HK1106283B - Natural gas liquefaction - Google Patents

Description

Natural gas liquefaction process

Background

The present invention relates to a process for treating natural gas or other methane-rich gas streams to produce a Liquefied Natural Gas (LNG) stream having high methane purity and a liquid stream containing primarily hydrocarbons heavier than methane.

Natural gas is typically recovered from wells drilled into subterranean reservoirs, the major portion of which is usually methane, i.e., methane comprising at least 50 mole% of the gas. Natural gas also contains smaller amounts of heavier hydrocarbons such as ethane, propane, butanes, pentanes, etc. as well as water, hydrogen, nitrogen, carbon dioxide and other gases, depending on the specifics of the subterranean reservoir.

Most natural gas is disposed of in gaseous form. The most common method of transporting natural gas from the wellhead to a gas processing facility and from there to natural gas customers is by high pressure gas transmission lines. In several instances, however, it has been found necessary and/or desirable to liquefy natural gas for transport and use. For example, in remote locations, there is often no pipeline facility that can conveniently deliver natural gas to the market. In this case, LNG having a greatly reduced specific volume relative to gaseous natural gas can greatly reduce the transportation cost of LNG by allowing the use of cargo ships and transportation trucks.

Other situations where natural gas is more suitable for liquefaction are when used as automotive fuel. In metropolitan areas, there are fleets of buses, taxis, and trucks powered by LNG if there is an economically available source of LNG. Due to the clean burning nature of natural gas, these LNG-fueled vehicles have greatly reduced air pollution compared to similar vehicles powered by gasoline or diesel engines that burn higher molecular weight hydrocarbons. Furthermore, if the purity of LNG is high (i.e., methane purity of 95 mole% or higher), the carbon content of methane: the hydrogen ratio is lower than all other hydrocarbon fuels and the amount of carbon dioxide (greenhouse gas) produced is greatly reduced.

The present invention relates broadly to a process for liquefying natural gas while co-producing as co-products a liquid stream consisting primarily of hydrocarbons heavier than methane, such as Natural Gas Liquids (NGLs) consisting of ethane, propane, butane and heavier hydrocarbon components, Liquefied Petroleum Gases (LPG) consisting of propane, butane and heavier hydrocarbon components, or condensates consisting of butane and heavier hydrocarbon components. The production of a co-product liquid stream has two important benefits: the LNG produced has high methane purity and the co-product liquid is a valuable product that can be used for many other purposes. Typical analysis results for natural gas streams to be treated according to the invention are (approximate mole percentages) 84.2% methane, 7.9% ethane and other C' s₂Component, 4.9% propane and other C₃Component, 1.0% isobutane, 1.1% n-butane, 0.8% pentane +, the balance nitrogen and carbon dioxide. Sometimes sulfur-containing gases are also present

Several methods are known for liquefying natural gas. See, for example, "LNG technology for offshore and medium-type installations" published in 3, 13-15, 2000 by Finn, Adrian j., grantl., Johnson and Terry r.tomlinson at 79, annual meeting group 429-450 of the gas processor association held in atlanta, georgia, and "optimization of minimum load LNG installation power systems" published in 2001 at 12-14, 3, 12-14, by Kikkawa, Yoshitsugi, Masaaki Ohishi and Noriyoshi Nozawa for reviewing several such methods. U.S. patent nos. 4445917, 4525185, 4545795, 4755200, 5291736, 5363655, 5365740, 5600969, 5615561, 5651269, 5755114, 5893274, 6014869, 6053007, 6062041, 6119479, 6125653, 6250105B1, 6269655B1, 6272882B1, 6308531B1, 6324867B1, 6347532B1, PCT patent application No. WO01/88447 and 10/161780 filed 6/4 of our co-pending patent application No. 2002 and 10/278610 filed 10/23 of our year 2002 also describe related methods. These processes typically include steps of purifying the natural gas (by means of removing water and problematic compounds such as carbon dioxide and sulfur compounds), cooling, condensing and expanding. The step of cooling and condensing the natural gas may be carried out in many different ways. The "cascade refrigeration" method employs heat exchange of natural gas with several refrigerants having successively lower boiling points, such as propane, ethane, and methane. Alternatively, this heat exchange operation may be carried out with a single refrigerant by evaporating the refrigerant at several different pressure levels. "multi-component refrigeration" employs heat exchange between natural gas and one or more refrigerant fluids composed of several refrigerant components instead of a plurality of single-component refrigerants. The natural gas expansion operation can be carried out both isenthalpically (e.g., by Joule-Thomson expansion) and isentropically (e.g., by work expansion turbine).

In the case of processes used to liquefy natural gas streams, it is often necessary to remove a large portion of the hydrocarbons heavier than methane before the methane-rich stream is liquefied. This de-hydrocarbon step is desirable for several reasons, including the need to control the heating value of the LNG stream and the value of the heavier hydrocarbon components themselves as products. Unfortunately, little attention has been paid to the benefits of the de-hydrocarbon step so far.

In accordance with the present invention, it has been discovered that careful integration of the de-hydrocarbon step into the LNG liquefaction process allows for the co-production of LNG and another heavier hydrocarbon liquid product with less energy usage than prior art processes. Although the present invention can be applied at low pressures, it is particularly advantageous to treat the feed gas at pressures in the range of 400-1500psia [2758-10342kPa (a) ], or higher.

