CN100449235C - natural gas liquefaction - Google Patents

Abstract

The present invention discloses a process for the liquefaction of natural gas (50) and the production of liquid streams comprising predominantly hydrocarbons heavier than methane (41). In the process, a liquefied natural gas stream (31 ) is partially cooled, expanded to intermediate pressure (14, 15), and supplied to a distillation column (19). The bottoms product (41) from the distillation column preferably contains a majority of any hydrocarbons heavier than methane which would reduce the purity of the liquefied natural gas (50). The residual gas stream (37) from the distillation column (19) is compressed to a higher intermediate pressure (16), cooled at pressure (60) to condense it, and then expanded to a lower pressure (61) to facilitate A liquefied natural gas stream is formed.

Description

natural gas liquefaction

technical field

The present invention relates to a process for processing natural gas or other methane-rich gas streams in order to produce liquefied natural gas (LNG) streams with high methane purity and liquid streams mainly comprising hydrocarbons heavier than methane. The applicant claims the benefit of an earlier U.S. Provisional Application No. 60/296,848, filed June 8, 2001, under § 35, United States Code, Section 119(e).

Background technique

Natural gas is usually obtained from wells drilled into underground reservoirs. Natural gas generally has a major proportion of methane, ie, methane comprises at least 50 mole percent of the natural gas. Natural gas also contains relatively small amounts of heavy hydrocarbons such as ethane, propane, butanes, pentane, etc., as well as water, hydrogen, nitrogen, carbon dioxide, and other gases, depending on the specific subterranean reservoir.

Most natural gas is handled in gaseous form. The most common way to transport natural gas from the wellhead to the natural gas processing plant and then to the natural gas user is in a high pressure gas transmission pipeline. In many cases, however, we have found it necessary and/or desirable to liquefy natural gas for transport or use. For example, in remote areas, there is often no pipeline infrastructure for the easy transportation of natural gas for sale. In such a case, the lower specific volume of LNG relative to gaseous natural gas can greatly reduce transportation costs since cargo ships and transport trucks can be used to transport LNG.

Another situation where natural gas liquefaction is desirable is when natural gas is used as a fuel for motor vehicles. In metropolitan areas, there are fleets of buses, taxis, and trucks that can be powered by LNG, if there is an economic source of LNG available. Due to the clean burning properties of natural gas, LNG fueled vehicles produce lower air pollution than comparable vehicles powered by gasoline and diesel engines that consume higher molecular weight hydrocarbons. Additionally, if the LNG is of high purity (i.e., has a methane purity of 95 mole percent or greater), the amount of carbon dioxide produced ( gases that cause the greenhouse effect) are lower.

The present invention generally relates to the liquefaction of natural gas and the production of hydrocarbons primarily comprising heavier than methane such as natural gas liquids (NGL) consisting of ethane, propane, butane-based and heavy hydrocarbon components, propane, butane-based and heavy hydrocarbon components Co-products of liquid streams consisting of liquefied petroleum gas (LPG) consisting of hydrocarbon components, or condensate consisting of butane-based and heavy hydrocarbon components). Generating a co-product liquid stream has two important benefits: the LNG produced is of high methane purity, and the co-product liquid is a useful product that can be applied to many other uses. A typical analysis of a natural gas stream processed in accordance with the present invention would be, in terms of approximate mole percentages, 84.2% methane, 7.9% ethane and other _C2 components, 4.9% propane and other _C3 components, 1.0% isobutane , 1.1% n-butane, 0.8% pentane plus the balance consisting of nitrogen and carbon dioxide. Sulfur containing gas is also sometimes present.

There are various known methods for liquefying natural gas. See, for example, Finn, Adrian J. , Grant L. Johnson, and Terry R. Tomlinson, "LNG Technology for Offshore and Mid-Scale Plants," and Proceedings of the 80th Annual Meeting of the American Natural Gas Processors Association, San Antonio, Texas, March 12-14, 2001, "Optimize the Power System of Baseload LNG Plant" by Kikkawa, Yoshitsugi, Masaaki Ohishi, and Noriyoshi Nozawa. No. 4,445,917; No. 4,525,185; No. 4,545,795; No. 4,755,200; No. 5,291,736; No. 6,014,869; No. 6,062,041; No. 6,119,479; No. 6,125,653; No. 6,250,105B1; No. 6,269,655B1; craft. These methods generally include the steps of natural gas being purified (by removing water and objectionable compounds such as carbon dioxide and sulfur compounds), cooling, condensing, and expanding. Cooling and condensation of natural gas can be achieved in a number of different ways. "Staging refrigeration" uses heat exchange between natural gas and several refrigerants with successively lower boiling points, such as propane, ethane, and methane. Alternatively, the heat exchange can be performed using a single refrigerant by evaporating one refrigerant at several different pressure levels. "Multicomponent refrigeration" uses heat exchange between natural gas and one or more refrigerant liquids consisting of several refrigerant components instead of multiple single component refrigerants. Expansion of natural gas can be performed both isenthalpically (eg using Joule Thomson expansion) and isentropically (eg using a working expansion turbine).

Regardless of the process used to liquefy a natural gas stream, removal of the majority of hydrocarbons heavier than methane is generally required prior to liquefaction of a methane-enriched gas stream. The reasons for this hydrocarbon removal step are various, including that the heating value of the LNG stream, and thus the values of these heavier hydrocarbon components as products, must be controlled within their own suitable ranges. Unfortunately, so far little attention has been focused on the effect of the hydrocarbon removal step.

Contents of the invention

In accordance with the present invention, we have discovered that careful integration of hydrocarbon removal steps into an LNG liquefaction process can produce both LNG and separated heavy hydrocarbon liquid products using less energy than prior art processes. Although applicable at lower pressures, the present invention is particularly advantageous when processing feed gas at pressures in the range of 400 to 1500 psig [2,758 to 10,342 kPa(a)] or higher.

The present invention provides a process for liquefying a natural gas stream comprising methane and heavy hydrocarbon components, wherein: (a) the natural gas stream is cooled under pressure so as to condense at least a portion thereof and form a condensed stream; and (b ) said condensed stream is expanded to a lower pressure to form a liquefied natural gas stream; said process is characterized by comprising the following processing steps: (1) said natural gas stream is processed in one or more cooling steps; (2) The cooled natural gas stream is expanded to an intermediate pressure; (3) the expanded cooled natural gas stream is directed to a distillation column where the expanded cooled natural gas stream is divided into and a volatile residual gas component of light components and a lower volatile component comprising most of said heavy hydrocarbon components; (4) said volatile residual gas component is cooled under pressure to facilitate condensation of at least its part and thereby form the condensed stream; (5) the condensed part is divided into at least two parts to constitute the condensed stream and the liquid stream; and (6) the liquid stream is introduced as its overhead feed to the distillation tower.

Description of drawings

For a better understanding of the invention, reference is made to the following examples and accompanying drawings, namely:

Fig. 1 is the flow chart of the natural gas liquefaction station suitable for the joint product of NGL involved in the present invention;

Fig. 2 is the pressure-enthalpy phase diagram of methane, is used to illustrate the present invention is better than the advantage of prior art process;

Fig. 3 is a flowchart of an alternative natural gas liquefaction station suitable for co-products of NGLs involved in the present invention;

Fig. 4 is a flowchart of an alternative natural gas liquefaction station suitable for LPG co-products according to the present invention;

Figure 5 is a flow diagram of an alternative natural gas liquefaction station suitable for co-products of condensate according to the present invention;

Figure 6 is a flow chart of an alternative natural gas liquefaction station suitable for co-products of liquid streams according to the present invention;

Fig. 7 is a flow chart of an alternative natural gas liquefaction station suitable for liquid flow co-products according to the present invention;

Fig. 8 is a flow chart of an alternative natural gas liquefaction station suitable for liquid flow co-products according to the present invention;

Fig. 9 is a flow chart of an alternative natural gas liquefaction station suitable for liquid flow co-products according to the present invention;

Figure 10 is a flow chart of an alternative natural gas liquefaction station suitable for co-products of liquid streams according to the present invention;

Figure 11 is a flow chart of an alternative natural gas liquefaction station suitable for co-products of liquid streams according to the present invention;

Figure 12 is a flow diagram of an alternative natural gas liquefaction station suitable for liquid stream co-products according to the present invention;

Figure 13 is a flow diagram of an alternative natural gas liquefaction station suitable for liquid stream co-products according to the present invention;

Figure 14 is a flow diagram of an alternative natural gas liquefaction station suitable for liquid stream co-products according to the present invention;

Figure 15 is a flow diagram of an alternative natural gas liquefaction station suitable for liquid stream co-products according to the present invention;

Figure 16 is a flow diagram of an alternative natural gas liquefaction station suitable for liquid stream co-products according to the present invention;

Figure 17 is a flow diagram of an alternative natural gas liquefaction station suitable for liquid stream co-products according to the present invention;

Figure 18 is a flow diagram of an alternative natural gas liquefaction station suitable for liquid stream co-products according to the present invention;

Figure 19 is a flow diagram of an alternative natural gas liquefaction station suitable for liquid stream co-products according to the present invention;

Figure 20 is a flow diagram of an alternative natural gas liquefaction station suitable for liquid stream co-products according to the present invention; and

Figure 21 is a flow diagram of an alternative natural gas liquefaction station suitable for liquid stream co-products according to the present invention.

Detailed ways

In the ensuing explanation of the above figures, a graph is provided in which the calculated flow rates for typical process operating conditions are summarized. In the graphs and graphs presented in the text, the numerical values of the flow rate (unit: mole/hour) have been rounded to the nearest whole number for the sake of convenience. The total flow rates shown in the charts include all non-hydrocarbon components and are therefore generally greater than the sum of the hydrocarbon component flow rates. Stated temperatures are approximate values rounded to the nearest degree. It should also be noted that for the purpose of comparing the processes shown in the figures, the process design calculations performed are based on the assumption that no heat leaks from the ambient into the process or from the process into the ambient. The quality of commercially available insulating materials makes this a very reasonable assumption, and one that would normally be made by a person of ordinary skill in the art.

For convenience, process parameters are reported in traditional British units and in International System of Units (SI) units. The molar flow rates given in the graphs can be considered as pounds molecules per hour or kilograms molecules per hour. Energy expenditure reported as horsepower (HP) and/or thousand British thermal units/hour (MBTU/Hr) corresponds to the specified molar flow rate in pounds molecules/hour. Energy consumption reported in kilowatts (kW) corresponds to the specified molar flow rate in kilomole/hour. Productivity reported in pounds per hour (Lb/Hr) corresponds to the stated molar flow rate in pounds molecules per hour. Productivity reported in kilograms per hour (kg/Hr) corresponds to the stated molar flow rate in kilomolecules/hour.

Example 1

Referring now to Figure 1, we begin with a description of the process to which the present invention relates in which it is desired to produce an NGL co-product comprising most of the heavy components of the methane and natural gas feedstream. In this simulated example of the invention, inlet gas enters the plant as stream 31 at 90°F [32°C] and 1285 psig [8,860 kPa(a)]. If the inlet gas contains carbon dioxide concentrations and/or sulfur compounds that would prevent the product stream from being in compliance, these compounds can be removed by appropriate pre-treatment of the feed gas (not shown). Additionally, feed streams are typically dehydrated to prevent hydroxide (ice) formation at low temperature conditions. Solid desiccants are usually used for this purpose.

