KR100877029B1 - Method and apparatus for liquefaction of natural gas streams - Google Patents

Abstract

A combined natural gas 50 liquefaction process is disclosed to produce a liquid stream containing heavy hydrocarbons that is superior to methane 41. In this process, the liquefied natural gas stream 31 is partially cooled, expanded to intermediate pressures 14 and 15 and supplied to the distillation column 19. The bottom product 41 from this distillation column optionally contains a large number of hydrocarbons that are heavier than methane, which can reduce the purity of the liquefied natural gas 50. Residual gas stream 37 from distillation column 19 is expanded to a high medium pressure (16), cooled under pressure (60) to condense the stream, and then expanded to low pressure (61) to liquefy the natural gas stream. Form.

Description

Method and apparatus for liquefaction of natural gas streams {NATURAL GAS LIQUEFACTION}

The present invention relates to a process for treating natural gas or other methane rich gas streams to produce liquefied natural gas (LNG) having a liquid stream containing high purity methane and heavy hydrocarbons predominant over methane.

Natural gas is generally recovered from gas wells that have been drilled into the underground reservoir. Natural gas is usually dominated by methane, ie methane contains at least 50 mole percent (mol%) of natural gas. Depending on the particular underground reservoir, natural gas also contains relatively small amounts of heavier hydrocarbons such as ethane, propane, butane, pentane and the like, as well as water, hydrogen, nitrogen, carbon dioxide and other gases.

Most natural gas is treated in gaseous form. The most common means for transporting natural gas from gas wellheads to gas processing plants to natural gas consumers is the high pressure gas delivery pipeline. However, under various circumstances it is known to liquefy natural gas for transportation or use. For example, at a distance, there is often no pipeline infrastructure that allows for convenient transportation to the natural gas trading market. In this case, very low cost LNG related to gaseous natural gas can be used to distribute LNG using offshore shipping and transport trucks, thereby significantly reducing transportation costs.

Another situation where liquefaction of natural gas is desirable is when natural gas is used as a vehicle fuel. In large metropolitan areas, a large number of buses, taxis, trucks, etc., can be powered by LNG if a practical source of LNG is available. Vehicles that use LNG as fuel generate significantly less air pollution due to the clean combustion characteristics of natural gas compared to similar vehicles powered by gasoline and diesel engines that burn high molecular weight hydrocarbons. If LNG is of high purity (i.e. methane purity of 95 mol% or more), the carbon and hydrogen to methane is lower than all other hydrocarbon fuels, resulting in very low carbon dioxide (greenhouse gas) emissions.

FIELD OF THE INVENTION The present invention relates generally to the liquefaction of natural gas, while liquefied petroleum gas (LPG) or butane consisting of natural gas liquids (NGL) consisting of ethane, propane, butane and heavy hydrocarbon components, propane, butane and heavy hydrocarbon components. The present invention relates to the production of a co-product of a liquid stream consisting mainly of hydrocarbons heavier than methane, such as condensate and heavy hydrocarbon components. The production of by-product liquid streams has two important advantages: The LNG produced has high purity methane, and the by-product liquid is a useful product that can be used for many other uses. The natural gas stream treated in accordance with the invention is approximately mole percent, along with the remainder consisting of nitrogen and carbon dioxide, at 84.2% methane, 7.9% ethane and other C ₂ components, 4.9% propane and other C ₃ components, 1.0% iso- Butane (iso-butane), 1.1% normal-butane, 0.8% pentane ⁺ are generally analyzed. Sulfur containing gas is also sometimes present.

Various methods for liquefying natural gas are known. For example, Finn, Adrian J., Grant L., for the investigation of many of these treatments. Grant L. Johnson, and Terry R. "Marine" in the Gas Processors Association's 79th Annual Meeting, 429-450 (March 13-March 15, 2000, Atlanta, Georgia, USA), published by Terry R. Tomlinson et al. And LNG Technology for Mid-Scale Plants "and the 80th Annual Meeting of the Association of Gas Processors announced by Kikawa Oshitsugi, Masaaki Ohishi and Noriyoshi Nozawa (from March 12, 2001 to March 14, See “Optimizing the Power System of a Base-Road LNG Plant” in San Antonio, Texas, USA. US Patent No. 4,445,917; No. 4,525,185; 4,545,795; 4,755,200; 4,755,200; 5,291,736; 5,291,736; 5,363,655; 5,363,655; 5,365,740; 5,365,740; 5,600,969; 5,600,969; 5,615,561; 5,615,561; 5,651,269; 5,651,269; 5,755,114; 5,755,114; 5,893,274; 5,893,274; No. 6,014,869; 6,062,041; 6,062,041; No. 6,119,479; 6,125,653; 6,125,653; 6,250,105 B1; 6,269,655 B1; 6,272,882 B1; 6,308,531 B1; 6,324,867 B1; And 6,347,532 B1 also describe related methods. These methods generally involve purifying, cooling, condensing and expanding natural gas (by removing water and problematic compounds such as carbon dioxide and sulfur compounds). Cooling and condensation of natural gas can be accomplished in many different ways. "Cascade refrigeration" employs heat exchange of natural gas by a plurality of refrigerants having successive low boiling points such as propane, ethane, and methane. Optionally, this heat exchange can be accomplished by evaporating a single refrigerant to a plurality of different pressure levels.

“Multi-component refrigeration” employs heat exchange of natural gas by one or more refrigerant fluids composed of multiple refrigerant components instead of multiple single component refrigerants. The expansion of the natural gas can be made isenthalpically (eg using Joule-Thompson expansion) and isentropically (eg using a work-expansion turbine).

Regardless of the method used to liquefy the natural gas stream, it is common to require the removal of a significant fraction of hydrocarbons heavier than methane before the methane-rich stream is liquefied. There are several reasons for this hydrocarbon removal step, which include the need to control the calorific value of the LNG stream and the calorific value of these heavy hydrocarbon components as products. Unfortunately, little attention has been paid to the efficiency of the hydrocarbon removal step so far.

According to the present invention, it can be seen that the careful integration of the hydrocarbon removal step into the LNG liquefaction process can produce both LNG and separated heavy hydrocarbon liquid products using much less energy than prior art treatments. Although the present invention can be applied under low pressure, it is particularly advantageous when the gas supply is treated in the range of 2,758 to 10,342 kPa (a) [400 to 1500 psia] or more.

1 is a flow chart of a natural gas liquefaction plant adapted to co-produce NGLs in accordance with the present invention.

2 is a pressure-enthalpy state diagram of methane used to illustrate the advantages of the present invention with respect to prior art treatments.

3 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce NGLs in accordance with the present invention.

4 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce LPG in accordance with the present invention.

5 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce condensate in accordance with the present invention.

6 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

7 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

8 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

9 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

10 is a flow diagram of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

11 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

12 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

13 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

14 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

15 is a flow diagram of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

16 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

17 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

18 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

19 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

20 is a flow chart of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

21 is a flow diagram of an optional natural gas liquefaction plant adapted to co-produce a liquid stream in accordance with the present invention.

In the following description of the figures, the table provides an overview of the flow rates calculated during typical processing conditions. In the table below, the values for flow rate (moles / hour) are shown for the convenience to be closest to the integer. The total stream flow rate shown in the table includes all of the non-hydrocarbon components and therefore is generally greater than the sum of the stream flow rates of the hydrocarbon components. The temperature indicated is a temperature value that is approximately an integer. The process design calculations performed to compare the processes shown in the figures are based on the assumption that there is no heat leakage from (or from) the process. Commercially available insulation materials make this assumption very reasonable and are generally made by skilled workers.

For convenience, the processing parameters are recorded in both traditional British and international units (SI). The molar flow rates given in the table may be interpreted as pound moles per hour or kilogram moles per hour. The energy consumptions reported in horsepower (HP) and / or thousands of British thermodynamic units (MBTU / Hr) per hour correspond to the molar flow rates stated in pound moles per hour. The energy consumption, reported in kilowatts (kW), corresponds to the molar flow rate stated in kilogram moles per hour. The production rate reported in pounds per hour (Lb / Hr) corresponds to the molar flow rate stated in pounds per hour. The production rate reported in kilograms per hour (kg / Hr) corresponds to the molar flow rate stated in kilograms per hour.                 

First embodiment

With reference to FIG. 1, an example of a treatment according to the invention is described where the production of NGL by-products containing mostly ethane and heavy hydrocarbons in a natural gas feed stream is desired. In this simulation of the invention, the inlet gas enters the plant as stream 31 at 32 ° C. [90 ° F.] and 8860 kPa (a) [1285 psia]. If the product stream contains carbon dioxide and / or sulfur compound concentrations that can interfere with meeting specifications, these compounds are removed by appropriate pretreatment of the feed gas (not shown). In addition, the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccants are usually used for this purpose.