For a better understanding of the present invention, reference is made to the following examples and accompanying drawings. Reference is made to the accompanying drawings

FIG. 1 is a flow diagram of a natural gas liquefaction plant suitable for the co-production of NGL in accordance with the present invention;

FIG. 2 is a pressure-enthalpy phase diagram of methane illustrating the advantage of the present invention over prior art processes; and

figures 3, 4,5, 6, 7 and 8 are flow diagrams of alternative natural gas liquefaction plants suitable for the co-production of liquid streams in accordance with the present invention.

In the following explanation of the above figures, a summary of the flow rates calculated for typical process conditions is provided. For convenience, the flow values (moles/hour) in these tables appear herein to have been rounded to the nearest whole number. The total stream flow shown in the table includes all non-hydrocarbon components and thus will generally be greater than the sum of the stream flows of hydrocarbon components. The temperatures indicated are approximate values which have been rounded off. It should also be noted that the process design for comparing the methods shown in the figures is calculated on the assumption that there is no heat loss from environment to process (or from process to environment). The quality of commercially available insulating materials makes this a reasonable assumption and is a commonly used assumption by those skilled in the art.

For convenience, process parameters are reported in the traditional English and International Systems (SI) units. The molar flow rates given in the tables can be interpreted as either pound moles per hour or kilogram moles per hour. The energy consumption reported in Horsepower (HP) and/or thousand british thermal units per hour (MBTU/Hr) corresponds to the stated molar flow rate in pound moles per hour. The energy consumption reported in kilowatts corresponds to the stated molar flow in kilogram moles per hour. The production rate reported in pounds per hour (Lb/Hr) corresponds to the stated pound flow in pound moles per hour. The production rate reported in kilograms per hour (Kg/Hr) corresponds to the stated molar flow in kilograms moles per hour.

Description of the invention

Referring now to fig. 1, we first describe a process in accordance with the present invention that is expected to produce NGL co-products comprising half the ethane and most of the propane and heavier components in a natural gas feedstream. In this simulation of the invention, inlet gas enters the apparatus as stream 31 at 90F [32℃ ] and 1285psia [8860kPa (a) ]. If the feed gas contains a concentration of carbon dioxide and/or sulfur compounds that makes the product stream out of specification, the feed gas is suitably pretreated to remove these species (not shown). In addition, the feed stream is typically treated to remove water to avoid hydrate (ice) formation under cryogenic conditions. Solid drying agents are typically used to remove water.

Inlet stream 31 is cooled in heat exchanger 10 by heat exchange with a refrigerant stream and flashed-44F [ -42 ℃ ] separator liquid (stream 39 a). Note that the heat exchanger 10 is in each case representative of a plurality of individual heat exchangers or a multi-pass heat exchanger or any combination thereof. (the decision whether to use more than one heat exchanger for the illustrated cooling apparatus depends on several factors including, but not limited to, inlet air flow, heat exchanger size, stream temperature, etc.). Cooled stream 31a enters separator 11 at 0F [ -18 ℃ C ] and 1278psia [8812kPa (a) ], separating the vapor (stream 32) from the condensed liquid (stream 33).

The vapor (stream 32) is separated from separator 11 into two streams 34 and 36, stream 34 comprising about 15% of the total vapor amount. In some cases, it may be more likely that stream 34 will be combined with a portion of the condensed liquid (stream 38) to form a combined stream 35, while stream 38 does not flow during this simulation. Stream 35 is passed through heat exchanger 13 in heat exchange relationship with refrigerant stream 71e and liquid distillation stream 40 to provide cooled and substantially condensed stream 35 a. The substantially condensed stream 35a is flash expanded at-109F [ -78 c ] through a suitable expansion device, such as expansion valve 14, to the operating pressure of fractionation column 19 (about 465psia [3206kpa (a)) ]. During expansion, a portion of the stream is vaporized, resulting in cooling of the entire stream. In the process illustrated in FIG. 1, expanded stream 35b exiting expansion valve 14 reaches a temperature of-125F < -87 > C and is then supplied to fractionation column 19 at an upper mid-point feed position in absorption section 19 a.

The remaining 85% of the vapor from separator 11 (stream 36) enters work expander 15 where mechanical energy is extracted from the portion of the high pressure feed. The expander 15 expands the vapor substantially isentropically to the operating pressure of the column. The work expansion process cools expanded stream 36a to a temperature of about-76 ° F [ -60 ℃ ]. Typical commercial expanders can recover about 80-85% of the theoretical work available in an ideal isentropic expansion process. The recovered work is typically used to drive a centrifugal compressor (e.g., device 16), which may be used, for example, to recompress the overhead gas (stream 49). Thereafter, the expanded and partially condensed stream 36a is supplied as a feed at the lower half feed point of the absorption section 19a of the distillation column 19. The remaining portion of stream 39 of the separator liquid (stream 33) is flash expanded via expansion valve 12 to slightly above the operating pressure of demethanizer 19 and stream 39 is cooled to-44F-42 c (stream 39a) and then used to cool the incoming inlet gas as previously described. Stream 39b, now 85 ° F [29 ℃ ], is then fed to the stripping section 19b of the demethanizer 19 at the second feed point of the lower half-column.

