UA76750C2 - Method for liquefying natural gas (versions) - Google Patents

Abstract

A process for liquefying natural gas (50) in conjunction with producing a liquid stream containing predominantly hydrocarbons heavier than methane (41) is disclosed. In the process, the natural gas stream to be liquefied (31) is partially cooled, expanded to an intermediate pressure (14, 15), and supplied to a distillation column (19). The bottom product (41) from this distillation column preferentially contains the majority of any hydrocarbons heavier than methane that would otherwise reduce the purity of the liquefied natural gas (50). The residual gas stream (37) from the distillation column (19) is compressed (16) to a higher intermediate pressure, cooled under pressure (60) to condense it, and then expanded (61) to low pressure to form the liquefied natural gas stream.

Description

Description of the invention

The invention relates to the process of processing a stream of natural gas or other methane-rich gas to obtain a stream of liquefied natural gas (LNG) containing high-purity methane and a liquid stream that mainly contains hydrocarbons heavier than methane | see conditional application 60/296 848 dated 06/08/20011.

Natural gas is usually extracted from wells drilled into underground reservoirs. The main component of this gas is methane, which is at least 5095 (mole) of gas. Depending on the specific reservoir, natural gas also contains smaller amounts of heavier hydrocarbons, such as ethane, propane, 70 butanes, pentanes, etc., as well as water, hydrogen, nitrogen, carbon dioxide, etc.

Of course, they will work with natural gas in gaseous form. The most common means of transporting natural gas from the well to gas processing enterprises and from there to natural gas consumers are high-pressure gas pipelines. However, in many cases it is necessary and/or desirable to liquefy natural gas for transportation or use. In remote locations, for example, 72 there is often no gas pipeline infrastructure that would enable conventional transportation of natural gas to market. In such cases, the much lower specific volume of LPG compared to natural gas significantly lowers the cost of transportation due to the use of cargo ships and cars.

Another factor in favor of liquefaction of natural gas is the possibility of its use as a fuel for vehicles. In large cities filled with large numbers of buses, taxis, cars and trucks, it is profitable to use CNG in them, if there is a sufficiently cheap source of it. Such vehicles that use LPG as fuel are less polluting due to the clean combustion of natural gas compared to similar vehicles with gasoline or diesel engines that burn higher molecular weight hydrocarbons. In addition, when using LPG of high purity (for example, methane with a purity of 9595 moles), a significantly smaller amount of carbon dioxide ("greenhouse" gas) is formed due to the lower carbon/hydrogen ratio for methane compared to other hydrocarbon fuels. Gee)

The invention generally relates to the liquefaction of natural gas with the simultaneous production of a liquid stream as a by-product, consisting mainly of hydrocarbons heavier than methane, for example, natural gas liquids (NGLs) containing ethane, propane, butanes and heavier hydrocarbon components , liquefied natural gas (LNG), or from condensates of butanes and heavier hydrocarbon components. The production of runoff liquid offers two important advantages: the methane in LPG is of high purity, and the runoff liquid is a valuable product that can be used for various purposes. The flow of natural gas to be treated according to the invention has the following composition (molar): 8295 methane, 7.995 phtane and other CO components, 4.995 propane and other C3 components, - 1.096 isobutane 1.190 normal butane 0.895 pentanes | the rest is nitrogen and carbon dioxide. Sulfur-containing gases are sometimes present. 3o There are many ways to liquefy natural gas | see for example Ripp, Aagiap 9., Sgapi Y. doppzop apa in

Tetpu K. Totiipzop, "IMO Tesppoiodu yog Obvzpoge apa Mia-ZsaIe Riapiv", Rgoseedipdz oi Zemepiu-Mipii Appiai

Sopmepip oi (She Savz Rgosevzzoge Avvzosiayop, pp. 429-450, AMTsapia, Seogdia, Magspy 13-15, 2000 and a review of such processes in Kikaja, Mozpiyividi, Mazaaki Opizpi apa Mopuozpi Mogama, "Oriitige (She Romzheg Zuviet oi Wazeioai "

IMO Riapg", Rgoseedipdz oi Eidpke(y Appia| Sopmepiop oi (ye baz Rgosevzvoge Avzosiaiop, Zap Apiopio, Tehav, Z 70 Magsp 12-14, 2001) and US patents 4,445,917, 4,525,185, 4,545,795, 4,755,200 , 5,291,736.5,363,655.5,365,740. 6,250,105 z" B1, 6,269,655 B1, 6,272,882 B1, 6,308,531 B1, 6,324,867 B1 and 6,347,532 VI. These methods usually include operations by which natural gas is purified (removing water and harmful compounds, for example carbon dioxide and sulfur compounds), cool, condense and expand. Cooling and condensation of natural gas 42 can be carried out in many ways. "Cascade cooling" uses the heat exchange of natural gas 7 with several refrigerants, the boiling point of which is successively reduced, for example, with propane ethane and

Gee! methane In another variant, this heat exchange is carried out with one refrigerant with evaporation of this refrigerant at different pressures. In case of "multicomponent cooling" heat exchange is used. natural gas with several liquid refrigerants, which contain several component refrigerants instead of several single-component refrigerants. The expansion of natural gas can be carried out both isenthalpically (using, for example, Joule-Thomson expansion), and isentropically (using, for example, a working expansion turbine).

Regardless of the method of liquefaction of the natural gas stream, a mandatory requirement is the removal of a significant portion of hydrocarbons heavier than methane before liquefaction of the methane-rich stream. Such removal of 29 is due to many reasons, including the need to control the calorific value of the LPG flow and

HF) calorific value of these heavier hydrocarbon components as independent products. Unfortunately, little attention has been paid to the efficiency of the hydrocarbon removal operation. o In accordance with the invention, it has been found that careful integration of the hydrocarbon removal operation into the natural gas liquefaction process can provide both LPG and a separate heavier hydrocarbon liquid product with significantly 60 lower energy costs compared to existing processes. Although the invention can be used at lower pressures, it is best to supply the gas to the treatment at a pressure of from 400 to 1500 psi. inch (from 2758 to 10342kPa(a) or higher.

For a better understanding of the invention, below are examples with drawings in which:

Fig. 1 is a flow diagram of a natural gas liquefaction plant adapted for the co-production of LPG according to the present invention,

Fig. 2 is a pressure-enthalpy phase diagram for methane, which illustrates the advantages of the invention compared to existing processes,

Fig. 3 is a flow diagram of another natural gas liquefaction plant adapted for co-production of LPG

According to the invention,

Fig. 4 is a flow diagram of another natural gas liquefaction plant adapted for LNG co-production according to the invention,

Fig. 5 - flow diagram of a natural gas liquefaction plant adapted for co-production of condensate according to the invention, 70 Fig. 6 - flow diagram of a natural gas liquefaction plant adapted for co-production of liquid flow according to the invention,

Fig. 7 is a flow diagram of another natural gas liquefaction plant adapted for the co-production of a liquid flow according to the invention,

Fig. 8 is a flow diagram of another natural gas liquefaction plant adapted for co-production of 7/5 liquid flow according to the invention,

Fig. 9 is a flow diagram of another natural gas liquefaction plant adapted for the co-production of a liquid flow according to the invention,

Fig. 10 is a flow diagram of another natural gas liquefaction plant adapted for the co-production of a liquid flow according to the invention,

Fig. 11 is a flow diagram of another natural gas liquefaction plant adapted for the co-production of a liquid flow according to the invention,

Fig. 12 is a flow diagram of another natural gas liquefaction plant adapted for the co-production of a liquid flow according to the invention,

Fig. 13 is a flow diagram of another natural gas liquefaction plant adapted for co-production of liquid flow components according to the invention,

Fig. 14 is a flow diagram of another natural gas liquefaction plant adapted for co-production i) liquid flow according to the invention,

Fig. 15 is a flow diagram of another natural gas liquefaction plant adapted for the co-production of a liquid flow according to the invention, Fig. 16 is a flow diagram of another natural gas liquefaction plant adapted for the co-production of a liquid flow according to the invention,

Fig. 17 is a flow diagram of another natural gas liquefaction plant adapted for co-production of M liquid flow according to the invention,

Fig. 18 - flow diagram of another natural gas liquefaction plant adapted for co-production of liquid flow according to the invention,

Fig. 19 is a flow diagram of another natural gas liquefaction plant adapted for co-production of a liquid flow according to the invention,

Fig. 20 is a flow diagram of another natural gas liquefaction plant adapted for co-production of a liquid flow according to the invention, and "

Fig. 21 is a flow diagram of another natural gas liquefaction plant adapted for co-production with a liquid flow according to the invention.

In the further description of the drawings, tables are given that give flow rates calculated for typical conditions;" process In these tables, the flow rate values (in mol/h) have been rounded to the nearest whole number for convenience. The total flow rates shown in the tables include all non-hydrocarbon components and are therefore greater than the sum of the hydrocarbon component flow rates. Temperatures are rounded to the nearest whole number. It should be noted that the calculations for the various processes illustrated by the drawings, performed for the purpose of their comparison, are based on the assumption of the absence of heat exchange with the process environment. The quality of industrial heat-insulating materials

Me. materials makes such an assumption justified. -I For convenience, process parameters are given both in traditional British units and in SI units. The molar flow rates given in the tables can be interpreted as lb-mol/hr. or as kg-mol/h. o Power consumption is given in both horsepower and thousands of British thermal units Productivity o is given in lb/hr, corresponding to flow rates in lb-mol/hr. Productivity given in kg/h corresponds to flow rates in kg-mol/h.