The feed stream 31 is cooled in the heat exchanger 10 by heat exchange with a refrigerant stream and a deboiler side reboiler liquid (stream 40) to -55 ° C [-68 ° F]. Note that the heat exchanger 10 is typically represented by one of a number of individual heat exchangers or a single multi-channel heat exchanger or a combination thereof (the determination of whether to use more than one heat exchanger for the indicated cooling service is as follows. But not limited to, depending on a number of factors including inlet gas flow rate, heat exchanger size, stream temperature, etc.). The cooled stream 31a is subjected to −34 ° C. [−30 ° F.] and 8,812 kPa (a) [1278psia] in a separation device 11 where steam (stream 32) is separated from condensed liquid (stream 33). Enter.

Steam (stream 32) is separated from the separation device 11 into two streams 34, 36. Stream 34 containing about 20% of the total vapor is combined with stream 33, the condensed liquid, to form stream 35. The combined stream 35 passes through the heat exchanger 13 in a heat exchange relationship with the refrigerant stream 71e, whereby the stream 35a cools and condenses. Stream 35a substantially condensed to -85 ° C [-120 ° F] is then operated at an operating pressure of the fractionation tower 19 (approximately 3206 kPa (a) [465 psia]) via a suitable expansion device such as expansion valve 14. Flash swell). A portion of the expanded stream is vaporized so that the entire stream is cooled. In the process shown in FIG. 1, the expanded stream 35b leaving the expansion valve 14 reaches a temperature of −86 ° C. [−122 ° F.], and the demethanation unit 19b of the fractionation tower 19 Supplied to the mid-point supply position.

The remaining 80% of the steam from the separation device 11 (stream 36) enters the work expansion device 15 where mechanical energy is extracted from this portion of the high pressure feed. Apparatus 15 expands the vapor from a pressure of about 8,812 kPa (a) [1278psia] to the tower operating pressure substantially isentropically, and the work expansion is at a temperature of approximately -75 ° C [-103 ° F]. Cooling the expanded stream 36a. Commercially available expansion devices can recover from about 80% to 85% of theoretically usable work with ideal isentropic expansion. The recovered work is often used to drive a centrifugal compressor 16 that can be used, for example, to recompress tower overhead gas (stream 38). Expanded and partially condensed stream 36a is sent as feed to the bottom-of-column feed point on distillation column 19.

The demethanation apparatus of the fractionation tower 19 is a prior art distillation column comprising a plurality of vertically spaced trays, one or more packed beds, or a combination of trays and packings. Although common in natural gas processing plants, a fractionation tower consists of two parts. The upper portion 19a is divided into the vapor and liquid portions, respectively, and the vapor from the lower distillation or demethanation unit 19b is combined with the vapor portion (if any) of the upper feed, so that the top of the tower is -93 ° C. Separator to form a cooling demethanizer overhead vapor (stream 37) exiting to [-135 [deg.] F.]. Lower demethanation 19b includes a tray and / or packing and provides the necessary contact between the liquid falling down and the vapor rising upwards. The demethanation unit also includes one or more reboilers (such as reboiler 20), which heats and vaporizes a portion of the liquid flowing down the column to provide stripping vapor flowing over the column. The liquid product stream 41 exits the bottom of the tower to 46 ° C. [115 ° F.] based on a typical specification of 0.020: 1 on a molar basis of the bottom product.

The demethanizer overhead steam (stream 37) is warmed in the heat exchanger 24 to 32 ° C. [90 ° F.], and a portion of the heated demethanizer overhead steam is taken out of the plant to produce fuel gas (stream (The amount of fuel gas to be taken out in this embodiment is mainly determined by the fuel required for the engine and / or turbine driving the gas compressor of the plant, such as refrigerant compressors 64, 66, 68, etc.). ). The remainder of the warmed demethanizer overhead steam (stream 38) is compressed by a compressor 16 driven by expansion devices 15, 61, 63. After cooling to 38 [deg.] C. [100 [deg.] F.] in the exhaust cooling unit 25, stream 38b cross-exchanges stream 37, which is a cold de-methanizer overhead vapor, at -86 [deg.] C. -123 [deg.] F.

Stream 38c then enters heat exchanger 60 and is further cooled by refrigerant stream 71d. After cooling to medium temperature, stream 38c is split into two portions. The first portion, stream 49, is further cooled to -160 [deg.] C. [-257 [deg.] F.) to condense and supercool the stream in heat exchanger 60, and then to the work expansion device 61 where mechanical energy is extracted from the stream. Enter. Apparatus 61 expands liquid stream 49 isotropically from a pressure of about 3,878 kPa (a) [562 psia] to an LNG storage pressure (107 kPa (a) [15.5 psia]) slightly above atmospheric pressure. The work expansion will cool the expanded stream 49a to a temperature of approximately -161 ° C. [-258 ° F.], which is then sent to the LNG storage tank 62 containing the LNG product (stream 50).

The other part of stream 38c, stream 39, is taken out of heat exchanger 60 to -107 [deg.] C. [-160 [deg.] F.), and through the appropriate expansion device such as expansion valve 17, Flashes up to working pressure In the process shown in FIG. 1, since there is no vaporization in the expanded stream 39a, the temperature of the stream is only slightly dropped, leaving the expansion valve 17 to a temperature of −107 ° C. [−161 ° F.]. The expanded stream 39a is then fed to the separator device 19a in the upper region of the fractionation tower 19. The liquid separated here becomes the top feed of the demethanation unit 19b.

Full cooling for streams 35 and 38c is provided by a closed cycle refrigeration loop. The working fluid for this cycle is a mixture of hydrocarbons and nitrogen, the composition of which is adjusted as necessary to provide the required refrigerant temperature while condensing at moderate pressure using commercially available coolants. In this case, since it is assumed that it has been condensed by cooling water, a refrigerant mixture composed of nitrogen, methane, ethane, propane, and heavy hydrocarbons is used in the simulation of the treatment of FIG. 1. The stream has a composition of 7.5% nitrogen, 41.0% methane, 41.5% ethane, and 10.0% propane in approximately mole percent with the remainder consisting of heavy hydrocarbons.

The refrigerant stream 71 leaves the exhaust cooling device 69 at 38 ° C. [100 ° F.] and 4,185 kPa (a) [607 psia]. The stream enters heat exchanger 10 and is cooled to −35 ° C. [−31 ° F.] and partially condensed by the partially warmed expanded refrigerant stream 71f and other refrigerant streams. In the simulation of FIG. 1, these different refrigerant streams are assumed to be propane refrigerants sold at three different temperature and pressure levels. The partially condensed refrigerant stream 71a then enters the heat exchanger 13 and is further cooled to -81 ° C [-114 ° F] by the partially warmed expanded refrigerant stream 71e and the refrigerant to be partially subcooled. (Stream 71b) is condensed. The refrigerant is further subcooled to -160 ° C [-257 ° F] in the heat exchanger 60 by the expanded refrigerant stream 71d. Subcooled liquid stream 71c enters work expansion device 63 and the device is mechanically driven from the stream such that it is expanded isotropically from a pressure of about 4,040 kPa (a) [586 psia] to a pressure of about 234 kPa (a) [34 psia]. Extract energy. A portion of the expanded stream is vaporized to cool the entire stream (stream 71d) to -164 ° C [-263 ° F]. The expanded stream 71d then reenters the heat exchangers 60, 13 and 10 and cools the vaporized and superheated streams 38c and 35 and the refrigerant (streams 71, 71a and 71b).

The superheated refrigerant vapor (stream 71g) leaves the heat exchanger 10 at 34 ° C. [93 ° F.] and is compressed in three stages to 4,254 kPa (a) [617 psia]. Each of the three compression stages (refrigerant compressors 64, 66, 68) is driven by an additional power source followed by a cooling device (exhaust cooling devices 65, 67, 69) to remove the heat of compression. The compressed stream 71 from the exhaust cooling apparatus 69 returns to the heat exchanger 10 to complete the cycle.

An overview of the stream flow rate and energy consumption of the process shown in FIG. 1 is described in the table below.

Table 1

(Figure 1)

Stream Flow Rate Overview- Kg moles / Hr [Lb.Moles / Hr]

Stream methane ethane Propane Bhutan + all 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  3,184 50 37,736   176     0     0 37,916

Recovery of NGL *

Ethane 95.06%

Propane 100.00%

Butane + 100.00%

Production rate 308,147 Lb / Hr [308,147 kg / Hr]

LNG products

Production rate 610,813 Lb / Hr [610,813 kg / Hr]                 

Purity * 99.52%

Low calorific value 912.3 BTU / SCF [33.99 MJ / ㎥]

power

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

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

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

Heat of auxiliary equipment

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

* (Based on unrounded flow)

The efficiency of LNG production processing is usually compared using the required “specific power consumption”, which is the ratio of the total cooling compression power to the total liquid production rate. Published non-powered information of the prior art for producing LNG ranges from 0.276 kW-Hr / kg [0.168 HP-Hr / Lb] to 0.300 kW-Hr / kg [0.182 HP-Hr / Lb]. It is believed that this was based on 340 days of operation at the LNG production plant. On the same basis, the non-power consumption of the embodiment of FIG. 1 of the present invention is 0.265 kW-Hr / kg [0.161 HP-Hr / Lb], which gives an efficiency improvement of 4% to 13% over the prior art treatment. do. In addition, the non-power consumption of prior art processing is not only NGL (C ₂ and heavy hydrocarbon) liquid streams, but LPG (C ₃ and heavy hydrocarbons) or relatively low recovery levels of condensate (C) as shown in this example of the invention. ₄ and heavier hydrocarbons) is based on co-producing only the liquid stream. Prior art processes require significantly more cooling power to co-produce NGL streams instead of LPG streams or condensate streams.