The demethanizer in fractionation column 19 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The fractionation column may be constructed of two sections, as is common in natural gas processing plants. The upper absorption (rectification) section 19a contains trays and/or packing to provide the necessary contact between the vapor portion of the ascending expanded stream 36a and the descending cold liquid to condense and absorb ethane, propane and heavier components; the lower stripping section 19b contains trays and/or packing to provide the necessary contact between the descending liquid and the ascending vapor. The stripping section also includes one or more reboilers (e.g., reboiler 20) that heat and vaporize a portion of the liquid flowing down the column to provide stripping vapor flowing up the column to strip the liquid product stream 41 of methane and lighter components. A liquid product stream 41 exits the bottom of the demethanizer 19 at 150 ℃ F. [66 ℃ C. ] according to the general specification requirement having a methane to ethane ratio of 0.020: 1 on a molar basis in the bottoms product. The overhead distillation gas stream 37, which contains primarily methane and lighter components, exits the demethanizer 19 at-108F [ -78 c ].

A portion of the distillation gas (stream 42) is withdrawn from the upper region of stripping section 19 b. This stream is cooled in heat exchanger 13 from-58F-50 c to-109F-78 c and partially condensed (stream 42a) by heat exchange with refrigerant stream 71e and liquid distillation stream 40. The operating pressure of reflux separator 22 (461psia [3182kPa (a)) ] is maintained slightly below the operating pressure of demethanizer 19. This provides the driving force required to pass the distillation gas stream 42 through heat exchanger 13 and thence to reflux separator 22 where the condensed liquid (stream 44) is separated from any uncondensed vapor (stream 43) in reflux separator 22. Stream 43 is combined with a distillation gas stream (stream 37) exiting from an upper region of the absorption section 19a of the demethanizer 19 to form a cold residue gas stream 47 at-108F < - > to-78 ℃.

The condensate, stream 44, is pumped to higher pressure by pump 23, whereupon the-109F-78℃ stream 44a is split into two portions. A portion is stream 45 which is sent to the upper region of absorption section 19a of demethanizer 19 and serves as a cooling liquid to contact the vapor flowing upwardly along the absorption section. Another portion is supplied as reflux stream 46 to an upper region of stripping section 19b of demethanizer 19.

A liquid distillation stream 40 is withdrawn from a lower region of absorption section 19a of demethanizer 19 and passed to heat exchanger 13 where it is heated while providing cooling for distillation gas stream 42, combined stream 35, and refrigerant (stream 71 a). The liquid distillation stream is heated from-79F [ -62 ℃ ] to-20F [ -29 ℃ ] as a partially vaporized stream 40a, before being fed as a mid-column feed to the stripping section 19b of the demethanizer 19.

The cold residue gas (stream 47) is heated to 94 ° F [34 ℃) in heat exchanger 24, and a portion (stream 48) is withdrawn as plant fuel gas. (the amount of fuel gas that must be removed is largely determined by the amount of fuel required to drive the engines and/or turbines in the plant, such as refrigerant compressors 64, 66 and 68 in this example). The remaining portion of the hot residue gas (stream 49) is compressed by compressor 16 driven by expanders 15, 61 and 63. After cooling to 100 ° F [38 ℃) in the exhaust gas cooler 25, stream 49b is further cooled to-93 ° F [ -69 ℃) in heat exchanger 24 by reverse exchange with the cold residue gas stream 47 (stream 49 c).

Stream 49c then enters heat exchanger 60 and is further cooled by expanded refrigerant stream 71d to-256F-160 c to condense and subcool it, thereby entering work expander 61 from which mechanical energy is extracted. Work expansion machine 61 expands liquid stream 49d substantially isentropically from about 638psia [4399kpa (a) ], to a LNG storage pressure slightly above atmospheric pressure (15.5psia [107kpa (a) ].

All cooling of stream 49c and a portion of the cooling of streams 35 and 42 is provided by a closed refrigeration loop. The working fluid for this refrigeration cycle is a mixture of hydrocarbons and nitrogen, the composition of which is adjusted as required to provide the refrigerant temperature while condensing at a reasonable pressure with the available cooling medium. In this case, it is assumed that the condensation is carried out with cooling water, so that a refrigerant mixture consisting of nitrogen, methane, ethane, propane and heavier hydrocarbons will be used in the simulation of the process of FIG. 1. The stream composition, in approximate mole percent, was 6.9% nitrogen, 40.8% methane, 37.8% ethane, and 8.2% propane, with the balance being heavier hydrocarbons.