Feed stream 31 is cooled in heat exchanger 10 by heat exchange with the refrigerant stream and the -68°F [-55°C] demethanizer column side reboiler liquid (stream 40). It should be noted that In all cases, heat exchanger 10 is represented as a plurality of individual heat exchangers or as a multi-pass heat exchanger, or any mixture thereof (the decision as to whether to use more than one heat exchanger as indicative of cooling maintenance will depend on the factors including (but not limited to) inlet gas flow rate, heat exchanger size, stream temperature, etc.). Cooling stream 31a enters separator 11 at -30°F [-34°C] and 1278 psi [8,812 kPa(a)] where steam (stream 32) is mixed with condensate ( Stream 33) phase separates.

The steam from separator 11 (stream 32 ) is split into two streams: 34 and 36 . Stream 34 , comprising approximately 20% of the total vapor, is combined with condensate, stream 33 , to form stream 35 . Combined stream 35 passes through heat exchanger 13 where it exchanges heat with refrigerant stream 71e to form cooled and condensed stream 35a. Fully condensed stream 35a at -120°F [-85°C] is then passed through a suitable expansion device ( Such as expansion valve 14) is rapidly expanded to the operating pressure of fractionation column 19 (approximately 465 psi [3,206 kPa(a)]). A portion of the flow is evaporated during expansion, resulting in cooling of the overall flow. In the process shown in Figure 1, the expanded stream 35b leaving the expansion valve 14 reaches a temperature of -122°F [-86°C] and is supplied at the midpoint feed point of the demethanizer section 19b of the fractionator 19 .

The remaining 80% of the steam from separator 11 (stream 36) goes to the working expansion machine 15 where the mechanical energy is extracted from this part of the high pressure delivery. The machine 15 expands the steam substantially isentropically from a pressure of 1278 psi [8,812 kPa(a)] to the column operating pressure, where the working expansion cools the expanded stream 36a to approximately -103°F [-75°C]. temperature. Commonly available commercial expansion agents are capable of recovering approximately 80-85% of the theoretically available work in ideal isentropic expansion. The recovered work is typically used to drive a centrifugal compressor (such as item 16), which can be used, for example, to recompress the overhead gas (stream 38). Expanded and partially condensed stream 36a is supplied as feed to distillation column 19 at the lower mid-column supply point.

The demethanizer in fractionation column 19 is a conventional distillation column comprising a plurality of vertically spaced trays, one or more packed beds, or some mix of trays and packages. As is often the case, in natural gas processing plants, fractionation columns can consist of two sections. The upper zone 19a is a separator in which the overhead feed is divided into respective vapor and liquid fractions, and in which vapor rising from the lower distillation or demethanizer zone 19b is combined with the vapor portion of the top feed, if any, so that A cooling demethanizer overhead vapor (stream 37) is formed which exits the column overhead at -135°F [-93°C]. The lower demethanizer zone 19b contains trays and/or packaging and provides the necessary contact between the downwardly falling liquid and the upwardly rising vapor. The lower demethanizer zone also includes one or more reboilers, such as reboiler 20, that heat and vaporize a portion of the liquid flowing down the column in order to provide stripping vapor that flows up the column. Liquid product stream 41 exits the bottom of the column at 115°F [46°C] based on the standard specification of a 0.020:1 ratio of methane to ethane in the bottoms product on a molar basis.

The demethanizer overhead steam (stream 37) is warmed to 90°F [32°C] in heat exchanger 24, and a portion of the warmed demethanizer overhead steam is withdrawn as plant gas fuel (stream 48 ) (primarily determined by the fuel used to drive the engines and/or turbines of the gas compressors in the plant, such as the refrigerant compressors 64, 66 and 68 in this example, the amount of gaseous fuel that must be extracted). Compressor 16, driven by expansion machines 15, 61 and 63, compresses the remaining demethanizer overhead vapor (stream 38). After stream 38b is cooled to 100°F [38°C] in discharge cooler 25, stream 38b is further cooled to -123° in heat exchanger 24 by cross-exchange with cooled demethanizer overhead steam, stream 37 F [-86°C].

Stream 38c then enters heat exchanger 60 and is further cooled by refrigerant stream 71d. After cooling to an intermediate temperature, stream 38c is split into two parts. The first portion, stream 49, is further cooled to -257°F [-160°C] in heat exchanger 60 to facilitate its condensation and supercooling, whereupon it enters working expansion machine 61 where to extract mechanical energy from the flow. The mechanism 61 expands the liquid stream 49 substantially isentropically from about 562 psi [3,878 kPa(a)] to the LNG storage pressure slightly above atmospheric pressure (15.5 psi [107 kPa(a)]). Working expansion cools expanded stream 49a to a temperature of approximately -258°F [-161°C], whereupon it is sent to LNG storage tank 62 for containing LNG product (stream 50).

Another substream 39 of stream 38c is withdrawn from heat exchanger 60 at -160°F [-107°C] and flash expanded to the operating pressure of fractionation column 19 by a suitable expansion device such as expansion valve 17 . In the process shown in Figure 1, there is no vaporization in expanded stream 39a, so its temperature drops only slightly to -161°F [-107°C] on exiting expansion valve 17. This expanded stream 39a is then supplied to the separation zone 19a in the upper region of the fractionation column 19 . The liquid separated therefrom becomes the overhead feed to the demethanizer zone 19b.

All cooling by convection 35 and 38c is provided by closed loop cooling loops. The working fluid for this cycle is a mixture of hydrocarbons and nitrogen, where the composition of the mixture must be adjusted to provide the required refrigerant temperature while condensing at a reasonable pressure with the available cooling medium. In this case, cooling water is assumed for condensation, so a refrigerant mixture consisting of nitrogen, methane, ethane, propane and heavy hydrocarbons is used in the simulated example of the Figure 1 process. In approximate mole percentages, the composition of the stream is: 7.5% nitrogen, 41.0% methane, 41.5% ethane, 10.0% propane, with the balance consisting of heavy hydrocarbons.

Refrigerant stream 71 exits discharge cooler 69 at 100°F [38°C] and 607 psig 4,185 kPa(a)]. Refrigerant stream 71 enters heat exchanger 10 and is cooled to -31°F [-35°C] and partially condensed by partially warmed expanding refrigerant stream 71f and other refrigerant streams. For the simulated example in Figure 1, it has been assumed that the other refrigerant stream is industrial grade quality propane refrigerant at three different temperature and pressure stages. Partially condensed refrigerant stream 71a then enters heat exchanger 13 for further cooling to -114°F [-81°C] by partially warmed expanded refrigerant stream 71f such that the refrigerant (stream 71b) Condensed and partially supercooled. The refrigerant is further subcooled in heat exchanger 60 to -257°F [-160°C] by expanded refrigerant stream 71d. The supercooled liquid stream 71c enters the working expansion machine 63 as the stream expands substantially isentropically from a pressure of about 586 psi [4,040 kPa(a)] to a pressure of about 34 psi [234 kPa(a) )], mechanical energy is extracted from the flow in the working expansion machine 63. During expansion, a portion of the stream is evaporated, causing the total stream to cool to -263°F [-164°C] (stream 71d). Expanded stream 71d then re-enters heat exchangers 60, 13 and 10 where cooling is provided to stream 38c, stream 35 and the refrigerant (streams 71, 71a and 71b) as expanded stream 71d is evaporated and overheated.

Superheated refrigerant vapor (stream 71 g) exits heat exchanger 10 at 93°F [34°C] and is compressed to 617 psi [4,254 kPa(a)] in three stages. Each of the three compression stages (refrigerant compressors 64, 66 and 68) is driven by auxiliary power and is followed by coolers (discharge coolers 65, 67 and 69) to remove the heat of compression. Compressed stream 71 from discharge cooler 69 is returned to heat exchanger 10 to complete the cycle.

The sum of the stream flow rates and energy consumption for the process shown in Figure 1 is listed in the chart below:

Chart I

(figure 1)

Flow Total-Lb.Moles/Hr[kg moles/Hr]

Stream methane ethane propane butane total number

31 40,977 3,861 2,408 1,404 48,656

32 32,360 2,675 1,469 701 37,209

33 8,617 1,186 939 703 11,447

34 6,472 535 294 140 7,442

36 25,888 2,140 1,175 561 29,767

37 47,771 223 0 0 48,000

39 6,867 32 0 0 6,900

41 73 3,670 2,408 1,404 7,556

48 3,168 15 0 0 0 3,184

50 37,736 176 0 0 37,916

NGL ^* Recovery rate in

Ethane 95.06%

Propane 100.00%

Butane series 100.00%

Productivity 308,147 Lb/Hr [308,147kg/Hr]

LNG products

Productivity 610,813 Lb/Hr [610,813kg/Hr]

Purity ^* 99.52%

Lower heating value 912.3 BTU/SCF [33.99Mj/m ³ ]

power

Refrigerant Compression 103,957 HP [170,904kW]

Propane Compression 33,815 HP [55,591kW]

Total compression 137,772 HP [226,495kW]

Effective calories

Demethanizer reboiler 29,364 MBTU/Hr [18,969kW]

^* (Based on flow rates without rounding)

The efficiency of an LNG production process is often compared using the required "specific power consumption", which is the ratio between the total refrigeration compression power and the total liquid production rate. The published information on the unit power consumption of LNG produced by the prior art process is expressed as between 0.168Hp-Hr/Lb [0.276kW-Hr/kg] to 0.182Hp-Hr/Lb [0.300kW-Hr/kg] The range is generally considered to be based on the fact that LNG product plants are online 340 days a year. Based on the same basis, the unit power consumption of the embodiment in FIG. 1 of the present invention is 0.161Hp-Hr/Lb [0.265kW-Hr/kg], which has an efficiency improvement of 4-13% over the prior art process. Also, it should be noted that the specific electricity consumption of the prior art process is based on the co-production of LPG ( _C3 and heavy hydrocarbons) or condensate ( _C4 and heavy hydrocarbons) only at lower recovery levels This is done on the basis of the co-production NGL ( _C2 and heavy hydrocarbons) stream as shown in this example of the invention. Prior art processes require relatively more refrigeration power if they co-produce an NGL stream rather than an LPG stream or a condensate stream.

There are two main factors that account for the increased efficiency of the present invention. The first factor can be understood by examining the thermodynamics of the liquefaction process when a high pressure gas flow such as that considered in this example is applied. Since the main component of this stream is methane, the thermodynamic properties of methane can be used for the purpose of comparing the liquefaction cycle used in the prior art process with the cycle used in the present invention. Figure 2 contains the pressure-enthalpy phase diagram for methane. In most prior art liquefaction cycles, all cooling of the gas stream is performed when the gas stream is at high pressure (path A-B) and then expanded (path B-C) to the pressure of the LNG storage vessel (slightly above atmospheric pressure). This expansion step may use a working expansion machine that is theoretically typically capable of recovering 75-80% of the work available in an ideal isentropic expansion. For purposes of simplicity, fully isentropic expansion is shown in Figure 2 for paths B-C. Even so, the enthalpy reduction provided by this working expansion is still rather small, since the isentropic lines in the liquid region of the phase diagram are approximately vertical.

A comparison is now made with the liquefaction cycle of the present invention. After partial cooling at high pressure (path A-A'), the gas stream is worked to expand (path A'-A") to an intermediate pressure (full isentropic expansion is shown for simplicity). At intermediate pressure (path A "-B') for the remainder of the cooling, after which the flow is expanded (path B'-C) to the pressure of the LNG storage vessel. The first working expansion step (path A'-A") of the present invention provides a significantly increased enthalpy reduction due to the less steep slope of the isentropic lines in the vapor region of the phase diagram. Therefore, the cooling required by the present invention The total amount (path A-A' combined with path A"-B') is less than the cooling required by the prior art process (path A-B), reducing the refrigeration (and thus refrigeration compression) required for the liquefied gas stream.

