NGL Extraction from Liquefied Natural Gas

Abstract

A process of extraction of natural gas liquids from a liquefied natural gas stream is disclosed.

Description

CROSS-REFERENCE TO RELATED APPLICATIONS

The present application claims priority to the provisional patent application 61/190,842 filed Sep. 3, 2008.

FIELD

The present invention generally relates to methods for extraction and recovery of natural gas liquids from a hydrocarbon gas stream. In particular, the methods of the present invention process liquefied hydrocarbon gas to separate ethane, propane, and other hydrocarbon liquids from a liquefied natural gas (LNG) stream, such as a LNG stream from storage prior to entry into a natural gas transportation pipeline.

BACKGROUND

Natural gas is an important energy source that is obtained from subterranean reservoirs, however, it is sometimes impractical or impossible to transport natural gas by pipeline from the wellhead where it is produced to the sites where it is needed, due to excessive distance or geographical barriers such as oceans. In such situations, liquefaction of natural gas offers an alternative way of transporting it.

Natural gas can be converted to liquefied natural gas (LNG) by cooling it to about −255° F., which reduces its volume to about 1/600th of its original value. This reduction in volume can make transportation more economical, for example the liquefied natural gas (LNG) can be transferred to a cryogenic storage tank located on an ocean-going ship for transportation. Once the ship arrives at its destination, the LNG can be offloaded to a storage facility that can contain the LNG in a liquid state in a cryogenic storage tank until it is converted back into a gas. Once it has been regasified, the natural gas can be transported by pipeline or other means to a location where it can be used as a fuel or a raw material for manufacturing other chemicals.

LNG is typically a mixture of various different gases including nitrogen, methane, ethane, propane, butane, and natural gasoline. Natural Gas Liquid (NGL) is a subset of these gases that make up LNG that can include ethane, propane, butane and natural gasoline. Either for economic gain or to meet increasingly stringent pipeline specifications, it is sometimes necessary to extract the NGLs from the LNG. For example, C₂, C₃ and C₄ hydrocarbons are valuable chemical intermediates and the C₃ and C₄ hydrocarbons can be of greater value when separated and utilized as a liquefied petroleum gas (LPG). C₅ and higher molecular weight hydrocarbons are valuable as blending stocks for motor fuels and for other purposes.

Current methods to extract the NGL's from a LNG stream associated with a LNG vaporization terminal typically involve pumping the LNG to an intermediate pressure, vaporizing the LNG in a large heat exchanger, distilling in a Demethanizer column to separate NGL's from the vaporized natural gas stream, compressing the Demethanizer overhead vapor, condensing the compressed Demethanizer overhead vapor by flowing back through the large heat exchanger to cross exchange thermal energy with the original LNG stream, collecting the condensed Demethanizer overhead into a cryogenic surge tank and then pumping the condensed Demethanizer overhead to a higher pressure before routing to the LNG vaporizer.

In view of the above, it would be desirable to have a LNG process to recover NGLs after the LNG vaporization, not requiring the LNG to be vaporized, recondensed and then pumped a second time before being routed to the LNG vaporizer, thus eliminating one or more high expense components such as a heat exchanger, a cryogenic surge tank and a second set of cryogenic LNG pumps. It would also be desirable to have a process with increased efficiency by eliminating the cryogenic compressor and/or any incremental cryogenic pump requirements.

SUMMARY

In an one embodiment of the invention a process for the separation of C₂+ or C₃+ hydrocarbons from a liquefied natural gas (LNG) stream is claimed that includes providing a first stream of liquid LNG having a pressure of at least 600 psig and passing the first stream through a heat exchanger to vaporize the LNG and form a second stream. The second stream is expanded in a turbo-expander to form a third stream having reduced pressure and temperature as compared to the second stream. A first portion of the third stream is fed to a distillation column at one or more feed trays, the column having a plurality of liquid recovery trays. A second portion of the third stream passes through the heat exchanger to be cooled by cross exchange with the first stream and form a fourth stream which is fed to the distillation column at a point higher than the first portion of the third stream inlet. A fifth stream of vapor hydrocarbon is withdrawn from an upper portion of the column having a reduced content of C₂+ or C₃+ hydrocarbons which is compressed in a turbo-compressor coupled to and powered by the turbo-expander. A sixth stream of liquid NGL hydrocarbon is withdrawn from a lower portion of the column having an increased content of C₂+ or C₃+ hydrocarbons.

The distillation column can be operated with a bottom temperature of about 0° F. to about 250° F. and a top temperature of about −180° F. to about −50° F. and a pressure of about 150 psig to about 550 psig.

The fourth stream can provide from 1% to 50% of the total feed to the distillation column, such as for example from 20% to 40% of the total feed to the distillation column.

