EA011599B1 - Configurations and methods of integrated ngl recovery and ng liquefaction - Google Patents

Abstract

Contemplated plants include a NGL recovery portion and a LNG liquefaction portion, wherein the NGL recovery portion provides a low-temperature and high-pressure overhead product directly to the LNG liquefaction portion. Feed gas cooling and condensation are most preferably performed using refrigeration cycles that employ refrigerants other than the demethanizer/absorber overhead product. Thus, cold demethanizer/absorber overhead product is compressed with the turbo-expansion and delivered to a liquefaction portion at significantly lower temperature and higher pressure without net compression energy expenditure.

Description

Field of Invention

The present invention relates to the extraction of natural gas condensate (PGA) and liquefaction of natural gas, and in particular to a combined plant for such processes.

BACKGROUND OF THE INVENTION

While global crude oil production is declining, natural gas production is still significant in many parts of the world. Natural gas is usually extracted from oil and gas production wells located onshore and on the high seas. Depending on the particular formation and gas-bearing formation, natural gas also contains relatively low amounts of non-methane hydrocarbons, including ethane, propane, ί-butane, η-butane, pentanes, hexane and heavier components, as well as water, nitrogen, dioxide carbon, hydrogen sulfide, mercaptans and other gases.

Natural gas from wellheads is usually processed to remove sulfur components, compressed and transported to consumers through high pressure pipelines. However, in remote locations lacking the necessary pipeline infrastructure, natural gas is usually transported by liquefying natural gas and transporting liquid gas (for example, using freight transport vessels for liquefied natural gas (LNG)). Unfortunately, liquefying natural gas is problematic because natural gas also contains aromatic compounds (e.g. benzene) and heavy hydrocarbons, which solidify when cooled to cryogenic temperatures. Consequently, most of the aromatic hydrocarbons must be removed to an extremely low concentration (typically less than one millionth per unit volume) to avoid freezing and clogging of cryogenic heat exchange equipment. Additionally, at least a portion of the lighter hydrocarbons, such as C2, C3 and C4, must be removed when LNG is imported into the North American natural gas market, which typically requires poorer natural gas. A typical North American pipeline gas contains mainly gaseous methane gas with a higher calorific value of 1050 to 1070 British thermal units / normal cubic foot. Extraction of non-methane components can be economically attractive, as these hydrocarbons can be sold as being in high demand compared to natural gas. For example, C2 is often used as a feedstock for petrochemical production, C3 and C4 are sold as CIS fuels and C5 + hydrocarbons can be further processed to be used for mixing with gasoline.

There are numerous plants and methods known in the art for recovering C2 and C3 + PGA from a natural gas feed. However, all past efforts have been focused on the removal of FGC hydrocarbons from natural gas using autonomous FGC extraction plants that operate independently of the natural gas liquefaction plants. These extraction processes mainly produce residual gas under low to medium pressure, which then requires compression and further cooling before liquefaction in the liquefaction plant. Typical examples of plants for extracting C2 and C3 + components from natural gas include those that use expander processes described in US Pat. Nos. 4,157,904 to Satre 11 and others, 4,251,249 to Sicbue, 4,671,039 to Wisk, 4,690,702 to Ragabo \\ bk | et al., 5,275,005 in the name of Satre 11 et al., 5,799,507 in the name of Apiop et al. or 5,890,378 in the name of Ratoe et al.

Other well-known high C2 recovery processes (e.g., US Pat. No. 6,116,050) require lowering the high pressure of a portion of the residual gas to the column to extract the PGA as methane-rich irrigation using a Joule-Thomson valve. While these processes improve C2 recovery, at least to some extent, the energy expended in recompressing the residual gas makes the process often uneconomical. To address some of these drawbacks, two-column plants can be implemented in which a high pressure absorber is in fluid communication with a lower-pressure distillation column to improve the efficiency of the extraction of PGC, as described in US Pat. No. 6,837,070. However, since these processes Freight One operates independently of liquefaction plants; they will generally require additional compression and cooling before liquefying the residual LNG gas.

In still other known configurations for the processing of PGC, a scrubber column is used in an LNG liquefaction plant to remove heavier components (C6 +). For example, a side stream from a cryogenic coil heat exchanger with spiral tubes is processed in a scrubber and fractionation unit, as shown in US Pat. No. 6,308,531 to Voebb et al. While such a process can be advantageously used to eliminate the formation of paraffin by removing C6 + and heavier components, it is not suitable for the removal of C2 + components, especially at high concentrations (65% or higher C2 recovery), and therefore will not be able to produce poor residual gas, which may be liquefied for the North America natural gas market. Still additionally known processes for the extraction of PGC, which are combined with the liquefaction of LNG, as disclosed by Poeb15 and others in patent No. 6662589, contain the doctrine that a liquid enriched in C2 can be used to absorb C3 in a high pressure absorber column. At the time when attempts are made to operate the absorber column at high pressure of the supplied gas

- 1 011599 (for example, 800 psi or higher) to reduce energy costs, it should be noted that the separation of Freight One suffers significantly due to a decrease in the relative volatility of the Freight One component, therefore, Freight One is produced with an excessive methane content. In addition, such process designs will typically be unable to achieve high recovery of C2 and C3 (e.g., greater than 60%).

