Natural gas liquefaction

Liquefaction of natural gas

Full description

Background of the invention:

This invention relates to a process for treating natural gas or other gas streams rich in methane to produce liquefied natural gas (LNG) containing high-purity methane and a liquid stream containing predominantly heavier hydrocarbons than methane. The applicants for this application claim the benefits mentioned in Title Deed No. 35 of the US Code of Law, Paragraph 119 (e) of the US Provisional Patent Application No. 60/296,848 filed on June 8, 2001.

Natural gas is usually extracted from wells drilled in underground reservoirs. It usually contains a large percentage of methane, meaning methane constitutes at least 50 molar percent of the gas. Depending on the specific underground reservoir, natural gas also contains relatively smaller amounts of heavier hydrocarbons such as ethane, propane, butanes, pentanes, and the like, as well as water, hydrogen, nitrogen, carbon dioxide, and gases. Other.

Most natural gas is used in the gaseous state. The most widely used means of transporting natural gas from the primary source to gas processing units and then to natural gas consumers is through high-pressure gas transmission pipelines. However, in several cases it has been found necessary and/or desirable to liquefy natural gas either for transportation or use. In remote places, for example, there is rarely a pipeline infrastructure that allows for convenient transportation of natural gas to the market. In such cases, the very low specific volume of LNG compared to normal gas can reduce transportation costs by enabling transportation of LNG using cargo ships and tanker trucks.

In another case, it is preferable to liquefy natural gas for use as vehicle fuel. In major metropolitan areas, there are convoys of buses, taxis and trucks that can...

Operated by liquefied natural gas if an economical source of liquefied natural gas is available. Such vehicles fueled with liquefied natural gas cause much less air pollution due to the clean combustion property of natural gas compared to similar vehicles powered by gasoline and diesel engines that burn higher molecular weight hydrocarbons. In addition, if natural gas has a high purity (i.e., methane purity of 9 molar percent or higher), the amount of carbon dioxide produced (greenhouse gas) is much lower due to the lower carbon-to-hydrogen ratio of methane compared to all other hydrocarbon fuels.

General description of the invention

The present invention relates generally to the liquefaction of natural gas with the production of a liquid stream containing mainly hydrocarbons heavier than methane in the form of a co-product such as natural gas liquids (NGL) consisting of ethane, propane, butane compounds and lesser hydrocarbon components, liquefied petroleum gas (LPG). ) consisting of propane, butane compounds and heavier hydrocarbon components or a condensate consisting of butane compounds and heavier hydrocarbon components. The process of producing a co-product in the form of a liquid stream has two important benefits, which are the high methane purity of the resulting liquefied natural gas and the value that characterizes the liquid co-product, as it can be used for many other purposes. A typical analysis of a natural gas stream to be processed in accordance with this invention is an approximate molar percentage of 84.2% for methane, 7.9% for ethane and other C2 components, 4.9% for propane and other C3 components, and 1.0% for isobutane. iso-butane, 1.1% for regular butane, 0.8% for pentane compounds and higher molecular weight compounds, with nitrogen and carbon dioxide representing the remaining percentage. There are also sometimes gases containing sulfur.

There are many known ways to liquefy natural gas. See, for example, what was reported by Adrian J. Finn, Grant L.Johnson Adrian, and Terry R. Tomlinson in a paper entitled “Liquefied Natural Gas Technology for Offshore and Medium-Scale Industrial Units” presented in lectures. Sessions of the 79th Annual Conference of the Gas Processors Association, pp. 429-0H4, Atlanta, Georgia, March 13-15, 2000. What came from Kikkawa, Yoshitsugi, Masaki

Masaaki Ohishi and Noriyoshi Nozawa in a paper entitled "Optimization of the electrical power network for a base-loaded natural gas unit", presented in the minutes of the 80th Annual Conference of the Gas Processors Association, San Antonio, Texas, in the period between March 12 and 14, 2001, to review a number of Of these operations. US Patent Nos. 5,445,917 also describe; 185, 4,525; 4,545,795; 4,755,200;5,291,736; 5,363,655; 5,365,740; 5,600,969;

٦١٥,٥٦١,5؛ ٥,٦٥١,٢٦٩؛٥,٧٥٥,١١٤؛٥,٨٩٣,٢٧٤؛ 6,14,869؛ 6,062,041

6,119,479; 6,125,653; 6,250,105 1B; 6,269,655 1B; 6,272,882 1B; 6,308,531 1B; 6,324,867 1B and 6,347,532 1B are closely related operations. These methods generally include steps in which natural gas is purified (by removing water and troublesome compounds such as carbon dioxide and sulfur compounds), cooled, condensed, and expanded. The cooling and condensation steps of natural gas can be accomplished in several different ways. The cascade cooling method uses the heat exchange of natural gas with several refrigerants that have successively lower boiling points, such as propane, ethane, and methane. Instead of that ; This heat exchange can be accomplished using a single refrigerant that is evaporated at several different pressures. The multi-component refrigeration method uses the heat exchange of natural gas with one or more refrigerants consisting of several refrigerants instead of several refrigerants with a single component. Expansion of natural gas can be achieved by enthalpy stability (eg, using Joule-Thomson expansion) and inertia stability (eg, using an operational expansion turbine).

Regardless of the method used to liquefy a natural gas stream, a significant portion of the hydrocarbons heavier than methane usually needs to be removed before the methane-rich stream can be liquefied. The reasons for this hydrocarbon removal step are many, including the need to control the calorific value of the LNG stream and the value of these heavier hydrocarbon components as products themselves. Unfortunately, to date, little attention has been paid to the effectiveness of the hydrocarbon removal step.

According to the present invention, it has been found that liquefied natural gas and a separate heavy hydrocarbon liquid product can be produced using much less energy than that used in prior art processes, by precisely including a hydrocarbon removal step in the natural gas liquefaction process. Although the present invention is suitable for application at low pressures, it is particularly useful for processing feed gases at absolute pressures ranging from 2758 to 10342 kPa (absolute) [1400 to 1500 psi] or higher.

Brief explanation of fees

To better understand the present invention, reference is made to the following examples and drawings:

Figure [1]: represents a diagram of the flow of operations in a natural gas liquefaction unit adapted to produce natural gas liquids in a contributory manner according to the present invention;

Figure [2]: represents a diagram of the pressures at which phase transformations occur for the methane used

As a function of enthalpy to illustrate the superiority of the present invention over prior art processes;

Figure [3]: represents a diagram of the process flow in a prepared alternative natural gas liquefaction unit

To jointly produce natural gas liquids according to the present invention;

Figure [4]: ​​represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce liquefied petroleum gas in a contributory manner according to the present invention;

Figure [5]: represents a diagram of the flow of operations in an alternative natural gas liquefaction unit

Adapted to produce co-condensate according to the present invention;

Figure [6]: represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [7]: represents a diagram of the process flow in a prepared alternative natural gas liquefaction unit

To contribute to producing a liquid stream according to the present invention;

Figure [8]: represents a schematic diagram of the operations in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [9]: represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [10]: represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [11]: represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [12]: represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [13]: represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [14]: represents a schematic diagram of the process flow in an alternative natural gas liquefaction unit configured to produce a liquid stream in a contributory manner according to the present invention;

Figure [15]: represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [16]: represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [17]: represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [18]: represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [19]: represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [20]: represents a diagram of the process flow in an alternative natural gas liquefaction unit

Adapted to produce a liquid stream in a contributory manner according to the present invention;

Figure [21]: represents a diagram of the process flow in a prepared alternative natural gas liquefaction unit

To contribute a liquid stream according to the present invention.

Detailed description:

In the following illustration of the figures above, tables are provided summarizing flow rates calculated for typical treatment process conditions. In the tables shown herein, flow rate values ​​(in moles per hour) have been rounded to the nearest integer for convenience. The total current rates shown in the tables include all non-hydrocarbon components and are therefore generally greater than the sum of the current flow equations for the hydrocarbon components. The indicated temperatures are approximate values, rounded to the nearest degree. It should also be noted that the process design calculations conducted for the purpose of comparing the processes depicted in the figures are based on the assumption that there is no heat leakage from the surrounding media into the treatment process or vice versa. The quality of commercially available insulating materials makes this assumption very reasonable and can usually be achieved by those skilled in the technique.