The second factor involved in the improved efficiency of the present invention is the superior performance of the hydrocarbon distillation system at lower operating pressures. The hydrocarbon removal step in most prior art processes is performed at high pressure, typically using scrubbers that use cold hydrocarbon liquid as an absorbent stream to remove heavy hydrocarbons from the incoming gas stream. Operating the scrubber at high pressure is not very efficient because it results in the co-absorption of most of the methane and ethane from the gas stream, which must then be removed from the absorbent liquid and cooled to LNG product part. In the present invention, the hydrocarbon removal step is carried out at intermediate pressures where the gas-liquid equilibrium is more favorable resulting in a very efficient recovery of the desired heavy hydrocarbons in the co-product liquid stream.

Example 2

If the specification of the LNG product allows more of the ethane contained in the feed gas to be recovered to the LNG product, then a more simplified embodiment of the invention can be employed. Figure 3 shows such an alternative embodiment. The inlet gas composition and conditions considered in the process presented in FIG. 3 are the same as in FIG. 1 . Accordingly, the process of FIG. 3 is comparable to the embodiment set forth in FIG. 1 .

In the simulated example of the Figure 3 process, the inlet gas cooling, separation and expansion shown in the NGL recovery zone are essentially the same as those used in Figure 1 . Inlet gas enters the plant as stream 31 at 90°F [32°C] and 1285 psi [8,860 kPa(a)] and passes in heat exchanger 10 with refrigerant flow and -35°F [- 37°C] The demethanizer side reboiler liquid (stream 40) is cooled by heat exchange. Cooling stream 31a enters separator 11 at -30°F [-34°C] and 1278 psi [8,812 kPa(a)] where steam (stream 32) is mixed with condensate ( Stream 33) phase separates.

The steam from separator 11 (stream 32 ) is split into two streams: 34 and 36 . Stream 34 , comprising approximately 20% of the total vapor, is combined with condensate, stream 33 , to form stream 35 . Combined stream 35 passes through heat exchanger 13 where it exchanges heat with refrigerant stream 71e to form cooled and substantially condensed stream 35a. Fully condensed stream 35a at -120°F [-85°C] is then flash-expanded to the operating pressure of fractionation column 19 (approximately 465 psi [3,206 kPa(a)] through a suitable expansion device, such as expansion valve 14. ). A portion of the flow is evaporated during expansion, resulting in cooling of the overall flow. In the process shown in FIG. 3 , expanded stream 35b leaving expansion valve 14 reaches a temperature of -122°F [-86°C] and is supplied to the separator zone in the upper zone of fractionation column 19 . The liquid separated there becomes the top feed to the demethanizer zone 19b in the lower zone of the fractionation column 19 .

The remaining 80% of the steam from separator 11 (stream 36) goes to the working expansion machine 15 where the mechanical energy is extracted from this part of the high pressure delivery. The machine 15 expands the steam substantially isentropically from a pressure of 1278 psi [8,812 kPa(a)] to the column operating pressure, where the working expansion cools the expanded stream 36a to approximately -103°F [-75°C]. temperature. The expanded and partially condensed stream 36a is supplied as feedstock to the distillation column 19 at the mid-column supply point.

Cooled demethanizer overhead vapor (stream 37) exits the column overhead at -123°F [-86°C]. Liquid product stream 41 exits the bottom of the column at 118°F [48°C] based on the standard specification of 0.020:1 methane to ethane ratio in the bottoms product on a molar basis.

The demethanizer overhead vapor (stream 37) is warmed to 90°F [32°C] in heat exchanger 24, and a portion of the warmed demethanizer overhead vapor (stream 48) is withdrawn as plant gas fuel. After stream 49b is cooled to 100°F [38°C] in discharge cooler 25, stream 49b is further cooled to -112°C in heat exchanger 24 by cross-exchange with cooling demethanizer overhead steam, stream 37. °F [-80°C].

Stream 49c then enters heat exchanger 60 and is further cooled to -257°F [-160°C] by refrigerant stream 71d to facilitate its condensation and supercooling, whereupon it enters working expansion machine 61 where Mechanical energy is extracted from the flow in the working expansion machine 61. The mechanism 61 expands the liquid stream 49d substantially isentropically from a pressure of about 583 psi [4,021 kPa(a)] to the LNG storage pressure slightly above atmospheric pressure (15.5 psi [107 kPa(a)]). Working expansion cools the expanded stream 49e to a temperature of approximately -258°F [-161°C], whereupon the expanded stream 49e is sent to the LNG storage tank 62 for containing the LNG product (stream 50).

Similar to the Figure 1 process, all cooling by convection 35 and 49c is provided by closed loop cooling loops. In approximate mole percentages, the composition of the working fluid stream used in the cycle of the Figure 3 process is: 7.5% nitrogen, 40.0% methane, 42.5% ethane, 10.0% propane, with the balance consisting of heavy hydrocarbons . Refrigerant stream 71 exits discharge cooler 69 at 100°F [38°C] and 607 psig 4,185 kPa(a)]. Refrigerant stream 71 enters heat exchanger 10 and is cooled to -31°F [-35°C] and partially condensed by partially warmed expanding refrigerant stream 71f and other refrigerant streams. For the simulated example in Figure 3, it has been assumed that the other refrigerant stream is industrial grade quality propane refrigerant at three different temperature and pressure stages. Partially condensed refrigerant stream 71a then enters heat exchanger 13 for further cooling to -121°F [-85°C] by partially warmed expanded refrigerant stream 71e such that the refrigerant (stream 71b) Condensed and partially supercooled. The refrigerant is further subcooled in heat exchanger 60 to -257°F [-160°C] by expanded refrigerant stream 71d. The supercooled liquid stream 71c enters the working expansion machine 63 as the stream expands substantially isentropically from a pressure of about 586 psi [4,040 kPa(a)] to a pressure of about 34 psi [234 kPa(a) )], mechanical energy is extracted from the flow in the working expansion machine 63. During expansion, a portion of the stream is evaporated, causing the total stream to cool to -263°F [-164°C] (stream 71d). Expanded stream 71d then re-enters heat exchangers 60, 13 and 10 where cooling is provided to stream 49c, stream 35 and the refrigerant (streams 71, 71a and 71b) as expanded stream 71d is evaporated and overheated.

Superheated refrigerant vapor (stream 71 g) exits heat exchanger 10 at 93°F [34°C] and is compressed to 617 psi [4,254 kPa(a)] in three stages. Each of the three compression stages (refrigerant compressors 64, 66 and 68) is driven by auxiliary power and is followed by coolers (discharge coolers 65, 67 and 69) to remove the heat of compression. Compressed stream 71 from discharge cooler 69 is returned to heat exchanger 10 to complete the cycle.

The sum of the stream flow rates and energy consumption for the process shown in Figure 3 is listed in the chart below:

Chart II

(image 3)

Flow Total-Lb.Moles/Hr[kg moles/Hr]

Stream methane ethane propane butane total number

31 40,977 3,861 2,408 1,404 48,656

32 32,360 2,675 1,469 701 37,209

33 8,617 1,186 939 703 11,447

34 6,472 535 294 140 7,442

36 25,888 2,140 1,175 561 29,767

37 40,910 480 62 7 41,465

41 67 3,381 2,346 1,397 7,191

48 2,969 35 4 0 3,009

50 37,941 445 58 7 38,456

NGL ^* Recovery rate in

Ethane 87.57%

Propane 97.41%

Butane series 99.47%

Productivity 296,175 Lb/Hr [296,175kg/Hr]

LNG products

Productivity 625,152 Lb/Hr [625,152kg/Hr]

Purity ^* 98.66%

Lower calorific value 919.7 BTU/SCF [34.27Mj/m ³ ]

power

Refrigerant Compression 96,560 HP [158,743kW]

Propane Compression 34,724 HP [57,086kW]

Total compression 131,284 HP [215,829kW]

Effective calories

Demethanizer Reboiler 22,117 MBTU/Hr [14,326kW]

^* (Based on flow rates without rounding)

Assuming that the LNG product factory produces on-line 340 days a year, the unit power consumption of the embodiment of Fig. 3 of the present invention is 0.153Hp-Hr/Lb [0.251kW-Hr/kg]. Compared with the prior art process, the implementation of Fig. 3 Example efficiency improvement is 10-20%. As previously noted for the Figure 1 embodiment, this efficiency increase is possible even when the present invention produces an NGL co-product rather than the LPG (or condensate co-product) produced by the prior art.

Compared with the embodiment of FIG. 1, the embodiment of FIG. 3 of the present invention requires approximately 5% less power per unit of liquid produced. Thus, for a given amount of available compression power, the Figure 3 embodiment can liquefy approximately 5% more natural gas than the Figure 1 embodiment by taking advantage of recovering less C2 _and heavier hydrocarbons in the NGL co-product. The choice between the embodiments of Fig. 1 and Fig. 3 of the present invention for a specific application is generally governed by the ratio of the economic value of the heavy hydrocarbons in the NGL product to their corresponding value in the LNG product or by the calorific value specification of the LNG product (Because the calorific value of the LNG produced by the embodiment of Fig. 1 is lower than the calorific value of the LNG produced by the embodiment of Fig. 3) dominate.

Example 3

If the specification of the LNG product allows all of the ethane contained in the feed gas to be recovered into the LNG product, or if there is no market for a liquid co-product containing ethane, then an invention such as that shown in Figure 4 can be used Alternative embodiments for the production of LPG co-product streams. The inlet gas composition and conditions considered in the process presented in FIG. 4 are the same as those in FIGS. 1 and 3 . Therefore, the process of FIG. 4 can be compared with the embodiments listed in FIGS. 1 and 3 .

In the simulated example of the Figure 4 process, inlet gas enters the plant as stream 31 at 90°F [32°C] and 1285 psig [8,860kPa(a)] and passes in heat exchanger 10 with refrigeration The agent stream is cooled by heat exchange with the -46°F [-43°C] flash separator liquid (stream 33a). Cooling stream 31a enters separator 11 at -1°F [-18°C] and 1278 psi [8,812 kPa(a)] where steam (stream 32) is mixed with condensate ( Stream 33) phase separates.

The steam (stream 32 ) from separator 11 enters the working expansion machine 15 where the mechanical energy is extracted from this part of the high pressure delivery. The machine 15 expands the steam substantially isentropically from a pressure of 1278 psi [8,812 kPa(a)] to a pressure of approximately 440 psi [3,034 kPa(a)] (operation of the separator/absorber column 18 pressure), where the working expansion cools the expanded stream 32a to a temperature of approximately -81°F [-63°C]. The expanded and partially condensed stream 32a is supplied to the absorption zone 18b in the lower region of the separator/absorber column 18 . The liquid portion of the expanded stream mixes with the liquid falling from the absorption zone and this mixed liquid stream 40 exits the bottom of the separator/absorber column 18 at -86°F [-66°C]. The vapor portion of the expanded stream rises upward through the absorption zone and contacts falling cooling liquid to facilitate condensation and absorption of _C3 components and heavies.