There are two main factors that account for the improved efficiency of the present invention. The first factor can be seen by examining the thermodynamics of the liquefaction process when applied to the high pressure gas stream as contemplated in this example. Since the main component of this stream is methane, the thermodynamic properties of methane can be used to compare the cycle used in the present invention with the liquefaction cycle employed in the prior art treatment. 2 includes a pressure-enthalpy state diagram of methane. In most liquefaction cycles of the prior art, all cooling of the gas stream is achieved where the stream is at high pressure (paths A-B), after which the stream is expanded (paths B-C) to the pressure of the LNG storage vessel (slightly higher than atmospheric pressure). . This expansion step can employ a work expansion device, which can typically recover about 75% to 80% of the theoretically available work with ideal isentropic expansion. For simplicity, the overall isentropic expansion is shown in Figure 2 by paths B-C. Even so, since the constant entropy line is nearly perpendicular in the liquid region of the state diagram, the enthalpy reduction provided by this work expansion is very small.

In contrast, the liquefaction cycle of the present invention will be described. After partial cooling to high pressure (paths A-A '), the gas stream is workpiece expanded to medium pressure (paths A'-A ") (again, for the sake of simplicity the total isentropic expansion). The remainder of the cooling is at medium pressure (paths A ″ -B ′), and the stream is then expanded to the pressure of the LNG storage vessel (paths B′-C). Since a constant entropy slope is not steep in the vapor region of the state diagram, a significantly greater enthalpy reduction is provided by the first work expansion step (path A′-A ″) of the present invention. Therefore, the total amount of cooling required for the present invention (sum of paths A-A 'and paths A-B') is less than the amount of cooling (paths A-B) required for the prior art processing, and the gas stream is liquefied. The amount of cooling (and thus cooling compression) required to do so is reduced.

A second factor describing the improved efficiency of the present invention is the excellent performance of the hydrocarbon distillation system at mastic copper pressure. Most of the hydrocarbon removal steps in the prior art processes are carried out at high pressures and use a scrub column which typically employs a cooling hydrocarbon liquid as the absorbent stream to remove heavy hydrocarbons from the inlet gas stream. Operating the scrub column at high pressure is not very efficient, as a result of which significant fractions of methane and ethane are co-absorbed from the gas stream, which must be continuously stripped from the sorbent liquid and cooled to be part of the LNG product. In the present invention, the hydrocarbon removal step is led to a medium pressure in which the vapor-liquid equilibrium is more advantageous, with the result that the desired heavy hydrocarbons of the by-product liquid stream are recovered very efficiently.

Second embodiment

If the specification for the LNG product allows more ethane contained in the feed gas recovered from the LNG product, a simpler embodiment may be employed in the present invention. 3 illustrates this alternative embodiment. The inlet gas composition and conditions considered in the processing shown in FIG. 3 are the same as in FIG. 1. Thus, the process of FIG. 3 can be compared with the embodiment shown in FIG.

In the simulation of the FIG. 3 process, the inlet gas cooling, separation, and expansion design for the NGL recovery is essentially the same as that used in FIG. 1. The inlet gas enters the plant as stream 31 at 32 ° C. [90 ° F.] and 8860 kPa (a) [1285 psia], and is -37 by the reboiler liquid (stream 40) on the refrigerant stream and on the demethanizer side. It is cooled in the heat exchanger 10 by heat exchange at [° C.-35 ° F.]. The cooled stream 31a is subjected to −34 ° C. [−30 ° F.] and 8,812 kPa (a) [1278 psia] in a separation device 11 where steam (stream 32) is separated from condensed liquid (stream 33). Enter.

Steam (stream 32) is separated from the separation device 11 into two streams 34, 36. Stream 34 containing about 20% of the total vapor is combined with stream 33, the condensed liquid, to form stream 35. The combined stream 35 is cooled by passing through a heat exchanger 13 of heat exchange associated with the refrigerant stream 71e to condense the stream 35a generally. The stream 35a, typically condensed to -85 [deg.] C. [-120 [deg.] F., is then operated at an operating pressure of the fractionation tower 19 (approximately 3206 kPa (a) [465 psia]) through a suitable expansion device such as an expansion valve 14. Into the flash. A portion of the expanded stream is vaporized so that the entire stream is cooled. In the process shown in FIG. 3, the expanded stream 35b leaving the expansion valve 14 reaches a temperature of -86 ° C. [-122 ° F.], and the separator section (in the upper region of the fractionation tower 19) separator section). The separated liquid here becomes the top feed of the methanation in the lower region of the fractionation tower 19.

The remaining 80% of the steam from the separation device 11 (stream 36) enters the work expansion device 15 where mechanical energy is extracted from this portion of the high pressure feed. Apparatus 15 expands the vapor from a pressure of about 8,812 kPa (a) [1278psia] to the tower operating pressure substantially isentropically, with the workpiece expansion at a temperature of approximately −75 ° C. [−103 ° F.]. 36a is cooled. Expanded and partially condensed stream 36a is sent as feed to the column-intermediate feed point on distillation column 19.

The cold demethanizer overhead steam (stream 37) exits the top of the fractionation tower 19 at -86 ° C [-123 ° F]. The liquid product stream 41 exits the bottom of the tower to 48 ° C. [118 ° F.] based on a typical specification of 0.020: 1 on a molar basis for the bottom product.

The demethanizer overhead steam (stream 37) is warmed in heat exchanger 24 to 32 ° C. [90 ° F.], which portion (stream 48) is then taken out to act as fuel gas for the plant. The remainder of the warmed demethanizer overhead vapor (stream 49) is compressed by compressor 16. After cooling to 38 [deg.] C. [100 [deg.] F.] in the exhaust cooling unit 25, stream 49b is exchanged at -80 [deg.] C. in heat exchanger 24 by cross-changing stream 37, which is a cooling demethanizer overhead vapor. Further cooled to -112 [deg.] F.

Stream 49c is then further cooled to -160 ° C [-257 ° F] by refrigerant stream 71d to enter heat exchanger 60 to condense and supercool the stream, after which mechanical energy is extracted from the stream. Enter the expansion device 61. The device 61 expands the liquid stream 49d isotropically from a pressure of about 4,021 kPa (a) [583psia] to an LNG storage pressure 107 kPa (a) [15.5 psia] slightly above atmospheric pressure. The work expansion cools the expanded stream 49e to a temperature of approximately −161 ° C. [−258 ° F.], which is then sent to the LNG storage tank 62 containing the LNG product (stream 50).

Similar to the process of FIG. 1, total cooling for streams 35 and 49c is provided by a closed cycle cooling loop. The stream used as the working fluid of the cycle of FIG. 3 has a composition of 7.5% nitrogen, 40.0% methane, 42.5% ethane, and 10.0% propane, in approximately mole percent, with the remainder consisting of heavy hydrocarbons. The refrigerant stream 71 leaves the exhaust cooling device 69 at 38 ° C. [100 ° F.] and 4,185 kPa (a) [607 psia]. The stream enters heat exchanger 10 and is cooled to −35 ° C. [−31 ° F.] and partially condensed by the partially warmed expanded refrigerant stream 71f and other refrigerant streams. In the simulation of FIG. 3, these different refrigerant streams are assumed to be propane refrigerants sold at three different temperature and pressure levels. The partially condensed refrigerant stream 71a then enters the heat exchanger 13 and is further cooled to -85 ° C [-121 ° F] by the partially warmed expanded refrigerant stream 71e and the refrigerant (stream 71 Condenses)) and partially cools. The refrigerant is further subcooled to -160 ° C [-257 ° F] in the heat exchanger 60 by the expanded refrigerant stream 71d. Subcooled liquid stream 71c enters work expansion device 63 and the device is mechanically driven from the stream such that it is expanded isotropically from a pressure of about 4,040 kPa (a) [586 psia] to a pressure of about 234 kPa (a) [34 psia]. Extract energy. A portion of the expanded stream is vaporized to cool the entire stream (stream 71d) to -164 ° C [-263 ° F]. The expanded stream 71d then reenters the heat exchangers 60, 13 and 10 and cools the vaporized and superheated streams 49c and 35 and the refrigerant (streams 71, 71a and 71b).                 

The superheated refrigerant vapor (stream 71g) leaves the heat exchanger 10 at 34 ° C. [93 ° F.] and is compressed in three steps to 4,254 kPa (a) [617 psia]. Each of the three compression stages (refrigerant compressors 64, 66, 68) is driven by an additional power source and follows a cooling device (exhaust cooling devices 65, 67, 69) for the removal of compressed heat. Compressed stream 71 from exhaust cooling apparatus 69 returns to heat exchanger 10 to complete the cycle.

An overview of the stream flow rate and energy consumption of the process shown in FIG. 3 is described in the table below.