Refrigerant stream 71 exits vent cooler 69 at 100 ° F [38 ℃ c ] and 607psia [4,185kpa (a) ]. Enters the heat exchanger 10 and is cooled to-15F-26 c and partially condensed with the partially warmed expanded refrigerant stream 71F and other refrigerant streams. For the simulation of fig. 1, it has been assumed that the other refrigerant streams are technical grade propane refrigerant at three different temperature and pressure levels. The partially condensed refrigerant stream 71a is then passed to heat exchanger 13 for further cooling to-109F-78 c with the partially warmed expanded refrigerant stream 71e, thereby condensing the refrigerant (stream 71 b). The refrigerant is subcooled in heat exchanger 60 to-256F-160 c with expanded refrigerant stream 71 d. Subcooled liquid stream 71c enters work expander 63 and mechanical work is extracted from this stream as the stream is expanded substantially isentropically from a pressure of about 586psia [4040kpa (a)) ] to a pressure of about 34psia [234kpa (a)) ]. During expansion, a portion of the stream is vaporized, cooling the entire stream to-262 ℃ F. [ -163 ℃ C. ] (stream 71 d). The expanded stream 71d then re-enters heat exchangers 60, 13 and 10 where it is vaporized and superheated while providing cooling to stream 49c, stream 35, stream 42 and the refrigerant (streams 71, 71a and 71 b).

Superheated refrigerant vapor (stream 71g) exits heat exchanger 10 at 93 ° F [34 ℃ c ] and is compressed to 617psia [4254kpa (a) ] in three stages. Each of these three compression stages (refrigerant compressors 64, 66 and 68) is driven by an auxiliary power supply and is followed by a chiller (exhaust chillers 65, 67 and 69) to remove the heat of compression. Compressed stream 71 is returned from exhaust cooler 69 to heat exchanger 10 to complete the cycle.

The flow rates and energy consumption for the process shown in FIG. 1 are summarized in the following table:

TABLE I

(FIG. 1)

Summary of flow Table pound moles/hour [ kilogram moles/hour ]

Logistics Methane Ethane (III) Propane Butane treasure Total of

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

32 38,538 3,336 1,847 830 44,556

33 2,439 525 561 574 4,100

34 5,781 501 277 125 6,683

36 32,757 2,835 1,570 705 37,873

40 3,896 2,170 1,847 829 8,742

42 8,045 1,850 26 0 9,922

43 4,551 240 1 0 4,792

44 3,494 1,610 25 0 5,130

45 1,747 805 12 0 2,565

46 1,747 805 13 0 2,565

37 36,393 1,970 11 0 38,380

41 33 1,651 2,396 1,404 5,484

47 40,944 2,210 12 0 43,172

48 2,537 137 1 0 2,676

50 38,407 2,073 11 0 40,496

Recoveries in NGL

Ethane 42.75%

99.53 percent of propane

Butane + 100.00%

Yield 246,263Lb/Hr [246,263kg/Hr ]

LNG product

Yield 679,113Lb/Hr [679,113kg/Hr ]

Purity 94.84%

Low heat value 946.0BTU/SCF [35.25MJ/m ]³]

Electric power

Refrigeration compressor 94,868HP [155,962kW ]

Propane compressor 25,201HP [41,430kW ]

Total 120,069HP [197,392kW ]

Heat of public works

Demethanizer reboiler 24,597MBTU/Hr [15,888kW ]

By (based on unscrupulous flow)

The efficiency of an LNG production process is typically compared with the required "specific energy consumption", i.e. the ratio between the total refrigeration compression power and the total liquid production rate. Published information on specific energy consumption for prior art processes for producing LNG is shown to range from 0.168HP-Hr/Lb [0.276kW-Hr/kg ] to 0.182HP-Hr/Lb [0.300kW-Hr/kg ], which is believed to be based on indications of 340 days of on-line production per year for LNG production plants. On the same basis, the specific energy consumption of the embodiment of FIG. 1 of the present invention is 0.139HP-Hr/Lb [0.229kW-Hr/kg ], which is 21-31% more efficient than the prior art process.

There are two main indicators that can illustrate the improvement in efficiency of the present invention. The first indicator can be understood by examining the thermodynamics of the liquefaction process when applying a high pressure gas stream as considered in this example. Because the major component of this stream is methane, the thermodynamic properties of methane can be used to compare the liquefaction cycle employed in prior art processes with the cycle employed in the present invention. Figure 2 contains a pressure-enthalpy phase diagram for methane. In most prior art liquefaction cycles, all cooling of the gas stream is performed while the gas stream is at an elevated pressure (paths a-B), after which the stream is expanded (paths B-C) to the pressure of the LNG storage tank (slightly above atmospheric pressure). The expansion step can utilize a work expander, and generally can recover about 75-80% of theoretically available work in an ideal isentropic expansion process. For simplicity, shown as a full isentropic expansion in path B-C of FIG. 2. Even so, the enthalpy drop provided by this process of work expansion is still quite small because the isentropic lines are near vertical in the liquid region of the phase diagram.

Now in comparison with the liquefaction cycle process of the present invention. After partial cooling at high pressure (path A-A '), the gas stream is work expanded (path A ' -A ') to an intermediate pressure. (also shown as full isentropic expansion for simplicity). The remaining cooling process is carried out at moderate pressure (path a "-B ') and then the stream is expanded (path B' -C) to LNG tank pressure. Because the isentropic lines are not steeply sloped in the vapor region of the phase diagram, the first work expansion step (path a' -a ") of the present invention provides a large enthalpy drop. Thus, the total amount of cooling required by the present invention (the sum of paths A-A 'and paths A "-B') is less than the cooling required by prior art methods (paths A-B), reducing the amount of refrigeration required to liquefy the gas stream (and the refrigeration compression process).