The turbo-compressor can provide at least 30% of the required compression for the fifth stream to reach a pipeline pressure and alternately can provide 50%, 80%, or 100% of the required compression. In embodiments the turbo-compressor can provide less than 100% of the required compression for the fifth stream to reach a pipeline pressure. The second stream entering the turbo-expander can be at a higher temperature than the fifth stream entering the turbo-compressor, and can be at least 50° F., at least 100° F., or at least 150° F. higher temperature than the fifth stream. The fifth stream before and after the turbo-compressor can have either a pressure of at least 700 psig, or a temperature greater than −50° F., or both

In an alternate embodiment of the present invention is a process for the separation of C₂+ or C₃+ hydrocarbons from a liquefied natural gas (LNG) stream. The process comprises providing a first stream of vaporized natural gas having a pressure of at least 800 psig and expanding the first stream in a turbo-expander to form a second stream having reduced pressure and temperature as compared to the first stream. The second stream is fed to a distillation column at one or more feed trays, the column having a plurality of liquid recovery trays. A third stream of liquid LNG having a pressure of at least 150 psig is provided and fed to the distillation column at a point higher than the second stream inlet. Hydrocarbon liquids are condensed in said recovery trays and a fourth stream of vapor hydrocarbon is removed from an upper portion of the column having a reduced content of C₂+ or C₃+ hydrocarbons. The fourth stream is compressed in a turbo-compressor coupled to and powered by the turbo-expander. A fifth stream of liquid NGL hydrocarbon is withdrawn from a lower portion of the column having an increased content of C₂+ or C₃+ hydrocarbons.

The distillation column can be operated with a bottom temperature of about 0° F. to about 250° F. and a top temperature of about −200° F. to about 0° F. and can be operated at a pressure of about 150 psig to about 550 psig. The second stream can provide from 1% to 50% of the total feed to the distillation column, such as for example from 20% to 40% of the total feed to the distillation column.

The turbo-compressor can provide at least 30% of the required compression for the fourth stream to reach a pipeline pressure and alternately can provide 50%, 80%, or 100% of the required compression. In embodiments the turbo-compressor can provide less than 100% of the required compression for the fifth stream to reach a pipeline pressure. The first stream entering the turbo-expander can be at a higher temperature than the fourth stream entering the turbo-compressor, and can be 50° F., 100° F., or 150° F. higher temperature than the fourth stream. The fourth stream before and after the turbo-compressor can have either a pressure of at least 700 psig, or a temperature greater than −50° F., or both.

BRIEF DESCRIPTION OF DRAWINGS

FIG. 1 is a simplified process flow diagram of a prior art method of extracting NGL from natural gas not associated with a LNG vaporization terminal.

FIG. 2 is a simplified process flow diagram of a prior art method of extracting NGL from a LNG stream associated with a LNG vaporization terminal.

FIG. 3 is a simplified process flow diagram of a method of extracting NGL from a LNG stream associated with a LNG vaporization terminal that is one embodiment of the present invention.

FIG. 4 is a simplified process flow diagram of a method of extracting NGL from a LNG stream associated with a LNG vaporization terminal that is one embodiment of the present invention.

DETAILED DESCRIPTION

To perform NGL extraction from natural gas not associated with a LNG vaporization terminal, the current state of the art includes a cryogenic turbo-expander process as shown in FIG. 1. The primary purpose of the turbo-expander in this design is to produce the necessary refrigeration for the process. Referring to FIG. 1, the process includes a heat exchanger 4, a turbo-expander/compressor 12, a Demethanizer 20 and a Recompressor 30. High pressure natural gas in stream 2 is chilled by heat exchange within heat exchanger 4 to a temperature of between −10° F. and −100° F. to form stream 6 that enters a separator 8. An overhead stream 10 from the separator 8 enters the expander side 11 of a turbo-expander/compressor 12, and exits as an expanded stream 14 that enters a Demethanizer 20. A bottoms stream 16 from the separator 8 can be expanded in valve 17 to form stream 18 that enters the Demethanizer 20, typically at a point lower than the inlet of stream 14. The turbo-expander 12 decompresses stream 10 and provides refrigeration to the Demethanizer 20. In the Demethanizer 20 methane vapor is distilled and exits as stream 24 and NGL liquid product can be recovered as a liquid from the bottom of the Demethanizer 20 as shown in stream 22. The Demethanizer overhead in stream 24 is then warmed through heat exchange with the inlet stream 2 to a temperature of between −10° F. to 80° F. to form stream 26 before entering the compressor side 13 of the expander/compressor 12. The compressor discharge stream 28 is then further compressed in the Recompressor 30 to pipeline pressure in stream 32. In this arrangement, the turbo-expander 11 produces the necessary cryogenic refrigeration for the process but typically only provides from 10% to 25% of the necessary recompression horsepower required to compress the residue gas stream 26 back to pipeline pressure. Thus there is a large recompression horsepower requirement associated with the process.