TABLE 2

(Figure 3)

Stream Flow Rate Overview- Kg moles / Hr [Lb.Moles / Hr]

Stream methane ethane Propane Bhutan + all 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

Recovery of NGL *

Ethane 87.57%

Propane 97.41%

Butane + 99.47%

Production rate 296,175 Lb / Hr [296,175 kg / Hr]

LNG products

Production rate 625,152 Lb / Hr [625,152 kg / Hr]

Purity * 98.66%

Low calorific value 919.7 BTU / SCF [34.27 MJ / ㎥]

power

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

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

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

Heat of auxiliary equipment

Demethanizer Reboiler 22,177 MBTU / Hr [14,326 kW]

* (Based on unrounded flow)

Assuming that the LNG production plant has been operating for 340 days per year, the specific power consumption of the embodiment of FIG. 3 of the present invention is 0.251 kW-Hr / kg [0.153 HP-Hr / Lb]. Compared to prior art processing, the embodiment of FIG. 3 gives an efficiency improvement of 10% to 20%. As mentioned earlier in the embodiment of FIG. 1, this efficiency improvement is possible in the present invention even though NGL by-products are produced rather than LPG or condensate by-products produced by prior art processing.

Compared to the embodiment of FIG. 1, the embodiment of FIG. 3 of the present invention requires about 5% less power per unit of liquid produced. Thus, at a given usable amount of compressive power, the embodiment of FIG. 3 can liquefy about 5% more natural gas than the embodiment of FIG. 1 by less recovery of C ₂ and heavy hydrocarbons of NGL by-products. The choice for a particular application between the embodiments of FIGS. 1 and 3 of the present invention is based on the monetary value of the heavy hydrocarbons of the NGL product or the calorific value for the LNG product (the implementation of FIG. 1) for the corresponding value of the LNG product. Generally because the calorific value of the LNG produced by the example is lower than the calorific value produced by the embodiment of FIG. 3).

Third embodiment

If the specification for the LNG product allows more ethane contained in the feed gas recovered from the LNG product, or if there is no market for ethane-containing liquid by-products, an alternative embodiment of the present invention as shown in Figure 4 is an LPG by-product. It may be employed to produce a stream. Influent gas compositions and conditions considered in the processing shown in FIG. 4 are the same as in FIGS. 1 and 3. Thus, the process of FIG. 4 can be compared with the embodiment shown in FIGS. 1 and 3.

In the simulation of the process of FIG. 4, the inlet gas enters the plant as stream 31 at 32 ° C. [90 ° F.] and 8860 kPa (a) [1285 psia], which is caused by the refrigerant stream and flashed separator liquid 33a. By heat exchange, it cools in the heat exchanger 10 to -43 degreeC [-46 degreeF]. The cooled stream 31a is subjected to -18 ° C [-1 ° F] and 8,812 kPa (a) [1278psia] in a separation device 11 where steam (stream 32) is separated from condensed liquid (stream 33). Enter.

Steam (stream 32) from separation device 11 enters work expansion device 15 which extracts mechanical energy from this portion of the high pressure feed. The device 15 expands the vapor substantially isentropically from a pressure of about 8,812 kPa (a) [1278psia] to a pressure of about 3,034 kPa (a) [440psia] (the operating pressure of the separator / absorber tower 18). The work expansion causes cooling of the expanded stream 32a to a temperature of approximately −63 ° C. [−81 ° F.]. Expanded and partially condensed stream 32a is fed to absorber 18b in the lower region of separator / absorber tower 18. The liquid portion of the expanded stream mixes with the liquid falling down from the absorbent portion, and the combined liquid stream 40 exits the bottom of the separator / absorber tower 18 at -66 ° C [-86 ° F]. The vapor portion of the expanded stream rises upwards through the absorbing portion and contacts the cooling liquid falling downwards to condense and absorb the C ₃ and heavy components.

Separator / absorber tower 18 is a prior art distillation column comprising a plurality of vertically spaced trays, one or more packed beds, or a combination of trays and packings. Although common in natural gas processing plants, the separator / absorber tower consists of two parts. The upper portion 18a is a tower in which any vapor contained in the top feed is separated from its corresponding liquid portion, and vapors rising from the lower distillation or absorber 18b are combined with the steam portion (if any) of the top feed. Separator to form a cold distillation stream 37 exiting the top. The lower absorber 18b includes a tray and / or packing and provides the necessary contact between the liquid falling downward and the vapor rising upward to condense and absorb the C ₃ and heavy components.

The combined liquid stream 40 from the bottom of the separator / absorber tower 18 is sent to the heat exchanger 13 by a pump 26 and the stream 40a is deethanized unit overhead (stream 42 And cooling of the refrigerant (stream 71a). The combined liquid stream is heated to −31 ° C. [−24 ° F.] and partially vaporizes stream 40b before the stream is supplied as a column-intermediate feed of deethanization apparatus 19. Separator liquid (stream 33) is passed to expansion valve 12 which cools stream 33 to -43 ° C. [-46 ° F.] (stream 33a) before the stream cools the inlet feed gas described above. By flash expansion slightly above the operating pressure of the deethanization apparatus 19. Stream 33b is currently 29 ° C. [85 ° F.], and then enters the middle-to-column bottom feed point of deethanization unit 19. In the deethanization apparatus, streams 40b and 33b are stripped of the methane and C ₂ components of the stream. The deethanation apparatus of tower 19 operating at about 3123 kPa (a) [453 psia] also includes prior art distillation comprising a plurality of vertically spaced trays, one or more packed beds, or a combination of trays and packings. It is a column. The deethanizer tower has two parts: any vapor contained in the top feed is separated from its corresponding liquid portion, and the vapors raised from the bottom distillation or deethanization 19b are vapors of the top feed (if Top separation device portion 19a, which is combined with (if present) to form a distillation stream 42 exiting the top of the tower; It comprises a tray and / or packing and consists of a lower deethanization 19b which provides the requisite contact between the liquid falling downward and the vapor rising upward. The deethanization section 19b also includes one or more reboilers (reboilers 20), which heat and vaporize a portion of the bottom liquid of the column to vaporize stream 41 which is a liquid product of methane and C ₂ components. Provide stripping vapor flowing over the column to remove. Typical specifications for bottom liquid products have a ratio of ethane and propane on a molar basis of 0.020: 1. The liquid product stream 41 exits the bottom of the deethanization apparatus at 101 ° C. [214 ° F.].

The operating pressure of the deethanizer 19 is maintained slightly above the operating pressure of the separator / absorber tower 18. This causes deethanizer overhead steam (stream 42) to flow to the top of separator / absorber tower 18 in a pressure stream through heat exchanger 13. In the heat exchanger 13, the deethanization unit overhead vapor of -28 ° C [-19 ° F] is combined with the liquid stream (stream 40a) combined from the bottom of the separator / absorber tower 18 to -67 ° C. Cool to [−89 ° F.] (stream 42a) and heat exchange with flashed refrigerant stream 71e which partially condenses. The partially condensed stream enters a reflux drum 22, which separates the condensed liquid (stream 44) from uncondensed vapor (stream 43). Stream 43 is combined with a distillation vapor stream (stream 37) leaving the upper region of separator / absorber tower 18 to form a cooling residual gas stream 47. The condensed liquid (stream 44) is pumped to high pressure by pump 23, after which stream 44a is divided into two parts. A portion of stream 45 is directed to the upper separator section of separator / absorber tower 18 and acts as a cooling liquid in contact with the vapor rising upward through the absorber. The other portion is fed to the deethanization unit 19 as a countercurrent stream 46 flowing at -67 ° C [-89 ° F] to the top feed point of the deethanization unit 19.

The cooling residual gas (stream 47) is warmed from −70 ° C. [−94 ° F.] to 34 ° C. [94 ° F.] in the heat exchanger 24, and the portion (stream 48) is then taken out and the fuel gas for the plant. Act as. The remainder of the warmed residual gas (stream 49) is compressed by the compression device 16. After cooling to 38 [deg.] C. [100 [deg.] F.] in the exhaust cooling apparatus 25, stream 49b is subjected to -61 [deg.] F. [-78 [deg.] F.] in the heat exchanger 24 by cross-exchange of stream 47 which is cooling residual gas. Is further cooled.

Stream 49c then enters heat exchanger 60 and is further cooled to -160 ° C [-255 ° F] by refrigerant stream 71d to condense and supercool the stream, after which mechanical energy is extracted from the stream. The work inflation device 61 enters. Apparatus 61 expands the liquid stream 49d substantially isotropically from a pressure of about 4,465 kPa (a) [648 psia] to an LNG storage pressure 107 kPa (a) [15.5 psia] slightly above atmospheric pressure. The work expansion will cool the expanded stream 49e to a temperature of approximately -160 ° C [-256 ° F], which is then sent to the LNG storage tank 62 containing the LNG product (stream 50).