Separator/absorber column 18 is a conventional distillation column comprising a plurality of vertically spaced trays, one or more packed beds, or some mix of trays and packaging. As is generally the case, in a natural gas processing plant, the separator/absorber column 18 may be constructed in two sections. The upper zone 18a is a separator in which all the vapor fraction contained in the top feed is separated from its corresponding liquid fraction, and in which the vapor rising from the lower distillation or absorption zone 18b is separated from the vapor fraction (if any) of the top feed ) are combined to form a cooled distillate stream 37, which is withdrawn from the top of the column. The lower absorption zone 18b contains trays and/or packaging and provides the necessary contact between downwardly falling liquid and upwardly rising vapor to facilitate condensation and absorption of _C3 components and heavies.

Mixed liquid stream 40 from the bottom of separator/absorber column 18 is sent by pump 26 to heat exchanger 13 where it provides for deethanizer overhead (stream 42) and refrigeration As the agent (stream 71a) cools it (stream 40a) is heated. The combined liquid stream is heated to -24°F [-31°C], partially vaporized stream 40b, before being supplied to the deethanizer 19 as mid column feed. Expansion valve 12 flash expands the separator liquid (stream 33) to slightly above the operating pressure of deethanizer 19, cooling stream 33 to -46°F before it provides cooling for the incoming feed gas as described above [-43°C] (stream 33a). Stream 33b at 85°F [29°C] enters deethanizer 19 at the low and mid column feed point. In deethanizer 19, methane and _C2 components are removed from streams 40b and 33b. The deethanizer in column 19, which operates at about 453 psig [3,123 kPa(a)], is also comprised of multiple vertically spaced trays, one or more packed beds, or some combination of trays and packages Traditional distillation tower. The deethanizer can also be composed of two sections: an upper zone 19a, in which all the vapor fraction contained in the overhead feed is separated from its corresponding liquid fraction, and in which the vapor rising from the lower distillation or deethanizer zone 19b is separated from the top The vapor portion of the feed, if any, combines to facilitate the formation of distillation stream 42, which exits overhead; and a lower deethanizer zone 19b comprising trays and/or packaging to provide the The necessary contact between a liquid and an upwardly rising vapor. Demethanization zone 19b also includes one or more reboilers, such as reboiler 20, which heat and vaporize a portion of the liquid at the bottom of the column in order to provide flow up the column to remove methane and _C2 components. Liquid product, stripping steam of stream 41. The standard specification for the bottoms liquid product has a ratio of methane to ethane of 0.020:1 on a molar basis. Liquid product stream 41 exits the bottom of the demethanizer at 214°F [101°C].

The operating pressure in demethanizer column 19 is maintained slightly higher than the operating pressure of separator/absorber column 18 . This allows the deethanizer overhead vapor (stream 42 ) to flow under pressure through heat exchanger 13 and into the upper region of separator/absorber column 18 . In heat exchanger 13, the deethanizer overhead at -19°F [-28°C] is in heat exchange relationship with the mixed liquid stream (stream 40a) from the bottom of separator/absorber column 18 and spreads rapidly throughout Refrigerant stream 71e, which is cooled to -89°F [-67°C] (stream 42a) and partially condensed. The partially condensed stream enters the reflux drum 22 where the condensed liquid (stream 44) is separated from the non-condensable vapor (stream 43). Stream 43 is combined with the distillation vapor stream (stream 37 ) leaving the upper region of separator/absorber column 18 to form cooled residual gas stream 47 . Pump 23 pumps the condensate (stream 44) to a higher pressure so that stream 44a is split into two parts. A portion, stream 45, is sent to the upper separator region of separator/absorber column 18 to serve as cooling liquid in contact with the vapor rising upward through the absorption region. The other portion is supplied to deethanizer 19 as counter stream 46 to the overhead feed point on demethanizer 19 at -89°F [-67°C].

The cooled residual gas (stream 47 ) is warmed from -94°F [-70°C] to 94°F [32°C] in heat exchanger 24 and a portion (stream 48 ) is withdrawn as gaseous fuel for the plant. The remainder of the warmed residual gas (stream 49 ) is compressed by compressor 16 . After stream 49b is cooled to 100°F [38°C] in discharge cooler 25, stream 49b is further cooled to -78°F [- 61°C].

Stream 49c then enters heat exchanger 60 and is further cooled to -255°F [-160°C] by refrigerant stream 71d to facilitate its condensation and supercooling, whereupon it enters working expansion machine 61 where Mechanical energy is extracted from the flow in the working expansion machine 61. The mechanism 61 expands the liquid stream 49d substantially isentropically from a pressure of about 648 psi [4,465 kPa(a)] to the LNG storage pressure slightly above atmospheric pressure (15.5 psi [107 kPa(a)]). Working expansion cools the expanded stream 49e to a temperature of approximately -256°F [-160°C], whereupon said expanded stream 49e is sent to LNG storage tank 62 for containing LNG product (stream 50).

Much of the cooling by convection 42 and all of the cooling by convection 49c is provided by closed loop cooling loops similar to the processes of Figures 1 and 3 . In approximate mole percentages, the composition of the working fluid stream used in the cycle of the Figure 4 process is: 8.7% nitrogen, 30.0% methane, 45.8% ethane, 11.0% propane, with the balance consisting of heavy hydrocarbons . Refrigerant stream 71 exits discharge cooler 69 at 100°F [38°C] and 607 psig 4,185 kPa(a)]. Refrigerant stream 71 enters heat exchanger 10 and is cooled to -17°F [-27°C] and partially condensed by partially warmed expanding refrigerant stream 71f and other refrigerant streams. For the simulated example in Figure 4, it has been assumed that the other refrigerant stream is industrial grade quality propane refrigerant at three different temperature and pressure stages. Partially condensed refrigerant stream 71a then enters heat exchanger 13 for further cooling to -89°F [-67°C] by partially warmed expanded refrigerant stream 71e, further condensing the refrigerant (stream 71b ). The refrigerant is completely condensed in heat exchanger 60 by expanded refrigerant stream 71d and then further subcooled to -255°F [-160°C]. The supercooled liquid stream 71c enters the working expansion machine 63 as the stream expands substantially isentropically from a pressure of about 586 psi [4,040 kPa(a)] to a pressure of about 34 psi [234 kPa(a) )], mechanical energy is extracted from the flow in the working expansion machine 63. During expansion, a portion of the stream is evaporated, causing the total stream to cool to -264°F [-164°C] (stream 71d). Expanded stream 71d then re-enters heat exchangers 60, 13 and 10 where cooling is provided to stream 49c, stream 42 and the refrigerant (streams 71, 71a and 71b) as expanded stream 71d is evaporated and overheated.

Superheated refrigerant vapor (stream 71 g) exits heat exchanger 10 at 90°F [32°C] and is compressed to 617 psi [4,254 kPa(a)] in three stages. Each of the three compression stages (refrigerant compressors 64, 66 and 68) is driven by auxiliary power and is followed by coolers (discharge coolers 65, 67 and 69) to remove the heat of compression. Compressed stream 71 from discharge cooler 69 is returned to heat exchanger 10 to complete the cycle.

The sum of the stream flow rates and energy consumption for the process shown in Figure 4 is listed in the chart below:

Chart III

(Figure 4)

Flow Total-Lb.Moles/Hr[kg moles/Hr]

Stream methane ethane propane butane total number

31 40,977 3,861 2,408 1,404 48,656

32 38,431 3,317 1,832 820 44,405

33 2,546 544 576 584 4,251

37 36,692 3,350 19 0 40,066

40 5,324 3,386 1,910 820 11,440

41 0 48 2,386 1,404 3,837

42 10,361 6,258 168 0 16,789

43 4,285 463 3 0 4,753

44 6,076 5,795 165 0 12,036

45 3,585 3,419 97 0 7,101

46 2,491 2,376 68 0 4,935

47 40,977 3,813 22 0 44,819

48 2,453 228 1 0 2,684

50 38,524 3,585 21 0 42,135

LPG ^* Recovery rate in

Propane 99.08%

Butane series 100.00%

Productivity 197,051 Lb/Hr [197,051kg/Hr]

LNG products

Productivity 726,918 Lb/Hr [726,918kg/Hr]

Purity ^* 91.43%

Lower heating value 969.9 BTU/SCF [36.14Mj/m ³ ]

power

Refrigerant Compression 95,424 HP [156,876kW]

Propane Compression 28,060 HP [46,130kW]

Total compression 123,484 HP [203,006kW]

Effective calories

Demethanizer Reboiler 55,070 MBTU/Hr [35,575kW]

^* (Based on flow rates without rounding)

Assuming that the LNG product factory is on-line for 340 days per year, the unit power consumption of the embodiment in Figure 4 of the present invention is 0.143Hp-Hr/Lb [0.236kW-Hr/kg]. Compared with the prior art process, the efficiency improvement of the embodiment of Fig. 4 is 17-27%.

Compared with the embodiment of Fig. 1 and Fig. 3, the embodiment of Fig. 4 of the present invention requires about 6% to 11% less power per unit of liquid produced. Thus, for a given amount of available compression power, the Figure 4 embodiment can liquefy approximately 6% more natural gas than the Figure ₁ embodiment or the Figure 3 implementation by taking advantage of recovering only C3 and heavier hydrocarbons as LPG co-products For example, about 11% of natural gas is liquefied. The choice between the Figure 4 and Figure 1 or 3 embodiments of the invention for a particular application is generally governed by the ratio of the economic value of ethane as part of the NGL product to their corresponding value in the LNG product or by the calorific value of the LNG product Specifications (since the calorific value of the LNG produced by the embodiment of Fig. 1 and Fig. 3 is lower than that of the LNG produced by the embodiment of Fig. 4) dictate.

Example 4

If the specification of the LNG product allows for all of the ethane and propane contained in the feed gas to be recovered into the LNG product, or if there is no market for liquid co-products containing ethane and propane, then a method such as that shown in Figure 5 can be used. An alternative embodiment of the invention is shown to facilitate the production of a condensate co-product stream. The inlet gas composition and conditions considered in the process presented in FIG. 5 are the same as those in FIGS. 1 , 3 and 4 . Therefore, the process of FIG. 5 can be compared with the embodiments listed in FIGS. 1 , 3 and 4 .

In the simulated example of the Figure 5 process, inlet gas enters the plant as stream 31 at 90°F [32°C] and 1285 psi [8,860kPa(a)] and The agent stream, the -37°F [-38°C] fire high pressure separator liquid (stream 33b), and the -37°F [-38°C] fire intermediate pressure separator liquid (stream 39b) are cooled by heat exchange. Cooling stream 31a enters high pressure separator 11 at -30°F [-34°C] and 1278 psig [8,812 kPa(a)] where steam (stream 32 ) is mixed with condensate (stream 33) Phase separation.

The steam (stream 32 ) from the high pressure separator 11 enters the working expansion machine 15 where the mechanical energy is extracted from this part of the high pressure delivery. The machine 15 expands the steam substantially isentropically from a pressure of 1278 psi [8,812 kPa(a)] to a pressure of approximately 635 psi [4,378 kPa(a)], wherein the working expansion cools the expanded stream 32a to A temperature of approximately -83°F [-64°C]. The expanded and partially condensed stream 32a enters the intermediate pressure separator 18 where the vapor (stream 42) is separated from the condensate (stream 39). Expansion valve 17 fires expansion of the intermediate pressure separator liquid (stream 39) to slightly above the operating pressure of deethanizer 19 before it enters heat exchanger 13 and before it provides support for residual gas stream 49 and refrigerant stream Cooling at 71a cools stream 39 to -108°F [-78°C] (stream 39a ) prior to cooling and thus into heat exchanger 10 to provide cooling of the incoming feed gas as described above. The -15°F [-26°C] stream 39c then enters the deethanizer 19 at the upper mid column feed point.