For ease, process variables are shown in traditional British units and in the International System of Units (SI). The molecular flow rates given in the tables can be expressed in pounds moles per hour or kilograms moles per hour. Energy consumption rates expressed in horsepower and/or thousand British thermal units per hour are similar to molecular flow rates expressed in pounds moles per hour. Consumption rates are similar. Energy expressed in kilowatts Molecular flow rates expressed in kilograms moles per hour. Production rates expressed in pounds per hour are similar to molecular flow rates expressed in pounds per hour. Production rates expressed in kilograms per hour are equivalent to molecular flow rates expressed in kilograms per hour.

Description of the invention

Example (1)

Referring now to Figure [1]; A process is illustrated in accordance with the present invention whereby it is desired to produce an NGL co-product containing a greater amount of ethane and heavier components in the natural gas feed stream. In this simulation of the process according to the present invention, the entry gas enters the unit at a temperature of 90°F [32°C] and an absolute pressure of 1285 psi [8860 kPa (absolute)] as a stream (31). If

If the inlet gas contains a concentration of carbon dioxide and/or sulfur compounds that may prevent the produced streams from meeting the required specifications, these compounds are removed by appropriate pretreatment of the feed gas (not shown). It also usually removes water from the feed stream to avoid the formation of water (snow) in rural conditions. A solid dryer is usually used for this purpose.

The feed stream (31) in the heat exchanger (10) is cooled by heat exchange with the cooling streams and demethanizer side reboiler fluids at (-68°F (minus 0°F)) (-55°C (minus °C)) (stream 40 ). It is noted that in both cases, the heat exchanger (10) is represented by either several single heat exchangers or a single multi-pass heat exchanger or any combination thereof. (The decision to use more than one heat exchanger for the aforementioned cooling services depends on several factors, including, but not limited to, the inlet gas flow rate, the size of the heat exchanger, current temperatures, etc.). The refrigerated stream (1 3a) enters a separator (11) at a temperature of -30 F (-34 C] and under an absolute pressure of 1278 psi (8812 kPa (absolute)), where the vapor (stream 32) is separated from the condensed liquid (stream 33). ).

The steam (stream 32) leaving the separator (11) is divided into two streams (34) and (36).

Stream (34), which contains about 20% of the total vapor, is mixed with the condensing liquid, stream (33), to form stream (35). The mixed stream (35) passes through the heat exchanger (13) where it exchanges heat with the refrigerant stream (71A), which causes the stream (35A) to cool and condense substantially. The essentially condensed stream (3A) is then expanded at -120 F [-85 C] ] through a suitable expansion device such as expansion valve (14), to the operating pressure (about 465 psi [absolute] 32.06 kPa (absolute)) of the fractionation tower (19). During expansion, part of the stream evaporates, which leads to cooling of the total stream. In the process shown in Figure [1], the expansion stream (35b) discharged from the expansion valve (34) reaches a temperature of -122 F -86 C) and is supplied to a feed location in the middle of the demethanation section (19b) of the fractionation tower (19).

The remaining amount of steam discharged from separator (11) (80% of the total steam) enters

(Current 36) An operational extension machine (15) where mechanical energy is extracted from this part of the current

High pressure feeding. The machine (15) expands the steam in a substantially constant inertial manner from an absolute pressure ranging from about 1278 psi [8812 kPa (absolute)] to the operating pressure of the tower, so that the operational expansion cools the expanding stream (36A) to a temperature of About -103 F [-75 C]. Typical commercially available expansion tools can recover 80 to 85% of the work theoretically available in a typical inertial expansion. The recovered work is often used to operate a centrifugal compressor (for example, tool 16) that can be used to recompress the upper gas of a tower (stream 38) on for example . The partially expanded and condensed stream (36a) is supplied as a feed stream to a distillation column (19) at a feed point down the middle of the column.

The demethanizer in the fractionation tower (9 1) is a traditional distillation column containing several trays spaced vertically), one or more packed layers, or one or more combinations of trays and packings. As is usually the case in natural gas processing units, the fractionation tower includes two sections: the upper section (19a), which is a separator where the upper feed stream is divided into its respective steam and liquid parts, and where the steam rising from the lower distillation or demethanization section (19b) is mixed. With the upper feed stream steam portion (if present) to form the upper cold steam of the demethanizer (stream 37) which exits the top of the tower at -135 F [-93 C]. The lower demethanization section (19b) contains trays and/or gaskets and ensures the necessary contact between the liquids flowing down and the vapors rising to the top. The demethanization section also includes one or more reboilers (eg reboiler 20) that heat and vaporize a portion of the liquids flowing down the column to supply the extraction vapors flowing up the column. The liquid product stream (41) exits the bottom of the tower at 115 F [46 C], based on a typical specification for a methane to ethane ratio of 1:0.020 on a molecular basis in the bottom product.

The upper steam of the demethanizer (stream 37) is heated to 90 F [32 C] in a heat exchanger (24), and a portion of the heated upper steam of the demethanizer is drawn to act as fuel gas (stream 48) for the unit. (The amount of fuel gas that must be drawn is generally determined by the fuel required for the engines and/or turbines that drive the gas compressors in the unit, such as compressors

Refrigerant 64, 66 and 68 in this example). The remaining part of the upper heated steam of the demethanizer (stream 38) is compressed by the compressor (16) driven by expansion machines (15), (61) and (63) 0 and after cooling to 100 F [38°C] in the vacuum cooler (25), the stream is cooled. (36b)

Further to -123 F [-86 C] in the heat exchanger (24) by cross-exchange with the upper cold steam of the demethanizer (stream 37).

Then the stream (38C) enters the heat exchanger (0 6) and is further cooled by the refrigerant stream (71D) 0. After cooling to an intermediate temperature, the stream (38C) is divided into two parts. The first part, the stream (49), is further cooled in the heat exchanger (60) to -257 F [-160 C] to condense and cool it minimally (below the dew point), where it then enters an operational expansion machine (61) where mechanical energy is extracted from the stream . The machine (61) expands the liquid stream (49) in an essentially inertial manner from a pressure of approximately 562 psi [3878 kPa (absolute)] to the LNG storage pressure (15.5 psi [107 kPa (abs.) (absolute)], which is slightly above atmospheric pressure, and the expansion cools the expanded stream (49A) to a temperature of about -258 F [-161 C], where it is then directed to the LNG tank (62), which contains the LNG product. (Current 50).

The current (39), the other part of the current (38C), is drawn from the heat exchanger (60) at a temperature of -160°F [-1.07°C] and expanded and radiated through a suitable expansion device, such as an expansion valve (17) to the tower operating pressure Retail (19). In the process shown in Figure [1], no evaporation of the expanding stream (39A) occurs, so its temperature is reduced only slightly to -161 F [-107 C] leaving the expansion valve (17). The expanding stream (39A) is then supplied to the separation section ( 19a) in the fold zone of the retail tower (19). The separated fluids inside it go to the demethanization section (19b) as an upper feed stream for it.

Complete cooling of streams (35) and (38C) is provided via a closed-cycle cooling loop.

The original fluid prepared for this cycle is a mixture of hydrocarbons and nitrogen, provided that the composition of the mixture is adjusted if necessary to provide the required cooling fluid temperature while the

Condensation at moderate pressure using the available cooling medium. In this case, assume an action

Condensation using cooling water, such that the refrigerant mixture consists of nitrogen and methane; ethane; Propane and heavier hydrocarbons in the process simulation according to Figure [1]. The composition of the stream is represented by an approximate molar percentage as follows: 7.5% for nitrogen, 41.0% for methane, 41.5% for ethane and 10.0% for propane, so that hydrocarbons are formed. Heavy duty rest of the installation.