Similar to the processing of FIGS. 1 and 3, multiple cooling for stream 42 and total cooling for stream 49c are provided by a closed cycle cooling loop. The stream used as working fluid for the cycle of the FIG. 4 process has a composition of 8.7% nitrogen, 30.0% methane, 45.8% ethane, and 11.0% propane, in approximately mole percent with the remainder consisting of heavy hydrocarbons. The refrigerant stream 71 leaves the exhaust cooling device 69 at 38 ° C. [100 ° F.] and 4,185 kPa (a) [607 psia]. The stream enters heat exchanger 10 and is cooled to -27 ° C [-17 ° F] and partially condensed by the partially warmed expanded refrigerant stream 71f and other refrigerant streams. In the simulation of FIG. 4, these different refrigerant streams are assumed to be propane refrigerants sold at three different temperature and pressure levels. The partially condensed refrigerant stream 71a then enters the heat exchanger 13 and is further cooled to -67 ° C [-89 ° F] by the partially warmed expanded refrigerant stream 71e and the refrigerant (stream 71b). More condensation). The refrigerant is condensed entirely and then supercooled to -160 ° C [-255 ° F] in the heat exchanger 60 by the expanded refrigerant stream 71d. The supercooled liquid stream 71c enters the work expansion device 63, which is in such an entropy expansion that it is generally isentropically expanded from a pressure of about 4,040 kPa (a) [586 psia] to a pressure of about 234 kPa (a) [34 psia]. Mechanical energy is extracted from the A portion of the expanded stream is vaporized to cool the entire stream (stream 71d) to -164 ° C [-264 ° F]. Expanded stream 71d then re-enters stream 49c, 42 and heat exchangers 60, 13, 10 that cool the vaporized superheated refrigerant (streams 71, 71a, 71b).

The superheated refrigerant vapor (stream 71g) leaves the heat exchanger 10 at 32 ° C. [90 ° F.] and is compressed in three stages to 4,254 kPa (a) [617 psia]. Each of the three compression stages (refrigerant compressors 64, 66, 68) is driven by an additional power source and follows a cooling device (exhaust cooling devices 65, 67, 69) for the removal of compressed heat. Compressed stream 71 from exhaust cooling apparatus 69 returns to heat exchanger 10 to complete the cycle.

An overview of the stream flow rate and energy consumption of the process shown in FIG. 4 is described in the table below.

TABLE 3

(Figure 4)

Stream Flow Rate Overview- Kg moles / Hr [Lb.Moles / Hr]

Stream methane ethane Propane Bhutan + all 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     One     0  2,684 50 38,524 3,585    21     0 42,135

Recovery of LPG *

Propane 99.08%

Butane + 100.00%

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

LNG products

Production rate 726,918 Lb / Hr [726,918 kg / Hr]

Purity * 91.43%

Low calorific value 969.9 BTU / SCF [36.14 MJ / ㎥]

power

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

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

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

Heat of auxiliary equipment

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

* (Based on unrounded flow)

Assuming that the LNG production plant has been operating for 340 days per year, the non-power consumption of the embodiment of FIG. 4 of the present invention is 0.236 kW-Hr / kg [0.143 HP-Hr / Lb]. Compared to prior art processing, the embodiment of FIG. 4 confers an efficiency improvement of 17% to 27%.

Compared to the embodiment of FIGS. 1 and 3, the embodiment of FIG. 4 of the present invention requires about 6% to 11% less power per unit of liquid produced. Thus, at a given usable amount of compressive power, the embodiment of FIG. 4 is about 6% more natural gas or the embodiment of FIG. 3 than the embodiment of FIG. 1 by recovering only C ₃ and heavy hydrocarbons of LPG by-products. About 11% more natural gas can be liquefied. The choice for a particular application between the embodiment of FIG. 4 and one of the embodiments of FIG. 1 or 3 of the present invention is based on the monetary value of the ethylene or the monetary value of the LNG product to the corresponding value of the LNG product. Generally required by one of the specifications (since the calorific value of the LNG produced by the embodiment of FIGS. 1 and 3 is lower than the calorific value produced by the embodiment of FIG. 4).

Fourth embodiment

If the specification for the LNG product permits more ethane and propane contained in the feed gas recovered from the LNG product, or if there is no market for liquid by-products containing ethane and propane, optional implementation of the invention as shown in FIG. An example may be employed to produce the condensate byproduct stream. Influent gas compositions and conditions considered in the processing shown in FIG. 5 are the same as those in FIGS. 1, 3 and 4. Thus, the process of FIG. 5 can be compared with the embodiment shown in FIGS. 1, 3 and 4.

In the simulation of the FIG. 5 process, the inlet gas enters the plant as stream 31 at 32 ° C. [90 ° F.] and 8,860 kPa (a) [1285 psia], flashed to a refrigerant stream, −38 ° C. [−37 ° F.]. It is cooled in the heat exchanger 10 by heat exchange with a high pressure separator liquid (stream 33b) and an intermediate pressure separator liquid (stream 39b) flashed at -38 ° C [-37 ° F]. The cooled stream 31a is supplied at −34 ° C. [−30 ° F.] and 8,812 kPa (a) [1278psia] in a high pressure separation unit 11 where steam (stream 32) is separated from condensed liquid (stream 33). Enter

Steam (stream 32) from separation device 11 enters work expansion device 15 which extracts mechanical energy from this portion of the high pressure feed. The device 15 expands the vapor isotropically, from a pressure of about 8,812 kPa (a) [1278 psia] to a pressure of about 4,378 kPa (a) [635 psia], with the work expansion being approximately -64 ° C. [-83 ° F.]. At this temperature, the expanded stream 32a is cooled. The expanded and partially condensed stream 32a enters the intermediate pressure separation device 18, and vapor (stream 42) is separated from the condensation liquid (stream 39). The intermediate pressure separator device (stream 39) cools stream 39 to −78 ° C. [−108 ° F.] before entering heat exchanger 13 (stream 39a) and residual gas stream 49 and The expansion valve 17 provided to the heat exchanger 10 to cool the refrigerant stream 71a to cool the inlet feed gas described above is flash expanded and heated slightly above the operating pressure of the depropanification apparatus 19. Stream 39c is currently -26 [deg.] C. [-15 [deg.] F.] and then enters the middle-to-column top feed point of the depropanification apparatus 19.

Stream 33, which is a liquid condensed from the high pressure separation device 11, cools stream 33 to -70 ° C [-93 ° F] (33a) before entering heat exchanger 13 and residual gas stream 49 ) And an expansion valve provided to the heat exchanger 10 to cool the refrigerant stream 71a to cool the inlet feed gas described above, which is slightly expanded above the operating pressure of the depropanification apparatus 19 and heated. Stream 33c is currently at 10 ° C. [50 ° F.], and then enters the middle-to-column bottom feed point of the depropanification apparatus 19. In a depropane apparatus, streams 39c and 33c remove its methane, C ₂ and C ₃ components. The depropanification apparatus of tower 19 operating at about 2,654 kPa (a) [385 psia] comprises a prior art distillation column comprising a plurality of vertically spaced trays, one or more packed beds, or a combination of trays and packings. to be. The depropanizer tower is equipped with two parts: any vapor contained in the top feed is separated from its corresponding liquid portion, and the vapors raised from the bottom distillation or depropanate 19b, if any, Top separation device portion 19a which is combined with to form a distillation stream 37 exiting the top of the tower; It comprises a tray and / or packing and consists of a lower depropane plate 19b which provides the requisite contact between the liquid falling down and the vapor rising upwards. The depropane plate 19b also includes one or more reboilers (reboilers 20), which heat and vaporize a portion of the bottom liquid of the column to produce a stream of liquid products of methane, C ₂ and C ₃ components ( Provide stripping vapor flowing over the column to remove 41). Typical specifications for bottom liquid products have a ratio of propane to butane on a molar basis of 0.020: 1. The liquid product stream 41 exits the bottom of the deethanization apparatus at 141 ° C. [286 ° F.].

The overhead distillation stream 37 leaves the depropane platen unit 19 at 2 ° C. [36 ° F.] and is cooled and partially condensed by the commercial propane refrigerant of the countercurrent condenser 21. Partially condensed stream 37a enters countercurrent drum 22 at −17 ° C. [2 ° F.], which separates condensed liquid (stream 44) from uncondensed vapor (stream 43). The condensed liquid (stream 44) is pumped by pump 23 to the top feed point of the depropane platen apparatus 19 as countercurrent stream 44a.

Uncondensed steam (stream 43) from countercurrent drum 22 is warmed to 34 ° C. [94 ° F.] in heat exchanger 24, and one portion (stream 48) is then taken out as fuel gas for the plant. Works. The remainder of the warmed vapor (stream 38) is compressed by the compression device 16. After cooling to 38 ° C. [100 ° F.] in the exhaust cooling unit 25, stream 38b is further cooled to −9 ° C. [15 ° F.] in heat exchanger 24 by cross-exchange of stream 43 which is cooling steam. Is cooled.