Example 1

Figure 1 illustrates a process according to the invention in which it is desirable to produce LPG as a co-product that contains most of the ethane and heavier components of the natural gas input stream. According to the simulation of the invention

F) gas enters the installation at 902E (322C) and 1285 lb/sq. inch (8860kPag)) (stream 31). If the inlet gas contains carbon dioxide and/or sulfur compounds in concentrations that do not meet the product stream requirements, these compounds are removed by proper inlet gas pretreatment (not shown). In addition, the inlet stream is usually dewatered to prevent ice formation during cooling. For this, usually use hard sieving agents.

The input flow 31 is cooled in the heat exchanger 10 through heat exchange with the refrigerant flows and liquids of the demethanization part of the reboiler at -689E (-5522) (flow 40). In all cases, heat exchanger 170 represents either multiple individual heat exchangers or a single multi-pass heat exchanger (the decision 65 to use multiple heat exchangers or a single multi-pass heat exchanger depends on many factors, including inlet gas flow rate, heat exchanger size, flow temperature, etc.).

Cooled stream Z1a enters separator 11 at -309E (-342C) and 127v8lb/sq. in. (8812kPa(a)) where the vapor (stream 32) is separated from the condensed liquid (stream 33).

The vapor (stream 32) from the separator 11 is separated into two streams, 34 and 36. Stream 34 (approximately 2095 of all vapors) is combined with the condensed liquid (stream 33) to form stream 35. The combined stream 35 passes through the heat exchanger 13 and undergoes heat exchange with flow 71 is a refrigerant, as a result of which cooling and significant condensation of flow ZbBa occur. The substantially condensed stream C5a at -12CPE (-852C) is flashed through a suitable expansion device, such as expansion valve 14, to the operating pressure (approximately 465 psi (3206bPa(a))) of distillation column 19. During expansion, a portion of the stream 70 evaporates, thereby cooling the entire stream.In the process of Fig.1, the expanded stream 356 exiting the expansion valve 14 reaches a temperature of -122 o (-862C) and enters the mid-entry point in the demethanization section 196 of the distillation column 19.

The remaining (8095) vapor from separator 11 (stream 36) enters the working expansion machine 15, in which mechanical energy is extracted from a portion of this high-pressure input stream. Machine 15 expands the vapor essentially 75 isentropically, starting at a pressure of about 1278 psi. inch (8812 kPas)) and to the working pressure of the column with the working cooling of the extended Zba flow to a temperature of approximately -1032E (-752С). Typical industrial expanders can provide 80-85956 of the work theoretically possible with an ideal isentropic expansion.

The resulting work is often used to drive a centrifugal compressor (eg, 16), which is used, for example, to recompress the column overhead gas (stream 38).

The expanded and partially condensed stream of Zba enters the distillation column 19 at the lower entry point in the middle of the column.

The demethanizer of the distillation column 19 is a conventional distillation column that contains a collection of vertically spaced plates, one or more packed layers, or a combination of packers and plates. In natural gas processing plants, the distillation column often consists of two sections. The first Ge section 19a is a separator in which the overhead stream is separated into vapor and liquid portions, and the vapors (5) rising from the lower distillation or demethanization section 190 are combined with the vapor portion (if any) of the overhead stream to form cold of the upper vapors of the demethanizer (stream 37) in the upper part of the column with a temperature of -1359E (-9322). The lower demethanization section 196 contains plates and/or packers and provides the necessary contact between the liquids falling down and the vapors rising up. The demethanization section (c) also contains one or more reboilers (e.g., 20) that heat and vaporize some of the liquid flowing down the column to create flash vapors rising upward. Liquid product stream 41 exits the bottom of the column at 1152 (462) according to the requirements for a molar ratio of methane/ethane and - (0.020:1) in the bottom product. "co

The upper vapors of the demethanizer (stream 37) are heated to 902E (322) in the heat exchanger 24 and part of these vapors

Zo is diverted for use as fuel gas (flow 48) for the installation (the amount of fuel gas to be diverted is determined mainly by the fuel demand for the engines and/or turbines driving the gas compressors of the installation, for example compressors 64, 6b 68 refrigerant). The remainder of the heated upper vapors of the demethanizer (stream 38) is compressed by compressor 16, which is driven by expansion machines 15, 61 « 63. After cooling to -100 eE (-382C) in the outlet cooler 25, stream 3865 is additionally cooled to -1239E (-862C) in the heat exchanger 24 through exchange with the cold overhead vapors of the demethanizer (stream 37). After that, the Zvs flow enters the heat exchanger 60 and is additionally cooled by the refrigerant flow 714. After cooling to an intermediate temperature, the stream Zvs is divided into two parts. The first part (stream 49) is further cooled in the heat exchanger 60 to -2572E (-1602С) for condensation and subcooling, after which it enters the working expansion machine 61, where the stream is extracted mechanical energy Machine 61 - 15 expands liquid stream 49 substantially isentropically from a pressure of approximately 562 psi (3878 kPa(a)) to an LPG storage pressure of 15.5 psi (107 kPa(a)). the expansion cools the expanded stream 49a to (o) a temperature of approximately -2582E (-1612C), after which it is directed to the CNG storage tank 62 (stream 50). -1 The rest of the stream Z8c (stream 39) exits the heat exchanger 60 at -1602E ( -1072С) and quickly evaporates

In the proper expansion device, for example, through the expansion valve 17, to the working pressure of the rectification 1 column 19. In this process (Fig. 1), the evaporation of the expanded stream Z9a does not occur, and therefore its temperature drops slightly to -1612Е (-1072С), on exits from valve 17 expansion. Next, the expanded flow Z9da enters the section 19a of the separator in the upper part of the rectification column 19. The liquid separated there becomes the upper input product of the section 1906 of the demethanizer. 5B All cooling of streams 35, З8s is carried out in a closed cycle. The working fluid for this cycle is a mixture of hydrocarbons and nitrogen, and the composition of the mixture is adjusted to ensure the desired temperature (F. of the refrigerant and condensation at an acceptable pressure using the available cooling medium. In this case, condensation is carried out by cooling water and therefore, when simulating the process of Fig. 1 uses a mixture of nitrogen, methane, ethane, propane and heavier hydrocarbons. The approximate molar composition of the stream is nitrogen - 7.595, methane - 41.095, ethane - 41.595, propane - 10.095 and the rest - heavier hydrocarbons.

Refrigerant stream 71 exits refrigerator 69 at a temperature of 1002 (3822) under a pressure of 700 psig. inch (4185 kPas(a)). It enters the heat exchanger 10 and is cooled to -312E (-359С) and partially condensed by the partially heated expanded flow 71 and other refrigerant flows. For the simulation of the embodiment of Fig. 1, it was assumed that these other refrigerant flows are commercial propane refrigerant flows at three different 65 temperatures and pressures. The partially condensed refrigerant flow 71a then enters the heat exchanger 13 for further cooling to -114 9 (-81205) by the partially heated expanded refrigerant flow 71e,

condensation and supercooling by the flow of 715 refrigerant. Next, the refrigerant in the heat exchanger 60 is supercooled to -2572E (-16022)3 by the flow 71a of the expanded refrigerant. The supercooled liquid stream 71c enters the working cooling machine 63 where mechanical energy is extracted from it with simultaneous substantially isentropic expansion from a pressure of approximately 586 psi. in. (4040kPa(a)) to about psig/sq. inch (234kPa(a)). During expansion, part of the flow evaporates, which cools the entire flow to -2639E (-16422) (flow 71 4) The expanded flow 714 again enters the heat exchangers 60, 13, 10 where it cools the flows Zvs, 35 and the refrigerant (flows 71, 71a, 715) with evaporation and overheating.

The superheated vapors of the refrigerant (flow 719) leave the heat exchanger 10 at 932E (342С) and are brought to a pressure of 6b17lb/sq in three stages 70. inch (4254kPa(a)). At each of these stages, the compressors 64, 66, 68 of the refrigerant are driven by an auxiliary power source with subsequent cooling (exhaust coolers 65, 67, 69) to remove the heat of compression. The stream 71 under pressure from the outlet cooler 69 returns to the heat exchanger 10 to complete the cycle.

The flow rates and energy consumption of the process of Fig. 1 are given in Table 1. 001000 v5yu o yavozhv yiv o |17o1vlvya vozyv | 06001000 from the state of Uni 01010100 F from the south

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Ф прдюант 0100 zv m pozh 001 « - с г» Teplodlyavlyasnietotev! 0011 -I b K.kd. of LPG production processes is usually compared to the desired "specific energy consumption", which is the ratio of full refrigerant compression capacity to full liquid production. Data from publications on this subject set limits for specific energy consumption from 0.168 hp-hr/lb ( 0.276bkW/kg) to g 20 0.182hp-hr/lb (0.300kW/kg) which is assumed to be based on a flow rate of 340 days per year for an LPG production plant. Accordingly, the specific energy consumption for embodying Figs. .1 is mee, 0.161 hp-hr/lb (0.265 kW/kg), which corresponds to a 4-1395 efficiency increase over existing solutions.

In addition, it should be noted that specific energy consumption in existing processes is based on co-production only

LNG (C3 and heavier hydrocarbons) or liquid condensate stream (C; and heavier hydrocarbons) with relatively low 22 yield, but not LPG (C" and heavier hydrocarbons), as shown in Fig.1. Existing processes need a lot

F! more cooling capacity to co-produce an LPG stream instead of an LNG or condensate stream.