Expansion valve 12 causes the condensate from high pressure separator 11, i.e., stream 33, to be flared to slightly above the operating pressure of deethanizer 19 before it enters heat exchanger 13 and before it is provided for residual gas stream 49 Cooling of stream 33 to -93°F [-70°C] (stream 33a ) is preceded by cooling of refrigerant stream 71a and thus into heat exchanger 10 to provide cooling of the incoming feed gas as described above. The 50°F [10°C] stream 33c then enters the deethanizer 19 at the lower mid column feed point. In the deethanizer, streams 39c and 33c are stripped of their methane, _C2 components, and _C3 components. Deethanizer 19, operating at approximately 385 psig [2,654 kPa(a)], is a conventional distillation column comprising a plurality of vertically spaced trays, one or more packed beds, or some mixture of trays and packages . The deethanizer can also consist of two sections: an upper zone 19a, in which any vapor contained in the top feed is separated from its corresponding liquid fraction, and in which vapor rising from the lower distillation or deethanizer zone 19b is separated from the top feed The vapor portion (if any) of the combined to facilitate the formation of distillation stream 37, the distillation stream 37 is withdrawn from the top of the column; Necessary contact with upwardly rising steam. Demethanization zone 19b also includes one or more reboilers, such as reboiler 20, which heat and vaporize a portion of the liquid at the bottom of the column in order to provide upward flow to the column for removal of methane, _C2 components, and The liquid product of the _C3 component, the stripping vapor of stream 41. The standard specification for the bottoms liquid product has a propane to butane ratio of 0.020:1 on a molar basis. Liquid product stream 41 exits the bottom of the demethanizer at 286°F [141°C].

The overhead distillate stream 37 exits the deethanizer 19 at 36°F [2°C] and is cooled and partially condensed by commercial grade quality propane refrigerant in the reflux condenser. Partially condensed stream 37a enters reflux drum 22 at 2°F [-17°C] where condensate (stream 44) is separated from non-condensable vapor (stream 43). Pump 23 pumps the condensate (stream 44) to the overhead feed point on demethanizer 19 as counter stream 44a.

The non-condensable vapor (stream 43)] from the reflux drum 22 is warmed to 94°F [34°C] in heat exchanger 24 and a portion (stream 48) is withdrawn as gaseous fuel for the plant. The remainder of the warmed vapor (stream 38 ) is compressed by compressor 16 . After stream 38b is cooled to 100°F [38°C] in discharge cooler 25, stream 38b is further cooled to 15°F [-9°C] in heat exchanger 24 by cross-exchange with cooling steam, stream 43 ].

Stream 38c is then mixed with intermediate pressure separator vapor (stream 42 ) to form cooled residual gas stream 49 . Stream 49 enters heat exchanger 13 and is cooled from -38°F [-39°C] to -102°F [-74°C] by separator liquid (streams 39a and 33a) as described above and by refrigerant stream 71e. ℃]. Partially condensed stream 49a then enters heat exchanger 60 and is further cooled to -254°F [-159°C] by refrigerant stream 71d to facilitate its condensation and supercooling, whereupon it enters working expansion machine 61, Mechanical energy is extracted from the flow in the working expansion machine 61 . The mechanism 61 expands the liquid stream 49d substantially isentropically from a pressure of about 621 psi [4,282 kPa(a)] to the LNG storage pressure slightly above atmospheric pressure (15.5 psi [107 kPa(a)]). Working expansion cools expanded stream 49c to a temperature of approximately -255°F [-159°C], whereupon it is sent to LNG storage tank 62 for containing LNG product (stream 50).

Much of the cooling of convection 49 and all of the cooling of convection 49a is provided by closed loop cooling loops similar to the processes of Figures 1, 3 and 4. In approximate mole percentages, the composition of the working fluid stream used in the cycle of the Figure 5 process is: 8.9% nitrogen, 34.3% methane, 41.3% ethane, 11.0% propane, with the balance consisting of heavy hydrocarbons . Refrigerant stream 71 exits discharge cooler 69 at 100°F [38°C] and 607 psig 4,185 kPa(a)]. Refrigerant stream 71 enters heat exchanger 10 and is cooled to -30°F [-34°C] and partially condensed by partially warmed expanding refrigerant stream 71f and other refrigerant streams. For the simulated example in Figure 5, it has been assumed that the other refrigerant stream is industrial grade quality propane refrigerant at three different temperature and pressure stages. Partially condensed refrigerant stream 71a then enters heat exchanger 13 for further cooling to -102°F [-74°C] by partially warmed expanded refrigerant stream 71e, further condensing the refrigerant (stream 71b ). The refrigerant is completely condensed in heat exchanger 60 by expanded refrigerant stream 71d and then further subcooled to -254°F [-159°C]. The supercooled liquid stream 71c enters the working expansion machine 63 as the stream expands substantially isentropically from a pressure of about 586 psi [4,040 kPa(a)] to a pressure of about 34 psi [234 kPa(a) )], mechanical energy is extracted from the flow in the working expansion machine 63. During expansion, a portion of the stream is evaporated, causing the total stream to cool to -264°F [-164°C] (stream 71d). Expanded stream 71d then re-enters heat exchangers 60, 13 and 10 where cooling is provided to stream 49a, stream 49 and the refrigerant (streams 71, 71a and 71b) as expanded stream 71d is evaporated and overheated.

Superheated refrigerant vapor (stream 71 g) exits heat exchanger 10 at 93°F [34°C] and is compressed to 617 psi [4,254 kPa(a)] in three stages. Each of the three compression stages (refrigerant compressors 64, 66 and 68) is driven by auxiliary power and is followed by coolers (discharge coolers 65, 67 and 69) to remove the heat of compression. Compressed stream 71 from discharge cooler 69 is returned to heat exchanger 10 to complete the cycle.

The sum of the stream flow rates and energy consumption for the process shown in Figure 5 is listed in the chart below:

Chart IV

(Figure 5)

Flow Total-Lb.Moles/Hr[kg moles/Hr]

Stream methane ethane propane butane total number

31 40,977 3,861 2,408 1,404 48,656

32 32,360 2,675 1,469 701 37,209

33 8,617 1,186 939 703 11,447

38 13,133 2,513 1,941 22 17,610

39 6,194 1,648 1,272 674 9,788

41 0 0 22 1,352 1,375

42 2,166 1,027 197 27 27,421

43 14,811 2,834 2,189 25 19,860

48 1,678 321 248 3 2,250

50 39,299 3,540 2,138 49 45,031

condensate ^* Recovery rate in

Butane series 95.04%

Pentane + 99.57%

Productivity 88,390 Lb/Hr [88,390kg/Hr]

LNG products

Productivity 834,183 Lb/Hr [834,183kg/Hr]

Purity ^* 87.27%

Lower calorific value 1033.8 BTU/SCF [38.52MJ/m ³ ]

power

Refrigerant Compression 84,974 HP [139,696kW]

Propane Compression 39,439 HP [64,837kW]

Total compression 124,413 HP [204,533kW]

Effective calories

Demethanizer Reboiler 52,913 MBTU/Hr [34,182kW]

^* (Based on flow rates without rounding)

Assuming that the LNG product factory is on-line for 340 days per year, the unit power consumption of the embodiment in Figure 5 of the present invention is 0.145Hp-Hr/Lb [0.238kW-Hr/kg]. Compared with the prior art process, the efficiency improvement of the embodiment of Fig. 5 is 16-26%.

Compared with the embodiment of Fig. 1 and Fig. 3, the embodiment of Fig. 5 of the present invention requires approximately 5% to 10% less power per unit of liquid produced. Compared with the embodiment of Fig. 4, the embodiment of Fig. 5 of the present invention requires substantially the same power per unit of liquid produced. Thus, for a given amount of available compression power, the FIG. ₅ embodiment can liquefy approximately 5% more natural gas than the FIG. The embodiment liquefies about 10% more natural gas or the same amount of natural gas as the embodiment shown in Fig. 4 liquefies. The choice between Figure 5 and the Figure 1, Figure 3 or Figure 4 embodiments of the invention for a particular application is generally governed by the ratio of the economic value of ethane and propane as part of the NGL or LPG product to their corresponding value in the LNG product Dominate or be governed by the calorific value specification of the LNG product (because the calorific value of the LNG produced by the embodiment of Fig. 1, Fig. 3 and Fig. 4 is lower than that of the LNG produced by the embodiment of Fig. 5).

other embodiments

It will be apparent to those of ordinary skill in the art that the present invention is applicable to all types of LNG liquefaction plants allowing co-products of NGL streams, LPG streams or condensate streams, as best adapted to the requirements of a given site. Furthermore, it should be understood that a variety of process formats may be used to recover the liquid co-product stream. For example, the Figure 1 and Figure 3 embodiments may be adapted to recover an LPG stream or a condensate stream as a liquid co-product stream rather than recovering an NGL stream as a liquid co-product stream as in Examples 1 and 2 previously described. The Figure 4 embodiment can be adapted to recover an _NGL stream comprising the majority of the C _components present in the feed gas or to recover a condensate stream comprising the majority of the C components and heavies present in the feed gas, Instead of producing an LPG co-product as in Example 3 above. The Figure 5 embodiment may be adapted for recovery of an NGL stream comprising the majority of the _C2 components present in the feed gas or for recovery of an LPG stream comprising the majority of the _C3 components present in the feed gas, rather than as in the preceding examples 4 as in the production of condensate co-products.

Figures 1, 3, 4 and 5 depict a preferred embodiment of the invention under specified process conditions. Figures 6 to 21 illustrate alternative embodiments of the invention that may be considered depending on the particular application. As shown in Figures 6 and 7, all or part of the condensate (stream 33) from separator 11 may be fed to fractionation column 19 at the separation lower mid-column feed location rather than heat exchanged with the inflow to A portion of the separator vapor (stream 34) from vessel 13 is mixed. Figure 8 shows an alternative embodiment of the invention which requires less equipment than the embodiment of Figures 1 and 6, although its specific power consumption is slightly higher. Likewise, Figure 9 shows an alternative embodiment of the invention which requires less equipment than the embodiments of Figures 3 and 7, again at the expense of slightly higher specific power consumption. Figures 10 to 14 show an alternative embodiment of the invention which requires less equipment than the embodiment of Figure 4, although its specific power consumption may be higher (it should be noted that in Figures 10 to 14 , distillation columns or systems such as deethanizer 19 include both reboiler absorber designs and reflux reboiler columns). Figures 15 and 16 show an alternative embodiment of the invention that combines the functions of the separator/absorber column 18 and deethanizer column 19 of the embodiments of Figures 4 and 10 to 14 into one Fractionation tower 19. Depending on the mass of heavy hydrocarbons in the feed gas and the feed gas pressure, the cooled feed stream 31a leaving the heat exchanger 10 may not contain any liquid (either because it is above its dew point, or because it is above its critical condensation pressure), The separator 11 shown in FIGS. 1 and 3 to 16 is therefore not required and the cooling feed flow can flow directly into a suitable expansion device, such as a working expansion machine 15 .

