] to the LNG storage pressure (107 kPa (a) [15.5 psia]), slightly above the atmospheric pressure. The work expansion cools the expanded current 49e to a temperature of approximately -160 ° C [-256 ° F], as a result of which is then directed to the LNG 62 storage tank, which holds the product of LNG (current 50).

Similar to the processes of Fig. 1 and Fig. 3, much of the cooling for current 42 and all cooling for current 49c is provided by a closed loop refrigeration circuit. The composition of the current used as the working fluid in the cycle for the process of Figure 4, in approximate mol percent, is 8.7% nitrogen, 30.0% methane, 45.8% ethane. and 11.0% propane, with the remaining heavier hydrocarbon compound. Refrigerant stream 71 leaves the discharge cooler 69 at 38 ° C [100 ° F] and 4,185 kPa (a) [607 psia]. It enters heat exchanger 10 and is cooled to -27 ° C [-17 ° F] and partially condensed by partially heated expanded refrigerant stream 71f and other refrigerant streams. For the simulation of figure 4, it has been assumed that these other refrigerant streams are commercial grade propellant refrigerants at three different levels of temperature and pressure. The partially condensed refrigerant stream 71a then enters the heat exchanger 13 for further cooling to -67 ° C [-89 ° F] by the partially heated expanded refrigerant stream 71 and further condenses the refrigerant (current 71b ). The refrigerant is fully condensed and then subcooled to -160 ° F [-255 ° F] in heat exchanger 60 via the expanded refrigerant stream 71 d. The subcooled liquid stream 71c enters a work expansion machine 63 where mechanical energy is extracted from the stream as it is substantially and isropropically expanded from a pressure of about 4,040 kPa (a) [586 psia]. ] to about 234 kPa (a) [34 psia]. During expansion, a portion of the current is vaporized, resulting in the total current cooling to -164 ° C [-264 ° F] (current 71 d). Expanded stream 71d then re-enters heat exchangers 60, 13, and 10, where it provides cooling to stream 49c, stream 42, and refrigerant (streams 71, 71a, and 71b) as it is vaporized and overheated. . The overheated refrigerant vapor (current 71 g) leaves heat exchanger 10 at 32 ° C [90 ° F] and is compressed in three stages to 4,254 kPa (a) [617 psia]. 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 refrigerators 65, 67, and 69) to remove compression heat. Compressed current 71 from discharge cooler 69 returns to heat exchanger 10 to complete the cycle.

A summary of current flow rates and energy consumption for the process illustrated in Figure 4 is presented in the following table: Table III (FIGURE 4) LPG * Propane Recovery 99.08% Butanes + 100.00% Production Rate 197.051 kg / hr [197.051 Lb / hr] LNG Product

Production Rate 726.918 kg / hr [726.918 Ig / hr] Purity * 91.43% Lowest Heating Value 36.14 MJ / m3 [969.9 BTU / SCF] Power Refrigerant Compression 156.876 kW [95.424 HP] propane 46,130 kW [28,060 HP] Total compression 203,006 kW [1233,484 HP] Utility heat Demethanizer referrer 35,575 kW [55,070 MBTU / Hr] * (Based on non-rounded flow rates) Assuming a current factor of 340 days per year for LNG production facility, the specific energy consumption for the embodiment of Figure 4 of the present invention is 0.143 HP-h / Lb [0.236 kW-h / kg]. Compared to prior art processes, the efficiency improvement is 17 - 27% for the embodiment of figure 4.

Compared to the embodiments of Figure 1 and Figure 3, the embodiment of Figure 4 of the present invention requires 6% to 11% less energy per unit of liquid produced. Thus, for a given amount of available compression energy, the embodiment of figure 4 could liquefy about 6% more natural gas than the embodiment of figure 1 or about 11% more natural gas than the embodiment of figure 3, in particular. due to the recovery of only C3 and heavier hydrocarbons as a co-product of LPG. The choice between the embodiment of Figure 4 versus the embodiments of Figure 1 or Figure 3 of the present invention for a particular application will generally be dictated by the monetary value of ethane as part of an NGL product versus its corresponding value in the NGL product. LNG or by specifying the heating value for the LNG product (since the heating value of the LNG produced by the embodiments of FIG. 1 and FIG. 3 is lower than that produced by the embodiment of FIG. 4).

Example 4 If the specifications for the LNG product allow all 0 ethane and propane contained in the feed gas to be recovered in the LNG product, or, if there is no market for a ethane and propane liquid by-product An alternative embodiment of the present invention, such as that shown in Figure 5, may be employed to produce a stream of condensed co-product. The inlet gas composition and conditions considered in the process shown in FIG. 5 are the same as those in FIGS. 1.3 and 4. Consequently, the process of FIG. 5 can be compared with the embodiments shown in FIGS. 1. , 3 and 4.