Stream 38c is then combined with intermediate pressure separator vapor (stream 42) to form a cooling residual gas stream 49. Stream 49 then enters heat exchanger 13 and from -39 [deg.] C. [-38 [deg.] F.] to -74 [deg.] C. by means of separator liquid (streams 39a, 33a) and refrigerant stream 71d as described above. Further cooled to [-102 ° F]. Partially condensed stream 49a then enters heat exchanger 60 and is further cooled to -159 [deg.] C. [-254 [deg.] F.] to condense by refrigerant stream 71d to overcool. Enter the work expansion device 61 extracted from the. Apparatus 61 expands liquid stream 49b isotropically from a pressure of about 4282 kPa (a) [621 psia] to an LNG storage pressure (107 kPa (a) [15.5 psia]) slightly above atmospheric pressure. The work expansion will cool the expanded stream 49c to a temperature of approximately -159 ° C. [-255 ° F.], which is then sent to an LNG storage tank 62 containing the LNG product (stream 50).

Similar to the processing of FIGS. 1, 3, and 4, multiple cooling for stream 49 and total cooling for stream 49a are provided by a closed cycle cooling loop. The stream used as working fluid for the cycle of the FIG. 5 process has a composition of 8.9% nitrogen, 34.3% methane, 41.3% ethane, and 11.0% propane, in approximately mole percent with the remainder consisting of heavy hydrocarbons. The refrigerant stream 71 leaves the exhaust cooling device 69 at 38 ° C. [100 ° F.] and 4,185 kPa (a) [607 psia]. The stream enters heat exchanger 10 and is cooled to −34 ° C. [−30 ° F.] and partially condensed by the partially warmed expanded refrigerant stream 71f and other refrigerant streams. In the simulation of FIG. 5, these different refrigerant streams are assumed to be propane refrigerants sold at three different temperature and pressure levels. The partially condensed refrigerant stream 71a then enters the heat exchanger 13 and is further cooled to -74 ° C [-102 ° F] by the partially warmed expanded refrigerant stream 71e and the refrigerant (stream 71b). More condensation). The refrigerant is condensed entirely and then subcooled to -159 ° C [-254 ° F] in the heat exchanger 60 by the expanded refrigerant stream 71d. The supercooled liquid stream 71c enters the work expansion device 63, which is expanded from the stream to isentropically expanded from a pressure of about 4,040 kPa (a) [586 psia] to a pressure of about 234 kPa (a) [34 psia]. Extract mechanical energy. A portion of the expanded stream is vaporized to cool the entire stream (stream 71d) to -164 ° C [-264 ° F]. Expanded stream 71d then re-enters heat exchanger 60, 13, 10 to cool streams 49a and 49 and the vaporized superheated refrigerant (streams 71, 71a and 71b).

The superheated refrigerant vapor (stream 71g) leaves the heat exchanger 10 at 34 ° C. [93 ° F.] and is compressed in three stages to 4,254 kPa (a) [617 psia]. Each of the three compression stages (refrigerant compressors 64, 66, 68) is driven by an additional power source and follows a cooling device (exhaust cooling devices 65, 67, 69) for the removal of compressed heat. Compressed stream 71 from exhaust cooling apparatus 69 returns to heat exchanger 10 to complete the cycle.

An overview of the stream flow rate and energy consumption of the process shown in FIG. 5 is described in the table below.

Table 4

(Figure 5)

Stream Flow Rate Overview- Kg moles / Hr [Lb.Moles / Hr]

Stream methane ethane Propane Bhutan + all 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 26,166 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

Recovery of Condensate *

Butane 95.04%                 

Pentane + 99.57%

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

LNG products

Production rate 834,183 Lb / Hr [834,183 kg / Hr]

Purity * 87.27%

Low calorific value 1033.8 BTU / SCF [38.52 MJ / ㎥]

power

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

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

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

Heat of auxiliary equipment

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

* (Based on unrounded flow)

Assuming that the LNG production plant has been operating for 340 days per year, the non-power consumption of the embodiment of FIG. 5 of the present invention is 0.238 kW-Hr / kg [0.145 HP-Hr / Lb]. Compared to the prior art processing, the embodiment of FIG. 5 gives an efficiency improvement of 16% to 26%.

In contrast to the embodiment of FIGS. 1 and 3, the embodiment of FIG. 5 of the present invention requires about 5% to 10% less power per unit of liquid produced. Compared to the embodiment of FIG. 4, the embodiment of FIG. 5 of the present invention requires essentially the same power as the power per unit of liquid produced. Thus, at the given usable amount of compressive power, the embodiment of FIG. 5 is approximately 10% more natural gas or the embodiment of FIG. 4 than the embodiment of FIG. 3 by recovery of only C ₄ and heavy hydrocarbons of the condensate by-product. The same amount of natural gas can be liquefied. The choice for a particular application between the embodiment of FIG. 5 and one of the embodiments of FIG. 1, FIG. 3 or FIG. 4 of the present invention is based on the monetary value of ethane and propane of NGL or LPG products over the corresponding value of LNG products. Or as generally required by one of the specifications of the calorific value for the LNG product (since the calorific value of the LNG produced by the embodiment of FIGS. 1, 3 and 4 is lower than the calorific value produced by the embodiment of FIG. 5). do.

Another embodiment

It will be appreciated by those skilled in the art that the present invention can be used in any type of LNG liquefaction plant to co-produce NGL streams, LPG streams, or condensate streams to suit the needs of a given plant location. It can also be appreciated that various treatment configurations can be employed to recover the liquid byproduct stream. For example, the embodiments of FIGS. 1 and 3 may allow recovery of LPG streams or condensate streams as liquid by-product streams rather than NGL streams as described above in the first and second embodiments. The embodiment of FIG. 4 recovers an NGL stream containing a significant fraction of the C ₂ components present in the feed gas rather than producing LPG by-products as described above in the third embodiment, or C ₄ present in the feed gas. It is possible to recover the condensate stream containing only components and heavy components. The embodiment of FIG. 5 recovers an NGL stream containing a significant fraction of the C ₂ component present in the feed gas rather than producing condensate by-products as described above in Example 4, or the C ₃ component present in the feed gas. It is possible to recover an LPG stream containing a significant fraction of.

1, 3, 4 and 5 show preferred embodiments of the invention for the indicated processing conditions. 6-21 illustrate alternative embodiments of the invention that may be considered for a particular application. As shown in FIGS. 6 and 7, all or part of the liquid condensed from the separator 11 (stream 33) is part of the separator vapor (stream 34) entering the heat exchanger 13. Rather than incorporating the ss, it may be fed to a fractionation tower 19 at a separate bottom column-intermediate feed location. FIG. 8 illustrates an alternative embodiment of the present invention that requires less equipment than the embodiment of FIGS. 1 and 6, although its non-power consumption is rather high. Similarly, FIG. 9 shows an alternative embodiment of the present invention that requires less equipment than the embodiment of FIGS. 3 and 7 with high non-power consumption costs. 10-14 illustrate alternative embodiments of the present invention that require less equipment than the embodiment of FIG. 4, although their non-power consumption may be high (de-ethanization, as shown in FIGS. 10-14). Note that a distillation column or system, such as apparatus 19, includes both the reboiler absorber tower design and countercurrent, reboiled tower design.). 15 and 16 show that a single fractionation column 19 also functions as the separator / absorber tower 18 and deethanization device 19 of the embodiment of FIGS. 4, 10-14. Depending on the amount of heavy hydrocarbons in the feed gas and feed gas pressure, the cooled feed stream 31a leaving the heat exchanger 10 may not contain any liquid (the stream is above the dew point, or Or because of exceeding the maximum critical pressure (cricondenbar)), the separation device 11 shown in FIGS. 1, 3 to 16 is not required, and the cooled feed stream is subjected to a suitable expansion such as the work expansion device 15. Can flow directly to the device.

The gas stream remaining after recovery of the liquid by-product stream (stream 37 of FIGS. 1, 3, 6-11, 13 and 14, before being fed to the heat exchanger 60 for condensation and subcooling), The arrangement of the streams 47 of FIGS. 4, 12, 15 and 16 and the stream 43 of FIG. 5 may be made in various ways. 1, 3 to 16, the stream is compressed to high pressure using energy from one or more work expansion devices, partially cooled in an exhaust cooling device, and then further by cross-exchange by the original stream. Is cooled. As shown in FIG. 17, it may be advantageous for many applications to compress the stream at high pressure, for example using an additional compression device 59 driven by an external power source. As shown by the equipment indicated by dashed lines in FIGS. 1, 3 to 16 (heat exchanger 24 and exhaust cooling apparatus 25), in many situations compression before entering heat exchanger 60. It may be advantageous to reduce the capital cost of the plant (increasing the cooling load of the heat exchanger 60 and increasing the power consumption of the refrigerant compressors 64, 66, 68) by reducing or eliminating the precooling of the circulated stream. have. In this case, the stream 49a leaving the compression device may also flow directly to the heat exchanger 24 as shown in 18 or directly to the heat exchanger 60 as shown in FIG. 19. If the work expansion device is not used to inflate any portion of the high pressure feed gas, a compression device driven by an external power source such as the compression device 59 shown in FIG. 20 may be used instead of the compression device 16. . In other situations no compression of the stream can be justified, so that the stream has a heat exchanger 60 as shown in FIG. 21 and equipment indicated by dashed lines in FIGS. 1, 3 to 16 (heat exchanger 24, compression). Directly by the apparatus 16 and the exhaust cooling apparatus 25. If heating the stream before the plant fuel gas (stream 48) is taken out is not included in the heat exchanger 24, an additional heater 58 may be used as shown in FIGS. 19-21 before the fuel gas is consumed. Similarly, it may be required to heat the fuel gas using an auxiliary plant stream or other process stream to provide the necessary heat. The selection conditions that must generally be evaluated for each application must all take into account factors such as gas consumption, plant size, desired by-product stream recovery levels and available equipment.