Currently known natural gas liquefaction processes generally include several stages in which natural gas is cooled and condensed using either a pure refrigeration component or one or more cycles with a mixture of refrigerants. The cascade refrigeration cycle cools and liquefies the feed gas by means of refrigerants with several pure components having successively lower boiling points, such as, for example, propane, ethylene, methane and nitrogen. The refrigerant mixture cycle uses a mixture of refrigerants and can therefore be configured to use a single compressor and heat exchanger, which simplifies the equipment. Alternatively, the feed gas may also be cooled by means of a refrigeration cycle for pre-cooling with propane or by expansion of natural gas or nitrogen using either Joule-Thomson butterfly valves or a turboexpander. Unfortunately, the most well-known stand-alone processes for liquefying natural gas using one or many refrigeration cycles (either cascade cooling or a cycle with a mixture of refrigerants) have relatively low efficiencies when C2 or C3 extracts are turned upstream than a natural gas liquefaction plant .

Thus, while numerous plant configurations and methods for extracting NGL and liquefying natural gas are known in the art, all or almost all of them suffer from various disadvantages. Thus, there is still a need for improved extraction of NGL and liquefaction of natural gas, and especially in plants in which the extraction of NGL and liquefaction of natural gas are combined.

SUMMARY OF THE INVENTION

According to the invention, an apparatus for producing liquefied natural gas is created, comprising an expander connected to a compressor and capable of driving it, a separator designed to receive a partially expanded expander of the steam portion of the natural gas feed stream, producing a cold overhead product stream and communicated with the compressor in cold the overhead product stream, while the compressor is capable of producing a compressed cryogenic product stream from the cold overhead product stream overhead, having a pressure of at least 700 psi and a temperature not exceeding -50 ° T, and a unit for liquefying natural gas in communication with the compressor outlet through a stream of compressed cryogenic product.

The separator may receive another expanded steam portion in a separate location.

The separator may be a demethanizer.

The installation may further comprise a deethanizer capable of producing product C3 + and product C2.

The installation may further comprise a pipeline for supplying at least a portion of the product C2 to the cold overhead product stream.

The separator may be an irrigated absorber. The demethanizer can provide an irrigation flow to the absorber or act at a lower pressure than the absorber.

The installation may further comprise a pipeline for supplying a cooled product of the bottom overhead of the absorber from the absorber to extract C2 and / or a pipe for supplying the heated product of the bottom overhead of the absorber from the absorber to exhaust C2.

According to the invention, a method for producing liquefied natural gas is created in which the steam part of the feed stream of natural gas is expanded in an expander and directed to a separator, a cold overhead product stream is produced in the separator and compressed in a compressor driven by an expander to produce compressed cryogenic steam with a pressure of at least 700 psi and a temperature not exceeding -50 ° T, liquefy the resulting compressed cryogenic vapor in the liquefaction unit.

A demethanizer can be used as a separator, to which the liquid part of the natural gas feed stream is additionally directed.

The product of the bottom overhead of the separator can be sent to the expander and the product C2 of the expander is selectively fed into the cold stream of the product of the overhead.

The separator can act as an irrigated absorber.

The overhead stream of the demethanizer can be supplied as an irrigation stream to the absorber, and the product of the bottom overhead of the absorber is fed to the demethanizer, operating at a lower pressure than the absorber.

The product of the bottom overhead of the absorber can be heated to divert C2 before entering the demethanizer or cooled to extract C2 before entering the demethanizer.

The deethanizer product can be fed to the bottom overhead of the demethanizer to recover C2 and C3 +.

According to the invention, a method for producing liquefied natural gas is also created in which the steam part of the feed stream of natural gas is expanded and directed to a separator,

- 20151599 in the separator, the cold overhead product, compresses the cold overhead product using the energy obtained by expanding the steam portion of the natural gas feed stream to produce cryogenic steam with a pressure of 700 to 900 psi and a temperature of -50 to -80 ° P, and liquefy the resulting cryogenic vapor.

An absorber or demethanizer can be used as a separator.

Various objectives, features, aspects and advantages of the present invention will become more apparent from the following detailed description of preferred embodiments of the invention.

Brief Description of the Drawings

FIG. 1 is a diagram of a known apparatus for the extraction of Freight One and liquefaction of natural gas

FIG. 2A is a diagram of an embodiment of a plant using a single column to produce a cold compressed overhead product and separating C2 and / or C3.

FIG. 2B is a diagram of an embodiment of a plant using two columns for producing a cold compressed overhead product and separating C2 and / or C3.