In the process simulation of Figure 5, the inlet gas enters the installation at 32 ° C [90 ° F] and 8,860 kPa (a) [1285 psia] as current 31 and is cooled in heat exchanger 10 by heat exchange with coolant streams liquid high pressure separator at -38 ° C (-37 ° F) (current 33b), and intermediate pressure separator liquids at -38 ° C [-37 ° F] (current 39b ). The cooled current 31a enters the high pressure separator 11 at -34 ° C [-30 ° F] and 8,812 kPa (a) [1278 psia], where vapor (current 32) is separated from condensed liquid (current 33). Steam (current 32) from the high pressure separator 11 enters the work expansion machine 15 where mechanical energy is extracted from that portion of the high pressure feed. Machine 15 substantially expands steam from a pressure of about 8,812 kPa (a) [1278 psia] to a pressure of about 4,378 kPa (a) [635 psia] with work expansion cooling the expanded current. 32a to a temperature of approximately -64 ° C [-83 ° F]. The expanded and partially condensed stream 32a enters the intermediate pressure separator 18, where steam (stream 42) is separated from condensed liquid (stream 39). The intermediate pressure separator liquid (current 39) is rapidly expanded to slightly above the operating pressure of the despruner 19 by expansion valve 17, cooling current 39 to -78 ° C [-108 ° F] (current 39a ) before it enters heat exchanger 13 and is heated while providing cooling to the waste gas stream 49 and refrigerant stream 71a and thereafter to the heat exchanger 10 to provide cooling to the incoming feed gas, as described above. Current 39c, now at -26 ° C [-15 ° F], then enters depranitizer 19 at an upper feed point of the median column. The condensed liquid, current 33, from the high pressure separator 11 is rapidly expanded to slightly above the operating pressure of the despraniser 19 through the expansion valve 12, cooling the current 33 to -70 ° C [-93 ° F] (current 33a) before it enters heat exchanger 13 and is heated while providing cooling for waste gas stream 49 and refrigerant stream 71a and thereafter for heat exchanger 10 to provide cooling for gas power supply as described above. Current 33c, now at 10 ° C [50 ° F], then enters deprananiser 19 at a lower mid-column feed point. In the depropanizer, streams 39c and 33c are extracted from their methane, C2 components and C3 components. The tower despropanizer, operating at about 2,654 kPa (a) [385 psia], is a conventional distillation column containing a plurality of vertically spaced trays, one or more filled beds, or a combination of trays and fillers. . The despraniser tower may also consist of two sections: an upper section 19a wherein any vapor contained in the upper feed is separated from its corresponding liquid portion and wherein the steam rising from the lower distillation section or the depranse section 19b is combined with the steam portion (if any) of the upper feed to form distillation stream 37, which exits from the top of the tower; and an undersized lower section 19b containing trays and / or fillers to provide the necessary contact between falling liquids and rising vapors. The despranization section 19b also includes one or more referrers (such as referrer 20) which heat and vaporize a portion of the liquid at the bottom of the column to provide extraction vapors circulating above the column to extract the product. liquid, stream 41, methane, C2 components and C3 components. A typical specification for the lower liquid product is to have a propane to butane ratio of 0.020: 1 on a volume basis. Liquid product stream 41 exits at the bottom of the deethaniser at 141 ° C [286 ° F]. Suspended distillation column 37 leaves the despraniser 19 at 2 ° C [36 ° F] and is partially cooled and condensed by commercial grade propane refrigerant in reflux condenser 21. Partially condensed stream 37a enters reflux drum 22 at -17 ° C [2 ° F], where the condensed liquid (stream 44) is separated from uncondensed steam (stream 43). Condensed liquid (stream 44) is pumped by pump 23 to a higher feed point in the deprangler 19 as reflux stream 44a. Uncondensed steam (stream 43) from reflux drum 22 is heated to 34 ° C [94 ° F] in heat exchanger 24 and a portion (stream 48) is then withdrawn to serve as combustible gas for the installation. The remainder of the heated steam (current 38) is compressed by the compressor 16.

After cooling to 38 ° C [100 ° F] in discharge chiller 25, current 38b is cooled once again to 9 ° C [15 ° F] in heat exchanger 24 by cross-exchanging with cold vapor, stream 43. Stream 38c then combines with intermediate pressure separator steam (stream 42) to form cold waste gas stream 49. Stream 49 enters heat exchanger 13 and is cooled to -39 ° C [-38 ° F] to -74 ° C [-102 ° F] by separator liquids (streams 39a and 33a) as described above and refrigerant stream 71 e. Partially condensed current 49a then enters heat exchanger 60 and is further cooled by refrigerant stream 71 d to -159 ° C [-254 ° F] to condense and subcool, as a result of which. it penetrates a work expansion machine 61 in which mechanical energy is extracted from the current. Machine 61 expands liquid stream 49b substantially and isentropically from a pressure of about 4,282 kPa (a) [621 psia] to the LNG storage pressure (107 kPa (a) [15.5 psia]), slightly above of atmospheric pressure. The work expansion cools the expanded current 49c to a temperature of approximately -159 ° C [-255 ° F], as a result of which is then directed to the LNG 62 storage tank, which holds the product of LNG (current 50).

Similar to the processes of FIG. 1, FIG. 3, and FIG. 4, much of the cooling to stream 49 and all cooling to stream 49a is provided by a closed loop refrigeration circuit. The composition of the current used as the working fluid in the cycle for the process of Figure 5, in approximate mol percent, is 8.9% nitrogen, 34.3% methane, 41.3% ethane and 11.0%. % propane, with the remainder being made up of heavier hydrocarbons. Refrigerant stream 71 leaves the discharge cooler 69 at 38 ° C [100 ° F] and 4,185 kPa (a) [607 psia]. It enters heat exchanger 10 and is cooled to -34 ° C [-30 ° F] and partially condensed by partially heated expanded refrigerant stream 71 f and other refrigerant streams.

For the simulation of Figure 5, these other coolant streams were assumed to be commercial grade propane refrigerants at three different levels of temperature and pressure. The partially condensed refrigerant stream 71a then enters the heat exchanger 13 for further cooling to -74 ° C [-102 ° F] by the partially heated expanded refrigerant stream 71 and further condensing the heat exchanger. refrigerant (current 71b). The refrigerant is fully condensed and then subcooled to -159 ° F [-254 ° F] in heat exchanger 60 via expanded refrigerant stream 71 d. The subcooled liquid stream 71c enters a work expansion machine 63 wherein mechanical energy is extracted from the stream as it is substantially expanded and isentropically of a pressure of about 586 psia [4,040 kPa (a )] to about 234 kPa (a) [34 psia]. During expansion, a portion of the current is vaporized, resulting in the total current cooling to -164 ° C [-264 ° F] (current 71 d). The expanded stream 71d then re-enters the heat exchangers 60, 13 and 10, where it provides cooling for stream 49a, stream 49 and refrigerant (currents 71, 71a and 71b) as It is vaporized and overheated. Overheated refrigerant vapor (current 71 g) leaves heat exchanger 10 at 34 ° C [93 ° F] and is compressed in three stages to 4,254 kPa (a) [617 psia]. 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 refrigerators 65, 67, and 69) to remove compression heat. Compressed current 71 from discharge cooler 69 returns to heat exchanger 10 to complete the cycle.

A summary of current flow rates and energy consumption for the process illustrated in figure 4 is presented in the following table: Table IV (FIGURE 5) Current Flow Summary - kg moles / h [Lb. Moles / h] Condensate Recovery * Butanes 95.04% Pentanes + 99.57% Production Rate 88.390 kg / hr [88.390 Lg / hr] LNG Product

Production Rate 834.183 kg / hr [834.183 Lb / hr] Purity * 87.27% Lowest Heating Value 38.52 MJ / m3 [103.8 BTU / SCF] Power Refrigerant Compression 139.696 kW [84.974 HP] propane 64,837 kW [39,439 HP] Total compression 204,533 kW [124,413 HP] Utility heat Demethanizer referrer 34,182 kW [52,913 MBTU / hr] * (Based on non-round flow rates) Assuming a current factor of 340 days a year for LNG production facility, the specific energy consumption for the embodiment of Figure 5 of the present invention is 0.238 kW-h / kg [0.145 HP-h / Lb]. of efficiency is 16 - 26% for the embodiment of figure 5.