According to the invention, the cooling of the inlet gas stream and the feed stream in the LNG production unit may be accomplished in various ways. In the processing of FIGS. 1, 3, 6-9 the inlet gas stream 31 is cooled and condensed by an external refrigerant stream and tower liquid from the fractionation tower 19. 4, 5, 10-14, flashed separator device liquid is used for this purpose along with the external refrigerant stream. 15 and 16, tower liquid and flashed separator device liquid are used for this purpose along with the external refrigerant stream. In FIGS. 17-21 only external refrigerant streams are used to cool the inlet gas stream 31. However, a cooling treatment stream can also be used to supply some cooling to the high pressure refrigerant (stream 71a) as shown in FIGS. 4, 5, 10 and 11. In addition, any stream having a lower temperature than the stream to be cooled may be used. For example, a side draw of steam may be withdrawn from the separator / absorber tower 18 or fractionation tower 19 and used for cooling. The specific arrangement of heat exchangers for the use and distribution of tower liquids and / or vapors for heat exchange treatment, for cooling the inlet and feed gases, as well as the selection of the treatment stream for a particular heat exchange facility, must be evaluated for each particular application. The choice of cooling source depends on a number of factors including, but not limited to, feed gas composition and condition, plant size, heat exchanger size, and potential cooling source temperature. It will be appreciated by those skilled in the art that any combination of cooling sources or cooling methods may be employed in combination to obtain the desired feed stream temperature.

In addition, additional external cooling to the inlet gas stream and the feed stream to the LNG production unit may be achieved in a variety of different ways. 1, 3 to 21, since the single composition refrigerant is used for the preliminary cooling of the multiple composition refrigerant stream, the boiling of the single composition refrigerant is assumed to be a high level of external cooling, and the vapor of the multi composition refrigerant Vaporizing is assumed to be a low level of external cooling. Optionally, both high and low levels of cooling can be achieved using a single composition refrigerant (ie cascade cooling) by successive low boiling points, or by using a single composition refrigerant up to a continuous low evaporation pressure. have. As another alternative, both high level cooling and low level cooling can be achieved using a multi-component refrigerant stream with each composition adjusted to provide the required cooling temperature. The choice of method of providing external cooling is not limited to the following, including a number of factors including feed gas composition, plant size, drive unit size of the compression device, heat exchanger size, enclosed heat sink temperature, and the like. Follow. It will be appreciated by those skilled in the art that any combination of the methods for providing external cooling described above may be employed in combination to obtain the desired feed stream temperature.

Condensed liquid stream leaving heat exchanger 60 (stream 49 of FIGS. 1, 6, and 8, stream 49d of FIGS. 3, 4, 7, 9-16, and 5, FIG. Subcooling of the stream 49b of FIGS. 19 and 20, the stream 49e of FIG. 17, the stream 49c of FIG. 18, and the stream 49a of FIG. 21) causes the stream to expand to the working pressure of the LNG storage tank 62. Reduce or eliminate the amount of flash steam that may be generated. This generally eliminates the need for flash gas compression, thereby reducing the non-power consumption for producing LNG. However, in many situations it may be advantageous to reduce the capital cost of the installation by reducing the size of the heat exchanger 60 and placing any flash that may be generated using flash gas compression or other means.

While individual stream expansions are shown in special expansion devices, optional expansion means may be employed where appropriate. For example, a work expanding or intermediate pressure backflow stream (stream 39 of FIGS. 1, 6, and 8) of a generally condensed feed stream (stream 35a of FIGS. 1, 3, 6, and 7) may be used. Requirements can also be guaranteed. The supercooled liquid stream (streams 49 of FIGS. 1, 6 and 8, and streams of FIGS. 3, 4, 7, 9 and 16) in which isenthalpy flash expansion leaves heat exchanger 60. 49d), stream 49b in FIGS. 5, 19 and 20, stream 49e in FIG. 17, stream 49c in FIG. 18, stream 49a in FIG. In order to avoid flash steam formation of expansion, one would need to either subcool more in the heat exchanger 60 or add flash steam compression or add other means for the placement of the resulting flash steam. Similarly, isoenthalpy flash expansion may be used in place of work cooling for the supercooled high pressure refrigerant stream (stream 71c of FIGS. 1, 3-21) leaving heat exchanger 60, resulting in compression of the refrigerant. Power consumption increases.

While what has been described as what is considered to be a preferred embodiment of the present invention, there are other modifications and variations that apply various conditions, forms of supply, or other requirements of the invention, without departing, for example, from the spirit of the invention as defined by the following claims. Other variations are possible to be appreciated by the skilled artisan.

According to the present invention, the hydrocarbon removal step into the LNG liquefaction process can be incorporated to produce both LNG and separated heavy hydrocarbon liquid products using much less energy than prior art processes.

Claims (116)