The refrigerant stream (71) exits from a vacuum cooler (69) at a temperature of 100 F [38 C] and under an absolute pressure of 607 lb/in2 [185 kPa (absolute)]. It enters the heat exchanger (10) and cools to -31 F [-35 C] and is partially condensed by the expanded and partially heated refrigerant stream (71 F) and by other refrigerant streams. To simulate the process according to Figure [1], assume that these other refrigerant streams are commercial-quality propane refrigerant at different values ​​of temperature and pressure. The partially condensed refrigerant stream (71A) then enters the heat exchanger (13) to further cool it to -4 1 1 F [-81 C] by the partially expanded and heated refrigerant stream (71A), which condenses the refrigerant and cools it minimally. Partial (Current 71B). The refrigerant is further cooled to -257 F [-0 6 1 C] in the heat exchanger (0 6) by the expanding refrigerant stream (71 C). The minimally cooled liquid stream (71d) enters the operational expansion machine (63), where mechanical energy is extracted from the stream when it expands adiabatically by thermal inertia substantially at a pressure ranging from about 586 lb/in2 absolute [4040 kPa (absolute)] to about 34 lb/in2 absolute. inch2 absolute [234 kPa (absolute)]. During expansion, part of the stream evaporates, cooling the total stream to -236 F [-164 C] (current 71 D). The expanding stream (71d) then enters the heat exchangers (60), (31) and (10) again where it provides cooling to the stream (38c), (stream 35) and the refrigerant (streams 71, 71a, 71b) as it evaporates and overheats. .

Overheated refrigerant vapor (current 71g) escapes from the heat exchanger (10)

At a temperature of 93°F [34°C] and pressure to 617 psi [4254 kPa (absolute)] in three stages. Each of the three compression stages (refrigerant compressors 64, 66 and 68) is driven by an additional power source and is attached to a chiller(s).

Vacuum 65, 67 and 69) to remove the heat of compression. The compressed stream (71) discharged from the vacuum cooler (69) returns to the heat exchanger (10) to complete the cycle.

The following table shows a summary of the current flow rates and energy consumption for the process shown in Figure [1].

Table (1)

(Figure [1])

Summary of current flow rates - lbmol/hour [kgmol/hour]

<img file="SA2015B1_D0001.tif" />

Recovery rates in natural gas liquids *natural gas 95.06% ethane

100.00% propane propane

Butane compounds and higher molecular weight compounds 100.00% Production rate 308,147 lbs/hour (308,147 kg/hour]

LNG producer natura gas liquefaction

Production rate 610813 lbs/hour [610813]

Purity* 99.52%

Minimum calorific value 912.3 BTU/standard foot [33.99 MJ/m3]

Ability

Refrigerant compression 103957 horsepower [170904 kW]

Propane compression / 33,815 horsepower [55,591 kW]

Total compression 137772 hp [495 226 kW]

Temperature of the utility facility

Demethanizer reboiler 29,364 thousand BTU/hour [18,969 kW]

* (depending on unrounded flow rates)

The efficiency of LNG production processes is usually compared using the “necessary specific power consumption,” which is the ratio of the total refrigerant compression capacity to the total liquid production rate. Published information on the specific power consumption of prior technology LNG production processes ranges from 0.168 horsepower-hours/pound [0.276 kW-hours/kg] to 0.182 horsepower-hours/pounds [0.300 kW-hours/kg], it is believed It depends on an operating factor of 340 days per year for the liquefied natural gas production unit. Based on this same basis, the specific power consumption according to the embodiment of the present invention in Figure [1] is 0.161 bhp-hour/lb [0.265 kW-h/kg], and this value provides an improvement in efficiency of 4 to 13% compared to Prior technical operations. In addition, it should be noted that the specific power consumption of prior art processes is based on the co-production only of a liquid LPG stream (heavier C3 hydrocarbons) or condensate product (C4 heavier hydrocarbons) at relatively low recovery levels and not a liquid stream Of natural gas liquids (C2 dicarbonate and heavier hydrocarbons) as shown in this example according to the present invention. Prior technology processes require significantly more cooling capacity for the co-production of a natural gas liquids stream rather than an LPG or condensate stream.

There are two main factors to which the improved effectiveness of the present invention is attributed. The first factor can be understood by studying the thermodynamics of the liquefaction process when it is applied

On a high-pressure gas stream such as the one considered in this example. Since the main component of this stream is methane, the thermodynamic properties of methane can be used for the purposes of comparing the liquefaction cycle used in the prior art processes versus the cycle used in the present invention. Figure [2] shows a diagram of the pressures at which phase transformations of methane take place as a function of enthalpy. In most liquefaction cycles

For the former technique, all cooling of the gaseous stream is carried out while the stream is at high pressure (Path A-B), where the stream is subsequently expanded (Path B-C) ​​to the pressure of the LNG storage vessel (a pressure slightly above atmospheric pressure). This expansion step may use an operational expansion machine that is usually capable of recovering theoretically available work in constant expansion - ideally thermal inertia of about 75 to 80%. For the purpose of simplicity, the expansion of the thermal inertia constant is shown completely in Figure [2] for path B-C. However, the drop in enthalpy provided by this operational expansion is quite small, owing to the lines of constant inertia being nearly vertical in the liquid region of the phase diagram.

This can be contrasted with the liquefaction cycle according to the present invention. The gas stream is expanded by operational expansion (Path A'-A'') to medium pressure after partial cooling at high pressure (Path A-A) (again, the expansion of the inertial constant is shown entirely for simplicity). The rest of the cooling takes place at pressure. The medium (path A''-B'), and then extends the stream (path B'-C) to the pressure of the LNG vessel. Because the static inertia lines are inclined less steeply in the vapor region of the phase diagram, a much greater enthalpy reduction is provided than the first operational expansion step (path A'-A&amp;) according to the present invention. Therefore, the total amount of cooling required for the present invention (the sum of Paths A-A' and A''-B') is less than the cooling required for prior art processes (Path A-B), which reduces the cooling (and thus the cooling pressure) required to liquefy the gaseous stream. .

The second factor to which the improved effectiveness of the present invention is attributed is the superior performance of hydrocarbon distillation systems at low operating pressures. The hydrocarbon removal step in most prior art processes is performed at high pressure, usually using a column

Washing uses a cold hydrocarbon liquid as the absorbent stream to remove heavier hydrocarbons from the incoming gaseous stream. Operating the scrubber column at high pressure is not considered very effective, because it leads to the co-absorption of a large portion of methane and ethane from the gas stream, from which the absorbed liquid must be removed and cooled to become part of the LNG product later. In the present invention, the hydrocarbon removal step is performed at an intermediate pressure where vapor-liquid equilibrium is much more desirable resulting in very efficient recovery of the desired heavier hydrocarbons in the co-product from the liquid stream.

Example (2)

If the LNG product specifications would allow a greater amount of the ethane in the feed gas to be recovered in the LNG product, a simpler embodiment of the present invention could be used. Figure [3] shows an alternative such embodiment. The inlet gas conditions and composition taken into account in the process represented in Figure [3] are similar to those in Figure [1]. Accordingly, the process described in Figure [3] can be compared with the embodiment shown in Figure [1].

In a simulation of the process described in Figure [3], the inlet gas cooling, separation and expansion scheme for the NGL recovery section is essentially similar to that used in Figure [1]. The inlet gas enters the unit at 9 0 F [32 C] and 1 285 lb/in2 absolute [8860 kPa (absolute)] in the form of a stream (31) and is cooled in a heat exchanger (0 1) by exchanging

Heat with refrigerant streams and side reboiler fluids for a demethanizer at -35 F

[-37m] (current 40). The cooled stream (31A) enters a separator (11) at -30 F [-34 C] and 1 278 lb/in2 absolute [2 881 kPa (absolute)] where the vapor (stream 32) is separated from the condensed liquid (stream 33).