FIG. 3 is a more detailed diagram of an installation option in accordance with FIG. 2A with a cascade refrigeration cycle and two cycles with a mixture of refrigerants to extract the PGA and liquefy natural gas.

FIG. 4 is a more detailed diagram of an installation option in accordance with FIG. 2B with a cascade refrigeration cycle and two cycles with a mixture of refrigerants for the extraction of PGC and liquefaction of natural gas.

FIG. 5 is a more detailed diagram of an installation option in accordance with FIG. 2B with two cascade refrigeration cycles and one cycle with a mixture of refrigerants for the extraction of PHC and liquefaction of natural gas.

FIG. 6 is a more detailed diagram of an installation option in accordance with FIG. 2B with two cascade refrigeration cycles and a cycle with a mixture of refrigerants / cascade cycle for the extraction of PGC and liquefaction of natural gas.

FIG. 7 is a more detailed diagram of an installation option in accordance with FIG. 2B with three cascade refrigeration cycles for the extraction of PGC and liquefaction of natural gas.

FIG. 8 is a graph depicting a composite heat curve for a natural gas liquefaction process.

Detailed description

The inventor has discovered that natural gas processing and liquefaction can be combined in various plants and methods in which the cold compression of lean natural gas in a compressor driven by a feed gas expander provides cold high-pressure natural gas that can be directly liquefied in a unit for liquefaction. Therefore, the need for clean energy to compress lean natural gas is neutral or even negative, while cooling the feed gas and condensation are achieved using different refrigeration cycles. Among other advantages, it is necessary to evaluate that the considered plants and methods provide the possibility of a combined process for the extraction of PHC and liquefaction of natural gas, in which 99% of propane and up to 85% of ethane can be extracted from the natural feed gas.

FIG. 1 illustrates a well-known stand-alone process for extracting CGC C2, which is combined with a stand-alone plant for liquefying natural gas. Here, the impurity-free and dried feed gas stream 1, typically supplied at about 1200 psi, is cooled in the feed gas heat exchanger 51 using the cold content of the overhead vapor of the column, side feed 22 of the reboiler, and external refrigerant 32. The liquid is then removed from cooled feed gas in the separator 52 and is sent to the column 58 Freight One, which acts as a demethanizer. The flash vapor from the separator 52 is divided into two parts, one part being cooled in the heat exchanger 54 to provide irrigation to the column and the other part is expanded in the turboexpander 64 to provide a cooled feed stream, which is sent to the lower section of the rectification column. It should be noted that the aforementioned autonomous gas subcooling process produces residual gas at a temperature approximately equal to the ambient temperature and approximately 450 psi. Such relatively low pressures and high temperatures are predominant due to the use of residual gas as a refrigerant for cooling the feed gas and supercooling the vapor portion of the feed gas and the pressure drops in the heat exchangers. Therefore, substantial re-compression in the re-compression compressor 100 and additional cooling (cooler not shown) of the residual gas are typically required before liquefaction in a plant, which significantly reduces process efficiency and economics.

On the contrary, the units under consideration presented here preserve essentially the entire cold content of the overhead product of the separator by direct supply of residual gas (the product of the overhead product of the separator) to the compressor without creating pressure drops in known heat transfer

- 3 011599 nicknames. Since the compressor is driven by expanding the vapor of the feed gas and the residual gas is much colder than previously known configurations, substantially higher compressor outlet pressures can be achieved at significantly lower temperatures. In addition, where the separator operates as an absorber, the pressure at the compressor outlet can be even higher. Therefore, it must be appreciated that in most of the plants and methods considered, the residual gas pressure is higher than 700 psi (typically between 700 and 900 psi) at a temperature lower than -50 ° P (typically from -50 to -80 ° P) can be achieved.

FIG. 2A depicts one embodiment of a plant in which the separator operates as a demethanizer, while FIG. 2B shows an embodiment of the apparatus in which the separator acts as an irrigated absorber and in which the demethanizer and deethanizer are then operated at a lower pressure to recover components C2 and / or C3 +. As shown in the figures, the C2 content in natural gas can be adjusted to a predetermined or desired concentration, or by combining the separated C2 with the product overhead of the column, as shown by dashed lines in FIG. 2A, or by controlling the temperature of the sludge product of the absorber, which is fed to the demethanizer (as shown by dashed lines for removal of C2) shown in FIG. 2B.

It must be recognized that the processes under consideration for combined extraction of Freight One and liquefaction significantly reduce the cost of equipment and energy costs for liquefying natural gas, while making it possible to fractionate Freight One into C2 and C3 + products. Such plants and processes will produce LNG, predominantly methane, which can be used and / or exported to North America, with heating values that meet gas pipeline standards.

In addition, it should be noted that the plants in question can work to produce LNG with variable ethane and propane content for non-US markets.