(a) the natural gas stream is cooled under pressure to condense at least a portion thereof to form a condensed stream, and (b) the condensed stream is expanded to low pressure to form a liquefied natural gas stream. In the process of liquefying a natural gas stream containing, (1) said natural gas stream is treated in one or more cooling stages; (2) the cooled natural gas stream is expanded to intermediate pressure and then sent to a mid-column feed position on the distillation column, the stream being a more volatile vapor distillation stream and Is separated into a relatively less volatile fraction containing the heavy hydrocarbon component as a major portion; (3) the vapor distillation stream exits the region of the distillation column below the expanded cooled natural gas stream and is sufficiently cooled to condense at least a portion thereof, thereby forming a vapor stream and a liquid stream; (4) at least a portion of the expanded cooled natural gas stream is intimately contacted with at least a portion of the liquid stream in the distillation column; (5) said vapor stream is combined with said more volatile vapor distillation stream to form a volatile residue gas fraction containing said methane and light components as main components; (6) said volatile residue gas fraction is cooled under pressure to condense at least a portion thereof, thereby forming said condensed stream, Method of liquefying a natural gas stream. (a) the natural gas stream is cooled under pressure to condense at least a portion thereof to form a condensed stream, and (b) the condensed stream is expanded to low pressure to form a liquefied natural gas stream. In the process of liquefying a natural gas stream containing, (1) said natural gas stream is treated in one or more cooling stages to partially condense; (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 medium pressure; (4) said expanded first vapor stream and said expanded first liquid stream are sent to a column-intermediate feed point on a distillation column, said streams being composed mainly of the more volatile vapor distillation stream and the heavy hydrocarbon component To separate into relatively less volatile fractions containing; (5) a vapor distillation stream exits the region of the distillation column below the expanded first vapor stream and is sufficiently cooled to condense at least a portion thereof to form a second vapor stream and a second liquid stream; (6) at least a portion of the expanded first vapor stream is in intimate contact with at least a portion of the second liquid stream in the distillation column; (7) said second vapor stream is combined with said more volatile vapor distillation stream to form a volatile residue gas fraction containing said methane and light components as main components; (8) said volatile residue gas fraction is cooled under pressure to condense at least a portion thereof and thereby form said condensed stream, Method of liquefying a natural gas stream. (a) the natural gas stream is cooled under pressure to condense at least a portion thereof to form a condensed stream, and (b) the condensed stream is expanded to low pressure to form a liquefied natural gas stream. In the process of liquefying a natural gas stream containing, (1) said natural gas stream is treated in one or more cooling stages; (2) said cooled natural gas stream is expanded to medium pressure and then sent to a column-intermediate feed point on a distillation column, said stream containing a relatively more volatile vapor distillation stream and said heavy hydrocarbon component as principal components; To separate into less volatile fractions; (3) a vapor distillation stream exits the region of the distillation column below the expanded cooled natural gas stream, and is cooled sufficiently to condense at least a portion thereof to form a vapor stream and a liquid stream; (4) a portion of the liquid stream is fed to the distillation column as another feed at a feed location in a region substantially the same as where the vapor distillation stream exits; (5) at least a portion of the expanded cooled natural gas stream is in intimate contact with at least a portion of the remaining portion of the liquid stream in the distillation column; (6) said vapor stream combines with said more volatile vapor distillation stream to form a volatile residue gas fraction containing said methane and light components as main components; (7) said volatile residue gas fraction is cooled under pressure to condense at least a portion thereof to form said condensed stream, Method of liquefying a natural gas stream. (a) the natural gas stream is cooled under pressure to condense at least a portion thereof to form a condensed stream, and (b) the condensed stream is expanded to low pressure to form a liquefied natural gas stream. In the process of liquefying a natural gas stream containing, (1) said natural gas stream is treated in one or more cooling stages to partially condense; (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 medium pressure; (4) said expanded first vapor stream and said expanded first liquid stream are sent to a column-intermediate feed point on a distillation column, said streams being composed mainly of the more volatile vapor distillation stream and the heavy hydrocarbon component To separate into relatively less volatile fractions containing; (5) a vapor distillation stream exits the region of the distillation column below the expanded first vapor stream and is sufficiently cooled to condense at least a portion thereof to form a second vapor stream and a second liquid stream; (6) a portion of the second liquid stream is supplied to the distillation column as another feed at a feed location in a region substantially the same as where the vapor distillation stream exits; (7) at least a portion of the expanded first vapor stream is in intimate contact with at least a portion of the remaining portion of the second liquid stream in the distillation column; (8) said second vapor stream is combined with said more volatile vapor distillation stream to form a volatile residue gas fraction containing said methane and light components as main components; (9) said volatile residue gas fraction is cooled under pressure to condense at least a portion thereof and thereby form said condensed stream, Method of liquefying a natural gas stream. (a) the natural gas stream is cooled under pressure to condense at least a portion thereof to form a condensed stream, and (b) the condensed stream is expanded to low pressure to form a liquefied natural gas stream. In the process of liquefying a natural gas stream containing, (1) said natural gas stream is treated in one or more cooling stages; (2) said cooled natural gas stream is expanded to medium pressure and then sent to a column-intermediate feed point on a distillation column, said stream containing the more volatile vapor distillation stream and said heavy hydrocarbon component as main components; To separate into relatively less volatile fractions; (3) a vapor distillation stream exits the region of the distillation column below the expanded cooled natural gas stream and is sufficiently cooled to condense at least a portion to form a vapor stream and a liquid stream; (4) at least a portion of the expanded cooled natural gas stream is in intimate contact with at least a portion of the liquid stream in the distillation column; (5) A liquid distillation stream exits the distillation column at a place above the region where the vapor distillation stream exits and the liquid distillation stream is heated and thereafter another feed at a place below the region where the vapor distillation stream exits. Sent back to the distillation column as water; (6) said vapor stream is combined with said more volatile vapor distillation stream to form a volatile residue gas fraction containing said methane and light components as main components; (7) said volatile residue gas fraction is cooled under pressure to condense at least a portion thereof and thereby form said condensed stream, Method of liquefying a natural gas stream. (a) the natural gas stream is cooled under pressure to condense at least a portion thereof to form a condensed stream, and (b) the condensed stream is expanded to low pressure to form a liquefied natural gas stream. In the process of liquefying a natural gas stream containing, (1) said natural gas stream is treated in one or more cooling stages to partially condense; (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 medium pressure; (4) said expanded first vapor stream and said expanded first liquid stream are sent to a column-intermediate feed point on a distillation column, said streams being composed mainly of the more volatile vapor distillation stream and the heavy hydrocarbon component To separate into relatively less volatile fractions containing; (5) a vapor distillation stream exits the region of the distillation column below the expanded first vapor stream and is sufficiently cooled to condense at least a portion to form a second vapor stream and a second liquid stream; (6) at least a portion of the expanded first vapor stream is in intimate contact with at least a portion of the second liquid stream of the distillation column; (7) A liquid distillation stream exits the distillation column at a location above the region where the vapor distillation stream exits, the liquid distillation stream is heated, and then another feed at a location below the region where the vapor distillation stream exits. Sent back to the distillation column as water; (8) said second vapor stream is combined with said more volatile vapor distillation stream to form a volatile residue gas fraction containing said methane and light components as main components; (9) said volatile residue gas fraction is cooled under pressure to condense at least a portion thereof and thereby form said condensed stream, Method of liquefying a natural gas stream. (a) the natural gas stream is cooled under pressure to condense at least a portion thereof to form a condensed stream, and (b) the condensed stream is expanded to low pressure to form a liquefied natural gas stream. In the process of liquefying a natural gas stream containing, (1) said natural gas stream is treated in one or more cooling stages; (2) said cooled natural gas stream is expanded to medium pressure and then sent to a column-intermediate feed point on a distillation column, said stream containing the more volatile vapor distillation stream and said heavy hydrocarbon component as main components; To separate into relatively less volatile fractions; (3) a vapor distillation stream exits the region of the distillation column below the expanded cooled natural gas stream, and is cooled sufficiently to condense at least a portion thereof to form a vapor stream and a liquid stream; (4) a portion of the liquid stream is fed to the distillation column as another feed at a feed location in a region substantially the same as where the vapor distillation stream exits; (5) at least a portion of the expanded cooled natural gas stream is in intimate contact with at least a portion of the remaining portion of the liquid stream in the distillation column; (6) a liquid distillation stream exits the distillation column at a place above the region where the vapor distillation stream exits, the liquid distillation stream is heated, and then another feed at a place below the region where the vapor distillation stream exits Sent back to the distillation column as water; (7) said vapor stream is combined with said more volatile vapor distillation stream to form a volatile residue gas fraction containing said methane and light components as main components; (8) said volatile residue gas fraction is cooled under pressure to condense at least a portion thereof and thereby form said condensed stream, Method of liquefying a natural gas stream. (a) the natural gas stream is cooled under pressure to condense at least a portion thereof to form a condensed stream, and (b) the condensed stream is expanded to low pressure to form a liquefied natural gas stream. In the process of liquefying a natural gas stream containing, (1) said natural gas stream is treated in one or more cooling stages to partially condense; (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 medium pressure; (4) said expanded first vapor stream and said expanded first liquid stream are sent to a column-intermediate feed point on a distillation column, said streams being composed mainly of the more volatile vapor distillation stream and the heavy hydrocarbon component To separate into relatively less volatile fractions containing; (5) a vapor distillation stream exits the region of the distillation column below the expanded first vapor stream and is sufficiently cooled to condense at least a portion to form a second vapor stream and a second liquid stream; (6) a portion of the second liquid stream is supplied to the distillation column as another feed at a feed location in a region substantially the same as where the vapor distillation stream exits; (7) at least a portion of the expanded first vapor stream is in intimate contact with at least a portion of the remaining portion of the second liquid stream in the distillation column; (8) A liquid distillation stream exits the distillation column at a place above the zone where the vapor distillation stream exits, the liquid distillation stream is heated, and then another feed at a site below the zone where the vapor distillation stream exits. Sent back to the distillation column as water; (9) said second vapor stream is combined with said more volatile vapor distillation stream to form a volatile residue gas fraction containing said methane and light components as main components; (10) said volatile residue gas fraction is cooled under pressure to condense at least a portion thereof and thereby form said condensed stream, Method of liquefying a natural gas stream. (a) the natural gas stream is cooled under pressure to condense at least a portion thereof to form a condensed stream, and (b) the condensed stream is expanded to low pressure to form a liquefied natural gas stream. In the process of liquefying a natural gas stream containing, (1) said natural gas stream is treated in one or more cooling stages; (2) the cooled natural gas stream is expanded to medium pressure; (3) the expanded cooled natural gas stream is sent to a distillation column, the stream having a relatively less volatile residue gas fraction containing the methane and light components as the main component and the heavy hydrocarbon component as the main component. To separate into volatile fractions; (4) the volatile residue gas fraction is cooled under pressure to condense at least a portion thereof and thereby form the condensed stream A process of liquefaction of a natural gas stream, characterized in that consisting essentially of processing steps. (a) the natural gas stream is cooled under pressure to condense at least a portion thereof to form a condensed stream, and (b) the condensed stream is expanded to low pressure to form a liquefied natural gas stream. In the process of liquefying a natural gas stream containing, (1) said natural gas stream is treated in one or more cooling stages to partially condense; (2) said partially condensed natural gas stream is separated to provide at least a vapor stream and a liquid stream; (3) said vapor stream is expanded to medium pressure; (4) said liquid stream is expanded to said intermediate pressure; (5) at least the expanded vapor stream and the expanded liquid stream are sent to a distillation column, the stream comprising the volatile residue gas fraction and the heavy hydrocarbon component as the main components containing the methane and light components as main components. To separate into relatively less volatile fractions containing; (6) said volatile residue gas fraction is cooled under pressure to condense at least a portion thereof and thereby form said condensed stream; A process of liquefaction of a natural gas stream, characterized in that consisting essentially of processing steps. The method according to any one of claims 1 to 10, Wherein the volatile residue gas fraction is compressed and then cooled under pressure to condense at least a portion thereof to thereby form the condensed stream. The method according to any one of claims 1 to 10, Wherein the volatile residue gas fraction is heated, compressed and then cooled under pressure to condense at least a portion thereof to thereby form the condensed stream. The method according to any one of claims 1 to 10, The volatile residue gas fraction contains the methane, light and C ₂ components as main components. The method according to any one of claims 1 to 10, The volatile residue gas fraction contains the methane, light component, C ₂ component and C ₃ component as main components. The method according to claim 11, The volatile residue gas fraction contains the methane, light and C ₂ components as main components. The method according to claim 12, The volatile residue gas fraction contains the methane, light and C ₂ components as main components. The method according to claim 11, The volatile residue gas fraction contains the methane, light component, C ₂ component and C ₃ component as main components. The method according to claim 12, The volatile residue gas fraction contains the methane, light component, C ₂ component and C ₃ component as main components. delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete delete

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