Compared to the embodiments of Figure 1 and Figure 3, the embodiment of Figure 5 of the present invention requires about 5% to 10% less energy per unit of liquid produced. Compared to the embodiment of figure 4, the embodiment of figure 5 of the present invention requires essentially the same energy per unit of liquid produced. Thus, for a given amount of available compression energy, the embodiment of figure 5 could liquefy about 5% more natural gas than the embodiment of figure 1, by about 10% more natural gas than the embodiment of figure 3. or approximately the same amount of natural gas as the embodiment of Figure 4, by virtue of the recovery of only C4 and heavier hydrocarbons as condensed co-product. The choice between the embodiment of figure 5 versus the embodiments of figure 3 or figure 4 of the present invention for a particular application will generally be dictated by the monetary values of ethane and propane as part of an NGL or LPG product. versus their corresponding values in the LNG product or by specifying the heating value for the LNG product (since the heating value of the LNG produced by the embodiments of figure 1, figure 3 and figure 4 is lower than that produced by the embodiment of figure 5).

Other Embodiments One of ordinary skill in the art will recognize that the present invention may be adapted for use with all types of LNG splitting facilities to allow the co-production of an NGL stream, an LPG stream or a LNG stream. as best suited to the needs at a given installation location. Further, it will be appreciated that a variety of process configurations may be employed for recovering the liquid by-product stream. For example, the embodiments of Figures 1 and 3 may be adapted to recover an LPG stream or a condensate stream as the liquid co-product stream rather than an NGL stream, as described above in Examples 1 and 2. The embodiment of Figure 4 may be adapted to recover an NGL stream containing a significant fraction of the C2 components present in the feed gas or to recover a condensate stream containing only the heavy and C4 components. present in the feed gas rather than producing a LPG co-product as described above for Example 3. The embodiment of Figure 5 may be adapted to recover an NGL stream containing a significant fraction of the C2 components present. in the feed gas, or recover a LPG stream containing a significant fraction of the C3 components present in the feed gas rather than producing a by-product as described above for Example 4.

Figures 1, 3, 4 and 5 represent preferred embodiments of the present invention for the indicated processing conditions. Figures 6 to 21 represent alternative embodiments of the present invention which may be considered for a particular application. As shown in figures 6 and 7, all or part of the condensed liquid (current 33) of separator 11 may be supplied to the fractionation tower 19 in a separate lower column mid-feed position rather than combine with the portion of the separator vapor (current 34) circulating to the heat exchanger 13. Figure 8 represents an alternative embodiment of the present invention that requires less equipment than the embodiments of figure 1 and figure 6, although your specific energy consumption is slightly higher. Similarly, Figure 9 represents an alternative embodiment of the present invention that requires less equipment than the embodiments of Figure 3 and Figure 7, again at the expense of a higher specific energy consumption. Figures 10 to 14 represent alternative embodiments of the present invention that may require less equipment than the embodiment of figure 4, although their specific energy consumptions may be higher. (Note that, as shown in figures 10 through 14, columns or distillation systems, such as the de-sanitizer 19, include recessed absorber tower designs and refluxed reclaimer tower designs). Figures 15 and 16 represent alternative embodiments of the present invention that combine the separator / absorber tower 18 and detaniser 19 functions in the embodiments of figures 4 and 10 to 14 into a single fractionation column 19. Depending on the amount of heavier hydrocarbons in the feed gas and feed gas pressure, the cooled feed stream 31 leaving the heat exchanger 10 may contain no liquid (because it is above its dew point or because it is above so that the separator 11 shown in figures 1 and 3 to 16 is not required and the cooled supply current can circulate directly to a suitable expansion device such as a work expansion machine 15. A arrangement of the remaining gas stream after recovery of the liquid by-product stream (stream 37 in figures 1, 3, 6 to 11, 13 and 14, stream 47 in figures 4 12, 15 and 16 and current 43 in FIG. 5) before it is supplied to the heat exchanger 60 for condensation and subcooling can be performed in various ways. In the processes of Figures 1 and 3 to 16, the current is heated, compressed to higher pressure using energy derived from one or more work expansion machines, partially cooled in a discharge chiller, then still cooled by exchange. crossed with the original chain. As shown in figure 17, some applications may favor current compression to higher pressure, using supplementary compressor 59 driven by an external power source, for example. As shown by the dashed equipment (heat exchanger 24 and discharge cooler 25) in figures 1 and 3 to 16, some circumstances may favor reducing the installation's capital cost by reducing or eliminating compressed current pre-cooling, before it enters heat exchanger 60 (at the expense of increasing the cooling load on heat exchanger 60 and increasing the power consumption of refrigerant compressors 64, 66, and 68). In such cases, the stream 49a leaving the compressor may flow directly to the heat exchanger 24 as shown in Figure 18 or flow directly to the heat exchanger 60 as shown in Figure 19. If the expansion machines work If they are not used to expand any portions of the high-pressure feed gas, a compressor driven by an external power source, such as compressor 59 shown in figure 20, may be used in place of compressor 16. Other circumstances may not justify any compression of the current at all, so that the current flows directly to the heat exchanger 60 as shown in figure 21 and by the dashed equipment (heat exchanger 24, compressor 16 and discharge cooler 25). 1 and 3 to 16. If heat exchanger 24 is not included to heat the stream before the fuel gas from the installation (stream 48) is removed, an aqua A supplemental burner 58 may be required to heat the combustible gas before it is consumed, using a utility or other process stream to provide the necessary heat, as shown in Figures 19 to 21.

Choices such as these should generally be evaluated for each application, as factors such as gas composition, installation size, desired co-product current recovery level and available equipment must all be considered.

In accordance with the present invention, cooling of the inlet gas stream and the feed stream to the LNG production section can be accomplished in many ways. In the processes of figures 1, 3 and 6 to 9, the gas inlet stream 31 is cooled and condensed by external refrigerant and tower liquid streams from the fractionation tower 19. In figures 4, 5 and 10 to 14, Vaporized separator liquids are used for this purpose together with external refrigerant streams.

In figures 15 and 16, tower liquids and vaporized separator liquids are used for this purpose in conjunction with external refrigerant streams. And in figures 17 through 21, only external refrigerant streams are used to cool the inlet gas stream 31. However, the process cold streams could also be used to provide some of the cooling to the high pressure refrigerant (current stream). 71a), as shown in Figures 4, 5, 10 and 11. Also, any current at a colder temperature than the current (s) being cooled may be used. For example, a side vapor aspiration from the separator / absorber tower 18 or fractionation tower 19 could be removed and used for cooling. The use and distribution of process heat exchange liquids and / or vapors from the process and the particular arrangement of inlet gas and feed gas cooling heat exchangers should be evaluated for each particular application, as well as the choice of process streams for specific heat exchange services. Selection of a cooling source will depend on a number of factors including, but not limited to, the composition and condition of the supply gas, installation size, heat exchanger size, potential cooling source temperature, etc. One skilled in the art will also recognize that any combination of the above cooling sources or cooling methods may be employed in combination to obtain the desired supply current (s) temperature (s).