A second indicator of the improved efficiency of the present invention is the superior operating performance of the hydrocarbon distillation system at lower operating pressures. The hydrocarbon removal step in most prior art processes is operated at elevated pressure, typically using a wash column, using cold hydrocarbon liquid as the absorption stream to remove heavier hydrocarbons from the incoming gas stream. It is not very efficient to operate the scrub column at high pressure as this would result in the co-absorption of most of the methane from the gas stream, which must be stripped from the absorbent liquid in a subsequent step and cooled to become part of the LNG product. In the present invention, the de-hydrocarbon step is carried out at moderate pressures, which are highly favorable to the gas-liquid equilibrium, allowing for very efficient recovery of the desired heavier hydrocarbons in the co-product liquid stream.

Other embodiments

Those skilled in the art will recognize that the present invention is applicable to all types of LNG liquefaction plants, co-producing NGL streams, LPG streams, or condensed liquid streams at a given plant in a manner that best meets the needs. Also, it will be appreciated that different process configurations may be employed to recover the liquid co-product stream. The invention is applicable to the recovery of C present in a feed gas containing a major proportion of C₂NGL stream of components, recovering only C present in the feed gas₃And heavier components of LPG or recovery of C only present in the feed gas₄And heavier components, rather than producing only a moderate proportion of the C present in the feed gas as described previously₂NGL co-products of components. When only partial recovery of C in the feed gas is desired₂Collecting substantially all C simultaneously₃And heavier components, the present invention is particularly advantageous over prior art processes, such as the embodiment of FIG. 1, whether C₂The reflux stream 45 maintains a high C level regardless of the degree of component recovery₃And (4) recovering the components.

According to the present invention, it is generally preferred that the absorption (rectification) section of the demethanizer is designed to contain a plurality of theoretical separation trays. However, as few as one theoretical plate can achieve the benefits of the present invention, and it is believed that even equivalent to one fractionation theoretical plate can achieve these benefits. For example, all or a portion of the pump pressure condensate (stream 44a) exiting reflux separator 22 and all or a portion of the substantially expanded condensate stream 35b from expansion valve 14 can be combined (e.g., in the expansion valve to demethanizer connection line), and if fully mixed, the vapor and liquid will mix together and separate according to the relative volatilities of the components of the total combined stream. For the purposes of the present invention, such an operation of mixing two streams will be considered to constitute an absorption section.

Figure 1 represents a preferred embodiment of the invention for the process conditions indicated. Fig. 3-8 illustrate alternative aspects of the present invention that are contemplated for specific applications. Depending on the amount of heavier hydrocarbons in the feed gas and the feed gas pressure, the cold feed stream 31a leaving the heat exchanger 10 may be free of any liquid (either because it is above its dew point, or because it is above its critical condensation pressure). In this case, the separator 11 shown in fig. 1 and 3-8 is not required, and the cold feed stream can be split into streams 34 and 36 and then directed to a heat exchanger (stream 34) and to a suitable expansion device (stream 36) such as a work expander 15.

As previously described, the distillation vapor stream 42 is partially condensed and the resulting condensate is used to absorb valuable C from the vapors rising along the absorber section 19a (FIGS. 1 and 4-8) or absorber 18 (FIG. 3) of demethanizer 19₃Components and heavier components. However, the present invention is not limited to this mode. It may also be beneficial to treat only a portion of the vapor or only a portion of the condensate as absorbent in this manner, for example, in cases where other designs dictate that a portion of the vapor or condensate bypass the absorption section 19a of the demethanizer 19. In some cases it is preferred that the distillation stream 42 be totally condensed rather than partially condensed in heat exchanger 13. Other cases may favor the distillation stream 42 being taken side-draw of all vapor from the fractionation column 19 rather than a partial vapor side-draw.

In the practice of the present invention, a slight pressure differential must be observed between the demethanizer 19 and reflux separator 22. If the distillation gas stream 42 passes through heat exchanger 13 and enters reflux separator 22 without any pressure push, reflux separator 22 must assume a situation where the operating pressure is slightly lower than the operating pressure of demethanizer 19. In this case, the liquid stream from the reflux separator can be pumped to the feed location of the demethanizer. An alternative is to provide a booster blower for the distillation gas stream 42 to raise the operating pressure in heat exchanger 13 and separator 22 sufficiently to allow the liquid stream 44 to be fed to demethanizer 19 without pumping.

It is not necessary that the high pressure liquid (stream 33 in fig. 1 and 3-8) be expanded and sent to the mid-column feed point of the distillation column. Alternatively, all or a portion of it can be combined with a portion of the separator vapor (stream 34) and passed to heat exchanger 13. (this scheme is illustrated by dashed stream 38 in FIGS. 1 and 3-8). Any remaining portion of the liquid may be expanded via a suitable expansion device, such as an expansion valve or expander, and fed to the mid-column feed point of the distillation column (stream 39b in fig. 1 and 3-8). Stream 39 in fig. 1 and 3-8 can also be used for cooling of the feed gas or other heat exchange means, either before or after expansion, and then sent to the demethanizer, similar to that shown by dashed stream 39a in fig. 1 and 3-8.