The steam (stream 32) resulting from the separator (11) is divided into two streams (34) and (36). Stream (34), which contains about 20% of the total vapor, is mixed with the condensed liquid, stream (33), to form stream (35). The mixed stream (35) passes through a heat exchanger (3 1) in a heat exchange relationship with the refrigerant stream (71A), which leads to significant cooling and condensation of the stream (35A). Then the condensing current expands substantially (35A) at -120 F [-85 C].

They proceed through a suitable expansion device, such as an expansion valve (14), to the operating pressure (of 465 psi absolute [approximately 3206 kPa (absolute))) of the splitting tower (19). During expansion, part of the stream evaporates, which leads to cooling of the total stream. In the process shown in Figure [3], the expansion stream (35B) leaving the expansion valve (4 1) reaches a temperature of -122 F [-86 C], and is supplied to the separation section in the upper area of ​​the fractionation tower (19). The separated fluids are the upper feed stream to the demethanization section in the lower zone

For the retail tower (19).

The remaining 80% of the steam produced from the separator (11) (stream 36) enters an operational expansion machine (5 1), where mechanical energy is extracted from this part of the high-pressure feed stream. The machine (15) expands the steam in a constant, essentially inertial manner from a pressure of about 1278 psi [2,881 kPa (absolute)] to the operating pressure of the tower while cooling the expanding stream (36A) by operational expansion to a temperature of -103 F. [-75 m] approx. The partially expanded and condensed stream (36A) is supplied as a feed stream to a distillation column (19) at a feed point in the middle of the column.

The cold upper steam of the demethanizer (stream 37) exits from the top of the fractionation tower (19).

At -123 F [-86 C]. The liquid product stream (41) exits the bottom of the tower at 118 F [48 C], based on the typical specification for a methane to ethane ratio of 0.020:1 on a molecular basis in the bottom product.

The upper steam of the demethanizer (stream 37) is warmed to 90 F [32 C] in a heat exchanger (4 2), and then a portion (stream 48) is calculated to serve as fuel gas for the unit. The remaining part of the heated core steam of the demethanizer (stream 49) is compressed using a compressor (16). After cooling to

100 F [38 C] in a vacuum cooler (25), stream (49B) is further cooled to -112 F [-80 C] in the heat exchanger (24) by cross-exchange with the cold aerosol vapor of the demethanizer, stream (37).

Then a stream (9 4C) enters a heat exchanger (60) and is further cooled by a refrigerant stream (71D) to -257 F [-160 C] to condense and cool it minimally, and then enters the

An operational expansion machine (61) that extracts mechanical energy from the current. The machine (1 6) extends a liquid stream (49 d) in a constant, essentially thermal inertia manner from a pressure of approximately 583 psi (4021 kPa (abs.)) to an LNG storage pressure (15.5 psi (1.07 Kilopascals (absolute)]) which is slightly above atmospheric pressure. The operational expansion cools the expanding stream (49°C) to a temperature of approximately -8°F [-161°C]. , where it is later directed to the liquefied natural gas storage tank (62) 1, which holds the liquefied natural gas product (stream 50).

Similar to the process shown in Figure [1], all cooling for streams (35) and (49C) is provided by a jacketed cooling loop. The composition of the stream used as the working fluid in the process cycle shown in Figure [3], as an approximate molar ratio, is 7.5% nitrogen, 40.0% methane, 42.5% ethane, and 0.0% propane, which constitutes Heavier hydrocarbons remaining percentage. The refrigerant stream (71) leaves vacuum cooler (69) at 100 F [38 C] and 6.07 lb/in2 absolute [4 1.85 kPa (absolute)]. It enters (the heat exchanger (10) and cools to -31 F [-35 C] and is partially condensed by the partially heated expanding refrigerant stream (71 F) and by) other refrigerant streams. For the simulation of the process shown in Figure [3], it was assumed that these other refrigerant streams were commercial-quality propane refrigerant at three different levels of temperature and pressure. Then the partially condensed refrigerant stream (71a) enters the heat exchanger (31) for the purpose of additional cooling to -121 F [-85°C] via the partially heated expanding refrigerant stream (71d), which leads to the condensation and cooling of the refrigerant minimally. Partial (Stream 71B). The refrigerant is further cooled minimally to -257 F [-160 C] in the heat exchanger (60) via the expanding refrigerant stream (71 C). The minimally cooled liquid stream (71 C) enters an operational expansion machine (63) where the mechanical energy is extracted from the stream when it expands adiabatically. Thermal inertia essentially extends from a pressure of about 586 psi [4040 kPa (absolute)] to about 34 psi [absolute]. 234 kPa (absolute)]. During expansion, part of the stream evaporates, which causes the total stream to cool to -263 F [-164 C] (the current

71d). Then the expanding stream (71D) enters again into the heat exchangers (60), (13) and (10), where it provides cooling for the stream (49C), stream (35) and the refrigerant (streams 71, 71A and 71B) when it evaporates and heats up properly. excessive.

The superheated refrigerant vapor (current 71g) leaves the heat exchanger (10) at 93°F [34°C] and is compressed in three stages to 617 lb/in2 absolute [4254 kPa (absolute)]. Each of the three compression stages (refrigerant compressors 64, 66 and 68) is managed by an additional power source and each stage is followed by a cooler (vacuum coolers 65, 67 and 69) to remove the heat of compression. The compressed stream (71) from the vacuum cooler (69) returns to the heat exchanger (10) to complete the cycle.

A summary of the current flow rates and power consumption for the process shown in Figure [3] is shown in the following table.

Table (2)

(Figure [3])

Summary of current flow rates - lbmol/hour [kgmol/hour]

<img file="SA2015B1_D0002.tif" />

Recovery rates in natural gas liquids *

Ethane 87.57%

Propane 97.41%

Butane compounds and higher molecular weight compounds 99.47%

Production rate: 29617 lbs/hour [296175 kg/hour]

Liquefied natural gas product natural gas liquefaction

Production rate 625152 lbs/hour [625152 kg/hour]

Purity. 98.66%

Minimum calorific value 9.7 1 9 Btu/standard ft3 [34.27 MJ/m3] Capacity

Refrigerant compression 96,560 hp (158,743 kW)

Propane compression 34,724 hp (57,086 kW)

Total compression 131,284 hp (215,829 kW)

Temperature of the utility facility

Reboilment stages in demethanizer 22,177 thousand Btu/hour (14,326 kW)

* (based on unrounded flow rates)

Assuming an operating factor of 34.0 days per year for an LNG production unit,

The specific power consumption for the embodiment of the present invention according to Figure (3) amounts to 153 horsepower-hours per pound [0.251 kilowatt-hours per kg]. Compared with the previous technical process, the improvement in effectiveness ranges from 10 to 20% for To embody according to Figure [3]. As previously noted for the embodiment according to Figure [1], this improvement in efficiency is possible using the present invention even if a co-product is produced from natural gas liquids instead of the co-product from liquefied petroleum gas or condensate produced using the prior art processes.

In comparison with the embodiment according to Fig. [1], the embodiment of the present invention according to Fig. [3] requires

To a lower capacity of about 5% per unit of liquid produced. Thus, for a given amount of available compression capacity, an embodiment of Fig. [3] can liquefy approximately 5% more natural gas compared to the embodiment of Fig. [1] by virtue of recovering less of the C2 dicarbonate components and heavier hydrocarbons in the co-product liquids. Natural gas . The choice between the embodiment of Fig. [1] and the embodiment of Fig. [3] according to the present invention for a particular application will usually depend on the monetary value of the heavier hydrocarbons in the NGL product versus their corresponding values ​​in the LNG product or on the specification of the calorific value of the LNG product (since . The calorific value of liquefied natural gas produced by embodiment according to Figure [1] is less than that produced by embodiment according to Figure [3])0

Example (3)

If the LNG product specifications would allow all of the ethane in the feed gas to be recovered in the LNG product or if there is no market for a liquid coproduct containing ethane, an alternative to the present invention such as that shown in Figure [4] may be used to produce a product stream. My contribution is from liquefied petroleum gas. The inlet gas conditions and composition taken into account in the process represented in Fig. [4] are similar to those in Figs. [1] and [3]. Accordingly, the process shown in Fig. [4] can be compared with the embodiments shown in Figs.