In addition, supplemental external cooling that is supplied to the inlet gas stream and supply current for the LNG production section can also be performed in many different ways. In Figures 1 and 3 to 21, the single component boiling refrigerant was assumed for the high level external refrigerant and the multi component vaporizing refrigerant was assumed for the low level external refrigerant with the single component refrigerant used. to pre-cool the multi-component refrigerant stream. Alternatively, both high-level and low-level cooling could be performed using successively lower boiling point single-component refrigerants (ie "cascade cooling"), or a single-component refrigerant at pressures. successively lower evaporation rates. Alternatively, both high-level and low-level cooling could be performed using multi-component refrigerant streams, with their respective compositions adjusted to provide the required cooling temperatures. Selection of the method for providing external refrigeration will depend on a number of factors, including but not limited to the same, composition and condition of the supply gas, installation size, compressor drive size, heat exchanger size, heat, ambient temperature of the heatsink, etc. One skilled in the art will also recognize that any combination of the methods of providing external cooling described above may be employed in combination to obtain the desired supply current temperature (s). The subcooling of the condensed liquid stream leaving heat exchanger 60 (stream 49 in figures 1, 6 and 8, stream 49d in figures 3,4, 7 and 9 to 16, stream 49b in figures 5,19 and 20 , stream 49e in Fig. 17, stream 49c in Fig. 18 and stream 49a in Fig. 21) reduces or eliminates the amount of instantaneous steam that may be generated during expansion of the stream to the operating pressure of the LNG storage tank 62. This generally reduces specific energy consumption to produce LNG by eliminating the need for instant gas compression. However, some circumstances may favor reducing the capital cost of the facility by reducing the size of the heat exchanger 60 and using flash gas compression or other means to dispose of any flash gas that may be generated.

Although individual current expansion is represented in particular expansion devices, alternative expansion means may be employed where appropriate. For example, conditions may ensure working expansion of the substantially condensed feed stream (stream 35a in figures 1,3, 6 and 7) or intermediate pressure reflow stream (stream 39 in figures 1, 6 and 8) . Further, isenthalic instantaneous expansion may be used in lieu of labor expansion for the subcooled liquid stream leaving heat exchanger 60 (stream 49 in figures 1, 6 and 8, stream 49d in figures 3, 4, 7 and 9 to 16, stream 49b in FIGS. 5, 19 and 20, stream 49e in FIG. 17, stream 49c in FIG. 18 and stream 49a in FIG. 21), but will require further cooling in the heat exchanger 60 to prevent vapor formation. instantaneous expansion or, in a different way, adding instantaneous steam compression or other means for disposal of the instantaneous vapor that results. Similarly, isenthalic instantaneous expansion may be used in place of labor expansion for the subcooled high-pressure refrigerant stream leaving heat exchanger 60 (current 71c, in Figures 1 and 3 to 21) with the resulting increase in energy consumption for refrigerant compression.

While what are believed to be preferred embodiments of the invention have been described, those skilled in the art will recognize that further and further modifications may be made thereto, for example, adapting the invention to various conditions, types of feed or the like. requirements without departing from the spirit of the present invention as defined by the appended claims.

Claims (52)