Using such plants and methods, high propane recovery (i.e., at least 95%) and high ethane recovery (up to 85%) from the relatively high pressure feed gas (e.g., between about 800 to 1,600 psi). inch) can be realized by operating the absorber at a higher pressure than the demethanizer. The compressor is then used to recycle the overhead stream of the demethanizer to the absorber, while the product of the bottom overhead of the absorber expands to provide cooling for the demethanizer. The overhead vapor from the absorber is compressed using energy (preferably exclusively) generated by the feed gas expander. Therefore, it is necessary to evaluate that the configurations and methods under consideration significantly reduce the energy consumption of the combined liquefaction plant. Further configurations related to certain aspects of the subject matter are disclosed in our co-pending application for US patent No. 10/478705, which is incorporated herein by reference.

While the subject matter of the invention is not limited, it is typically preferred that the cooling processes for both the extraction of PGC and the liquefaction of the residual gas can be configured to use a combination of one or more refrigeration evaporation cycles to provide cooling in at least three ranges temperature: the first temperature range from 10 to -35 ° R for pre-cooling the feed gas, the second temperature range from -60 to -160 ° R to produce irrigation demethanizer or absorber and the third dia temperature range from -180 to -270 ° Р for gas liquefaction.

FIG. 3 shows a more detailed diagram of an installation option in which the separator operates as a demethanizer (see also FIG. 2A). Regarding the feed gas streams, it is contemplated that numerous natural and artificial feed gas streams are suitable for use in connection with the teachings presented herein, and particularly preferred feed gas streams include natural gas streams, refinery gas and synthetic gas streams from hydrocarbon materials, such as for example, naphtha, coal, oil, brown coal, etc. Therefore, the pressure of the feed gas flows in question can vary significantly. However, it is generally preferable that the corresponding supply gas pressures for plant configurations in accordance with the subject invention be generally in the range of 800 to 1600 psi and that at least a portion of the feed gas is expanded in the turboexpander to provide cooling and / or energy for re-compression of the residual gas. The total mass balances illustrating the gas composition and flow rate for the approximate feed gas and products are shown in table. 1 below.

- 4 011599

Table 1

Here, the feed gas stream 1 enters the apparatus at approximately 1200 psi and 120 ° E and is cooled in the heat exchanger 51 typically from 10 to -30 ° E to form stream 2 using multiple cooling streams, including a liquid stream 5 from the separator 52, side stream 22 of the reboiler from the demethanizer 61, steam 70 of flash vapor from the LNG storage vessel 69, and a propane refrigerant stream 32. The propane refrigerant is typically generated in a cascade propane refrigeration unit 101, evaporating at least three different pressures to provide cooling for the heated stream 33. It should be noted that various heat exchangers (for example, fin and plate heat exchangers or spiral tube heat exchangers) can be used to achieve a convergence of temperatures, which provides high thermodynamic efficiency, as shown in the combined composite curves in FIG. 8.

The cooled feed gas stream 2 is separated in the separator 52, forming a gaseous part 3 and a liquid part 4. The pressure of the liquid part 4 is reduced in the Joule-Thomson valve 53, forming a stream 5, and it is selectively heated into the stream 6 by the heat content of the supplied gas before entering demethanizer 61 (in which re-boiling is carried out by reboiler 63). The gaseous part 3 from the separator 52 is divided into two parts. One part, stream 7, is sent to a heat exchanger 54 to provide irrigation to the absorber, and the other part, stream 8, expands in a turboexpander 64 to form a cooled steam stream 10, typically at a temperature of from -80 to -100 ° E, and to generate energy to drive the residual gas compressor 65. Chilled steam 10 is supplied to a demethanizer 61, which operates at pressures from 400 to 650 psi, most typically at 450 psi. It must be appreciated that the ratio of flow rates of stream 8 to stream 3 can be adjusted so as to adapt it to the desired recovery level C2 and / or correspond to the desired production rates for product C2. The demethanizer 61 is irrigated by the top irrigation stream 12 (formed from stream 9 through the Joule-Thomson valve 55) from the heat exchanger 54. The irrigation stream is preferably cooled in the heat exchanger 54 to a temperature of from about -125 to -155 ° E using a refrigerant mixture stream 72 and 74 agents (via stream 73 and the Joule-Thomson valve 91), which is produced from stream 72 by means of a compressed mixture of refrigerants from the refrigeration unit 102. The thus heated refrigerant 75 is then returned to the refrigeration unit 102.