[1]and [3].

In the simulation shown in Figure [4], an inlet gas enters the unit at 0 9 F (32 C) and 1 285 lb/pack2 absolute [8860 kPa (in the form of a stream (31)) and is cooled in a heat exchanger (10) by exchanging Heat with refrigerant and liquid streams from a flash separator at -46°F [-43°C] (stream 33A) and the coolant stream (31A) enters

(1 1) At -1 F [-18 C] and 1278 lb/in2 absolute [8812 kPa (absolute)] where the vapor (stream 32) separates from the condensing liquid (stream 33).

The steam (stream 32) resulting from the separator (11) enters an operational expansion machine (15) where - the mechanical energy z this part of the high-pressure feed stream. The machine works

(15) The steam expands adiabatically by thermal inertia substantially from a pressure of about 1278 psi [8812 kPa (absolute)] to a pressure of about 440 psi (3034 kPa (absolute)] (separation/absorption tower operating pressure 8 1) With the expanded stream (32a) cooled by operational extension to a temperature of approximately -81'F [-63°C], the expanded and partially condensed stream (32a) is supplied to the absorption section (18b) in the lower area of ​​the tower.

Separation/absorption (18). The liquid portion of the expanding stream mixes with the liquids falling downward from the absorption section, and the mixed liquid stream (40) emerges from the bottom of the separation/absorption tower (18) at -86 F [-66 C] - and the vapor portion rises The stream expands upward through the absorption section and comes into contact with the cold liquid falling downward to condense and absorb the C3 and heavier components.

The separation/absorption tower (18) is a conventional distillation column containing several vertically spaced trays, one or more packed layers, or a specific combination of trays and packings. As is often the case in natural gas processing units, the separation/absorption tower can consist of two sections. The upper section (18a) is a separator, which separates any vapor present in the upper feed stream from its corresponding liquid portion, and where the steam rising from the lower distillation or absorption section (18b) mixes with the vapor portion (if present) of the upper feed stream to form the distillation stream. The cold (37) that emerges from the top of the tower. The lower absorption section (18b) contains trays and/or gaskets and provides the necessary contact between falling liquids and upward vapors for condensation and absorption of C3 and heavier components.

The mixed liquid stream (40) is directed from the bottom of the separation/absorption tower (18) to a heat exchanger (13) via a pump (26), where it is heated (stream 40a) when the upper vapor of the demethanizer (stream 42) and the refrigerant (stream 71a) cool. The mixed liquid stream is heated to -24 F [-31 C], which partially evaporates stream (40b) before it is supplied to a feed point in the middle of

Deethane column (19). The separator fluid (current 33) expands and develops pressure

The operating pressure of the deethane (19) is slightly increased by an expansion valve (2 1), which cools the stream (33) to -46 F [-43 C] (stream 33 A) before the incoming feed gas is cooled as previously described. Then current (33B), currently at 85 F [-29 C], enters the extractor. Ethane (19) is at a lower feed point in the middle of the column. In deethane, methane and C2 carbonate components are removed from streams (40b) and (33b). The deethane in the tower (19), which operates at about 453 lb/in2 absolute [3123 kPa (absolute)], is also a conventional distillation column containing several vertically spaced trays, one or more packed layers, or some combination of trays and packings. . The deethane tower may consist of two sections, also an upper separation section (19a), where any steam present in the upper feed stream is separated from the corresponding liquid portion, and where the steam rising from the lower distillation or deethane section (9 1b) is mixed with the steam portion (if present). from the upper feed stream to form a distillation stream (42) that exits from the top of the tower; A lower deethane section (19b) contains trays and/or gaskets to provide the necessary contact between the liquids falling to the bottom and the vapors rising to the top. The deethane section (19b) also includes one or more reboilers (such as reboiler 02) that heat and vaporize a portion of the liquid at the bottom of the column to supply stripper vapors that flow to the top of the column to remove methane and C2 carbonate components from the liquid product. (Current 41). A typical specification for the bottom liquid product is to have an ethane to propane ratio of 1:0.20 on a molecular basis. The liquid product stream (41) exits the bottom of the deethane at 214 F [101 C].

The operating pressure in the deethane (19) is maintained so that it is slightly higher than the operating pressure of the separation/absorption tower (8 1). This allows equal steam flow to the deethane (stream 42) under pressure through the heat exchanger (13) and then into the upper section of the separation/absorption tower (18). In the heat exchanger (13), the steam is directed to the deethane dehydrogenase at -19 F [-28 C] in a heat exchange relationship with the mixed slurry stream (current 0 4 A) from the bottom of the separation/absorption tower (18) and the flash refrigerant stream (71 A), which The current cools to -

89 F [-67 C] (Stream 42A) and partially condenses it, and the partially condensed stream enters a cylindrical reflux vessel (22), where the condensed liquid (Stream 44) is separated from the non-condensable steam (Stream 43), and the stream (43) is mixed with the steam distillation stream (Stream 37) Leaving the upper area of ​​the separation/absorption tower (18) to form a cold residual gas stream (47) 0, the condensed liquid (stream 44) is pumped to a higher pressure via a pump (23) and then the stream (44a) is divided into two parts. One of the two parts, current (45), is directed to the equal section of the separator of the separation/absorption tower

(18) To act as the cold liquid that comes into contact with the vapors rising to the top through the absorption section and supplies the other part to the deethane (9 1) in the form of a return stream (46) and flows to an upper feed point in the deethane (9 1) at - 89 F[-67m].

The cold residual gas (stream 47) is warmed from 4 -9 F [-0 7 C] to 4 9 F [34 C] in a heat exchanger (24), and then part (stream 48) to act as fuel gas for the unit. The remaining heated gas (stream 49) is compressed through a compressor (16). After cooling to 100 F [38 C] in a vacuum cooler (25), stream (49b) is further cooled to -78 F [-61 C] in the heat exchanger (24) by exchanging the cooled residual gas, stream (47).

Then a stream (49 C) enters a heat exchanger (60) and is further cooled by a refrigerant stream (71 C) to -255 F [-0 16 C] to condense and cool it minimally, and it later enters an operational expansion machine (1 6) where the energy is extracted Mechanical current. The machine (1 6) extends a liquid stream (9 4d) in a constant, essentially inertial manner from a pressure of about 648 psi [465 kPa (absolute)] to the LNG storage pressure (15.5 psi). [1.07 kPa (absolute)]), which is slightly more than atmospheric pressure. The operational expansion cools the expanded stream (49 AH) to a temperature of approximately -256 F [-160 C]. , where it is then directed to the liquefied natural gas storage tank (62)

Which holds the liquefied natural gas producer (Stream 50).

Similarly, according to Figure [1] and Figure [3], most of the cooling is provided to the stream (42).

All cooling of the current (49 C) is by a closed-cycle cooling loop. The composition of the current used is

The working fluid in the process cycle according to Figure [4], as an approximate molar percentage, is 8.7% nitrogen, 30.0% methane, 45.8% ethane and 11.0% propane, with heavier hydrocarbons making up the rest of the percentage. . The refrigerant stream (71) leaves vacuum cooler (69) at 100 F [38°C] and 607 lb/in2 absolute [4185 kPa (absolute)]. It enters the heat exchanger (10) and cools to -17 F [-27 C] and is partially condensed by the expanded and partially heated refrigerant stream (71 F) and by other refrigerant streams. For the simulation of the process shown in Figure [4], it has been assumed that these other refrigerant streams are commercial-quality propane refrigerant at three different levels of temperature and pressure. Then the partially condensed refrigerant (71a) enters the heat exchanger (13) for additional cooling to -89°F (-67°C] via an expanded and partially heated refrigerant stream (71a), which additionally condenses the refrigerant (71b stream). The refrigerant is completely condensed and then cooled down to -255 F [-0 16 C] in the heat exchanger (60) through the expanding refrigerant stream (1 7 C). The minimally cooled liquid stream (17C) enters an operational expansion machine (63) where mechanical energy is extracted from the stream as it expands adiabatically and thermal inertia is essentially from a pressure of about 586 lb/in2 absolute [4040 kPa (absolute)]. During expansion, part of the current evaporates, cooling the total stream to -264 F (-164 C) (current 71 D). Then the expanding stream (71D) again enters the heat exchangers (60), (13) and (10) where it supplies cooling to the stream (49C), stream (42) and the refrigerant (streams 71, 71A, and 71B) as it evaporates. And overheating.