According to the invention, the vapor feed can be split in several ways. In the processes of fig. 1 and 3-8, the vapor diversion is performed after cooling and separation of any formed liquid. However, the splitting of the high pressure gas may be performed before any feed gas is cooled or after the gas is cooled and before any separation stages. In some cases, the gas split operation may be performed within a separator.

Fig. 3 shows a case where the fractionation column is constructed in two vessels, i.e., an absorption column 18 and a stripping column 19. In this case, the overhead vapor (stream 53) from stripper 19 can be split into two portions. A portion (stream 42) is sent to heat exchanger 13 to generate reflux for absorber column 18 as previously described. Any remaining portion (stream 54) flows to the lower section of absorber column 18 to contact the substantially expanded condensed stream 35b and the reflux (stream 45). Pump 26 is used to feed the liquid from the bottom of absorption column 18 (stream 51) to the top of stripping column 19 so that the two columns effectively function as a distillation system. Determining whether the fractionation column is constructed as a single vessel (e.g., demethanizer 19 in fig. 1 and 4-8) or as multiple vessels depends on several factors such as the size of the plant, the distance of the production facilities, etc.

In some cases it may be preferred to withdraw the cold liquid distillation stream 40 leaving the absorption section 19a in fig. 1 and 4-8 or the absorber column 18 in fig. 3 for heat exchange, while in other cases it may not be supported at all for withdrawing stream 40 for heat exchange, and thus stream 40 in fig. 1 and 3-8 is shown in dashed lines. When the present invention is operated to recover most of the ethane in the feed gas without reducing the ethane recovery in the demethanizer 19, although only a portion of the liquid from the absorption section 19a is used for heat exchange, more refrigeration is sometimes obtained than in the case of a conventional side reboiler with the liquid from the stripping section 19 b. This is because the liquids in the absorption section 19a of the demethanizer 19 can be utilized at cooler temperatures than the stripping section 19 b. This feature is also achieved when the fractionation column 19 is constructed in two vessels, as shown by the dashed stream 40 in fig. 3. When the liquid from absorber column 18 is pumped as shown in fig. 3, the liquid leaving pump 26 (stream 51a) can be split into two portions, one portion (stream 40) for heat exchange and then sent to the mid-column feed position (stream 40a) above stripper column 19. Any remaining portion (stream 52) becomes the top feed to stripper 19. As shown by the dashed stream 46 in fig. 1 and 3-8, it is preferred in this case to divide the liquid stream from reflux pump 23 (stream 44a) into at least two streams so that a portion (stream 46) can be fed to the stripping section of fractionation column 19 (fig. 1 and 4-8) or to stripping column 19 (fig. 3) to increase the liquid flow in this portion of the distillation system and improve the rectification of stream 42, while the remaining portion (stream 45) is fed to the top of absorption section 19a (fig. 1 and 4-8) or to absorption column 18 (fig. 3).

After recovery of the liquid co-product stream (stream 47 in fig. 1 and 3-8), the residual gas stream can be disposed of in a number of ways before it is fed to heat exchanger 60 for condensing and subcooling operations. In the process of fig. 1, the stream is heated, compressed to higher pressure using energy from one or more work expanders, partially cooled in a vent cooler, and then further cooled by reverse heat exchange with the feed stream. As shown in fig. 4, some applications prefer to compress this stream to a higher pressure using, for example, an auxiliary compressor 59 driven by an external power source. As shown by the dashed line equipment (heat exchanger 24 and vent cooler 25) in fig. 1, some circumstances favor reducing the capital cost of the equipment by reducing or eliminating pre-cooling of the compressed stream prior to entering heat exchanger 60 (at the expense of increasing the cooling load on heat exchanger 60 and increasing the energy consumption of refrigerant compressors 64, 66 and 68). In this case, stream 49a leaving the compressor may flow directly into heat exchanger 24 as shown in FIG. 5, or directly into heat exchanger 60 as shown in FIG. 6. If no work expander is used to expand any portion of the high pressure feed gas, compressor 16 may be replaced by a compressor driven by an external power source, such as compressor 59 shown in FIG. 7. Other situations may not warrant any compression of the stream, and the stream flows directly into heat exchanger 60 as shown in fig. 8 and bypasses the dashed line equipment (heat exchanger 24, compressor 16, and discharge cooler 25) in fig. 1. If operation of the heat exchanger 24 to heat the stream is not included before the plant fuel gas (stream 48) is withdrawn, an auxiliary heater 58, using a utility stream or other process stream to provide the required heat, is necessary to warm the fuel gas prior to its being flared, as shown in FIGS. 6-8. Such options typically must be evaluated for each application, and factors such as gas composition, plant size, desired co-product stream recovery, and available equipment must all be considered.