The demethanizer 61 produces an overhead steam stream 28 at a temperature of from about -120 to -140 ° E and a low heat stream 14 at a temperature of from about 20 to 80 ° E. The overhead steam is compressed by the residual gas compressor 65 to form a discharge stream 29, typically at a pressure of about 700 to 900 psi and a temperature of from -50 to -80 ° E. It is necessary to particularly evaluate that the compression of cryogenic vapor is energy efficient and results in a high degree of compression in the compressor, which significantly reduces the cooling flow for liquefaction (using a third temperature level). In addition, it should be noted that the compression of the overhead product does not require clean energy, since the compressor is connected to the expander 64. Thus, by using a relatively high pressure of the supplied gas (for example, approximately 1000 psi) and compression of the cold overhead product separator, cooled residual gas can be fed into the liquefaction unit without consuming pure compression energy at a higher pressure and at a lower temperature than would be possible in other known processes PGC division. The cold compressed residual gas 29 is then further cooled and condensed in a heat exchanger 67 to a temperature of from about -255 to -265 ° E using a mixture of refrigerants 79 operating at a temperature of from -250 to -270 ° E. The refrigerant 79 is produced by the refrigeration unit 103 for mixing after the compressed stream 76 is cooled in the heat exchangers 54 and 67 (to form a stream 78), and expands in the Joule-Thomson cycle through the valve 92. The heated stream 80 is then returned to the refrigeration unit 103 .

The pressure of the liquefied residual gas 81 is further reduced to a pressure of approximately 16.0 psi through the Joule-Thomson valve 90 to form a stream 82, which is stored in the LNG storage vessel 69. The LNG product is discharged as stream 30, optionally combined with a Joule-Thomson-extended stream 15 of product C2 and exported to the terminal for loading vessels, in a storage container, or for other use. In some cases, and depending on the composition of the natural gas and temperature, substantial quantities of light gas can be released from the heat exchanger-liquefier, which can be used as a cooling source in subsequent heat exchangers to form fuel gas 71, which typically compresses to pressure in the fuel manifold.

As indicated above, a portion of ethane product stream 15 may be directed from the deethanizer 59 to an LNG storage tank for mixing with lean LNG to produce heavier and richer LNG, which may be required to adapt to various LNG markets. The deethanizer 59 receives the overhead product from the demethanizer and re-boils with a reboiler 34 to form the C3 + overhead product, which is discharged as liquid 23, for storage or further processing. The deethanizer overhead condenser 62 provides cooling for the overhead product C2. One part of the overhead product is provided as a deethanizer irrigation stream 18 from the separator drum 68 to the column through the pump, while the other part 19 is sent for storage or other use as stream 17.

Most preferably, the first column (demethanizer) separately receives the first and second parts of the supplied gas vapor, the first part of the supplied gas vapor being cooled by cooling the first level and the second part being cooled by cooling the second level, which provides irrigation to the demethanizer. In such configurations, it should be noted that the flow control unit (typically automated and using a controller programmable according to the desired product composition and / or feed gas composition) controls at least one of the first and second parts of the feed gas steam to produce the desired levels ethane recovery, from 10 to 85% of the feed gas, while maintaining a high C3 recovery (98% or higher).

At least a portion of the product of the bottom overhead of the demethanizer is fed to the deethanizer, which fractionates the product of the sludge of the demethanizer into ethane overhead and the product of the bottom overhead C3 +. Thus, it must be recognized that the considered methods and configurations enable the production of C2 at variable costs by mixing at least a portion of the liquid C2 of the overhead LNG stream. It must be recognized further that mixing significantly simplifies the operation of the installation for the extraction of Freight One and makes it possible to maintain the same process mode (temperature and pressure), regardless of the net C2 production rates.

FIG. 4 is a more detailed diagram of an embodiment in which the separator has a configuration such as an absorber that operates at a higher pressure than the downstream demethanizer and deethanizer (see also FIG. 2A). As regards the feed gas streams, the same assumptions apply as discussed for the configurations used in accordance with FIG. 3 above. The total mass balances illustrating the gas composition and flow rate for the approximate feed gas and products are shown in table. 1 above.

In general, the absorber receives the expanded feed gas and the irrigation stream, which are produced from the overhead steam from the demethanizer after the overhead steam is compressed and cooled by cooling the second level. In such configurations, the demethanizer column is in fluid communication with the absorber and receives the column feed stream and operates at a pressure that is at least 50-100 psi less, more preferably 100-300 psi .inch less than the working pressure of the absorber. Therefore, most typically, the feed gas has a pressure of 900 to 1,600 psi, expands in a turboexpander, and

- 6 011599 is fed to the absorber. The sludge product from the absorber expands to a pressure in the differential pressure range of 50 to 350 psi (relative to the demethanizer) and is thus cooled by the Joule-Thomson effect to a temperature of from -90 to -130 ° E. The cooled and expanded product stream of the lower overhead forms a rectification stream, which is fed to the demethanizer to extract C2. In the demethanizer, boiling is carried out by means of heat content from the supplied gas and optionally an external heat source while controlling the methane content in the sludge product at approximately 1.5 mol% (or, unless otherwise required, to meet the desired technical requirements for the product).

More specifically, as shown in FIG. 4, feed gas stream 1 enters the installation at a pressure of approximately 1200 psi and a temperature of 120 ° E and is cooled in a heat exchanger 51 to a temperature typically from 10 to -30 ° E, forming stream 2, using numerous cooling flows, including a liquid stream 5 from a separator 52, a reboiler side stream 22 from a demethanizer 61, flash vapor 70 from a natural gas storage vessel 69, and a propane refrigerant stream 32 of a refrigeration unit 101.