A process for liquefying a natural gas stream 31 containing methane and heavier hydrocarbon components, wherein: (a) said natural gas stream is cooled under pressure to condense at least a portion thereof and form a stream condensate 49; and (b) said condensed stream 49 is expanded 61 to less pressure to form said liquefied natural gas stream 50; characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10; (2) said cooled natural gas stream 31a, 32 is expanded 15 to an intermediate pressure 32a; (3) said cooled expanded natural gas stream 32a is directed to a distillation column 19 wherein said stream is separated into a volatile fraction of residual gas 37 containing a major portion of said methane and lighter components and a relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbon components; and (4) said volatile waste gas fraction 37, 38c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream; (5) said condensed portion is divided into at least two portions to thereby form said condensed stream 49 and a liquid stream 39; and (6) said liquid stream 39a is directed to said distillation column 19 as a top feed thereto. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 for partial condensation thereof 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thus provide at least one vapor stream 32 and a first liquid stream 33; (3) said vapor stream 32 is expanded 15 to an intermediate pressure 32a; (4) said first liquid stream 33 is expanded 12 to said intermediate pressure 33a; (5) at least said expanded vapor stream 32a and said first expanded liquid stream 33a are directed to a distillation column 19 wherein said streams are separated into said volatile waste gas fraction 37 containing a major portion of said methane and lighter components and said relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbon components; (6) said volatile waste gas fraction 37, 38c is pressure-cooled 60 to condense at least a portion thereof; (7) said condensed portion is divided into at least two portions to thereby form said condensed stream 49 and a second liquid stream 39; (8) said second stream of liquid 39a is directed to said distillation column 19 as a top feed thereto. Process according to Claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10; (2) said cooled natural gas stream 31a, 32 is divided into at least one first gas stream 34 and a second gas stream 36; (3) said first gas stream 34, 35 is cooled 13 to substantially condense it all 35a and then expanded 14 to said intermediate pressure 35b; (4) said second gas stream 36 is expanded 15 to said intermediate pressure 36a; (5) said first substantially expanded condensed gas stream 35b and said second expanded gas stream 36a are directed to a distillation column 19 wherein said streams are separated into a volatile fraction of residual gas 37 containing a portion said methane and lighter components and a relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbon components; and (6) said volatile fraction of residual gas 37, 49c is pressure cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thus provide at least one vapor stream 32 and one liquid stream 33; (3) said gas stream 32 is divided into at least one first gas stream 34 and a second gas stream 36; (4) said first gas stream 34, 35 is cooled 13 to substantially condense it all 35a and then expanded 14 to an intermediate pressure 35b; (5) said second gas stream 36 is expanded 15 to said intermediate pressure 36a; (6) said liquid stream 33 is expanded 12 to said intermediate pressure 33a; (7) said first substantially condensed expanded gas stream 35b, said second expanded gas stream 36a and said expanded liquid stream 33a are directed to a distillation column 19 wherein said streams are separated into said volatile fraction of residual gas 37 containing a major portion of said methane and lighter components and said relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbon components; and (8) said volatile waste gas fraction 37, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thereby provide a vapor stream 32 and a liquid stream 33; (3) said vapor stream 32 is divided into at least one first gas stream 34 and a second gas stream 36; (4) said first gas stream 34 is combined with at least a portion of said liquid stream 33, thereby forming a combined stream 35; (5) said combined stream 35 is cooled 13 to substantially condense it all 35a and then expanded 14 to an intermediate pressure 35b; (6) said second gas stream 36 is expanded 15 to said intermediate pressure 36a; (7) any remaining portion of said liquid stream 33 is expanded 12 to said intermediate pressure 33a; (8) said substantially condensed expanded combined stream 35b, said second expanded gas stream 36a and said remaining portion of said liquid stream 33a are directed to a distillation column wherein said streams are separated into said volatile fraction residual gas 37 containing a major portion of said methane and lighter components and said relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbon components; and (9) said volatile waste gas fraction 37, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to Claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10; (2) said cooled natural gas stream 31a, 32 is divided into at least one first gas stream 34 and a second gas stream 36; (3) said first gas stream 34, 35 is cooled 13 to substantially condense it all 35a and then expanded 14 to an intermediate pressure 35b; (4) said second gas stream 36 is expanded 15 to said intermediate pressure 36a; (5) said first substantially condensed expanded gas stream 35b and said second expanded gas stream 36a are directed to a distillation column 19 wherein said streams are separated into said volatile waste gas fraction 37 containing a portion said major methane and lighter components and said relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbon components; (6) said volatile waste gas fraction 37, 38c is pressure-cooled 60 to condense at least a portion thereof; (7) said condensed portion is divided into at least two portions to thereby form said condensed stream 49 and a liquid stream 39; and (8) said liquid stream 39a is directed to said distillation column 19 as a top feed thereto. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thus provide at least one vapor stream 32 and a first liquid stream 33; (3) said vapor stream 32 is divided into at least one first gas stream 34 and a second gas stream 36; (4) said first gas stream 34, 35 is cooled 13 to substantially condense it all 35a and then expanded 14 to an intermediate pressure 35b; (5) said second gas stream 36 is expanded 15 to said intermediate pressure 36a; (6) said first liquid stream 33 is expanded 12 to said intermediate pressure 33a; (7) said first substantially condensed expanded gas stream 35b, said second expanded gas stream 36a and said first expanded liquid stream 33a are directed to a distillation column 19 wherein said streams are separated into said volatile fraction residual gas 37 containing a major portion of said methane and lighter components and said relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbon components; (8) said volatile fraction of residual gas 37, 38c is pressure-cooled 60 to condense at least a portion thereof; (9) said condensed portion is divided into at least two portions to thereby form said condensed stream 49 and a second liquid stream 39; and (10) said second liquid stream 39a is directed to said distillation column 19 as a top feed thereto. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thereby provide a vapor stream 32 and a first liquid stream 33; (3) said vapor stream 32 is divided into at least one first gas stream 34 and a second gas stream 36; (4) said first gas stream 34 is combined with at least a portion of said liquid stream 33, thereby forming a combined stream 35; (5) said combined stream 35 is cooled 13 to substantially condense it all 35a and then expanded 14 to an intermediate pressure 35b; (6) said second gas stream 36 is expanded 15 to said intermediate pressure 36a; (7) any remaining portion of said liquid stream 33 is expanded 12 to said intermediate pressure 33a; (8) said substantially condensed expanded combined stream 35b, said second expanded gas stream 36a and said remaining portion of said liquid stream 33a are directed to a distillation column 19 wherein said streams are separated at said volatile fraction of residual gas 37 containing a major portion of said methane and lighter components and said relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbon components; (9) said volatile waste gas fraction 37, 38c is pressure-cooled 60 to condense at least a portion thereof; (10) said condensed portion is divided into at least two portions to thereby form said condensed stream 49 and a liquid stream 39; (11) said second liquid stream 39a is directed to said distillation column as a top feed thereto. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10; (2) said cooled natural gas stream 31a, 32 is expanded 15 to an intermediate pressure 32a; (3) said natural gas stream 32a is separated 18 to thus provide at least one vapor stream 42 and one liquid stream 39; (4) said liquid stream 39 is expanded 17 to a lower intermediate pressure 39a; (5) said expanded liquid stream 39a is directed to a distillation column 19 wherein said stream is separated into a more volatile vapor distillation fraction 37 and a relatively less volatile fraction 41 containing a major portion of said distillation components. heavier hydrocarbons; (6) said more volatile steam distillation stream 37, 38c is combined with said vapor stream 42 to form a volatile waste gas fraction 49 containing a major portion of said methane and lighter components; and (7) said volatile waste gas fraction 49a is pressure cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49b. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thereby provide a first vapor stream 32 and a first liquid stream 33; (3) said first vapor stream 32 is expanded 15 to an intermediate pressure 32a; (4) said first steam stream 32a is separated 18 to thus provide a second steam stream 42 and a second liquid stream 39; (5) said second liquid stream 39 is expanded 17 to a lower intermediate pressure 39a; (6) said first liquid stream 33 is expanded 12 to said lower intermediate pressure 33a; (7) said second expanded liquid stream 39c and said first expanded liquid stream 33c are directed to a distillation column 19 wherein said streams are separated into a more volatile steam distillation stream 37 and a relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbon components; (8) said more volatile steam distillation stream 37, 38c is combined with said second steam stream 42 to form a volatile fraction of residual gas 49 containing a major portion of said methane and lighter components; and (9) said volatile waste gas fraction 49a is pressure cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49b. Process according to Claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10; (2) said cooled natural gas stream 31a, 32 is expanded 15 to an intermediate pressure 32a and then directed to a contact device 18, thereby forming a volatile fraction of residual gas 37 containing a main portion. said methane and lighter components and a first liquid stream 40; (3) said first liquid stream 40b is directed to a distillation column 19 wherein said stream is separated into a more volatile steam distillation stream 42 and a relatively less volatile fraction 41 containing a major portion of the hydro components. - heavier carbides; (4) said more volatile steam distillation stream 42 is cooled enough to condense at least a portion thereof, thereby forming a second liquid stream 42a; (5) at least a portion of said cooled expanded natural gas stream 32a is intimately contacted with at least part of said second liquid stream 42a in said contact device 18; and (6) the volatile fraction of residual gas 37, 49c is cooled 60 under pressure to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thus provide at least one vapor stream 32 and a first liquid stream 33; (3) said vapor stream 32 is expanded 15 to an intermediate pressure 32a and then directed to a contact device 18 thereby forming a volatile fraction of residual gas 37 containing a major portion of said methane and further components. light and a second stream of liquid 40; (4) said first liquid stream 33 is expanded 12 to said intermediate pressure 33a; (5) said second liquid stream 40b and said first expanded liquid stream 33b are directed to a distillation column 19 wherein said streams are separated into a more volatile steam distillation stream 42 and a relatively small fraction. less volatile 41 containing a major portion of the heavier hydrocarbon components; (6) said more volatile steam distillation stream 42 is sufficiently cooled 13 to condense at least a portion thereof, thereby forming a third liquid stream 42a; (7) at least a portion of said expanded vapor stream 32a is intimately contacted with at least part of said third liquid stream 42a in said contact device 18; and (8) said volatile waste gas fraction 37, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10; (2) said cooled natural gas stream 31a, 32 is expanded 15 to an intermediate pressure 32a and then directed to a contact device 18, thereby forming a first steam stream 37 and a first flow stream. liquid 40; (3) said first liquid stream 40b is directed to a distillation column 19 wherein said stream is separated into a more volatile steam distillation stream 42 and a relatively less volatile fraction 41 containing a major portion of said stream components. heavier hydrocarbons; (4) said more volatile steam distillation stream 42 is sufficiently cooled 13 to condense at least a portion thereof 42a, thereby forming a second steam stream 43 and a second liquid stream 44; (5) a portion 46 of said second liquid stream 44 is directed to said distillation column 19 as a top feed thereto; (6) at least a portion of said expanded cooled natural gas stream 32a is intimately contacted with at least part of the remaining portion 45 of said second liquid stream 44 in said contact device 18; (7) said first vapor stream 37 is combined with said second vapor stream 43 to form a volatile fraction of residual gas 47 containing a major portion of said methane and lighter components; and (8) said volatile waste gas fraction 47, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to Claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thus provide at least a first vapor stream 32 and a first liquid stream 33; (3) said first vapor stream 32 is expanded 15 to an intermediate pressure 32a and then directed to a contact device 18, thereby forming a second vapor stream 37 and a second liquid stream 40; (4) said first liquid stream 33 is expanded 12 to said intermediate pressure 33a; (5) said second liquid stream 40b and said first expanded liquid stream 33b are directed to a distillation column 19 wherein said streams are separated into a more volatile steam distillation stream 42 and a relatively small fraction. less volatile 41 containing a major portion of the heavier hydrocarbon components; (6) said more volatile steam distillation stream 42 is sufficiently cooled to condense at least a portion thereof 42a, thereby forming a third vapor stream 43 and a third liquid stream 44; (7) a portion 46 of said third liquid stream 44 is directed to said distillation column 19 as a top feed thereto; (8) at least a portion of said first expanded steam stream 32a is intimately contacted with at least part of the remaining portion 45 of said third liquid stream 44 in said contact device 18; (9) said second vapor stream 37 is combined with said third vapor stream 43 to form a volatile fraction of residual gas 47 containing a major portion of said methane and lighter components; and (10) the volatile fraction of residual gas 47, 49c is cooled 60 under pressure to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10; (2) said natural gas stream 31a, 32 is expanded 15 to an intermediate pressure 32a and then directed to a contact device 18, thereby forming a volatile fraction of residual gas 37 containing a major portion of the said methane and lighter components and a first stream of liquid 40; (3) said first liquid stream 40a is heated 13 and then directed 40b to a distillation column 19 wherein said stream is separated into a more volatile steam distillation stream 42 and a relatively less fraction volatile 41 containing a major portion of the heavier hydrocarbon components; (4) said more volatile steam distillation stream 42 is cooled enough to condense at least a portion thereof, thereby forming a second liquid stream 42a; (5) at least a portion of said cooled expanded natural gas stream 32a is intimately contacted with at least part 45 of said second liquid stream 42a, 44a in said contact device 18; and (6) the volatile fraction of residual gas 37, 49c is cooled 60 under pressure to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thereby provide a vapor stream 32 and a first liquid stream 33; (3) said vapor stream 32 is expanded 15 to an intermediate pressure 32a and then directed to a contact device 18 thereby forming a volatile fraction of residual gas 37 containing a major portion of said methane and further components. light and a second stream of liquid 40; (4) said second liquid stream 40a is heated 13; (5) said first liquid stream 33 is expanded 12 to said intermediate pressure 33a; (6) said second heated liquid stream 40b and said first expanded liquid stream 33b are directed to a distillation column 19 wherein said streams are separated into a more volatile steam distillation stream 42 and a relatively less volatile fraction 41 containing a major portion of the heavier hydrocarbon components; (7) said more volatile steam distillation stream 42 is sufficiently cooled 13 to condense at least a portion thereof, thereby forming a third liquid stream 42a; (8) at least a portion of said expanded vapor stream 32a is intimately contacted with at least part 45 of said second liquid stream 44a in said contact device 18; and (9) said volatile waste gas fraction 37, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10; (2) said cooled natural gas stream 31a, 32 is expanded 15 to an intermediate pressure 32a and then directed to a contact device 18, thereby forming a first steam stream 37 and a first flow stream. liquid 40; (3) said first liquid stream 40a is heated 13 and then directed 40b to a distillation column 19 wherein said stream is separated into a more volatile steam distillation stream 42 and a relatively less fraction volatile 41 containing a major portion of the heavier hydrocarbon components; (4) said more volatile steam distillation stream 42 is cooled sufficiently to condense at least a portion thereof 42a, thereby forming a second steam stream 43 and a second liquid stream 44; (5) a portion 46 of said second liquid stream 44a is directed to said distillation column as a top feed thereto; (6) at least a portion of said first cooled expanded natural gas stream 32a is intimately contacted with at least part 45 of the remaining portion of said second liquid stream 44a in said contact device 18; (7) said first vapor stream 37 is combined with said second vapor stream 43 to form a volatile fraction of residual gas 47 containing a major portion