The operation of cooling the inlet gas stream and the feed stream to the LNG production section according to the present invention may be carried out in many ways. In the processes of fig. 1 and 3-8, inlet gas stream 31 is cooled and condensed by an external refrigerant stream and flashed separator liquid. However, some cooling may also be provided to the high pressure refrigerant (stream 71a) with the cold process stream. Also, any stream that is cooler in temperature than the stream to be cooled may be utilized. For example, the vapor may be taken from a side stream of the fractionation column 19 in FIGS. 1 and 4 to 8 or the absorption column 18 in FIG. 3 and used for the cooling operation. The use and distribution of column liquids and/or gases for process heat exchange and the specific layout of the heat exchangers for cooling the inlet and feed gases must be evaluated for each application and for the process streams selected for the specific heat exchange equipment. The selection of the cooling source will depend on several factors including, but not limited to, feed gas composition and conditions, plant size, heat exchanger size, potential cooling source temperature, and the like. One skilled in the art will also recognize that any combination of the above cooling sources or cooling methods may be used in combination to achieve the desired feed stream temperature.

Also, supplemental external refrigeration may be provided to the inlet gas stream and the feed stream to the LNG production section in a number of different ways. In fig. 1 and 3-8, boiling of a single component refrigerant has been assumed for high level external refrigeration operations and evaporation of a multi-component refrigerant has been assumed for low level external refrigeration operations, the single component refrigeration operation being used to pre-cool the multi-component refrigerant stream. Alternatively, the high-level cooling and low-level cooling operations may be performed simultaneously using a plurality of single-component refrigerants having sequentially lower boiling points (i.e., "cascade refrigeration") or one single-component refrigerant at sequentially lower evaporation pressures. Another alternative is to use a multi-component refrigerant stream whose composition has been adjusted to provide the desired cooling temperature for both high and low grade cooling operations. The choice of which method to provide external refrigeration operation will depend on several factors including, but not limited to, feed gas composition and conditions, plant size, compressor driver size, heat exchanger size, ambient heat rejection temperature, etc. One skilled in the art will also recognize that any combination of the above methods of providing external refrigeration may be used in combination to achieve the desired feed stream temperature.

Subcooling the condensed liquid stream exiting heat exchanger 60 (stream 49d in fig. 1 and 3, stream 49e in fig. 4, stream 49c in fig. 5, stream 49b in fig. 6 and 7, and stream 49a in fig. 8) reduces the amount of flash vapor generated during expansion of the stream to the operating pressure of LNG storage tank 62. This typically reduces the specific energy consumption of the LNG production process by eliminating the need for a flash gas compression process. However, some circumstances may favor reducing the capital cost of the equipment by reducing the size of the heat exchanger 60 and disposing of any flash gas that may be generated using flash gas compression or other means.

Although the expansion of the individual streams is shown in some specific expansion devices, additional means of expansion may be substituted as appropriate. For example, conditions may permit work expansion of the substantially condensed feed stream (stream 35a in fig. 1 and 3-8). Also, the process of work expanding the subcooled liquid stream leaving heat exchanger 60 ((stream 49d in FIGS. 1 and 3, stream 49e in FIG. 4, stream 49c in FIG. 5, stream 49b in FIGS. 6 and 7, and stream 49a in FIG. 8) can be replaced with an isenthalpic flash expansion process, but with either more deep subcooling in heat exchanger 60 being required to avoid flash vapor formation during expansion or to augment the flash vapor compression process or other means for disposing of the flash vapor produced.

It should also be recognized that the relative feed amounts provided by each branch from which the vapor stream is split depends on several factors, including gas pressure, feed gas composition, the amount of heat that can be economically extracted from the feed, the hydrocarbon components to be recovered into the liquid co-product, and the amount of power available. More feed to the top of the column increases recovery, but less energy is recovered from the expander, thus increasing the power requirements of the recompression process. Increasing the bottom feed reduces power consumption but also reduces product recovery. The relative position of the feed in the column may depend on the inlet gas composition or other factors such as the desired recovery and the amount of liquid formed by the inlet gas cooling process. Also, depending on the relative temperatures and amounts of the various streams, two or more streams, or portions thereof, may be combined and the combined stream fed to the column at the feed location.

While there has been described what are believed to be the preferred embodiments of the invention, those skilled in the art will recognize that other modifications may be made thereto, for example, to adapt the invention to various conditions, types of materials, or other requirements, without departing from the spirit of the invention as defined by the following claims.

Claims (23)

1. A process for liquefying a natural gas stream containing methane and heavier hydrocarbon components comprising

(a) Said natural gas stream is cooled under pressure to condense at least a part of it and form a condensed stream; and

(b) said condensed stream is expanded to lower pressure to form a liquefied natural gas stream; it is characterized in that

(1) The natural gas stream is treated in one or more cooling steps;

(2) said cooled natural gas stream is divided into at least a first stream and a second stream;

(3) said first stream is cooled to condense substantially all of it and is thereafter expanded to intermediate pressure;

(4) said second stream is expanded to said intermediate pressure;

(5) said expanded first stream and said expanded second stream are passed to a distillation column to separate said streams into a more volatile vapor distillation stream and a less volatile fraction containing a major portion of said heavier hydrocarbon components; and

(6) a vapor distillation stream is withdrawn from a region of said distillation column below said expanded second stream and cooled sufficiently to condense at least a part of it, thereby forming a residual vapor stream and a reflux stream;

(7) passing the reflux stream to the distillation column as its top feed, wherein the reflux stream is the only source of reflux in the process;

(8) said residual vapor stream is combined with said more volatile vapor distillation stream to form a volatile residue gas fraction containing a major portion of said methane and components lighter than methane;

(9) said volatile residue gas fraction is cooled under pressure to condense at least a part of it, thereby forming said condensed stream.