Two main factors determine the improvement of the efficiency. according to the invention. The first factor is related to the thermodynamics of the high-pressure gas flow liquefaction process considered in this example. Since the main component of this flow is methane, thermodynamic properties of methane can be used to compare the existing liquefaction cycle with the cycle of the invention. Fig. 2 contains the pressure-enthalpy phase diagram for methane. In most existing cycles, all cooling of the gas flow occurs under high pressure (path A-B), after which the flow expands (path B-C), gaining pressure in the methane storage tank (slightly above atmospheric). A work expansion machine can be used in this expansion operation, usually capable of extracting 75-8095 of the work theoretically available in an ideal isentropic expansion. For simplicity, a fully isentropic expansion is illustrated along the B-C path. The decrease in enthalpy by this work expansion is small because the lines of constant entropy are almost vertical in the liquid region of the phase diagram.

This contrasts with the liquefaction cycle according to the invention. After partial cooling under high pressure (path A-A"), the gas flow undergoes operational expansion (path A"-A"), acquiring an intermediate pressure (for simplification, a fully isentropic expansion is considered). Further cooling is carried out at an intermediate pressure (path A"-B ") after which the flow expands, gaining pressure in the LPG storage tank. Since the lines of constant entropy are less steep in the gas region of the phase diagram, the first work expansion operation provided by the invention gives a significantly greater reduction in enthalpy. Therefore, the complete cooling provided by the invention (the sum of the paths A-A' and A"-B) is less than the cooling in the existing processes (path A-B), and this reduces the cooling (and therefore the cooling pressure) required to liquefy the gas stream. 70 The second factor determining the improvement of the efficiency according to the invention there is better functioning of hydrocarbon distillation systems at lower operating pressures. The hydrocarbon removal operation in most existing processes is performed at high pressure using a scrubber column and a cold hydrocarbon liquid as an in-stream absorbent to remove heavier hydrocarbons from the input gas stream. The operation of the scrubber column under high pressure is not very efficient, because it causes co-absorption of a significant part of methane and /5 ethane from the gas stream, which must then be cleaned of liquid absorbent and cooled to become part of LPG. According to the invention, the hydrocarbon removal operation is performed at an intermediate pressure when the vapor-liquid equilibrium is more favorable and this provides a very effective removal of the desired heavier hydrocarbons in the co-product liquid stream.

Example 2

If LPG requirements require a higher ethane content in the input gas to be stored in the final

ZNG, the embodiment of the invention can be simplified. Fig. 3 illustrates such an embodiment. The input gas composition and process conditions for Fig. 3 are similar to those provided for Fig. 1. Accordingly, the process of Fig. 3 can be compared with the process of Fig. 1.

When modeling the process of Fig. 3 cooling of the input gas, the separation and the expansion scheme for the RNG receiving section are essentially similar to those used for Fig. 1. high school

Gas enters the plant at 902 (3222) and 1285 lb/sq. inch (8860OkPats(a)) (stream 31) and is cooled in the heat exchanger 10 through heat exchange with refrigerant flows and liquids of the demethanization side of the reboiler at about -359E (-372С) (stream 40). The cooled Zt1a stream enters separator 11 at -309E (-342C) and 1278 lb/sq. inch (8812 kPa)) where the vapor (stream 32) is separated from the condensed liquid (stream 33).

Steam (stream 32) from separator 11 is divided into two streams, 34 and 36. Stream 34 (approximately 2090 of all steam) av! combines with the condensed liquid (flow 33) to form flow 35. The combined flow 35 passes through the heat exchanger 13 and undergoes heat exchange with flow 71, which is a refrigerant, as a result of which cooling and significant condensation of the ZbBa flow occur. The substantially condensed ZbBa stream at -120 2 (-85922) is rapidly vaporized through a suitable expansion device, such as expansion valve 14, to the operating pressure (approximately 465 psi (3206 bkPa(a))) of distillation column 19. During the expansion, a part of the y-o stream evaporates thereby cooling the entire stream. In the process of Fig. 3, the expanded stream 350 leaving the y- expansion valve 14 reaches a temperature of -122 PE (-8622) and enters the separator section of the upper part of the rectification column 19 The liquids separated here become the input streams of the deethanization section in the lower part of the rectification column 19.

The remaining (8095) vapor from separator 11 (stream 36) enters the working expansion machine 15, in which mechanical energy is extracted from a portion of this high-pressure input stream. Machine 15 expands the vapor essentially isentropically, starting at a pressure of about 1278 psig. inch (8812 kPas)) and to the working pressure of the column with the working and cooling of the expanded Zba flow to a temperature of approximately -103 RE (-7592). The expanded and partially » condensed Zba stream enters the rectification column 19 at the inlet point in the middle of the column.

Demethanizer overhead vapors (stream 37) are heated to 902E (322) in heat exchanger 24 and a portion of these vapors (stream 48) are diverted for use as plant fuel gas. The rest of the heated upper vapors - demethanizer (stream 49) is compressed by compressor 16. After cooling to -100 РЕ (-382С) in the outlet

F of the refrigerator 25 stream 496 is additionally cooled to -1122E (-802C) in the heat exchanger 24 due to mutual exchange with the cold upper vapors of the demethanizer (stream 37). it. After that, the stream 49c enters the heat exchanger 60 and is additionally cooled by the refrigerant stream 714 to sl 20 2579E (-1602С) for condensation and subcooling, after which it enters the working expansion machine 61, where mechanical energy is extracted from the stream. Machine 61 expands liquid stream 494 substantially isentropically from a pressure of approximately 583 psi. inch (4021 kPas(a)) to LPG storage pressure of 15.5 lb/sq. inch (107kPa(a)) slightly above atmospheric. The working expansion cools the expanded stream 49e to a temperature of approximately -2589°E (-1612C), after which it is directed to the CNG storage tank 62 (stream 50).

Similar to the process of Fig. 1, all cooling of streams 35, 49c is carried out in a closed cooling circuit. (F) Approximate molar composition of the stream used as a working fluid in the process cycle of Fig. 3Z nitrogen - 7.590, methane - 40.095, ethane - 42.595, propane - 10.095 and the rest - heavier hydrocarbons.

Refrigerant stream 71 exits refrigerator 69 at a temperature of 1002 (3822) under a pressure of 700 psig. in inch (4185 kPas(a)) It enters the heat exchanger 10 and is cooled to -312E (-352С) and partially condensed by the partially heated expanded flow of the 71st and other refrigerant flows. For the simulation of FIG. 3, these other refrigerant streams were assumed to be commercial propane refrigerant streams at three different temperatures and pressures. The partially condensed flow 71a of the refrigerant then enters the heat exchanger 13 for further cooling to -121 EE (-859C) by the partially heated expanded flow 71e of the refrigerant, 65 Condensation and subcooling by the flow 715 of the refrigerant. Next, the refrigerant in the heat exchanger 60 is supercooled to -257 2E (-1602С) by the flow 714 of the expanded refrigerant. Supercooled liquid stream

71c enters the working cooling machine 63, where mechanical energy is extracted from it with simultaneous substantially isentropic expansion from a pressure of approximately 586 psi. in. (4040 kPas(a)) to about 34 lb/sq. inch (234kPa(a)). During expansion, part of the stream evaporates, which cools the entire stream to -2639E (-16422) (stream 71 4) Expanded stream 714 again enters heat exchangers 60, 13, 10, where it cools streams 49c, 35 and the refrigerant (streams 71, 71a , 715) with evaporation and overheating.

Superheated refrigerant vapor (stream 719) exits heat exchanger 10 at 932E (342C) and is brought to a pressure of 617 psi in three stages. in. (4254kPa(a)) At each of these stages, the refrigerant compressors 64, 66, 68 are driven by an auxiliary power source with subsequent cooling (refrigerating outlet 70 coolers 65, 67, 69) to remove the heat of compression. The stream 71 under pressure from the outlet cooler 69 returns to the heat exchanger 10 to complete the cycle.

The flow rates and energy consumption of the process of Fig. 3 are given in Table 2.

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If we assume a flow factor of 340 days per year for an LPG production plant, then the specific energy consumption for the implementation of Fig. 3 will be 0.153 hp-h/lb (0.265 kW/kg), which corresponds to an increase - k.k.d. . on 10-2095. As already noted, the invention provides an increase in efficiency, even with the co-production of LPG, rather than LNG or condensate, as in existing processes.

Compared to the embodiment of Fig. 1, the embodiment of Fig. 3 requires approximately 595 less energy per unit of liquid produced. Therefore, with the given available power for compression, the embodiment of Fig. 3 gives approximately 595 more liquefied s 20 of natural gas compared to the embodiment of Fig. 1 due to obtaining a smaller amount of C»5o and heavier hydrocarbons in the RNG co-product. The choice between the embodiments of Fig. 1 and Fig. 3 in specific cases is determined either by the cost of heavier (ze) hydrocarbons in the RNG product compared to their corresponding cost in the CNG product or by the calorific value of CNG (since the calorific value of CNG given by the embodiment of Fig. 1 is lower than calorific value of CNG given by the embodiment of Fig. 3).

Example C 22 If LPG requirements allow all ethane from the input gas to be kept in the final LPG or if there is no demand for

Gee! liquid ethane-containing co-product, another embodiment of the invention illustrated in Fig. 4 can be used to obtain a LNG co-product stream. The composition of the input gas and the process conditions for Fig. Accordingly, the process of Fig. 4 can be compared with the processes of Fig. 1 and 3.