The propane refrigerant is produced from the heated stream 33 by means of a cascade propane refrigeration unit, evaporating at least three different pressure levels.

The cooled feed gas stream 2 is separated in the separator 52, forming a gaseous part 3 and a liquid part 4. The pressure of the liquid part is reduced in the Joule-Thomson valve 53, forming a stream 5, and it is selectively heated to stream 6 by the heat content of the supplied gas before entering the demethanizer 61. The gaseous part 3 of the separator 52 is divided into two parts. One part, stream 7, is sent to a heat exchanger 54 to provide irrigation to the absorber, and the other part, stream 8, expands in a turboexpander 64 to form a cooled steam stream 10 at a temperature typically from -80 to -100 ° E to generate energy to bring into the action of the compressor 65 of the residual gas. Chilled steam 10 is supplied to an absorber 58, which operates at a pressure significantly higher than 450 psi, typically 500 to 700 psi, and most typically at 600 psi. The ratio of the flow rates of the steam stream 8 and the steam stream 3 can be adjusted by changing to achieve a specific level of extraction of C2 and / or to meet the desired performance of the product C2. In the table. 2 below illustrates approximately the effect of the ratio of the flow rates of steam stream 8 and steam stream 3 on the recovery of C2 and C3.

table 2

The absorber 58 is irrigated by two cold streams, the first irrigation (irrigation of the upper part) being carried out by flow 27 (through 56 and 11) from the demethanizer 61 and the second irrigation is carried out by flow 12 (through 9 and 55) from the heat exchanger 54. The irrigation flows are cooled to a temperature from about -125 to -155 ° E through a refrigerant mixture stream 74, which is produced by means of a compressed refrigerant mixture from a refrigeration unit 102, which is cooled in a heat exchanger 54 and cooled by a valve I'm 91 Joule-Thomson.

The absorber produces a stream of 28 overhead steam at a temperature of from about -120 to -140 ° E and a stream of 14 overhead at a temperature of from about -100 to -110 ° E. The overhead steam is compressed by the residual gas compressor 65 using the energy produced by the expander 64, forming the discharge stream 29 at a pressure of typically about 900 psi and a temperature of from -70 to -80 ° E. It must be particularly appreciated that cryogenic vapor compression is thermodynamically more efficient, resulting in a high compression ratio in the compressor, which reduces the cooling flow for liquefaction. The residual gas is cooled and condensed in a heat exchanger 67 to a temperature of from about -255 to -265 ° E using a mixture of refrigerants 79 operating at a temperature of from -180 to -270 ° E, which is produced by a refrigeration unit 103 for mixing, after how the compressed stream 76 is cooled in the heat exchangers 54, 67, and expanded in a Joule-Thomson cycle through the valve 92. The pressure of the liquefied residual gas is further reduced to a pressure in the stream 82 at approximately 16.0 psi through the Joule-Thomson valve 90, and flash vapor is stored in a natural gas storage vessel 69. The natural gas product is discharged as stream 30 and discharged to storage or transport, in some cases, depending on the composition of the natural gas and temperature, a significant amount of light gas 70 is released from the heat exchanger-liquefier, which can be

- 7 011599 recovered as fuel gas after its cold content is recovered. If desired, a portion of ethane product stream 15 may be directed from deethanizer 59 to LNG storage or transport. In these cases, poor LNG can be converted into heavier and richer LNG.

The absorber overhead product stream 14 is preferably expanded in a Joule Thomson valve (or other expansion device) 60 to a pressure that is approximately 50-350 psi less than the pressure in the absorber and enters as a cooled stream 20 into the demethanizer at a temperature between -90 to -130 ° E. The demethanizer re-boils using a reboiler 63 and produces a bottom product 25, which is then fed to the deethanizer 59. The overhead product 24 of the demethanizer is then sent back to the absorber as an irrigation stream 11. Finally, the overhead product 24 is re-compressed to form a stream 26 (to a pressure above the pressure in the absorber) by means of a compressor 66 and cooled in a heat exchanger 54 to form a stream 27 that expands into the irrigation stream 11. Alternatively (not shown in Fig. 4, 2B), and especially where C2 is discharged, the product of the bottom of the absorber is expanded in the Joule-Thomson cycle, heating in countercurrent with the stream 1 of the supplied gas. The stream thus heated is additionally heated in the overhead condenser of the demethanizer and then fed to the demethanizer as a feed stream.

The deethanizer 59 has a configuration such as a reflux column using a reboiler 34 to separate C2 from the C3 + components, the C3 + components being discharged from the column as stream 23. The overhead product C2 is condensed in the overhead condenser 62 and separated in the drum 68. One part 18, product C2 is pumped back through the pump into the column as irrigation, while the other part 19 is discharged for mixing with LNG or storage / transport through stream 17. As for the remaining components and process conditions, the same with Most considerations apply to similar components as described in FIG. 3 above.