of said methane and lighter components; and (8) said volatile waste gas fraction 37, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thus provide at least a first vapor stream 32 and a first liquid stream 33; (3) said first vapor stream 32 is expanded 15 to an intermediate pressure 32a and then directed to a contact device 18, thereby forming a second vapor stream 37 and a second liquid stream 40; (4) said second liquid stream 40a is heated 13; (5) said first liquid stream 33 is expanded 12 to said intermediate pressure 33a; (6) said second liquid stream 40b and said first expanded liquid stream 33b are directed to a distillation column 19 wherein said streams are separated into a more volatile steam distillation stream 42 and a relatively small fraction. less volatile 41 containing a major portion of the heavier hydrocarbon components; (7) said more volatile steam distillation stream 42 is sufficiently cooled 13 to condense at least a portion thereof 42a, thereby forming a third vapor stream 43 and a third liquid stream 44; (8) a portion 46 of said third liquid stream 44a is directed to said distillation column as a top feed thereto; (9) at least a portion of said first expanded steam stream 32a is intimately contacted with at least part 45 of the remaining portion of said third liquid stream 44a in said contact device 18; (10) said second vapor stream 37 is combined with said third vapor stream 43 to form a volatile fraction of residual gas 47 containing a major portion of said methane and lighter components; and (11) said volatile waste gas fraction 47, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10; (2) said cooled natural gas stream 31a, 32 is expanded 15 to an intermediate pressure 32a and then directed to a column medium feed position in a distillation column 19 wherein said stream is separated into a more volatile steam distillation stream 37 and a relatively less volatile fraction 41 containing a major portion of the heavier hydrocarbon components; (3) a steam distillation stream 42 is withdrawn from a region of said distillation column 19 below said expanded cooled natural gas stream 32a and sufficiently cooled 13 to condense at least a portion thereof 42a, thereby forming a vapor stream 43 and a liquid stream 44; (4) at least a portion of said cooled expanded natural gas stream 32a is intimately contacted with at least part of said liquid stream 44a, 45 in said distillation column 19; and (5) said vapor stream 43 is combined with said more volatile steam distillation stream 37 to form a volatile fraction of residual gas 47 containing a major portion of said methane and lighter components; and (6) said volatile fraction of waste gas 47, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to Claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thus provide at least a first vapor stream 32 and a first liquid stream 33; (3) said first vapor stream 32 and said first liquid stream 33 are expanded 15, 12 to an intermediate pressure 32a, 33a; (4) said first expanded vapor stream 32a and said first expanded liquid stream 32b are directed to column medium feed positions in a distillation column 19 wherein said streams are separated into a distillation stream. more volatile vapor 37 and a relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbons; (5) a steam distillation stream 42 is withdrawn from a region of said distillation column 19 below said first expanded steam stream 32a and sufficiently cooled 13 to condense at least part of it 42a, thereby forming a second stream. vapor 43 and a second stream of liquid 44; (6) at least a portion of said first expanded steam stream 32a is intimately contacted with at least part of said second liquid stream 44a, 45 in said distillation column 19; (7) said second vapor stream 43 is combined with said more volatile steam distillation stream 37 to form a volatile waste gas fraction 47 containing a major portion of said methane and lighter components; and (8) said volatile waste gas fraction 47, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10; (2) said cooled natural gas stream 31a, 32 is expanded 15 to an intermediate pressure 32a and then directed to a column medium feed position in a distillation column 19 wherein said stream is separated in a more volatile steam distillation stream 37 and a relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbons; (3) a steam distillation stream 42 is withdrawn from a region of said distillation column 19 below said expanded natural gas stream 32a and sufficiently cooled to condense at least part of it 42a, thereby forming a steam stream 43 and a liquid stream 44; (4) a portion 46 of said liquid stream 44a is supplied to said distillation column 19 as another feed to it at a feed location in substantially the same region as said steam distillation stream 42. is withdrawn; (5) at least a portion of said expanded chilled natural gas stream 32a is intimately contacted with at least part of the remaining portion 45 of said liquid stream 44a in said distillation column 19; (6) said vapor stream 43 is combined with said more volatile steam distillation stream 37 to form a volatile fraction of residual gas 47 containing a major portion of said methane and lighter components; and (7) said volatile waste gas fraction 47, 49c is pressure cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to Claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thereby provide a first vapor stream 32 and a first liquid stream 33; (3) said first vapor stream 32 and said first liquid stream 33 are expanded 15, 12 to an intermediate pressure 32a, 33a; (4) said first expanded vapor stream 32a and said first expanded liquid stream 33b are directed to column medium feed positions in a distillation column 19 wherein said streams are separated into a distillation stream. more volatile vapor 37 and a relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbons; (5) a steam distillation stream 42 is withdrawn from a region of said distillation column 19 below said first expanded steam stream 32a and sufficiently cooled 13 to condense at least part of it 42a, thereby forming a second stream. vapor 43 and a second stream of liquid 44; (6) a portion 46 of said second liquid stream 44a is provided to said distillation column 19 as another feed to it, at a feed location in substantially the same region as said steam distillation stream 42 is withdrawn; (7) at least a portion of said first expanded steam stream 32a is intimately contacted with at least part 45 of the remaining portion of said liquid stream 44a in said distillation column 19; (8) said second vapor stream 43 is combined with said more volatile steam distillation stream 37 to form a volatile waste gas fraction 47 containing a major portion of said methane and lighter components; and (9) said volatile fraction of waste gas 47, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10; (2) said cooled natural gas stream 31a, 32 is expanded 15 to an intermediate pressure 32a and then directed to a column medium feed position in a distillation column 19 wherein said stream is separated in a more volatile steam distillation stream 37 and a relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbons; (3) a steam distillation stream 42 is withdrawn from a region of said distillation column 19 below said expanded natural gas stream 32a and sufficiently cooled to condense at least part of it 42a, thereby forming a steam stream 43 and a liquid stream 44; (4) at least a portion of said cooled expanded natural gas stream 32a is intimately contacted with at least part of said liquid stream 44a in said distillation column 19; (5) a liquid distillation stream 40 is withdrawn from said distillation column 19 at a location above the region where said steam distillation stream 42 is withdrawn, as a result of which said liquid distillation stream 40 is heated 10 and then redirected 40a to said distillation column 19 as another feed thereto at a location below the region where said steam distillation stream 42 is withdrawn; (6) said vapor stream 43 is combined with said more volatile steam distillation stream 37 to form a volatile fraction of residual gas 47 containing a major portion of said methane and lighter components; and (7) said volatile waste gas fraction 47, 49c is pressure cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thereby provide a first vapor stream 32 and a first liquid stream 33; (3) said first vapor stream 32 and said first liquid stream 33 are expanded 15, 12 to an intermediate pressure 32a, 33a; (4) said first expanded vapor stream 32a and said first expanded liquid stream 33b are directed to column medium feed positions in a distillation column 19 wherein said streams are separated into a distillation stream. more volatile vapor 37 and a relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbons; (5) a steam distillation stream 42 is withdrawn from a region of said distillation column 19 below said first expanded steam stream 32a and sufficiently cooled 13 to condense at least part of it 42a, thereby forming a second stream. vapor 43 and a second stream of liquid 44; (6) at least a portion of said first expanded steam stream 32a is intimately contacted with at least part of said second liquid stream 44a in said distillation column 19; (7) a liquid distillation stream 40 is withdrawn from said distillation column 19 at a location above the region where said steam distillation stream 42 is withdrawn, as a result of which said liquid distillation stream 40 is heated 10 and then redirected to said distillation column 19 as another feed thereto at a location below the region where said steam distillation stream 42 is withdrawn; (8) said second vapor stream 