When simulating the process of Fig. 4, gas enters the plant at 902E (322C) and 1285 lb/sq. inch (8860kPa(a)) 60 (stream 31) and is cooled in the heat exchanger 10 through heat exchange with refrigerant flows and fast separation fluids at -462E (-43223 (stream ЗЗа). The cooled stream Z1a enters the separator 11 at -19E (-182С ) and 127 psi (8812 kPa)) where the vapor (stream 32) is separated from the condensed liquid (stream 33).

The steam (flow 32) from the separator 11 enters the working expansion machine 15, in which mechanical energy is extracted from part of the high-pressure flow. Machine 15 expands the vapors essentially isentropically starting from a pressure of about 1278 psig. in. (8812kPa(a)) and to a pressure of about 440 psig. in. (3034kPa(a)) (the operating pressure of the separation/absorption column 18) and reduces the temperature of the expanded stream 32a to approximately -8121E (-632C).

The expanded and partially condensed stream 32a enters the absorption section 1806 at the bottom of the separation/absorption column 18). The liquid portion of the expanded stream mixes with the liquid falling down from the absorption section, and the combined liquid stream 40 exits the bottom of the separation/absorption column 18 at -86beE (-662C). The vapor part of the expanded flow rises up through the absorption section and contacts the cold liquid falling down for condensation and absorption of C3-components and heavier components.

The separation/absorption column 18 is a conventional distillation column that contains a collection of vertically spaced plates, one or more packed beds, or a combination of packers and plates. In 70 natural gas processing plants, the separation/absorption column 18 often consists of two sections. The upper section 1843 is a separator in which the vapors of the upper stream are separated from the liquid portion and the vapors rising from the lower section 1860 of distillation or absorption are combined with the vapor portion (if any) of the upper stream to form a cold distillation stream 37, which exits the upper part of the column

The lower, absorption section 1865 contains plates and/or packers and provides the necessary contact between the falling liquids 75 and the rising vapors for condensation and absorption of C3 components and heavier components.

The combined liquid stream 40 from the bottom of the separation/absorption column 18 is directed by the pump 26 to the heat exchanger 13 (stream 40a), where it is heated by cooling the upper stream 42 of the deethanizer and the refrigerant (stream 71a).

The combined liquid stream is heated to -242E (-312C2)3, partially vaporizing stream 4065 before feeding it as a mid-column stream to the deethanizer 19. The separator liquid (stream 33) is rapidly vaporized in the expansion valve 20 12, gaining a pressure just above the deethanizer operating pressure and a temperature of -462E (-432C) (stream ZZa) before it cools the inlet gas as described above. Stream 330, which now has a temperature of 852 (2922), enters the deethanizer 19 at the lower entry point in the middle of the column. In the deethanizer, streams 406 and 33p are stripped of methane and CO components. A deethanizer in column 19 which operates at about 453 lb/sq. inch (3123 kPad(a)), is also a conventional distillation column that contains a collection of vertically spaced 25 plates, one or more packed beds, or a combination of packers and plates. The deethanization column Ge) can consist of two sections of the upper separator section 19a, where the vapors of the upper stream are separated from the liquid part and the vapors rising from the lower section 196 of distillation or deethanization are combined with the vapor part (if any) of the upper stream to form a distillation stream 42, which exits the top of the column, and a lower deethanization section 195 that includes plates and/or packers to provide contact between the falling liquids and the rising vapors. The deethanization section 196 also includes one or more reboilers (for example, 20), which heat and evaporate part of the liquid at the bottom of the column to create stripping vapors that rise up to clean the liquid product (stream 41) - methane and -

Co-components. According to typical requirements, the bottom product should have an ethane/propane molar ratio of 0.020:1. "I

Stream 41 of the liquid product comes from the bottom of the deethanizer at a temperature of 2142E (10120). 3o The working pressure in the deethanizer 19 is maintained at a level slightly higher than the working pressure of the column 18 in separation/absorption. This allows the overhead vapors (stream 42) of the deethanizer to pass under pressure through the heat exchanger 13 and from there into the upper section of the separation/absorption column 18. In the heat exchanger 13, the upper vapors of the deethanizer at a temperature of -199E (-282C) undergo heat exchange with the combined liquid flow 40a, which comes from the bottom of the separation/absorption column 18 and the flow 71 is a rapidly evaporated refrigerant that cools this flow to -899E (-672C) (stream 42a) and partially condenses it. The partially condensed stream enters the collector 22 of the irrigation fraction, where the condensed liquid (stream 44) is separated from non-condensed vapors (stream 43). The stream 43 is combined with the stream 37 of distillation vapors coming from the upper part of the separation/absorption column 18, with the formation of a stream 47 of cold dry gas. Condensed liquid (flow 44) under higher pressure is supplied by pump 23, after which flow 44a is divided into two parts. One of them (stream 45) enters - 15 the upper separator section of the separation/absorption column 18 and serves as a cold liquid in contact with the vapors rising up through the absorption section. The second part enters the deethanizer 19 as counterflow 46, which (o) flows to the upper entry point of the deethanizer 19 at -892E (-672C). -1 Cold dry gas (stream 47) is heated from -8992E (-672C) to 942E (342C) in the heat exchanger 24, after which the 5th part (stream 48) is selected as fuel gas for the installation. The rest of the heated dry gas (stream 49) 1 is compressed by the compressor 16. After cooling to 100 o (3822) in the outlet cooler 25, the stream 495 o is additionally cooled to -782E (-612С) in the heat exchanger 24 through heat exchange with cold dry gas (stream 47) .

The stream 49c enters the heat exchanger 60 and is cooled by the refrigerant stream 714 to -2552E (-1602С) for condensation and subcooling, after which it enters the working expansion machine 61, which extracts mechanical energy from the stream. Machine 61 expands liquid stream 4940 substantially isentropically from a pressure of approx.

F, 64 in 8 pounds/sq. inch (4465 kPa(a)) to LPG storage pressure - 15.5 lb/sq. inch (107kPa(a)), i.e. slightly above atmospheric. The working expansion cools the expanded stream 49e to a temperature of approximately -256 sg (-1602C), after which it enters the CNG storage tank 62 (stream 50). 60 Similar to the processes of Fig. 1 and 3, all cooling of streams 42, 49c is carried out in a closed cooling circuit. The approximate molar composition of the stream used as a working fluid in the process cycle of Fig. 4 is nitrogen - 8,795, methane - 30,095, ethane - 45,895, propane - 11,095 and the rest - heavier hydrocarbons. Refrigerant stream 71 exits cooler 69 at a temperature of 1009 (3822) at a pressure of 7 psig. inch (4185kPa(a)). It enters the heat exchanger 10 and cools down to -31 9 (-35922) and is partially condensed by the partially heated expanded flow of the 71st and other refrigerant flows. For the simulation of Figure 4, these other refrigerant streams were assumed to be commercial propane refrigerant streams at three different temperatures and pressures. The partially condensed flow 71a of the refrigerant then enters the heat exchanger 13 for further cooling to -121 2E (-859С3 is a partially heated expanded flow 71 of the refrigerant, for condensation and subcooling by the flow 715 of the refrigerant. Further, the refrigerant in the heat exchanger 60 is subcooled to -2552E (-1602С) flow Expanded Refrigerant 714 The supercooled liquid stream 71c enters the working cooling machine 63 where mechanical energy is extracted from it with simultaneous substantially isentropic expansion from a pressure of about 586 psi (4040 kPas(a)) to about 34 psi (234 kPa(a) )). During expansion, part of the flow evaporates, which cools the entire flow to -2649E (-1642С) (flow 71 a). The expanded flow 71 again enters the heat exchangers 60, 13, 10, where it cools 70 flows 49c, 35 and the refrigerant (flows 71, 71a, 715) during evaporation and overheating.

Superheated refrigerant vapor (stream 719) exits heat exchanger 10 at 902 (322C) and is brought to a pressure of 617 psi in three stages. in. (4254kPa(a)) At each of these stages, the refrigerant compressors 64, 66, 68 are driven by an auxiliary power source, followed by cooling (exhaust coolers 65, 67, 69) to remove the heat of compression. The stream 71 under pressure from the outlet cooler 69 75 returns to the heat exchanger 10 to complete the cycle.

Flow rates and energy consumption of the process of Fig. 4 are given in table 3. 2 tu zbe; call! /7o sov sch o 1018 0mzh0zv av lozvy | already L6v00000000lvtva yaz dvz 317701 la ovotv vv | "66 | o go o zo zvo |za) v100po yu v | zv vv10000 5 same |z) t 00000000 m v/m iv 0101000 F in zvvj | zv 20100000 az z Receive 010100 in "

Prdani 1 o Z s- x" poet 00001000 v -

F Teplodlya you are the owner! 11111 -

Fo s If we accept the flow rate of 340 days per year for the installation of LPG production, then the specific energy consumption for the implementation of Fig. 3 will be 0.143 hp-h/lb (0.236 kW/kg), which corresponds to an increase in efficiency . on 17-2790.