Therefore, from a different perspective, the absorber in the plants under consideration receives the liquid part of the natural gas supply and the second vapor part of the natural gas supply, and the pressure of the second part decreases through the turboexpander. Preferred absorbers produce a overhead product that expands, cools, and is fed to a demethanizer to absorb C2 + components. Preferred products of the bottom overhead of the demethanizer are subsequently fractionated in the deethanizer into the overhead liquid C2 and the bottom product C3 +. In a still further considered aspect, the absorber produces a overhead product, which is predominantly methane at a cryogenic temperature (-100 ° E or lower), which is further compressed using energy generated by expansion in the feed gas turbine expander. Such configurations produce high pressure cryogenic steam at a temperature of from -75 to -100 ° E and a pressure of from 800 to 900 psi or higher, which subsequently liquefies to form LNG using third-level cooling.

While the configurations in accordance with FIG. 4 are generally preferred, it should be noted that numerous alternative cooling methods and configurations for the first, second and / or third cooling stages are also considered suitable here. For example, FIG. 5 illustrates an embodiment of a plant in which a third temperature range of a refrigerant from -180 to -270 ° E is provided through a cascade cycle 103 with a methane refrigerant, operating at least three pressure levels. Alternatively, and depending on the composition and pressure of the residual gas, a pure refrigerant component, such as methane, may also be suitable. FIG. 6 illustrates another embodiment in which a cascaded propane pre-cooling cycle 104 is added to the outlet of the mixing refrigeration unit 102. Such an alternative refrigeration unit is particularly suitable when a very high ethane recovery is required or when the feed gas contains a very high ethane and propane content. FIG. 7 illustrates another alternative embodiment in which a propane cascade refrigerant, ethylene cascade refrigerant, and methane refrigerant are used to extract the PGA and liquefy natural gas.

Thus, the absorber in the considered installations and methods is configured to separately receive the first and second parts of the supplied gas vapor and the overhead of the demethanizer, the first part of the supplied gas vapor and the overhead of the demethanizer column providing irrigation for the absorber. In such configurations, the flow control unit controls at least one of the first and second parts of the feed gas steam to produce the desired ethane recovery levels, from 10 to 85% of the feed gas, while maintaining a high C3 recovery (98% or higher) . It is further considered that at least a part of the demethanizer sludge product is fed to a deethanizer, which fractionates the product of the bottom overhead of the demethanizer into the overhead product of ethane and the product of the bottom product C3 +. Thus, a preferred configuration can provide C2 variable capacities by mixing at least a portion of the excess C2 liquid of the LNG overhead. It must be specifically recognized that this mixing stage simplifies the operation of the installation for the extraction of Freight One and makes it possible to maintain the same process conditions (temperature and pressure), regardless of the net C2 capacity.

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In preferred installations, at least three temperature ranges are provided by one or more refrigeration evaporation cycles: a first cooling temperature range of 10 to -35 ° P for pre-cooling the feed gas, a second temperature range of -60 to -160 ° P for irrigation of the first column and a third temperature range from -180 to -270 ° P for gas liquefaction. In general, it is preferable that the refrigerant in the refrigeration circuits under consideration contain one, two or more hydrocarbon components and may further include nitrogen, halocarbons and / or other refrigerants. Consider refrigeration cycles may also include combinations of refrigeration cycles, and in particular combinations of cycles of multicomponent mixtures of refrigerants, a single component cascade cycle, a gas expander cycle, and a propane pre-cooling cycle. For example, it is contemplated that cooling in a first temperature range of 10 to -35 ° P utilizes propane pre-cooling or cascade cooling and cools at least one portion of the feed gas and refrigerant of a second temperature level. Cooling at a second temperature level from -60 to -160 ° P can then use a cycle with a mixture of refrigerants or cascade cooling using a pure component, such as ethylene, to cool the irrigation of the absorber, and cooling at a third temperature level from -180 to -270 ° P can use a cycle with a mixture of refrigerants or cascade cooling using a pure component, such as methane, to liquefy the residual gas. Other preferred refrigeration cycles include pressure reducing devices, such as turbo expanders and Joule-Thomson valves. As for temperature levels, (combinations) of refrigeration cycles and cooling medium, it should be noted that they can be adjusted as required in order to achieve the lowest energy consumption in the cooling and liquefaction processes.