The superheated refrigerant vapor (71 g stream) leaves the heat exchanger (10) at 0.9 F (32 C) and is compressed in three stages to 617 lb/in2 absolute [4254 kPa (absolute)]. Each of the three stages of compression is driven (compressors Refrigerant 66, 64 and 68) using an additional power source, and each stage is followed by a cooler (vacuum coolers 65, 67, and 69) to remove the heat of compression, and the compressed stream (71) returns from the vacuum cooler (69) to the heat exchanger (10) to complete the cycle. .

A summary of the current flow rates and energy consumption for the process shown in Figure [4] is shown in the following table:

Table (3)

(Figure [4])

Summary of current flow rates - lbmol/hour [kgmol/hour]

<img file="SA2015B1_D0003.tif" />

Recovery rates in liquefied petroleum gas, natural gas* propane, 99.08%

Butane compounds and higher compounds 0 0.0 0 1% molecular weight 197051 lb/hour [197051 kg/hour]

LNG product production rate

Production rate 726918 lbs/hour [726918 kg/hour)

Purity. 91.43%

The minimum calorific value is 969.9 BTU/standard ft3 [36.14

MJ/m3]

Ability

Refrigerant compression 95,424 hp [156,876 kW]

Propane compression 28,060 hp [46,130 kW]

Total compression 123484 horsepower [06 30 20 kW]

Temperature of the utility facility

Dehydrator reboiler 55,070 thousand BTU/hour methane kW! 35,575 kW]

* (based on unrounded flow rates)

Assuming an operating factor of 34.0 days per year for an LNG production unit,

The specific power consumption rate for the embodiment of the present invention according to Figure [4] is 0.143 horsepower-hours/pound [0.236 kilowatt-hours/kg]. Compared with previous technology processes, the percentage of improvement in effectiveness according to the embodiment of Figure [4] ranges from 17-27%.

In comparison with the two embodiments according to Figures [1] and [3], the embodiment of the present invention requires

For Figure [4] to a lower capacity of 6% to 11% per unit of liquid produced. Thus, for the amount

Given the available compression capacity, an embodiment of Fig. [4] can liquefy approximately 6% more natural gas compared to the embodiment of Fig. [1], or a greater quantity of natural gas.

About 11% compared to - Figure [3] - recovery of C3 hydrocarbons and heavier hydrocarbons only in the form of a co-product of liquefied petroleum gas. In general, the choice between an embodiment of Fig. {4] and an embodiment of Fig. [1] or Fig. [3] according to the present invention for a particular application depends either on the monetary value of ethane as part of the natural gas liquefaction product versus its corresponding value in the liquefied natural gas product or on a specification of the calorific value of the product. LNG (since the calorific value of LNG produced by the embodiment of Fig. [1] and the embodiment of Fig. [3] is less than that produced by the embodiment of Fig.

[4]).

Example (4)

If the LNG product specifications would allow all of the ethane and propane in the feed gas to be recovered in the LNG product or if there is no market for a liquid coproduct containing ethane and propane, an alternative embodiment of the present invention may be used such as that shown in Figure [5]. To produce a condensed coproduct stream. The inlet gas conditions and composition taken into account in the process represented in Figure [5] are similar to those in Figures [1], [3] and [4]. Accordingly, the process of Figure [5] can be compared with the embodiments shown in Figures [1], [3] and [4].

In a simulation of the process according to Figure [5], an inlet gas enters the unit at 90 F [32 C] and 1285 psi [8860 kPa (absolute)] in the form of stream (31) and is cooled in a heat exchanger (10) by heat exchange with streams Refrigerants, liquids from a high-pressure flash separator at -37°F [-38°C] (Stream 33B) and liquids from a medium-pressure flash separator at -37°F [-38°C] (Stream 39B). The cooled stream (31A) enters a high-pressure separator (11) at -30 F [-34 C] and 1278 psi [8812 kPa (absolute)] where the vapor (stream 32) is separated from the condensed liquid (stream 33).

The steam (stream 32) from the high-pressure separator (1 1) enters an operational expansion machine (15), where mechanical energy is extracted from this part of the high-pressure feed stream. The machine (15) works to expand the steam with constant thermal inertia, essentially due to a pressure of 15

About 1278 psi [8812 kPa (absolute)] to a pressure of about 635 psi [4378 kPa (absolute)] and the expansion stream (32a) is cooled by operational expansion to a temperature of approximately -83°F [-64°C] and the expansion and condensing stream enters Partially (32a) to a medium pressure separator (18) where the vapor (stream 42) is separated from the condensed liquid (stream 39). The liquid resulting from the medium pressure separator (stream 39) is flash expanded to a pressure slightly higher than the operating pressure of the depropanizer (9 1) by an expansion valve (17), which cools (stream 39) to -108 F [-78 C] (stream 39 A). Before entering the heat exchanger (13), it is then heated when it cools to the remaining gas stream (19) and the refrigerant stream (71a), then it enters the heat exchanger (10) to cool the incoming feed gas as previously described. (Current 39C), which is present at -15 F [-26 C], enters the depropane (19) at a feed point.

Above the middle of the column.

The condensed liquid (stream 33), resulting from the high-pressure separator (11), is flash expanded to a pressure slightly higher than the operating pressure of the depropanizer (19) by an expansion valve (12), which cools the stream (33) to -93 F [-70 C]. ] (Stream 33A) before entering the heat exchanger (13), then it heats up when it cools the remaining gas stream (49) and the refrigerant stream (71A), then it enters the heat exchanger (10) to cool the incoming feed gas as previously described. The current (33C), which is at this point at 50F [10C], enters the depropanate (91) at a feed point lower than the middle of the column. In depropane, methane removes the C2 dicarbonate components and the C3 tricarbonate components from streams (39c) and (33c). The tower depropanizer (19), which operates at about 385 lb/in2 absolute [2654 kPa (absolute)], is a conventional distillation column containing several vertically spaced trays, one or more packed layers, or some combination of trays and packings. The depropane tower may consist of two sections: an upper separation section (19a) where the steam in the upper feed stream is separated from its liquid portion, and where the steam rising from the lower distillation or depropanation section (19b) is mixed with the steam portion (if present) of the feed stream. The upper one forms the distillation stream (37) that exits from the top of the tower; And a lower section to remove

Propane (9 1b) contains trays and/or gaskets to provide comfortable contact between the liquids falling to the bottom and the vapors rising to the top. The depropanization section (19b) also includes one or more reboilers (such as the reboiler (20)) that heat and vaporize a portion of the liquid at the bottom of the column to supply stripping vapors that flow to the top of the column to remove the liquid product, (stream 41), Consisting of methane, dicarbonate components C2 and tricarbonate components C3 - A typical specification for the bottom liquid product is to obtain a ratio of propane to butane compounds of 1:0.02 on a volume basis 0 The liquid product stream (1 4) exits the bottom of the depropane at 286F [141AD].

The upper distillation stream (37) leaves the depropanate (19) at 36°F [2°C] and is cooled and partially condensed by a commercial-quality propane refrigerant in the reflux condenser (21). The partially condensed stream (37A) enters the cylindrical reflux vessel (22) at 2 F [-17 C], where the condensed liquid (stream 44) is separated from the non-condensable vapor (stream 43). The condensed liquid (stream 4 4) is pumped using a pump (23) to an upper feed point in the depropane (9 1) in the form of a reference stream (44a).