Compared to the embodiments of Fig. 1 and C, the embodiment of Fig. 4 requires approximately 6-1196 less energy per unit

GF) of the produced liquid. Therefore, with the given available power for compression, the embodiment of Fig. 4 can give approximately 690 times more liquefied natural gas compared to the embodiment of Fig. 1 due to obtaining C from and heavier hydrocarbons as a co-product of LNG. The choice between the embodiments of Fig. 1 and Fig. 3 in specific cases is determined either by the cost of heavier hydrocarbons in the RNG product compared to their corresponding cost in the LPG product, or by the calorific value of 60 ZHng (since the calorific value of LPG given by the embodiments of Fig. 1 and 3 is lower for the calorific value of CNG, which is given by the embodiment of Fig. 4).

Example 4

If LPG requirements require that all ethane and propane of the input gas be retained in the final LPG or if there is no demand for a liquid ethane- and propane-containing co-product for the co-product condensate stream, the embodiment of the invention illustrated in Fig. 5 may be applied. The input gas composition and process conditions for Fig. 5 are similar to those provided for Fig. 1, 2 and 4. Accordingly, the process of Fig. 5 can be compared with the processes of Fig. 1, 2 and 4.

When simulating the process of Fig. 5, gas enters the plant at 909E (3222) and 1285 lb/sq. inch (886OkPa(a)) (stream 31) and is cooled in heat exchanger 10 by heat exchange with refrigerant streams, high pressure flash separation fluids at -372E (-382C) (stream 335) and intermediate pressure flash separation fluids at -372E (-382C ) (thread 395). Cooled stream Z1a enters separator 11 at -309E (-3422) 11278 lb/sq. inch (8812kPas)) where the vapor (stream 32) is separated from the condensed liquid (stream 33).

Steam (flow 32) from the high-pressure separator 11 enters the working expansion machine 15, in which mechanical energy is extracted from part of the high-pressure flow. Machine 15 expands the vapor essentially isentropically, 70 starting at a pressure of about 1278 psig. in. (8812 kPa(a)) and to a pressure of about bZB5 psig. inch (4378kPac(a)) and lowers the temperature of the expanded flow 32a to about -832E (-64C). The expanded and partially condensed stream 32a enters the intermediate pressure separator 18, where the vapor (stream 42) is separated from the condensed liquid (stream 39). The liquid of the intermediate pressure separator (stream 39) is quickly evaporated in the expansion valve 17 with the acquisition of a pressure slightly higher than the working pressure of the depropanizer 19 with 75 cooling of the stream 39 to -108 9 (-782C) (stream ZOa), after which it enters the heat exchanger 13 and is heated, cooling the dry gas stream 49 and the refrigerant stream 71a, and from there enters the heat exchanger 10 to cool the inlet gas, as described above. the middle point of the column.

The condensed liquid (stream 33) from the high-pressure separator 11 is quickly evaporated by the expansion valve 17 and acquires a pressure slightly higher than the working pressure of the depropanizer 19, cooling to -932E (-702C) (stream ЗЗа) before entering the heat exchanger 13, where it is heated, cooling the dry gas flow 49 and the refrigerant flow 71a. Then it enters the heat exchanger 10 to cool the incoming gas. The ZZc stream, which now has a temperature of 509 (1022), enters the depropanizer 19 at the lower entry point in the middle of the column. In the depropanizer, streams 405 and 335 are stripped of methane, CO components, and C3 components. A depropanizer in column 19, which operates at about 385:F?Uunt/sq. inch (2654 kPas(a)), is a conventional distillation column which (3) contains a collection of vertically spaced plates, one or more packed beds, or a combination of packers and plates. The depropanization column may consist of two sections of an upper separator section 192 where a pair of upper of the flow are separated from the liquid part and vapors rising from the lower section 196 of distillation or depropanization are combined with the vapor part (if any) of the upper flow with the formation of the distillation flow 37, which comes from the upper part of the column, and from the lower section 19r deethanizer, which includes plates and/or packers to provide contact between the liquids falling down and the vapors rising up. The depropanization section 195 also includes one or more reboilers in (e.g., 20) that heat and vaporize some of the liquid at the bottom columns for the creation of selectable vapors (He) rise up to clean the liquid product (flow 41) - methane, C o-components and Cz--com ponents According to 39 typical requirements, the bottom product should have a propane/butane molar ratio of 0.020:1. Stream 41 of the liquid in the product comes from the bottom part of the deethanizer at a temperature of 2862E (14120).

The upper distillation stream 37 leaves the depropanizer 19 at 36beE (222) and is cooled and partially condensed with a commercial propane refrigerant in the dephlegmator 21. The partially condensed stream 37a at 22E (-172C) enters the collector 22 of the irrigation fraction, where the condensed liquid (stream 44) is separated from uncondensed vapors (stream 43). Condensed liquid (flow 44) is supplied by pump 23 to the upper inlet point c of the depropanizer 19 as a counterflow 44a. :z» Uncondensed vapor (stream 47) from collector 22 of the irrigation fraction is heated to 94 EE (342C) in heat exchanger 24, after which part (stream 48) is selected as fuel gas for the installation. The rest of the heated dry gas (flow 49) is compressed by the compressor 16. After cooling to 100 RE (3822) in the outlet cooler 25, the flow 3865 is further cooled to 152E (-9a27) in the heat exchanger 24 through heat exchange with cold vapors (flow 43).

Next, the flow Zvs is combined with the flow 42 of vapors of the intermediate pressure separator with the formation of the flow 49 -I of cold dry gas. Flow 49 enters the heat exchanger 13 and is cooled from -389 (-399223 to -1029E (-74905) by the separator fluids (flows ZOa, ZZa), as described above, and flow 71 is the refrigerant. The partially condensed flow 49a then enters the heat exchanger 60 for further cooling to -254 oE (-159905) 2 condensation and supercooling by the flow 71 and refrigerant, after which it enters the working cooling machine 61, where mechanical energy is extracted from it with a simultaneous substantially isentropic expansion from a pressure of approximately 621 psi ( 4282 kPas)) to LPG storage pressure - 15.5 pounds/sq. inch (107kPacda)), slightly above atmospheric. The working expansion cools the expanded flow to -255 2 (-15922), after which it flows from o into the LPG storage tank 62 (flow 50).

As in the processes of Figs. 1, 3 and 4, almost all cooling of flows 49, 49c is carried out in a closed cooling circuit. The approximate molar composition of the stream used as a working fluid in the process cycle of Fig. C is nitrogen - 8.995, methane - 34.395, ethane - 41.395, propane - 10 195 and the rest - heavier hydrocarbons. Refrigerant stream 71 bo exits refrigerator 69 at a temperature of 1009 (3822) under a pressure of 7 psig. inch (4185kPa(a)). It enters the heat exchanger 10 and is cooled to -30 9 (-3422) and is partially condensed by the partially heated expanded flow of the 71st and other refrigerant flows. For the simulation of Fig. 5, it was assumed that these other refrigerant flows are commercial propane refrigerant flows at three different temperatures and pressures. refrigerant,

condensation and supercooling by the flow of 715 refrigerant. Further, the refrigerant in the heat exchanger 60 is completely condensed and supercooled to -254 9 (-15990) by the flow 714 of the expanded refrigerant.

The supercooled liquid stream 71c enters the working cooling machine 63 where mechanical energy is extracted from it with simultaneous substantially isentropic expansion from a pressure of approximately 58 bpsi. inch (4040 kPas(a)) to about psi inch (234kPa(a)). During the expansion, part of the stream evaporates, which cools the entire stream to -264 2 (-16492) (stream 71 a) The expanded stream 71 a again enters the heat exchangers 60, 13, 10, where it cools streams 49a and 49 and the refrigerant (streams 71 , 71a, 715) with evaporation and overheating.

Superheated refrigerant vapor (stream 719) exits heat exchanger 10 at 932 (342C) and is brought to a pressure of 617 psi in three stages. inch (4254 kPas(a)) At each of these stages, the refrigerant compressors 64, 66, 68 are driven by an auxiliary power source with subsequent cooling (exhaust coolers 65 67 69) to remove the heat of compression. The stream 71 under pressure from the outlet cooler 69 returns to the heat exchanger 10 to complete the cycle.

The flow rates and energy consumption of the process of Fig. 5 are given in Table 4.

Fo zo Odevuunt 01000 yu v

Ф прдан 0 зв М пеужист 1 " - з - г" Teplodlyavlyasnyetyutev 011 -I b If we accept the operating factor of flows of 340 days per year for the CNG production installation, then the specific energy consumption for the implementation of Fig. 3 will be 0.145 hp-h. /lb (0.238kW/kg), which corresponds to an increase - k.k.d. at 16-26905. c 20 Compared with the embodiments of Fig. 1 and 3, the embodiment of Fig. 5 requires approximately 5-1095 less energy per unit of produced liquid. Compared to the embodiment of Fig. 4, the embodiment of Fig. 5 requires approximately the same amount of energy per unit s of produced liquid. Therefore, with the given available power for compression, the embodiment of Fig. 5 gives approximately 595 more liquefied natural gas compared to the embodiment of Fig. 3 due to obtaining only C and heavier hydrocarbons in the condensed co-product. The choice between the embodiments of Fig. 1, Fig. 3 or Fig. 4 in specific cases is determined by 22 or the cost of heavier hydrocarbons as part of LPG and LNG compared to their corresponding cost in the LPG product,

GF! or the calorific value of CNG (since the calorific value of CNG given by embodiments of Fig. 1, C and 4 is lower than the calorific value of CNG given by embodiment of Fig. 5) o Other embodiments

It is clear that the invention can be adapted for use with all types of LPG liquefaction plants for 60 co-production of LPG, LNG or condensate flows, according to the conditions of the location of the plant. It is also clear that various process variants can be used to produce a liquid co-product stream. For example, the embodiments of Figures 1 and 3 can be adapted to produce an LNG or condensate stream as a liquid co-product instead of an LPG stream as described in Examples 1 and 2. Fig. 4 can be adapted to produce an RNG stream with a high content of Co components of the input gas instead of an LNG co-product, as described in Example 3. The embodiment of Fig. 5 can be adapted to produce an RNG stream with a high content

Co-components of the input gas or LNG with a high content of C3-components of the input gas instead of a co-product

LNG as described in example 4.