43 is combined with said more volatile steam distillation stream 37 to form a volatile waste gas fraction 47 containing a major portion of said methane and lighter components; and (9) said volatile fraction of waste gas 47, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10; (2) said cooled natural gas stream 31a, 32 is expanded 15 to an intermediate pressure 32a and then directed to the column medium feed position in a distillation column 19 wherein said stream is separated into a more volatile steam distillation stream 37 and a relatively less volatile fraction 41 containing a major portion of the heavier hydrocarbon components; (3) a steam distillation stream 42 is withdrawn from a region of said distillation column 19 below said expanded cooled natural gas stream 32a and sufficiently cooled 13 to condense at least a portion thereof 42a, thereby forming a vapor stream 43 and a liquid stream 44; (4) a portion 46 of said liquid stream 44a is supplied to said distillation column 19 as another feed to it at a feed location in substantially the same region as distillation vapor stream 42 is provided. withdrawal; (5) at least a portion of said cooled expanded natural gas stream 32a is intimately contacted with at least part 45 of the remaining portion of said liquid stream 44a in said distillation column 19; (6) a liquid distillation stream 40 is withdrawn from said distillation column 19 at a location above the region where said steam distillation stream 42 is withdrawn, as a result of which said liquid distillation stream 40 is heated 10 and then redirected 40a to said distillation column 19 as another feed thereto at a location below the region where said steam distillation stream 42 is withdrawn; (7) said vapor stream 43 is combined with said more volatile steam distillation stream 37 to form a volatile fraction of residual gas 47 containing a major portion of said methane and lighter components; and (8) said volatile waste gas fraction 47, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thereby provide a first vapor stream 32 and a first liquid stream 33; (3) said first vapor stream 32 and said first liquid stream 33 are expanded 15, 12 to an intermediate pressure 32a, 33a; (4) said first expanded vapor stream 32a and said first expanded liquid stream 33b are directed to column medium feed positions in a distillation column 19 wherein said streams are separated into a distillation stream. more volatile vapor 37 and a relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbons; (5) a steam distillation stream 42 is withdrawn from a region of said distillation column 19 below said first expanded steam stream 32a and sufficiently cooled 13 to condense at least part of it 42a, thereby forming a second stream. of steam and a second stream of liquid 44; (6) a portion 46 of said second liquid stream 44a is provided to said distillation column 19 as another feed to it, at a feed location in substantially the same region as said steam distillation stream 42 is withdrawn; (7) at least a portion of said first expanded steam stream 32a is intimately contacted with at least part 45 of the remaining portion of said second liquid stream 44a in said distillation column 19; (8) a liquid distillation stream 40 is withdrawn from said distillation column 19 at a location above the region where said steam distillation stream 42 is withdrawn, as a result of which said liquid distillation stream 40 is heated 10 and then redirected 40a to said distillation column 19 as another feed thereto at a location below the region where said steam distillation stream 42 is withdrawn; (9) said second vapor stream 43 is combined with said more volatile steam distillation stream 37 to form a volatile waste gas fraction 47 containing a major portion of said methane and lighter components; and (10) said volatile waste gas fraction 47, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to Claim 1, characterized in that it consists essentially of processing steps (1) through (4) and does not include processing steps (4) and (5). Process according to Claim 1, characterized in that: (1) said natural gas stream 31 is treated in one or more cooling steps 10 to partially condense it 31a; (2) said partially condensed natural gas stream 31a is separated 11 to thus provide at least one vapor stream 32 and one liquid stream 33; (3) said vapor stream 32 is expanded 15 to an intermediate pressure 32a; (4) said liquid stream 33 is expanded 12 to said intermediate pressure 33a; (5) at least said expanded vapor stream 32a and said expanded liquid stream 33a are directed to a distillation column 19 wherein said streams are separated into a volatile fraction of residual gas 37 containing a major portion of said methane. and lighter components and a relatively less volatile fraction 41 containing a major portion of said heavier hydrocarbon components; and (6) said volatile waste gas fraction 37, 49c is pressure-cooled 60 to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to any one of Claims 1, 3 to 5, 11 to 26 and 28, characterized in that said volatile fraction of residual gas 37, 49 is compressed 16 and then cooled 60 under pressure. to condense at least a portion thereof and thereby form said condensed stream 49d. Process according to claim 2 or 7, characterized in that: (1) said volatile fraction of residual gas 37, 38 is compressed 16 and then cooled 60 under pressure to condense at least a portion thereof; and (2) said condensed portion is divided into at least two portions to thereby form said condensed stream 49 and said liquid stream 39. Process according to Claim 3, 8 or 9, characterized in that: (1) said volatile fraction of residual gas 37, 38 is compressed 16 and then cooled 60 under pressure to condense. at least a portion thereof; and (2) said condensed portion is divided into at least two portions to thereby form said condensed stream 49 and said second liquid stream 39. Process according to Claim 9, characterized in that said more volatile steam distillation stream 37, 38 is compressed 16 and then combined with said steam stream 42 to form said volatile fraction. waste gas 49 containing a major portion of said methane and lighter components. Process according to Claim 10, characterized in that said more volatile steam distillation stream 37, 38 is compressed 16 and then combined with said second steam stream 42 to form said fraction. volatile waste gas 49 containing a major portion of said methane and lighter components. Process according to any one of Claims 1, 3 to 5 and 11 to 26 and 28, characterized in that said volatile fraction of residual gas 37 is heated 24, compressed 16 and then cooled. 60 under pressure to condense at least a portion thereof and thereby form said condensed stream 49d. A process according to claim 1 or 6, characterized in that: (1) said volatile fraction of residual gas 37 is heated 24, compressed 16 and then cooled 60 under pressure to condense by least a portion thereof; and (2) said condensed portion is divided into at least two portions to thereby form said condensed stream 49 and said liquid stream 39. A process according to claim 2, 7 or 8, characterized in that: (1) said volatile fraction of waste gas 37 is heated 24, compressed 16 and then cooled 60 under pressure to condense at least a portion thereof; and (2) said condensed portion is divided into at least two portions to thereby form said condensed stream 49 and said second liquid stream 39. Process according to claim 9, characterized in that said more volatile steam distillation stream 37, 43 is heated 24, compressed 16, cooled 25, 24 and then combined with said flow stream. steam 42 to form said volatile fraction of residual gas 49 containing a major portion of said methane and lighter components. Process according to Claim 10, characterized in that said more volatile steam distillation stream 37, 43 is heated 24, compressed 16, cooled 25, 24 and then combined with said second stream. of vapor 42 to form said volatile waste gas fraction 49 containing a major portion of said methane and lighter components. Process according to any one of claims 1 to 28, 32, 33, 37 and 38, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C2 components. Process according to any one of claims 1 to 28, 32, 33, 37 and 38, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C2 components and C3 components. Process according to Claim 29, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C 2 components. Process according to claim 30, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C 2 components. Process according to Claim 31, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C 2 components. Process according to Claim 34, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C 2 components. Process according to Claim 35, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C 2 components. Process according to claim 36, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C 2 components. A process according to claim 29, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C2 components and C3 components. Process according to Claim 30, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C2 components and C3 components. Process according to claim 31, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C2 components and C3 components. Process according to Claim 34, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C2 components and C3 components. Process according to Claim 35, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C2 components and C3 components. Process according to Claim 36, characterized in that said volatile fraction of residual gas 37 contains a major portion of said methane, lighter components and C2 components and C3 components.