The non-condensing steam (stream 43) leaving the cylindrical reflux vessel (22) is warmed to 94°F [34°C] in a heat exchanger (24), then - part (stream 48) to act as fuel gas in the industrial unit, and the rest of the heated steam (stream 38) is compressed using a compressor. (16). After cooling to 100 F [38 C] in a vacuum cooler (25), the stream (38 B) is further cooled to 15 F [-9 C] in the heat exchanger (24) by cross-exchange with cold steam (stream 43).

Then (stream 38C) mixes with the steam resulting from the medium-pressure separator (stream 42) to form a cool residual gas stream (49) 0 and the stream (49) enters the heat exchanger (13) and cools from -38 F [-39 C] to -2 0 1 F [-74 C] by the separator fluids (streams 39A and 33A) as described above and by the refrigerant stream (71A). The partially condensed stream (49A) enters the heat exchanger (60) and is further cooled by the refrigerant stream (71D) to -254F.

[-159 m] to partially condense it, and then it enters an operational extension machine (61) where mechanical energy is extracted from the current. The machine (61) extends the liquid stream (49b) in a constant thermal inertial manner from an edge pressure of 1.62 lb/in2 absolute [4282 kPa (absolute)] to the storage pressure of liquefied natural gas (1505 lb/in2 absolute [1.07 Kilopascals (absolute)], that is, slightly higher than atmospheric pressure

The expanded stream (49C) is cooled to a temperature of about -255°F [-159°C] and is then directed to an LNG storage tank (62) that holds the LNG product (stream 50).

Similar to the operations in Figure [1], Figure [3] and Figure [4], most of the cooling of stream (49) and all of the cooling of stream (49a) is provided by a closed-cycle training loop. The composition of the stream used as a working fluid in the process cycle is according to Figure [5], as an approximate molar ratio, consisting of 8.9% nitrogen, 34.3% methane, 41.3% ethane and 11% propane, with the heavier hydrocarbons making up the remainder. percentage . The refrigerant stream (71) leaves the vacuum cooler (69) at 100 F [38°C] and 607 lb/in2 absolute [4185 kPa (absolute)]. It enters the heat exchanger (10) and cools to -30°F [-34°C] and is partially condensed by the expanded and partially heated refrigerant stream (71°F) and by other refrigerant streams. For the simulation in Figure [5], assume that these other refrigerant streams are commercial-quality propane refrigerants at three different levels of temperature and pressure. Then the partially condensed refrigerant stream (71A) enters the heat exchanger (13) for additional cooling to -102 F [-74 C] by the partially heated and expanded refrigerant stream (71A), which further condenses the refrigerant (71B stream). The refrigerant is completely condensed and then partially cooled to -4°F [-159°C] in the heat exchanger (60) by an expanding refrigerant stream (71°C). The partially cooled liquid stream (71c) enters the operating expansion machine (63), where mechanical energy is extracted from the stream when it is adiabatically stressed. Thermal inertia essentially extends from a pressure of about 586 psi (absolute) [4040 kPa (absolute)] to about 34 psi (absolute)] to about 34 psi (absolute). Absolute [234 kPa (absolute)] During expansion a portion evaporates

From the current they lead to the cooling of the total current to -264 F [-164 C] (current 71 D) 0 Then the expanding current (71 D) enters the heat exchangers (60), (13), (10) where it supplies cooling to the current (49 A), Current (49) and the refrigerant (currents 71, 71a, and 71b), which leads to its evaporation and excessive heating.

The superheated refrigerant vapor (current 71g) leaves the heat exchanger (01) at 93°F [34] and is compressed in three stages to 617 psi [absolute] 4254 kPa (absolute)]. Each of the three compression stages (refrigerant compressors 64, 66, and 68) is managed by an additional power source, and each stage is followed by a cooler (vacuum coolers 65, 67, and 69) to remove the heat of compression. The compressed stream (71) from the vacuum cooler (69) returns to the heat exchanger (10) to complete the cycle.

Table (4)

(Figure [5])

Summary of current flows - lbmol/hour [kgmol/hour]

<img file="SA2015B1_D0004.tif" />

Recovery rates in condensed compounds*

Butane compounds butanes 95.04%

Pentanes and higher compounds 99.57%

Molecular weight 88390 lb/h [88390 kg/h]

Production rate

Natural gas producer, liquefied

Production rate 834,183 lb/hour [834,183 kg/hour]

Purity. 87.27%

Minimum calorific value 33.8 01 BTU/standard ft3[38.52

MJ/m3]

Ability

Refrigeration compression 84,974 hp [139,696 kW]

Propane compression 39,439 hp [64,837 kW]

Total compression 124413 horsepower [204533 kW]

Temperature of the utility facility

Demethanizer reboiler 52,913 thousand BTU/hour [34,182

kilowatt]

* (based on unrounded flow rates)

Assuming an operating factor of 340 days per year for an LNG production unit,

The specific power consumption rate for the embodiment of the present invention according to Figure [5] is 0.145 horsepower-hours/pound [0.238 kW-hours/kg]. Compared to previous technology processes, the percentage of improvement in effectiveness according to the embodiment of Figure [5] ranges from 16-26%.

In comparison with the two embodiments according to Figure [1] and Figure [3], the embodiment of the present invention according to Figure [5] requires 5% to 10% less capacity per unit of liquid produced. In comparison with the embodiment of Fig. [4], the embodiment of the present invention according to Fig. [5] requires essentially the same capacity per unit of liquid produced. Thus, for a given amount of available compression capacity, an embodiment of figure [5] can liquefy about 5% more natural gas than with an embodiment of figure [1], and about 10% more natural gas than with an embodiment of figure [3]. Or almost. same

The amount of natural gas is as in - Figure (4] - thanks to the recovery of C4 hydrocarbons and heavier hydrocarbons only in the form of a condensed co-product. In general, the choice between the embodiment of Figure [5] and the embodiments of the present invention according to Figure [1], Figure [3] or Figure [4] for a particular application depends on the monetary values ​​of ethane and propane as part of the natural gas or liquefied petroleum gas liquefaction product versus their corresponding values ​​in the product. LNG or on the calorific value specification of the LNG product (since the calorific value of LNG produced by the embodiments according to Figures [1], [3] and [4] is less than that produced by the embodiment according to Figure [5]).

Other incarnations

Those familiar with the technology will realize that the present invention can be adapted for use with all types of natural gas liquefaction units to enable the co-production of a natural gas liquid stream, an LPG stream or a condensate stream, as best suited to the needs at a particular location of the unit. He will also realize that different process configurations can be used to recover the liquid coproduct stream. For example, the two embodiments can be modified according to Figure [1] and Figure [3] to recover a liquefied petroleum gas stream or a condensate product stream as a liquid coproduct stream instead of a liquid natural gas stream as previously described in Examples (1) and (2). An embodiment of Fig. [4] can be configured to recover a liquid natural gas stream containing a significant portion of the C2 components present in the feed gas or to recover a condensate product stream containing only the C4 components and the heavier components present in the feed gas, rather than producing a co-product from Liquefied petroleum gas, as previously described in Example (3). The embodiment can be configured in accordance with Figure [5] to recover a liquid natural gas stream containing a significant portion of the C2 components present in the feed gas or to recover a liquefied petroleum gas stream containing a significant portion of the C3 components present in the feed gas instead of producing a co-product. Condenser as previously described in the example (;).