The liquid co-product stream (stream 37 in FIGS. 1, 3, 6 to 11, 13 and 14; stream 47 in FIGS. ; and stream 43) in Figure 5) After recovery, the disposal of the residual gas stream can be performed in a number of ways. In the process of Figures 1 and 3 through 16, the stream is heated using energy from one or more working expansion machines, compressed to a higher pressure, cooled in a discharge cooler, and passed through the Laterally exchanged for further cooling. As shown in Figure 17, some applications may prefer to compress the stream to a higher pressure using, for example, an auxiliary compressor 59 driven by an external power source. As shown in Figure 1 and the dashed line devices of Figures 3 through 16 (heat exchanger 24 and discharge cooler 25), some circumstances may favor Cooling (at the expense of increased load on heat exchanger 60 and increased power consumption of refrigerant compressors 64, 66 and 68) reduces equipment cost. In this case, stream 49a leaving the compressor may flow directly into heat exchanger 24 as shown in FIG. 18 , or directly into heat exchanger 60 as shown in FIG. 19 . Compressor 16 may be replaced by a compressor driven by an external power source, such as compressor 59 shown in FIG. 20, if no working expansion machinery is used for any partial expansion of the high pressure feed gas. Other cases may not adjust any compression of the stream at all, so the stream flows directly into the device (heat exchanger 24, compressor 16 and discharge cooler 25) in Fig. In the heat exchanger 60 shown. If heat exchanger 24 to heat plant fuel gas (stream 48) prior to extraction of the stream is not included, auxiliary heater 58 may be required to warm the fuel gas prior to combustion, using useful stream or other The process stream supplies the required heat as shown in Figures 19-21. Selections such as these must generally be evaluated for each application, and factors such as gas composition, plant size, desired level of recovery of co-product streams, and available equipment must all be considered.

According to the invention cooling of the inlet gas stream and feeding of said stream to the LNG production area can be performed in a number of ways. In the processes of FIGS. 1 , 3 and 6 to 9 , the inlet gas stream 31 is cooled and condensed by an external refrigerant stream and column liquid from fractionation column 19 . In the processes of Figures 4, 5 and Figures 10 to 14, a flash separator liquid is used with an external refrigerant flow for this purpose. In Figures 15 and 16, column liquid and flash separator liquid are used together with an external refrigerant stream for this purpose. Whereas in Figures 17 to 21 only the external refrigerant flow is used to cool the inlet gas flow 31 . However, the cooling process stream can also be used to provide some cooling to the high pressure refrigerant (stream 71a ), such as shown in FIGS. 4 , 5 , 10 and 11 . Additionally, any stream that is cooler than the stream being cooled can be utilized. For example, sidedraw vapor from separator/absorber column 18 or deethanizer column 19 may be extracted and used for cooling. The use and distribution of column liquid and/or vapor for heat exchange must be estimated for each specific application, As well as the specific arrangement of heat exchangers for inlet gas and feed gas cooling, and the selection of process streams for specific heat exchange services. The choice of cooling station will depend on a number of factors including, but not limited to, feed gas composition and conditions, plant size, heat exchanger size, potential cooling source temperature, and the like. It will be understood by those of ordinary skill in the art that any combination of the above cooling sources or cooling methods may be used in a mixed fashion to achieve the desired feedstream temperature.

Furthermore, auxiliary external cooling of the inlet gas stream and the feedstock stream supplied to the LNG production zone can also be performed in a number of different ways. In Fig. 1 and Figs. 3 to 21, it has been assumed that a single-component refrigerant is boiled for high-level external refrigeration, and it has been assumed that a multi-component refrigerant is evaporated for low-level external refrigeration, where a single The sub-refrigerants are used to pre-cool the multi-component refrigerant stream. Alternatively, both high-stage cooling and low-stage cooling may be performed using a single-component refrigerant with progressively lower boiling points (ie, staged refrigeration), or with single-component refrigerants at progressively lower evaporation pressures. Alternatively, both high and low stage cooling may be performed using a multicomponent refrigerant stream with higher components of the multicomponent refrigerant stream adjusted to provide the necessary cooling temperature. The choice of method to provide external refrigeration will depend on factors including, but not limited to, feed gas composition and conditions, plant size, compressor drive size, heat exchanger size, ambient heat sink temperature, etc. . Those of ordinary skill in the art will appreciate that any of the combinations described above to provide external refrigeration can be used in combination to achieve the desired feedstream temperature.

The condensate 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 49b in Figure 17 Stream 49e of FIG. 18; stream 49c in FIG. 18; and stream 49a in FIG. This generally reduces the specific electricity consumption for LNG production by eliminating the need for flash gas compression. However, some circumstances may favor reducing the cost of the plant by reducing the size of the heat exchanger 60, and using flash gas compression or otherwise disposing of any flash gas that may be produced.

Although independent flow expansion is shown in a particular expansion device, alternative expansion devices may be used where appropriate. For example, conditions may warrant working expansion of a substantially condensed feed stream (stream 35a in Figures 1, 3, 6 and 7) or an intermediate pressure reflux stream (stream 39 in Figures 1, 6 and 8). In addition, isenthalpic flash expansion can be used instead of the supercooled liquid stream leaving heat exchanger 60 (stream 49 in FIGS. 1, 6 and 8; stream 49d in FIGS. 3, 4, 7 and 9 to 16; , stream 49b in 19 and 20; stream 49e in FIG. 17; stream 49c in FIG. 18; and stream 49a in FIG. Flash steam is formed during expansion, or additional flash steam expansion or other means for disposing of the flash steam produced is required. Likewise, isenthalpic flash expansion can be used instead of working expansion of the supercooled high-pressure refrigerant stream (stream 71c in FIGS. 1 and 3-21 ) exiting heat exchanger 60, resulting in increased refrigerant compression. required power consumption.

While the preferred embodiment of the present invention has been described, those of ordinary skill in the art will appreciate that other and further modifications may be made thereto without departing from the spirit of the invention as defined in the following claims, such as , in order to adapt the invention to various conditions, raw material types or other requirements.

Claims (52)