Fig. 1, 3, 4 and 5 illustrate preferred embodiments of the invention according to the specified conditions Fig. 6-21 illustrate other embodiments of the invention that may be suitable for other applications. As shown in Fig. b, 7, all the Condensed liquid or its part (flow 33) from the separator 11 can enter the rectification column 19 at separate entry points in the lower part of the middle of the columns without combining with the vapor part of the separator (flow 34), which passes to the heat exchanger 13. Fig. 8 illustrates another embodiment of the invention, which requires less equipment than the embodiment of Fig. 1, b, although the specific energy consumption is slightly higher. Similarly, Fig. 9 illustrates another embodiment of the invention that requires less equipment than the embodiment of Figs. 3, 7, but again at the expense of greater energy consumption Figs. 10-14 illustrate another embodiment of the invention that may require less equipment , than the embodiment of Fig. 4, but with higher energy consumption (it should be noted that, as shown in Fig. 10-14, distillation columns or systems such as deethanizer 19 include elements of both a reboiler absorption column and a dephlegmation reboiler column). Fig. 15, 16 illustrate other embodiments of the invention that combine the functions of the separator/absorption column 18 and the deethanizer 19 of the embodiments of Figs. 4 and 10 - 14 in a single rectification column 19. Depending on the amount of heavier hydrocarbons in the input gas and the pressure of the input gas, the cooled input stream Z1a , coming out of the heat exchanger 10, may not have liquid (since it is in a state above the dew point above or above the cryocondensator), as a result of which the need for the separator 11 disappears (Fig. 1, 3-16) and the cooled input stream Z31ia can pass directly to a suitable expansion device, for example, to the working expansion machine 15.

Treatment of the gas stream after removal of the liquid co-product stream (stream 37 in Fig. 1, 3, 6-11, 13, 14, stream 47 in Fig. 4, 12, 15, 16 and stream 43 in Fig. 5) before entering the heat exchanger 6bО for condensation and supercooling can be done in many ways. In the processes of Fig. 1, 3-16, the flow is heated, compressed to a higher pressure using energy from one or more working expansion machines, partially cooled in the outlet cooler and then further cooled through heat exchange with the original flow. As shown in Fig.17, in some cases it is advisable to compress the flow to a higher pressure, for example, with an auxiliary compressor 59, which operates from an external energy source. As shown by dashed lines i), the use of certain equipment (heat exchanger 24 and outlet cooler 25 in Fig. 1, 3-16) can help reduce capital costs by reducing or eliminating pre-cooling of the compressed flow before feeding it to the heat exchanger 60 (by increasing the cooling (reducing the load on the heat exchanger 60 and increasing energy consumption by compressors 64, 66, 68 of the refrigerant). In such cases, flow 49a from the compressor can directly enter the heat exchanger 24 (Fig. 18) or into the heat exchanger b6O (Fig. 193. If for the expansion of any part of the high-pressure input gas not M, working expansion machines are used, instead of the compressor 16, it can a compressor may be used, for example, compressor 59 (Fig. 20), which is driven by an external energy source. Other circumstances may not justify compressing the flow at all, and then this flow passes directly into the heat exchanger 60 (Fig. 21) and through the equipment marked with dashed lines (heat exchanger 24, compressor 16 and outlet cooler 25 Fig. 1 and C - 16). If the heat exchanger 24 is not provided for heating the flow 48 before the removal of the fuel gas of the installation, there may be a need for an auxiliary heater 58 for heating the fuel gas before consumption using the auxiliary flow or another process flow "to provide the necessary heat (Fig. 19-21). In general, such a choice must be made in each case, taking into account such factors as the composition of the gas, the size of the installation, the desired level of formation of the co-product flow, and the available equipment. ;" According to the invention, cooling of the inlet gas flow and the inlet flow of the LPG production section can be carried out in many ways. In the processes of Fig. 1, 3, 6-9, the incoming flow of gas 31 is cooled and condensed

The external flows of the refrigerant and the lower liquids from the rectification column 19. In Fig. 4, 5, 10-14, for this purpose, rapidly evaporated liquids of the separator are used together with the external flows of the refrigerant. In Fig. 15, 16, lower liquids and evaporated liquids of the separator together with external flows are used for this

Me, refrigerant. In Fig. 17-21, only the external flows -I of the refrigerant are used to cool the incoming gas stream Z1. However, the cold streams of the process can also be used to cool the refrigerant at high pressure (stream 7Ta) (Fig. 4, 5, 10, 11). In addition, any stream with a temperature lower than the temperature of the stream to be cooled can be used. For example, a part of the steam from the separation/absorption column 18 or the rectification column 19 can be diverted and used for cooling.

The use and distribution of column fluids and/or vapors for heat exchange and the location of heat exchangers, as well as the selection of streams for a particular heat exchange, must be evaluated for each specific application. The choice of cooling source depends on many factors including (but not limited to) the composition and condition of the inlet gas, the size of the plant, the size of the heat exchanger, the potential temperature of the cooling source, etc. It is clear that for

F) to achieve the desired temperatures of the inlet stream, any combination of the specified cooling sources and cooling methods can be used.

In addition, auxiliary external cooling for the inlet gas flow and the inlet flow of the LPG production section can be implemented in many ways. In Fig. 1, 3-21, the boiling single-component refrigerant is used for external cooling of a high level and the evaporation of a multi-component refrigerant for external cooling of the lower level, and the single-component refrigerant is used for pre-cooling the flow of a multi-component refrigerant. In another variant, both low- and high-level cooling can be provided using single-component 65 refrigerants with successively decreasing boiling points ("cascade cooling") or single-component refrigerants under successively reduced pressures. Another option involves the use of flows of a multicomponent refrigerant, the composition of which is adjusted to provide the desired cooling temperatures of both high and low levels. The choice of external cooling method depends on many factors, including (but not limited to) the composition and condition of the inlet gas, the size of the installation, the size of the discharge compressor, the size of the heat exchanger, the external temperature of the heat sink, etc. It is clear that any combination of the specified cooling sources and cooling methods can be used to achieve the desired inlet flow temperatures.

Subcooling of the flow of condensed liquid leaving the heat exchanger 60 (flow 49 in Fig. 1, 6, 8 flow 494 in Fig. 3, 4, 7, 9-16 flow 496 in Fig. 5, 19, 20 flow 49e in Fig. 17 flow 49c in Fig. 18 and flow 49a 7/o in Fig. 21) reduces the amount of vapors that can quickly form during the expansion of the flow to the working pressure of the LPG storage tank 62. This reduces the total specific energy consumption when obtaining CNG due to the elimination of the need for rapid gas compression. However, circumstances may favor lower capital costs by reducing the size of the heat exchanger 60 and using rapid gas compression to remove any vaporization gas that may occur.

Although it has been shown that the expansion of individual streams is carried out by separate devices, other means of expansion can be used if necessary. For example, conditions may justify operational expansion of a substantially condensed input flow (flow ZbBa in Fig. 1, 3, 6, 7) or flow of intermediate pressure phlegm (flow 39 in Fig. 1, 6, 8). In addition, instead of the operational expansion of the flow of supercooled liquid leaving the heat exchanger 60 (flow 49 in Fig. 1, 6, 8, flow 494 in Fig. 3, 4, 7, 9-16, flow 49r in Fig. 5, 19, 20, stream 49e of FIG. 17, stream 49c of FIG. 18, and stream 49a of FIG. 21), isenthalpic rapid expansion may be used, but this will necessitate or greater subcooling in the heat exchanger 60 to prevent rapid expansion evaporation. , or the addition of rapid vapor compression or other means of removing rapidly generated vapors. Similarly, isenthalpic rapid expansion can be used instead of work expansion for a supercooled, HIGH pressure refrigerant stream EXITING heat exchanger 60 (stream 71c in Fig. 1, 3-21), with an accompanying increase in refrigerant compression energy consumption.

The above description of the preferred embodiments will enable any person skilled in the art to use the invention by making appropriate modifications and changes in accordance with the concepts and principles of the invention. The scope of the invention is not limited by the given embodiments and is determined by the given new principles and features defined by the Formula about the invention.