Figures [1], [3], [4] and [5] represent preferred embodiments of the present invention for the processing conditions shown. Figures [6] to [21] depict alternative embodiments of the present invention that may be considered for a particular application. As shown in Figures [6] and [7], all or part can be supplied

of condensing liquid (stream 33) from the separator (11) to a splitting tower (19) at a separate feed location lower than the middle of the reigns instead of mixing with a portion of the separator vapor (stream 34) flowing to the heat exchanger (13). Figure [8] depicts an alternative embodiment of the present invention that requires less equipment than the two embodiments according to Figures [1] and [6], although its power consumption rate is slightly higher. Likewise, Figure [9] depicts an alternative embodiment of the present invention that requires less equipment than the two embodiments according to Figure [3] and eats [7] 0 also at the expense of increasing the specific power consumption rate. Figures [10] to [41] depict alternative embodiments of the present invention requiring less equipment than the embodiment according to. For Figure [4], even though their specific power consumption rates are higher. (Noting that, as shown in Figures [10] to [14], distillation columns and systems such as deethane (9 1) include designs for an absorption tower in which re-boiling is made and designs of a re-boiling tower in which they are re-boiled.) Figures [15] and [6 1] depict two alternative embodiments of the present invention that combine the functions of the separation/absorption tower (18) and deethane (9 1) in embodiments of Figure [4] and Figures [10] through [14] into a single fractionation column (9 1 ). Depending on the amount of heavier hydrocarbons in the feed gas and the pressure of the feed gas, the cooled feed stream (31a) leaving the heat exchanger (10) may not contain any liquid (because it exists at a point higher than its dew point or because it exists at a higher pressure than the symbiosis has), and therefore does not need the separator (1 1) shown in Figure [1] and Figures [3] to [16], and the cooled feed stream may flow directly to a suitable expansion device such as an operational expansion machine (15).

The gaseous stream can be eliminated after recovering the liquid coproduct stream (Stream 37) in Figures [1], [3], [6] to [1 1], [3 1] and [14], (Stream 47) in Figures [4], [12], [H1], [16] and (current 43) in Figure [5]) before being supplied to the heat exchanger (60) for condensation and partial cooling in several ways. In the processes according to Figure [1] and Figures [3] and [16], the stream is heated and then compressed to a higher pressure using energy obtained from one or more operational extension machines, partially cooled in a vacuum cooler, and then further cooled by cross-exchange with the original stream. As shown in Figure [17], it may be preferable in some applications to compress the current to a higher pressure using an additional compressor (59), driven, for example, by a power source.

External capacity. As shown by the equipment shown with dotted lines (heat exchanger (4 2) and vacuum cooler (25)) in Figure [1] and Figures [3] to [16], it may be preferable in some cases to reduce the capital cost of the facility by reducing or eliminating Pre-cooling the compressed stream before it enters the heat exchanger (60) (at the expense of increasing the cooling load on the heat exchanger (60) and increasing the consumption rate in the refrigerant compressors (64), (66) and (68)). In these cases, the current (49A) leaving the compressor may flow directly to the heat exchanger (24) as shown in Figure [81], or flow directly to the heat exchanger (60) as shown in Figure [91] if no expansion machines are used. In order to operate any part of the high-pressure feed gas, a compressor driven by an external power source can be used, such as the compressor (59) shown in Figure [20], instead of the compressor (16). In other cases, it is not justified to perform any compression of the stream at all. Therefore, the current flows directly to the heat exchanger (60) as shown in Figure [1 2] and with the equipment planned with crossed lines (heat exchanger 4 2, compressor 6 1, and vacuum cooler 25) shown in Figure [1] and figures from [3] to [16]. If a heat exchanger (24) is not included to heat the stream before withdrawing the unit’s fuel gas (stream 48), an additional heater (58) may be needed to warm the fuel gas whose consumption is reduced, using the utility stream or another process stream to supply the necessary heat, as is Shown in Figures [9 1] to [1 2]. A choice of such should usually be evaluated for each application, taking into account factors such as gas composition, unit size, recovery level of the desired coproduct stream and available equipment.

According to the present invention, the inlet gas stream and feed stream to the LNG production section can be cooled in several ways. In the processes according to Figures [1], [3] and [6] to [9], the inlet gas stream (31) is cooled and condensed by material streams External cooling and fluids from the retail tower (9 1). In Figures [4], [5] and [10] to [14], fluids from a flash separator are used for this purpose with the external refrigerant streams. In Figures [15] and [61], fluids from the tower and fluids from the flash separator are used for this purpose with the external refrigerant streams. In Figures (71) to [21], only external refrigerant streams are used to cool the inlet gas stream

(31). However, cold process streams can also be used to supply some cooling to the high-pressure refrigerant (stream 71a), such as those shown in Figures [4], [5], [10] [11]. Also, any stream at temperature can be used Cooler than the current(s) being cooled. For example, a side stream of steam can be drawn from the separation/absorption tower (18) or fractionation tower (19) and used for cooling. The use and distribution of fluids and/or vapors from the tower to perform the heat exchange, and the specific arrangement of heat exchangers to cool the inlet gas and feed gas, should be evaluated for each specific application, as well as the selection of process streams for specific heat exchange services. The choice of cooling source depends on several factors, including , but not limited to, feed gas conditions and composition, unit size, heat exchanger size, potential cooling source temperature...etc. Those skilled in technology will also realize that any combination of the above cooling sources or cooling methods can be used to obtain the desired feed stream temperature(s).

Furthermore, additional external cooling supplied to the inlet gas stream and feed stream to the LNG production section can also be achieved in several different ways. It was assumed in Figure [1] and Figures [3] to (1 2) that the single-component refrigerant was boiled for high-level external cooling, and it was assumed that the multi-component refrigerant was evaporated for low-level external cooling, so that the single-component refrigerant was used for preliminary cooling of the multi-component refrigerant stream. Alternatively, both high-level cooling and low-level cooling can be achieved using single-component refrigerants with successively lower boiling points (cascade cooling) or a single single-component refrigerant at significantly lower evaporating pressures Sequential. Alternatively, high-level cooling and low-level cooling can be achieved using multi-component refrigerant streams that adjust their own compositions to provide the necessary cooling temperatures. Testing of the external cooling supply method depends on several load factors, such as, but not limited to, the conditions and composition of the feed gas, the size of the unit, the size of the compressor drive, the size of the heat exchanger, the temperature of the surrounding heat sink...etc. A technical person will also realize that it is possible Use any combination of the external cooling supply methods described above to obtain the desired supply stream temperature(s).

Partial cooling of the liquid stream leaving the heat exchanger results in (0 6) ((stream 49) in Figures [1], [6] and [8], (stream 49d) in Figures [3], [4], [7] and [9] to [16] (Stream 49b) in Figures [5], [19], and [20], (Stream 49e) in Figure [17], (Stream 49c) in Figure [18], and (Stream 49a) in Figure Figure [21]) reduces or eliminates the amount of flash steam that may be generated during the expansion of the stream to the operating pressure of the LNG storage tank (62). This generally reduces the specific power consumption rate of LNG production by eliminating the need for flash gas compression. However, in some cases it may be preferable to reduce the capital cost of the facility by reducing the size of the heat exchanger (60) and using flash gas compression or other means to eliminate any flash gas that may form.

Although it is described as an independent current extension in certain extension devices, suitable alternative extension means can be used. For example, conditions may permit operational extension of the supply stream

Essentially the condenser ((stream H3A) in Figures [1], [3], [6] and [7]) or the medium pressure reflux stream ((Stream 39) in Figures [1], [6] and [8]) Constant enthalpy flash expansion can also be used instead of operational expansion of the partially cooled liquid stream that leaves

Heat exchanger (60) ((Current 49) in Figures [1], [6] and [8], (Current 49D) in Figures [3], [4], [7] and [9] to [16], (Current 49B) in Figures [0], [19] and [20], (Current 49E) in Figure [7 1], (Current 49C) in Figure [8 1] and (Current 49A) in Figure [1 2]. ), but it will require either more partial cooling in the heat exchanger (60) to avoid the formation of flash steam in the expansion process, or increasing the compression of the flash steam or other methods to get rid of the resulting flash steam. Likewise, constant enthalpy flash expansion can be used instead of operational expansion of the partially cooled high-pressure refrigerant stream leaving the heat exchanger (60) ((stream 71C) in Figures [1] and Figures [3] to [21]) with the result being an increase an average

Power consumption required to compress the refrigerant.

While he has described what he believes to be preferred embodiments of the invention, those familiar with the art will recognize that other and additional modifications can be made to it, for example to adapt the invention to conditions

for various types of supply currents or other requirements without deviating from the principle of the present invention as defined in the following claims.