As for the remaining components and the process mode of FIG. 5-7, the same considerations apply to the same components as described in FIG. 3 above. It should be noted that all components in the considered configurations (for example, heat exchangers, pumps, valves, compressors, expanders, irrigated absorbers, demethanizers, deethanizers, etc.) are commercially available and are suitable for use in accordance with the doctrines presented here. Further, it is generally contemplated that configurations in accordance with the subject invention can be widely applied to gas plants where a high recovery of propane and ethane is desired and the feed gas is available at pressures of more than 800 psi. In addition, such configurations produce high pressure cryogenic steam rich in methane, which will advantageously reduce the cost of equipment and operation when combined with a natural gas liquefaction plant. Tab. 3 below illustrates the temperature and pressure of the residual gas from the installation for the extraction of Freight One and the energy savings of the combined plants under consideration compared with previously known stand-alone plants based on 70 mol.% Ethane extraction. The energy savings of the plant configurations in question are approximately 10% compared to known plants that can be used to produce an equivalent amount of additional LNG.

Table 3

Autonomous installations for extraction of Freight One and liquefaction of natural gas Combined plants for the extraction of Freight One and liquefaction of natural gas Temperature residual gas, ° G 120 -80 to -60 Residual Gas Pressure psi 550 790-900 Total energy consumption for extraction of Freight One and liquefaction, MW 360 320

- 9 011599

Thus, specific embodiments and applications for the combined extraction of Freight One and liquefaction of natural gas are disclosed. It should be apparent, however, to those skilled in the art that many more modifications, other than those already described, are possible here without going beyond the concepts of the invention. The subject matter of the invention therefore should not be limited, with the exception of the essence of the attached claims. In addition, in interpreting both the description and the claims, all terms should be interpreted in the broadest possible manner consistent with the context. In particular, the terms “comprises” and “comprising” should be interpreted as referring to elements, components or steps in a non-exclusive manner, indicating that said elements, components or steps may be present, used, or combined with other elements, components or steps that not explicitly mentioned. In addition, where the definition or use of a term in a reference that is incorporated herein by reference is incompatible or contrary to the definition of the term provided herein, the definition of the term provided herein is applied and the definition of the term is not applied in the reference.

Claims (18)

CLAIM 1. Installation for producing liquefied natural gas, containing an expander connected to a compressor and capable of driving it, a separator designed to accept a partially expanded expander of the steam portion of the natural gas feed stream, producing a cold overhead product stream and communicated with the compressor through a cold stream overhead product, while the compressor is capable of producing a compressed cryogenic overhead product stream from a cold overhead product stream having at least 700 psi and a temperature not exceeding -50 ° E, and a natural gas liquefaction unit in communication with the compressor outlet through a stream of compressed cryogenic product. 2. The apparatus of claim 1, wherein the separator is capable of receiving another expanded steam portion at a separate location. 3. The apparatus of claim 1, wherein the separator is a demethanizer. 4. The apparatus of claim 3, further comprising a deethanizer capable of producing a product C3 + and a product C2. 5. The apparatus of claim 4, further comprising a pipeline for supplying at least a portion of the product C2 to the cold overhead product stream. 6. The apparatus of claim 1, wherein the separator is an irrigated absorber. 7. The apparatus of claim 6, wherein the demethanizer is capable of providing an irrigation flow to the absorber. 8. The apparatus of claim 6, wherein the demethanizer is capable of operating at a lower pressure than the absorber. 9. The apparatus of claim 6, further comprising a conduit for supplying a cooled product of the bottom overhead of the absorber from the absorber to extract C2 and / or a conduit for supplying a heated product of the bottom overhead of the absorber from the absorber to exhaust C2. 10. A method of producing liquefied natural gas, in which the steam part of the feed stream of natural gas is expanded in the expander and sent to the separator, a cold overhead product stream is produced in the separator and compressed in a compressor driven by the expander to produce compressed cryogenic steam with a pressure of at least 700 psi and a temperature not exceeding -50 ° E, the resulting compressed cryogenic vapor is liquefied in a liquefaction unit. 11. The method according to claim 10, in which a demethanizer is used as a separator, to which the liquid part of the natural gas feed stream is additionally directed. 12. The method according to claim 10, in which the product of the bottom overhead of the separator is sent to the expander, and the product C2 of the expander is selectively fed into the cold stream of the product of the overhead. 13. The method of claim 10, wherein the separator acts as an irrigated absorber. 14. The method according to claim 11, in which the overhead stream of the demethanizer is supplied as an irrigation stream to the absorber, and the product of the lower overhead of the absorber is fed to the demethanizer, operating at a lower pressure than the absorber. 15. The method according to 14, in which the product of the lower overhead of the absorber is heated to exhaust C2 before entering the demethanizer or cooled to extract C2 before entering the demethanizer. 16. The method according to 14, in which the product of the bottom overhead demethanizer is fed to the deethanizer to extract C2 and C3 +. 17. A method of producing liquefied natural gas, in which the steam part of the feed stream of natural gas is expanded and sent to the separator, a cold overhead product is produced in the separator, the cold overhead product is compressed using the energy obtained as a result of expansion of the steam part of the feed stream of natural gas obtaining cryogenic steam with a pressure of from 700 to 900 psi and a temperature of from -50 to -80 ° E, and liquefied the cryogenic vapor. 18. The method according to 17, in which an absorber or demethanizer is used as a separator.