Claims (1)

The formula of the invention - (Ce) 35 1. The method of liquefaction of a flow of natural gas containing methane and heavier hydrocarbon components, in which: (a) the specified flow of natural gas is cooled under pressure to condense at least part of it and form a condensed flow and (B) ) the specified condensed flow is expanded to a lower pressure to form the specified flow of "liquefied natural gas, and w-v s (1) the specified flow of natural gas is processed in one or more stages of cooling for its partial condensation, :z" (2) the flow of the specified partially the condensed gas is separated to obtain a vapor stream and a liquid stream, (3) said vapor stream is divided into at least a first gas stream and a second gas stream, -AND (4) said first gas stream is cooled to substantially complete condensation and then expanded to an intermediate pressure, (22) (5) the specified second gas flow is expanded to the specified intermediate pressure, -And (6) the specified liquid flow is expanded to the specified intermediate pressure, (7) the specified expanded significantly condensed first stream, the specified expanded second gas 1 stream and the expanded liquid stream are introduced into the distillation column, where the specified streams are separated into a volatile dry gas fraction containing the main part of the specified methane and more light components, and a relatively less volatile fraction containing the main part of the specified heavier hydrocarbon components, (8) the specified volatile fraction of dry gas is cooled under pressure to condense at least part of it and obtain the specified condensed stream. 2. The method of liquefaction of a flow of natural gas containing methane and heavier hydrocarbon components, u (F. to which: GI (a) the specified flow of natural gas is cooled under pressure to condense at least part of it and the formation of a condensed flow and in (B) the specified condensed flow is expanded to a lower pressure to form the specified flow of liquefied natural gas, and (1) the specified the flow of natural gas is processed in one or more stages of cooling for its partial condensation, (2) the flow of the specified partially condensed gas is separated to obtain a vapor stream and a liquid b/Stream, (3) the specified vapor stream is divided into at least the first gas stream and the second gas stream , (4) said first gas stream is combined with at least a portion of said liquid stream to form a combined stream, (5) said combined stream is cooled to substantially complete condensation and then expanded to an intermediate pressure, (6) said second gas stream the stream is expanded to said intermediate pressure, (7) any residue of said liquid stream is expanded to said intermediate pressure, (8) said expanded substantially condensed combined stream, said expanded second gas stream and said residue of said liquid stream are introduced into a distillation column , where said streams 7/0 are separated into a volatile fraction of dry gas containing the main part of said methane and lighter components, and a relatively less volatile fraction containing a main part of said heavier hydrocarbon components, (9) said volatile fraction of dry gas cooled under pressure to condense at least part of it and obtain the specified condenso bathy stream C. A method of liquefying a natural gas stream containing methane and heavier hydrocarbon components, in which: (a) said natural gas stream is cooled under pressure to condense at least a portion thereof and form a condensed stream and (B) said condensed stream is expanded to a lower pressure to form the specified flow of 2o Liquefied natural gas, and (1) the specified flow of natural gas is processed in one or more stages of cooling for its partial condensation, (2) the flow of the specified partially condensed natural gas is separated to obtain a vapor flow and a first liquid flow, c ( 3) the specified vapor stream is divided into at least the first gas stream and the second gas stream, (4) the specified first gas stream is cooled to substantially complete condensation and then expanded to i) intermediate pressure, (5) the specified second gas stream is expanded to the specified intermediate pressure , (6) the specified first liquid flow is expanded to the specified of intermediate pressure, o z o (7) the specified expanded substantially condensed first gas stream, the specified expanded second gas stream and the expanded first liquid stream are introduced into a distillation column, where the specified streams o are separated into a volatile fraction of dry gas containing the main part of the specified methane and more light components, and on a relatively less volatile fraction containing the main part of the indicated heavier hydrocarbon components, d) (8) the indicated volatile fraction of the dry gas is cooled under pressure to condense at least part of it, (9) the indicated condensed portion split into at least two parts to form the said condensed stream and the second liquid stream and (10) said second liquid stream is introduced into said distillation column as the upper input stream. " 4. The method of liquefaction of a flow of natural gas containing methane and heavier hydrocarbon components, in which: (a) the specified flow of natural gas is cooled under pressure to condense at least part of it and;" formation of a condensed stream and (B) said condensed stream is expanded to a lower pressure to form said stream of Liquefied Natural Gas, and (1) said stream of natural gas is processed in one or more stages of cooling to partially condense it, Me. (2) the flow of the specified partially condensed natural gas is separated to obtain a vapor stream and a first liquid stream, (3) the specified vapor stream is divided into at least a first gas stream and a second gas stream, o (4) the specified first gas stream is combined with at least a portion of said first liquid stream o to form a combined stream, (5) said combined stream is cooled to substantially complete condensation and then expanded to an intermediate pressure, (6) said second gas stream is expanded to said intermediate pressure, (7 ) any residue of the first liquid stream is expanded to the specified intermediate pressure, Ф) (8) the specified expanded significantly condensed combined stream, the specified expanded second gas stream and the specified residue of the first liquid stream are introduced into the distillation column, where the specified streams are divided into volatile fraction of dry gas, containing the main part of the specified methane and lighter components, i.e. essentially a less volatile fraction containing a major portion of said heavier hydrocarbon components, (9) said volatile fraction of the dry gas is cooled under pressure to condense at least a portion thereof, (10) said condensed portion is separated into at least two portions to form thereby 65 said condensed stream and the second liquid stream and (11) said second liquid stream is introduced into said distillation column as the upper input stream. 5. A method of liquefying a stream of natural gas containing methane and heavier hydrocarbon components, in which: (a) said stream of natural gas is cooled under pressure to condense at least part of it and form a condensed stream and (B) said condensed stream is expanded to a lower pressure to form the specified stream of liquefied natural gas, and (1) the specified stream of natural gas is processed in one or more cooling stages for its partial condensation, 70 (2) the specified stream of partially condensed natural gas is separated to obtain a vapor stream and a first liquid stream, (3 ) the specified vapor stream is expanded to an intermediate pressure, after which it is introduced into the contact device, resulting in a volatile dry gas fraction containing the main part of the specified methane and lighter components, and a second liquid stream, (4) the specified first liquid stream is expanded to the specified intermediate pressure, (5) specified second liquid flow and zazna said expanded first liquid stream is introduced into a distillation column where said streams are separated into a distillation stream of more volatile vapors and a relatively less volatile fraction containing the bulk of said heavier hydrocarbon components, (6) said distillation stream of more volatile vapors is cooled to a level, sufficient to condense at least a portion thereof, thereby obtaining a third liquid stream, (7) at least a portion of said expanded vapor stream is brought into intimate contact with at least a portion of said third liquid stream in said contact device, and (8) said volatile dry gas fraction is cooled under pressure , to condense at least part of it and obtain the specified condensed flow. with 6. A method of liquefying a stream of natural gas containing methane and heavier hydrocarbon components, in which: (a) said stream of natural gas is cooled under pressure to condense at least part of it and form a condensed stream and (B) said condensed stream is expanded to a lower of pressure to form the specified flow of liquefied natural gas, and (1) the specified flow of natural gas is processed in one or more stages of cooling for its partial condensation, M (2) the specified flow of partially condensed natural gas is separated to obtain the first vapor flow and the first of a liquid stream, ise) (3) the specified first vapor stream is expanded to an intermediate pressure, after which it is introduced into the contact device, resulting in a second vapor stream and a second liquid stream, (4) the specified first liquid stream is expanded to a specified intermediate pressure, (5) the specified second liquid stream and the specified expanded first liquid stream are introduced into d distillation column, where the specified streams are divided into a distillation stream of more volatile vapors and a relatively less volatile fraction containing the main part of the specified heavier hydrocarbon components, from c (6) the specified distillation stream of more volatile vapors is cooled to a level sufficient for condensation of at least its parts, and thus receive a third vapor and a third liquid flow, ;" (7) a portion of said third liquid stream is introduced into said distillation column as an upper inlet stream, 8) at least a portion of said first expanded vapor stream is brought into intimate contact with -And at least a portion of the remainder of said third liquid stream in said contact device, (9) said the second steam stream is combined with the third steam stream to obtain a volatile fraction Me. dry gas containing the main part of the specified methane and lighter components, and -And (10) the specified volatile fraction of the dry gas is cooled under pressure to condense at least part of it and obtain the specified condensed stream. o 7. A method of liquefying a flow of natural gas containing methane and heavier hydrocarbon components, in which: (a) the specified flow of natural gas is cooled under pressure to condense at least part of it and form a condensed flow and 5B (B) the specified condensed flow is expanded to a lower pressure to form the specified stream of liquefied natural gas, and F) (1) said stream of natural gas is processed in one or more stages of cooling for its partial condensation, (2) said stream of partially condensed natural gas is separated to obtain a vapor stream and of the first liquid stream, (3) the specified vapor stream is expanded to an intermediate pressure, after which it is introduced into the contact device, resulting in a volatile fraction of dry gas containing the main part of the specified methane and lighter components, and the second liquid stream, (4) the specified second the liquid stream is heated, 65 (5) the specified first liquid stream is expanded to the specified about intermediate pressure, (6) said second liquid stream and said expanded first liquid stream are introduced into a distillation column, where said streams are separated into a distillation stream of more volatile vapors and a relatively less volatile fraction containing the main part of said heavier hydrocarbon components, (7 ) said distillate stream of more volatile vapors is cooled to a level sufficient to condense at least a portion thereof to form a third liquid stream, (8) at least a portion of said expanded vapor stream is brought into intimate contact with at least a portion of said third liquid stream in said contact device and (9) the specified volatile fraction of the dry gas is cooled under pressure to condense at least part of it and obtain the specified condensed stream. 