2. The process of claim 1 wherein the cooled natural gas stream is partially condensed;

said partially condensed natural gas stream is separated to provide a vapor stream and a liquid stream;

said vapor stream is divided into at least said first stream and said second stream;

said liquid stream is expanded to said intermediate pressure to form an expanded liquid stream; and

said expanded first stream, said expanded second stream, and said expanded liquid stream are passed to said distillation column.

3. The process of claim 2 wherein said liquid stream is expanded to said intermediate pressure and heated; and

said expanded first stream, said expanded second stream, and said heated expanded liquid stream are passed to said distillation column.

4. The process of claim 2, wherein said first stream is combined with at least a portion of said liquid stream (38) to form a combined stream;

said combined stream is cooled to condense substantially all of it and is thereafter expanded to said intermediate pressure;

expanding the remaining portion of said liquid stream to said intermediate pressure to form an expanded remaining portion of said liquid stream; and

said expanded combined stream, said expanded second stream, and said expanded remaining portion of said liquid stream are passed to a distillation column.

5. The process of claim 4 wherein said expanded remaining portion of said liquid stream is heated to form a heated expanded remaining portion of said liquid stream; and

said expanded combined stream, said expanded second stream, and said heated expanded remaining portion of said liquid stream are passed to a distillation column.

6. The process of any of claims 1-5, wherein a liquid distillation stream is withdrawn from the distillation column at a location above the region where the vapor distillation stream is withdrawn, followed by heating the liquid distillation stream, and then reintroduced into the distillation column as another feed at a location below the region where the vapor distillation stream is withdrawn.

7. The process of any of claims 1-5, wherein said reflux stream is divided into at least a first portion and a second portion, said first portion being subsequently sent to said distillation column as its overhead stream, said second portion being supplied as another feed to said distillation column at a feed position in substantially the same region as said vapor distillation stream is withdrawn.

8. The process of claim 6 wherein said reflux stream is divided into at least a first portion and a second portion, said first portion being subsequently fed to said distillation column as an overhead stream thereof, said second portion being fed to said distillation column as another feed at a feed position in substantially the same region as said vapor distillation stream is withdrawn.

9. The process of any of claims 1-5 wherein said volatile residue gas fraction is compressed and then cooled under pressure to condense at least a part of it, thereby forming said condensed stream.

10. The process of claim 6 wherein said volatile residue gas fraction is compressed and then cooled under pressure to condense at least a part of it, thereby forming said condensed stream.

11. The process of claim 7 wherein said volatile residue gas fraction is compressed and then cooled under pressure to condense at least a part of it, thereby forming said condensed stream.

12. The process of claim 8 wherein said volatile residue gas fraction is compressed and then cooled under pressure to condense at least a part of it, thereby forming said condensed stream.

13. The process of any of claims 1-5 wherein said volatile residue gas fraction is heated, compressed, and then cooled under pressure to condense at least a part of it, thereby forming said condensed stream.

14. The process of claim 6 wherein said volatile residue gas fraction is heated, compressed, and then cooled under pressure to condense at least a part of it, thereby forming said condensed stream.

15. The process of claim 7 wherein said volatile residue gas fraction is heated, compressed, and then cooled under pressure to condense at least a part of it, thereby forming said condensed stream.

16. The process of any one of claims 8, 10 to 12 wherein said volatile residue gas fraction is heated, compressed, and then cooled under pressure to condense at least a part of it, thereby forming said condensed stream.

17. The process of any of claims 1-5 wherein said volatile residue gas fraction comprises a major portion of said methane, components lighter than methane and is selected from the group consisting of C₂Component (A) and (C)₂Component + C₃The heavier hydrocarbon component of the component.

18. The process of claim 6 wherein said volatile residue gas fraction comprises a major portion of said methane, components lighter than methane and is selected from the group consisting of C₂Component (A) and (C)₂Component + C₃The heavier hydrocarbon component of the component.

19. The process of claim 7 wherein said volatile residue gas fraction comprises a major portion of said methane, components lighter than methane and is selected from the group consisting of C₂Component (A) and (C)₂Component + C₃The heavier hydrocarbon component of the component.

20. The process of claim 9 wherein said volatile residue gas fraction comprises a major portion of said methane, components lighter than methane and is selected from the group consisting of C₂Component (A) and (C)₂Component + C₃The heavier hydrocarbon component of the component.

21. The process of claim 13 wherein said volatile residue gas fraction comprises a major portion of said methane, components lighter than methane and is selected from the group consisting of C₂Component (A) and (C)₂Component + C₃The heavier hydrocarbon component of the component.

22. The process of claim 16 wherein said volatile residue gas fraction comprises a major portion of said methane, components lighter than methane and is selected from the group consisting of C₂Component (A) and (C)₂Component + C₃The heavier hydrocarbon component of the component.

23. The process of any of claims 8, 10 to 12, 14 to 15 wherein said volatile residue gas fraction comprises a major portion of said methane, components lighter than methane and selected from the group consisting of C₂Component (A) and (C)₂Component + C₃The heavier hydrocarbon component of the component.