1. Processes for the liquefaction of natural gas streams containing methane and heavy hydrocarbon components, wherein: (a) said natural gas stream is cooled under pressure so as to condense at least a portion thereof and form a condensed stream; and (b) said condensed stream is expanded to a lower pressure to form a liquefied natural gas stream; The process is characterized in that it comprises the following processing steps: (1) said natural gas stream is processed in one or more cooling steps; (2) said cooled natural gas stream is expanded to intermediate pressure; (3) said expanded cooled natural gas stream is directed to a distillation column where said expanded cooled natural gas stream is separated into a volatile residual gas component comprising most of said methane and light components and a lower volatility component comprising a majority of said heavy hydrocarbon components; (4) said volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form said condensed stream; (5) said condensing portion is divided into at least two portions thereby constituting said condensed stream and a liquid stream; and (6) The liquid stream is introduced into the distillation column as its top feed. 2. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps to facilitate its partial condensation; (2) said partially condensed natural gas stream is separated to provide at least one vapor stream and a first liquid stream; (3) the vapor stream is expanded to an intermediate pressure; (4) said first liquid stream is expanded to said intermediate pressure; (5) At least said expanded vapor stream and said expanded first liquid stream are directed into a distillation column in which said expanded vapor stream and said expanded first stream are divided into a volatile residual gas component of said methane and light components and a less volatile component comprising a majority of said heavy hydrocarbon components; (6) said volatile residual gas component is cooled under pressure to facilitate condensation of at least a portion thereof; (7) said condensing portion is divided into at least two portions thereby constituting said condensed stream and a second liquid stream; and (8) The second liquid stream is introduced into the distillation column as its overhead feed. 3. The process of claim 1, wherein the natural gas stream is not expanded to intermediate pressure, the process being characterized by: (1) said natural gas stream is processed in one or more cooling steps; (2) said cooled natural gas stream is divided into at least a first gaseous stream and a second gaseous stream; (3) said first gaseous stream is cooled to condense substantially all of it before being expanded to an intermediate pressure; (4) said second gaseous stream is expanded to said intermediate pressure; (5) said expanded substantially condensed first gaseous stream and said expanded second gaseous stream are introduced into a distillation column in which said expanded substantially condensed first gaseous stream and said the expanded second gaseous stream is divided into a volatile residual gas component comprising a majority of said methane and light components and a less volatile component comprising a majority of said heavy hydrocarbon components; and (6) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 4. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps to facilitate its partial condensation; (2) said partially condensed natural gas stream is separated to provide a vapor stream and a liquid stream; (3) said vapor stream is divided into at least a first gaseous stream and a second gaseous stream; (4) said first gaseous stream is cooled to condense substantially all of it before being expanded to an intermediate pressure; (5) said second gaseous stream being expanded to said intermediate pressure; (6) said liquid flow is expanded to said intermediate pressure; (7) said expanded substantially condensed first gaseous stream, said expanded second gaseous stream and said expanded liquid stream are introduced into a distillation column in which said expanded substantially condensed The first gaseous stream, the expanded second gaseous stream and the expanded liquid stream are separated into a volatile residual gas component comprising most of the methane and light components and a group comprising most of the heavy hydrocarbons lower volatile content; and (8) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 5. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps to facilitate its partial condensation; (2) said partially condensed natural gas stream is separated to provide a vapor stream and a liquid stream; (3) said vapor stream is divided into at least a first gaseous stream and a second gaseous stream; (4) said first gaseous stream mixes with at least a portion of said liquid stream to form a mixed stream; (5) the mixed stream is cooled to condense substantially all of it before being expanded to intermediate pressure; (6) said second gaseous stream being expanded to said intermediate pressure; (7) any remaining portion of said flow is expanded to said intermediate pressure; (8) said expanded substantially condensed mixed stream, said expanded second gaseous stream and said remainder of said liquid stream are directed into a distillation column in which said expanded substantially condensed The mixed stream, the expanded second gaseous stream and the liquid stream are divided into a volatile residual gas component comprising most of the methane and light components and a relatively heavy hydrocarbon component comprising most of the heavy hydrocarbon components. low volatile content; and (9) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 6. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is processed in one or more cooling steps; (2) said cooled natural gas stream is divided into at least a first gaseous stream and a second gaseous stream; (3) said first gaseous stream is cooled to condense substantially all of it before being expanded to an intermediate pressure; (4) said second gaseous stream is expanded to said intermediate pressure; (5) said expanded substantially condensed first gaseous stream and said expanded second gaseous stream are introduced into a distillation column in which said expanded substantially condensed first gaseous stream and said the expanded second gaseous stream is divided into a volatile residual gas component comprising a majority of said methane and light components and a less volatile component comprising a majority of said heavy hydrocarbon components; (6) said volatile residual gas component is cooled under pressure to facilitate condensation of at least a portion thereof; (7) said condensing portion is divided into at least two portions thereby constituting said condensed stream and a liquid stream; and (8) The liquid stream is introduced into the distillation column as its top feed. 7. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps to facilitate its partial condensation; (2) said partially condensed natural gas stream is separated to provide a vapor stream and a first liquid stream; (3) said vapor stream is divided into at least a first gaseous stream and a second gaseous stream; (4) said first gaseous stream is cooled to condense substantially all of it before being expanded to an intermediate pressure; (5) said second gaseous stream being expanded to said intermediate pressure; (6) said first liquid stream is expanded to said intermediate pressure; (7) said expanded substantially condensed first gaseous stream, said expanded second gaseous stream and said expanded first liquid stream are introduced into a distillation column in which said expanded substantially condensed The condensed first gaseous stream, said expanded second gaseous stream and said expanded first liquid stream are separated into a volatile residual gas component comprising a majority of said methane and light components and a component comprising a majority of said heavy Lower volatile components of high-quality hydrocarbon components; (8) said volatile residual gas component is cooled under pressure to facilitate condensation of at least a portion thereof; (9) said condensing portion is divided into at least two portions thereby constituting said condensed stream and a second liquid stream; and (10) The second liquid stream is introduced into the distillation column as its overhead feed. 8. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps to facilitate its partial condensation; (2) said partially condensed natural gas stream is separated to provide a vapor stream and a first liquid stream; (3) said vapor stream is divided into at least a first gaseous stream and a second gaseous stream; (4) said first gaseous stream mixes with at least a portion of said first liquid stream to form a mixed stream; (5) the mixed stream is cooled to condense substantially all of it before being expanded to intermediate pressure; (6) said second gaseous stream being expanded to said intermediate pressure; (7) any remaining portion of said first liquid stream is expanded to said intermediate pressure; (8) said expanded substantially condensed mixed stream, said expanded second gaseous stream, and said remainder of said first liquid stream are directed into a distillation column in which said expanded The substantially condensed mixed stream, said expanded second gaseous stream and said first liquid stream are separated into a volatile residual gas component comprising a majority of said methane and light components and a component comprising a majority of said heavy hydrocarbons lower volatility of components; (9) said volatile residual gas component is cooled under pressure to facilitate condensation of at least a portion thereof; (10) said condensing portion is divided into at least two portions thereby constituting said condensed stream and a second liquid stream; and (11) The second liquid stream is introduced into the distillation column as its overhead feed. 9. The process of claim 1 wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is processed in one or more cooling steps; (2) said cooled natural gas stream is expanded to intermediate pressure; (3) said expanded cooled natural gas stream is separated to provide a vapor stream and a liquid stream; (4) the liquid stream is expanded to a lower intermediate pressure; (5) The expanded liquid stream is directed to a distillation column where the expanded liquid stream is divided into a more volatile steam distillation stream and a steam distillation stream containing most of the heavy hydrocarbon components lower volatile components; (6) said more volatile steam distillation stream is mixed with said steam stream to form a volatile residual gas component comprising a majority of said methane and lights; and (7) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 10. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is processed in one or more cooling steps so as to condense at least a portion thereof; (2) said partially condensed natural gas stream is separated to provide a first vapor stream and a first liquid stream; (3) said first vapor stream is expanded to an intermediate pressure; (4) said expanded first vapor stream is separated to provide a second vapor stream and a second liquid stream; (5) said second liquid stream is expanded to a lower intermediate pressure; (6) said first liquid stream is expanded to said lower intermediate pressure; (7) The expanded second liquid stream and the expanded first liquid stream are introduced into a distillation column in which the expanded second liquid stream and the expanded first liquid stream is separated into a more volatile steam distillation stream and a less volatile component comprising the majority of said heavy hydrocarbon components; (8) said more volatile steam distillation stream is mixed with said second steam stream to form a volatile residual gas component comprising a majority of said methane and lights; and (9) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 11. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is processed in one or more cooling steps; (2) said cooled natural gas stream is expanded to an intermediate pressure before being introduced into a contacting device to form a volatile residual gas component and a first liquid stream comprising a majority of said methane and lights; (3) The first liquid stream is directed to a distillation column where it is divided into a more volatile steam distillation stream and a steam distillation stream containing most of the heavy hydrocarbon components lower volatile components; (4) said more volatile steam distillation stream is cooled sufficiently to condense at least a portion thereof to form a second liquid stream; (5) at least a portion of said expanded cooled natural gas stream is in intimate contact with at least a portion of said second liquid stream in said contacting device; and (6) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 12. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps so as to partially condense it; (2) said partially condensed natural gas stream is separated to provide a vapor stream and a first liquid stream; (3) said vapor stream is expanded to an intermediate pressure and then introduced into a contacting device, thereby forming a volatile residual gas component and a second liquid stream comprising a majority of said methane and light components; (4) said first liquid stream is expanded to said intermediate pressure; (5) The second liquid stream and the expanded first liquid stream are introduced into a distillation column in which the second liquid stream and the expanded first liquid stream are divided into more a volatile steam distillation stream and a less volatile component comprising a majority of said heavy hydrocarbon components; (6) said more volatile steam distillation stream is cooled sufficiently to condense at least a portion thereof to form a third liquid stream; (7) at least a portion of said expanded vapor stream is in intimate contact with at least a portion of said third liquid stream in said contacting device; and (8) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 13. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is processed in one or more cooling steps; (2) said cooled natural gas stream is expanded to an intermediate pressure and then introduced into a contacting device to form a first vapor stream and a first liquid stream; (3) The first liquid stream is directed to a distillation column where it is divided into a more volatile steam distillation stream and a steam distillation stream containing most of the heavy hydrocarbon components lower volatile components; (4) said more volatile steam distillation stream is cooled sufficiently to condense at least a portion thereof, thereby forming a second vapor stream and a second liquid stream; (5) a portion of the second liquid stream is introduced into the distillation column as its overhead feed; (6) at least a portion of said expanded cooled natural gas stream is in intimate contact with at least a portion of the remainder of said second liquid stream in said contacting device; (7) said first vapor stream is mixed with said second vapor stream to form a volatile residual gas component comprising a majority of said methane and lights; and (8) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 14. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps so as to partially condense it; (2) said partially condensed natural gas stream is separated to provide a first vapor stream and a first liquid stream; (3) said first vapor stream is expanded to an intermediate pressure and then introduced into a contacting device to form a second vapor stream and a second liquid stream; (4) said first liquid stream is expanded to said intermediate pressure; (5) The second liquid stream and the expanded first liquid stream are introduced into a distillation column in which the second liquid stream and the expanded first liquid stream are divided into more a volatile steam distillation stream and a less volatile component comprising a majority of said heavy hydrocarbon components; (6) said more volatile steam distillation stream is cooled sufficiently to condense at least a portion thereof, thereby forming a third vapor stream and a third liquid stream; (7) a portion of the third liquid stream is introduced into the distillation column as its top feed; (8) at least a portion of said expanded first vapor stream is in intimate contact with at least a portion of a remainder of said third liquid stream in said contacting device; (9) said second vapor stream is mixed with said third vapor stream to form a volatile residual gas component comprising a majority of said methane and lights; and (10) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 15. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is processed in one or more cooling steps; (2) said cooled natural gas stream is expanded to an intermediate pressure before being introduced into a contacting device to form a volatile residual gas component and a first liquid stream comprising a majority of said methane and lights; (3) The first liquid stream is heated and then directed to a distillation column where it is divided into a more volatile steam distillation stream and a steam distillation stream containing most of the heavy Lower volatility of hydrocarbon components; (4) said more volatile steam distillation stream is cooled sufficiently to condense at least a portion thereof to form a second liquid stream; (5) at least a portion of said expanded cooled natural gas stream is in intimate contact with at least a portion of said second liquid stream in said contacting device; and (6) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 16. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps so as to partially condense it; (2) said partially condensed natural gas stream is separated to provide a vapor stream and a first liquid stream; (3) said vapor stream is expanded to an intermediate pressure before being introduced into a contacting device, thereby forming a volatile residual gas component comprising a majority of said methane and lights and a second liquid stream; (4) the second liquid flow is heated; (5) said first liquid stream is expanded to said intermediate pressure; (6) The heated second liquid stream and the expanded first liquid stream are introduced into a distillation column in which the heated second liquid stream and the expanded first liquid stream is separated into a more volatile steam distillation stream and a less volatile component comprising the majority of said heavy hydrocarbon components; (7) said more volatile steam distillation stream is cooled sufficiently to condense at least a portion thereof, thereby forming a third liquid stream; (8) at least a portion of said expanded vapor stream is in intimate contact with at least a portion of said third liquid stream in said contacting device; and (9) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 17. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is processed in one or more cooling steps; (2) said cooled natural gas stream is expanded to an intermediate pressure and then introduced into a contacting device to form a first vapor stream and a first liquid stream; (3) The first liquid stream is heated and then directed to a distillation column where it is divided into a more volatile steam distillation stream and a steam distillation stream containing most of the heavy Lower volatility of hydrocarbon components; (4) said more volatile steam distillation stream is cooled sufficiently to condense at least a portion thereof, thereby forming a second vapor stream and a second liquid stream; (5) a portion of the second liquid stream is introduced into the distillation column as its overhead feed; (6) at least a portion of said expanded cooled natural gas stream is in intimate contact with at least a portion of the remainder of said second liquid stream in said contacting device; (7) said first vapor stream is mixed with said second vapor stream to form a volatile residual gas component comprising a majority of said methane and lights; and (8) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 18. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps so as to partially condense it; (2) said partially condensed natural gas stream is separated to provide a first vapor stream and a first liquid stream; (3) said first vapor stream is expanded to an intermediate pressure and then introduced into a contacting device to form a second vapor stream and a second liquid stream; (4) the second liquid flow is heated; (5) said first liquid stream is expanded to said intermediate pressure; (6) The heated second liquid stream and the expanded first liquid stream are introduced into a distillation column in which the heated second liquid stream and the expanded first liquid stream is separated into a more volatile steam distillation stream and a less volatile component comprising the majority of said heavy hydrocarbon components; (7) said more volatile steam distillation stream is cooled sufficiently to condense at least a portion thereof, thereby forming a third vapor stream and a third liquid stream; (8) a portion of the third liquid stream is introduced into the distillation column as its top feed; (9) at least a portion of said expanded first vapor stream is in intimate contact with at least a portion of a remainder of said third liquid stream in said contacting device; (10) said second vapor stream is mixed with said third vapor stream to form a volatile residual gas component comprising a majority of said methane and lights; and (11) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 19. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is processed in one or more cooling steps; (2) The cooled natural gas stream is expanded to intermediate pressure before being directed to a mid-column feed location on the distillation column where the cooled natural gas stream is split into more volatile steam distillation streams and a lower volatility component comprising a majority of said heavy hydrocarbon components; (3) a vapor distillation stream is withdrawn from a region of said distillation column below said expanded cooled natural gas stream and cooled sufficiently to condense at least a portion thereof to form a vapor stream and a liquid stream; (4) at least a portion of said expanded cooled natural gas stream is in intimate contact with at least a portion of said liquid stream in said distillation column; (5) said steam stream is mixed with said more volatile steam distillation stream to form a volatile residual gas component comprising a majority of said methane and lights; and (6) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 20. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps so as to partially condense it; (2) said partially condensed natural gas stream is separated to provide a first vapor stream and a first liquid stream; (3) said first vapor stream and said first liquid stream are expanded to an intermediate pressure; (4) The expanded first vapor stream and the expanded first liquid stream are introduced into a mid-column feed location on the distillation column where the expanded first vapor stream and the expanded The first liquid stream of is divided into a more volatile steam distillation stream and a less volatile component comprising most of said heavy hydrocarbon components; (5) a steam distillation stream is withdrawn from a region of said distillation column below said expanded first vapor stream and cooled sufficiently to condense at least a portion thereof to form a second vapor stream and a second liquid stream; (6) at least a portion of said expanded first vapor stream is in intimate contact with at least a portion of said second liquid stream in said distillation column; (7) said second vapor stream is mixed with said more volatile steam distillation stream to form a volatile residual gas component comprising a majority of said methane and lights; and (8) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 21. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is processed in one or more cooling steps; (2) The cooled natural gas stream is expanded to intermediate pressure before being directed to a mid-column feed location on the distillation column where the cooled natural gas stream is split into more volatile steam distillation streams and a lower volatility component comprising a majority of said heavy hydrocarbon components; (3) a vapor distillation stream is withdrawn from a region of said distillation column below said expanded cooled natural gas stream and cooled sufficiently to condense at least a portion thereof to form a vapor stream and a liquid stream; (4) a portion of said liquid stream is supplied as another feedstock thereof to said distillation column at a feed location in substantially the same region as said steam distillation stream is extracted; (5) at least a portion of said expanded cooled natural gas stream is in intimate contact with at least a portion of the remainder of said liquid stream in said distillation column; (6) said steam stream is mixed with said more volatile steam distillation stream to form a volatile residual gas component comprising a majority of said methane and lights; and (7) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 22. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps so as to partially condense it; (2) said partially condensed natural gas stream is separated to provide a first vapor stream and a first liquid stream; (3) said first vapor stream and said first liquid stream are expanded to an intermediate pressure; (4) The expanded first vapor stream and the expanded first liquid stream are introduced into a mid-column feed location on a distillation column where the expanded first vapor stream and the the expanded first liquid stream is divided into a more volatile steam distillation stream and a less volatile component comprising a majority of said heavy hydrocarbon components; (5) a steam distillation stream is withdrawn from a region of said distillation column below said expanded first vapor stream and cooled sufficiently to condense at least a portion thereof to form a second vapor stream and a second liquid stream; (6) a portion of said second liquid stream is supplied to said distillation column as another feedstock to said distillation column at a feed location in substantially the same region as said steam distillation stream is extracted; (7) at least a portion of said expanded first vapor stream is in intimate contact with at least a portion of a remainder of said second liquid stream in said distillation column; (8) said second vapor stream is mixed with said more volatile steam distillation stream to form a volatile residual gas component comprising a majority of said methane and lights; and (9) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 23. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is processed in one or more cooling steps; (2) The cooled natural gas stream is expanded to intermediate pressure before being directed to a mid-column feed location on the distillation column where the cooled natural gas stream is split into more volatile steam distillation streams and a lower volatility component comprising a majority of said heavy hydrocarbon components; (3) a vapor distillation stream is withdrawn from a region of said distillation column below said expanded cooled natural gas stream and cooled sufficiently to condense at least a portion thereof to form a vapor stream and a liquid stream; (4) at least a portion of said expanded cooled natural gas stream is in intimate contact with at least a portion of said liquid stream in said distillation column; (5) The liquid distillation stream is extracted from the distillation column at a position above the region where the steam distillation stream is extracted, so that the liquid distillation stream is heated, and then used as another feedstock for the distillation column at the location where the steam distillation stream is extracted is introduced into the distillation column at a location below the steam distillation stream region; (6) said steam stream is mixed with said more volatile steam distillation stream to form a volatile residual gas component comprising a majority of said methane and lights; and (7) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 24. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps so as to partially condense it; (2) said partially condensed natural gas stream is separated to provide a first vapor stream and a first liquid stream; (3) said first vapor stream and said first liquid stream are expanded to an intermediate pressure; (4) The expanded first vapor stream and the expanded first liquid stream are introduced into a mid-column feed location on a distillation column where the expanded first vapor stream and the the expanded first liquid stream is divided into a more volatile steam distillation stream and a less volatile component comprising a majority of said heavy hydrocarbon components; (5) a steam distillation stream is withdrawn from a region of said distillation column below said expanded first vapor stream and cooled sufficiently to condense at least a portion thereof to form a second vapor stream and a second liquid stream; (6) at least a portion of said expanded first vapor stream is in intimate contact with at least a portion of said second liquid stream in said distillation column; (7) The liquid distillation stream is extracted from the distillation column at a position above the region where the steam distillation stream is extracted, so that the liquid distillation stream is heated, and then used as another feedstock for the distillation column at a location where the steam distillation stream is extracted is introduced into the distillation column at a location below the steam distillation stream region; (8) said second vapor stream is mixed with said more volatile steam distillation stream to form a volatile residual gas component comprising a majority of said methane and lights; and (9) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 25. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is processed in one or more cooling steps; (2) The cooled natural gas stream is expanded to intermediate pressure before being directed into a mid-column feed location on a distillation column where the cooled natural gas stream is split into more volatile steam distillation stream and a lower volatility component comprising most of said heavy hydrocarbon components; (3) a vapor distillation stream is withdrawn from a region of said distillation column below said expanded cooled natural gas stream and cooled sufficiently to condense at least a portion thereof to form a vapor stream and a liquid stream; (4) a portion of said liquid stream is fed into said distillation column as another feedstock to said distillation column at a feed location in substantially the same region as said steam distillation stream is extracted; (5) at least a portion of said expanded cooled natural gas stream is in intimate contact with at least a portion of the remainder of said liquid stream in said distillation column; (6) The liquid distillation stream is extracted from the distillation column at a position above the area where the steam distillation stream is extracted, whereupon the liquid distillation stream is heated, and then used as another feedstock thereof at a location where the steam distillation stream is extracted is introduced into the distillation column at a position below the flow region; (7) said steam stream is mixed with said more volatile steam distillation stream to form a volatile residual gas component comprising a majority of said methane and lights; and (8) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 26. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps so as to partially condense it; (2) said partially condensed natural gas stream is separated to provide a first vapor stream and a first liquid stream; (3) said first vapor stream and said first liquid stream are expanded to an intermediate pressure; (4) The expanded first vapor stream and the expanded first liquid stream are introduced into a mid-column feed location on a distillation column where the expanded first vapor stream and the the expanded first liquid stream is divided into a more volatile steam distillation stream and a less volatile component comprising a majority of said heavy hydrocarbon components; (5) a steam distillation stream is withdrawn from a region of said distillation column below said expanded first vapor stream and cooled sufficiently to condense at least a portion thereof to form a second vapor stream and a second liquid stream; (6) a portion of said second liquid stream is supplied to said distillation column as another feed to said distillation column at a feed location in substantially the same region as said steam distillation stream is extracted; (7) at least a portion of said expanded first vapor stream is in intimate contact with at least a portion of a remainder of said second liquid stream in said distillation column; (8) The liquid distillation stream is extracted from the distillation column at a position above the region where the vapor distillation stream is extracted, so that the liquid distillation stream is heated, and then used as another feedstock for the distillation column at a location where the steam distillation stream is extracted is introduced into the distillation column at a location below the steam distillation stream region; (9) said second vapor stream is mixed with said more volatile steam distillation stream to form a volatile residual gas component comprising a majority of said methane and lights; and (10) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 27. The process according to claim 1, characterized in that it consists essentially of process steps (1)-(4) and does not include process steps (5) and (6). 28. The process of claim 1, wherein said natural gas stream is not expanded to intermediate pressure, said process comprising the processing steps of: (1) said natural gas stream is treated in one or more cooling steps so as to partially condense it; (2) said partially condensed natural gas stream is separated to provide at least a vapor stream and a liquid stream; (3) the vapor stream is expanded to an intermediate pressure; (4) said liquid flow is expanded to said intermediate pressure; (5) at least said expanded vapor stream and said expanded liquid stream are directed to a distillation column in which said expanded vapor stream and said expanded liquid stream are divided into a volatile residual gas component of methane and said light components and a less volatile component comprising a majority of said heavy hydrocarbon components; and (6) The volatile residual gas component is cooled under pressure so as to condense at least a portion thereof and thereby form the condensed stream. 29. The process of claim 1, 3, 4, 5, 11, 12, 13, 14, 15, 16, 17, 18, 19, 20, 21, 22, 23, 24, 25, 26 or 28 , characterized in that said volatile residual gas component is compressed and thereafter cooled under pressure in order to condense at least a portion thereof and thereby form said condensed stream. 30. The process of claim 2 or 7, wherein (1) said volatile residual gas component is compressed and then cooled under pressure to facilitate condensation of at least a portion thereof; and (2) The condensation part is divided into at least two parts, thereby forming the condensation stream and the liquid stream. 31. The process of claim 3, 8 or 9 wherein, (1) said volatile residual gas component is compressed and then cooled under pressure to facilitate condensation of at least a portion thereof; and (2) The condensed part is divided into at least two parts, thereby forming the condensed stream and the second liquid stream. 32. The process of claim 9, wherein the more volatile steam distillation stream is compressed and then mixed with the vapor stream to form a The volatile residual gas components. 33. The process of claim 10, wherein the more volatile steam distillation stream is compressed and then mixed with the second vapor stream to form a The volatile residual gas composition of the components. 34. The process of claim 1, 3, 4, 5, 11, 12, 13, 14, 15, 16, 17, 18, 19, 20, 21, 22, 23, 24, 25, 26 or 28 , characterized in that said volatile residual gas component is heated, compressed and thereafter cooled under pressure in order to condense at least a portion thereof and thereby form said condensed stream. 35. The process of claim 1 or 6 wherein, (1) the volatile residual gas component is heated, compressed, and then cooled under pressure so as to condense at least a portion thereof; and (2) The condensation part is divided into at least two parts, thereby forming the condensation stream and the liquid stream. 36. The process of claim 2, 7 or 8 wherein, (1) the volatile residual gas component is heated, compressed, and then cooled under pressure so as to condense at least a portion thereof; and (2) The condensed part is divided into at least two parts, thereby forming the condensed stream and the second liquid stream. 37. The process of claim 9, wherein the more volatile steam distillation stream is heated, compressed, cooled, and then mixed with the vapor stream to form a The volatile residual gas composition of light components. 38. The process of claim 10, wherein said more volatile steam distillation stream is heated, compressed, cooled, and then mixed with said second vapor stream to form a The volatile residual gas components of methane and lights. 39. According to claims 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, 11, 12, 13, 14, 15, 16, 17, 18, 19, 20, 21, 22, 23 , 24, 25, 26, 27, 28, 32, 33, 37 or 38 described process, it is characterized in that, described volatile residual gas component comprises most of described methane, light components and C ₂ components . 40. According to claims 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, 11, 12, 13, 14, 15, 16, 17, 18, 19, 20, 21, 22, 23 , 24, 25, 26, 27, 28, 32, 33, 37 or the described process of 38, it is characterized in that, described volatile residual gas component comprises most of described methane, light components, C ₂ components and _C3 components. 41. The process of claim 29, wherein said volatile residual gas components comprise a majority of said methane, lights and _C2 components. 42. The process of claim 30, wherein said volatile residual gas components comprise a majority of said methane, lights and _C2 components. 43. The process of claim 31 wherein said volatile residual gas constituents comprise a majority of said methane, lights and _C2 components. 44. The process of claim 34, wherein said volatile residual gas components comprise a majority of said methane, lights and _C2 components. 45. The process of claim 35, wherein said volatile residual gas components comprise a majority of said methane, lights and _C2 components. 46. The process of claim 36, wherein said volatile residual gas constituents comprise a majority of said methane, lights and _C2 components. 47. The process of claim 29, wherein said volatile residual gas components comprise a majority of said methane, lights, _C2 components, and _C3 components. 48. The process of claim 30, wherein said volatile residual gas components comprise a majority of said methane, lights, _C2 components, and _C3 components. 49. The process of claim 31 wherein said volatile residual gas components comprise a majority of said methane, lights, _C2 components and _C3 components. 50. The process of claim 34, wherein said volatile residual gas components comprise a majority of said methane, lights, _C2 components, and _C3 components. 51. The process of claim 35, wherein said volatile residual gas components comprise a majority of said methane, lights, _C2 components, and _C3 components. 52. The process of claim 36, wherein said volatile residual gas components comprise a majority of said methane, lights, _C2 components, and _C3 components.

Images