70 8. A method of liquefying a stream of natural gas containing methane and heavier hydrocarbon components, in which: (a) said stream of natural gas is cooled under pressure to condense at least part of it and form a condensed stream and (B) said condensed stream is expanded to a lower of pressure to form said stream of 7/5 Liquefied natural gas, wherein (1) said stream of natural gas is processed in one or more stages of cooling to partially condense it, (2) said stream of partially condensed natural gas is separated to obtain a first vapor stream and a first liquid stream flow, (3) said first vapor stream is expanded to an intermediate pressure, then injected into the contact device, resulting in a second vapor stream and a second liquid stream, (4) said second liquid stream is heated, (5) said first liquid stream is expanded to of specified intermediate pressure, (6) specified heated second liquid stream and specified and the expanded first liquid stream is introduced into a distillation column, where said streams are separated into a distillation stream of more volatile vapors and a relatively less volatile fraction containing the main part of said heavier hydrocarbon components, and) (7) said distillation stream of more volatile vapors is cooled to a level sufficient to condense at least a portion thereof, thereby producing a third vapor stream and a third liquid stream, (8) a portion of said third liquid stream is introduced into said distillation column as an overhead feed stream, (9) at least a portion of said expanded first vapor stream introduced into close contact with at least a part of the remainder of the specified third liquid stream in the specified contact device and (10) the specified second vapor stream is combined with the specified third vapor stream to obtain a volatile fraction of dry gas containing the main part of the specified methane and more light components, and ise) (11) the specified volatile fraction dry gas is cooled under pressure to condense at least part of it and obtain the specified condensed flow. 9. A method of liquefaction of a stream of natural gas containing methane and heavier hydrocarbon components, in which: (a) the specified stream of natural gas is cooled under pressure to condense at least part of it and "the formation of a condensed stream and in c (B) the specified condensed stream is expanded to a lower pressure for the formation of the specified flow Х of liquefied natural gas, and у? (1) said stream of natural gas is processed in one or more stages of cooling for its partial condensation, (2) said stream of partially condensed natural gas is separated to obtain a first vapor-I stream and a first liquid stream, (3) said first vapor stream and said the first liquid stream is expanded to an intermediate pressure, Me, (4) the specified expanded first vapor stream and the expanded first liquid stream are introduced into the inlet point -I in the middle of the distillation column, where these streams are separated into a distillation stream of more volatile vapors and a 5o Relatively less volatile fraction , containing the main part of said heavier hydrocarbon 1 components, o (5) said distillation vapor stream is taken from the zone of said distillation column located below said expanded first vapor stream, and cooled to a level sufficient to condense at least a portion thereof, and thereby the second vapor stream and the second liquid stream, 5B (6) each less than a part of the specified expanded first vapor stream is brought into close contact with at least a part of the specified second liquid stream in the specified distillation column, Ф) (7) the specified second vapor stream is combined with the specified distillation stream of more volatile vapors to obtain a volatile dry gas fraction, containing the main part of the specified methane and lighter components, and 60 (8) the specified volatile fraction of the dry gas is cooled under pressure to condense at least part of it and obtain the specified condensed stream. 10. A method of liquefying a stream of natural gas containing methane and heavier hydrocarbon components, in which: (a) said stream of natural gas is cooled under pressure to condense at least part of it and form a condensed stream and (B) said condensed stream is expanded to a lower pressure to form the specified flow of liquefied natural gas, and (1) the specified flow of natural gas is processed in one or more stages of cooling for its partial condensation, (2) the specified flow of partially condensed natural gas is separated to obtain the first vapor flow and the first liquid flow, ( 3) said first vapor stream and said first liquid stream are expanded to an intermediate pressure, (4) said expanded first vapor stream and expanded first liquid stream are introduced into an inlet point in the middle of the distillation column, where these streams are separated into a distillation stream of more volatile vapors and 7 /0 Relatively less volatile fraction containing the main part of said heavier hydrocarbon components, (5) said distillation vapor stream is withdrawn from the zone of said distillation column downstream of said first vapor stream and cooled to a level sufficient to condense at least a portion thereof, thereby producing a second vapor stream and a second liquid stream . at least part of the remainder of the specified second liquid stream in the specified distillation column, (8) the specified second vapor stream is combined with the specified distillation stream of more volatile vapors to obtain a volatile dry gas fraction containing the main part of the specified methane and lighter components, and ( 9) dry the specified volatile fraction gas is cooled under pressure to condense at least part of it and obtain the specified condensed flow. 11. The method of liquefaction of a flow of natural gas containing methane and heavier hydrocarbon components in the YAKOMU system: (a) the indicated flow of natural gas is cooled under pressure to condense at least part of it and i) the formation of a condensed flow and (B) the indicated condensed flow expanded to a lower pressure to form the specified flow of liquefied natural gas, and (1) the specified flow of natural gas is processed in one or more stages of cooling for its partial condensation, which) (2) the specified flow of partially condensed natural gas is separated to obtain the first steam of the flow and the first liquid flow, (3) the specified first vapor flow and the specified first liquid flow are expanded to an intermediate pressure, ise) (4) the specified expanded first vapor flow and the expanded first liquid flow are introduced into the inlet point c in the middle of the distillation column, where these streams are separated into a distillation stream of more volatile vapors and a relatively less volatile fraction , containing the main part of the specified heavier hydrocarbon components, (5) the specified distillation vapor stream is taken from the zone of the specified distillation column, located below the specified expanded first vapor stream, and cooled to a level sufficient to condense at least a part of it, and thereby obtain a second vapor stream and a second liquid stream, (6) at least a portion of said expanded first vapor stream is introduced into intimate contact with;" at least part of the specified second liquid stream in the specified distillation column, (7) the specified liquid distillation stream is withdrawn from the specified distillation column at a place located above the sampling zone of the specified distillation vapor stream, after which the specified -I distillation liquid stream is heated and reintroduced into the specified distillation column as another input stream in a place located below the zone of selection of the distillation vapor stream, Me, (8) the specified second vapor stream is combined with the specified distillation stream of more volatile vapors to -I obtain a volatile dry gas fraction containing the main part of the specified methane and lighter components, and (9) the specified volatile fraction of the dry gas is cooled under pressure to condense at least part of it and obtain the specified condensed stream. 12. A method of liquefying a flow of natural gas containing methane and heavier hydrocarbon components, in which: (a) the indicated flow of natural gas is cooled under pressure to condense at least part of it and form a condensed flow and (F, (B) the indicated condensed flow is expanded to a lower pressure to form said stream of liquefied natural gas, wherein (1) said stream of natural gas is processed in one or more cooling stages to partially condense it, (2) said stream of partially condensed natural gas is separated to produce a first vapor stream and a first liquid stream, (3) said first vapor stream and said first liquid stream are expanded to an intermediate pressure, (4) said expanded first vapor stream and expanded first liquid stream are introduced into the inlet point 65 in the middle of the distillation column, where these streams are separated into a distillation stream more of volatile vapors and on a relatively less volatile fraction containing gol a portion of said heavier hydrocarbon components, (5) said distillation vapor stream is withdrawn from the zone of said distillation column located below said expanded first vapor stream and cooled to a level sufficient to condense at least a portion thereof, thereby obtaining a second vapor stream and a second liquid stream, (6) a portion of said second liquid stream is introduced into said distillation column as another input stream at a location substantially within the sampling zone of said distillation vapor stream, (7) at least a portion of said expanded first vapor stream is introduced into a tight contact with at least a portion of the remainder of said second liquid stream in said distillation column, 70 (8) said distillation liquid stream is withdrawn from said distillation column at a location upstream of said distillation vapor stream withdrawal point, whereupon said distillation liquid stream is heated and reintroduced into said distillation column as another inlet stream at a location downstream of the distillation vapor stream sampling zone, (9) said second vapor stream is combined with said distillation stream of more volatile vapors to /5 obtain a volatile dry gas fraction, which contains a major portion of said methane and lighter components, and (10) said volatile dry gas fraction is cooled under pressure to condense at least a portion thereof and obtain said condensed stream. 13. The method according to any one of claims 1, 2, 5 - 12, which is characterized by the fact that the specified volatile fraction of dry gas 2o is compressed and then cooled under pressure to condense at least part of it and obtain the specified condensed stream. 14. The method according to any one of claims 1, 2, 5 - 12, which is characterized in that the specified volatile fraction of dry gas is heated, compressed and then cooled under pressure to condense at least part of it and obtain the specified condensed stream. with 15. The method according to paragraphs C or 4, characterized in that (1) said volatile dry gas fraction is heated, compressed and then cooled under pressure to i) condense at least a portion thereof, and (2) said condensed portion is split into at least two portions to produce said condensed stream and the specified second liquid stream. о зо 16. The method according to any of claims 1 - 15, which is characterized by the fact that the indicated volatile fraction of dry gas contains the main part of the indicated methane, lighter components and Co-components. at 17. The method according to any one of claims 1 - 15, which is characterized by the fact that the specified volatile fraction of dry gas contains M the main part of the specified methane, lighter components, Co-components and C3-components. 18. The method according to claim 13, which is characterized by the fact that the specified volatile fraction of dry gas contains the main part of the specified methane, lighter components and Co-components. uh- 19. The method according to claim 14, which is characterized by the fact that the specified volatile fraction of dry gas contains the main part of the specified methane, lighter components and Co-components. 20. The method according to claim 15, which is characterized by the fact that the specified volatile fraction of dry gas contains the main part of the specified methane, lighter components and Co-components. " 21. The method according to claim 13, which is characterized by the fact that the indicated volatile fraction of dry gas contains the main part of the mentioned methane, lighter components, Co-components and C3-components. 22. The method according to claim 14, which is characterized by the fact that the indicated volatile fraction of dry gas contains the main part of the indicated methane, lighter components, Co-components and C3-components. 23. The method according to claim 15, which is characterized by the fact that the specified volatile fraction of dry gas contains the main part of the specified methane, lighter components, Co-components and C3-components. -I (22) -I c 50 (42) F) ime) 60 b5