To extract the NGL that may be present in a LNG stream associated with a LNG vaporization terminal, a cryogenic process is typically used that includes a Demethanizer distillation vessel where the methane vapor is distilled over the top of the vessel and the NGL liquid product is recovered as a liquid from the bottom of the vessel. The Demethanizer, however, generally has a maximum operating pressure of approximately 550 psig at which point the gas composition in the Demethanizer approaches its critical temperature and pressure. When the Demethanizer reaches these critical conditions, distillation becomes problematic, thus, the Demethanizer will generally operate at a pressure equal to or less than approximately 550 psig.

Given that most natural gas takeaway pipelines operate at between 800 and 1400 psig, a large amount of compression is generally required to compress the gas from the reduced Demethanizer pressure to the required pipeline pressure, which can be a substantial operating cost. Thus, the challenge in designing new NGL extraction processes for LNG has been to find ways of using the abundant cryogenic refrigeration to reduce or eliminate this large gas compression requirement. The prior art technology for NGL extraction from natural gas not associated with a LNG vaporization terminal as shown in FIG. 1 is not desirable since the refrigeration produced by the expander is not required and there is a large recompression horsepower requirement.

FIG. 2 shows a prior art process of extracting NGL from a LNG stream associated with a LNG vaporization terminal that is a “Condense and Pump to Vaporizer” approach. In this approach, the Demethanizer 46 overhead gas stream 56 is compressed in compressor 58 and then cooled and condensed at an intermediate pressure through heat exchange in heat exchanger 42 against the inlet liquid LNG stream 40. The condensed Demethanizer overhead stream 62 then flows to a surge drum 64, is pumped to the required higher pressure and sent to LNG vaporizers via stream 70. This process by being able to condense the Demethanizer overhead using the refrigeration in the LNG liquid and then pumping to pipeline pressure versus compressing the Demethanizer overhead greatly reduces the recompression horse power requirement. However for deep ethane recovery, some cryogenic compression (either by compressor or expander booster compressor) is necessary to ensure the complete condensing of the Demethanizer overhead vapor.

Disclosed are methods for NGL extraction from LNG wherein all or the vast majority of the power required to compress the Demethanizer overhead gas to pipeline pressure is provided by a turbo-expander. Typically turbo-expanders are used in the NGL extraction processes to produce cryogenic refrigeration. In this application with LNG, since there is an abundance of cryogenic refrigeration, the turbo-expander purpose is to provide the power to compress the Demethanizer overhead gas to pipeline pressure. This process also has the advantage of being located after the LNG vaporization versus before the LNG vaporization as in the Condense and Pump designs represented above in FIG. 2.

FIG. 3 illustrates one embodiment of the invention that is a process 100 wherein all the power required to compress the Demethanizer overhead vapor to pipeline pressure is produced by a turbo-expander. Liquid LNG 102 is pumped 104 to the required high pressure in Stream 106 (for example 500 psig to 2500 psig and −255° F.) before being heated and vaporized in exchanger 110 to form Stream 112. The process exchanger, 110, can comprise one or more exchangers that can be configured in parallel, in series, or combinations thereof. The vaporized natural gas in Stream 112 (generally 600 psig to 1,500 psig, and from 0° F. to 100° F.) is then expanded/decompressed in the turbo-expander 113 end of an expander Compressor 114 to a pressure at or near the Demethanizer 130 operating pressure. The expanded vapor 116 exiting the turbo-expander 113 can then be split into two streams, Stream 118 and Stream 122. In embodiments the pressure of Stream 118 and Stream 122 can range from the Demethanizer 130 operating pressure to a pressure above the Demethanizer operating pressure by 100 psig, 50 psig, or 25 psig. In embodiments the flow through Stream 106 can range from 100 MMSCF to 1,000 MMSCF; or from 200 MMSCF to 800 MMSCF; or from 300 MMSCF to 600 MMSCF.

Stream 118 can be further cooled by heat exchange with Stream 106 in exchanger 110, at least partially condensed and subcooled to form Stream 120 which can have a temperature ranging from −100° F. to −250° F. and then flow through control valve 150 to the Demethanizer 130 as a reflux stream 124. In one embodiment Stream 124 can have a temperature of from −150° F. to −250° F. In another embodiment Stream 124 can have a temperature of from −200° F. to −230° F. Stream 124 can represent between 5% to 50% of the total flow to the Demethanizer. Stream 122, representing the remaining 50% to 95% of the flow, can pass through control valve 152 forming stream 126 that is routed to a lower point in the Demethanizer 130 than is Stream 124. In an embodiment Stream 124 can represent approximately 10% to 30% of the total flow to the Demethanizer 130 while Stream 126 represents approximately 70% to 90% of the Demethanizer feed. In an embodiment Stream 124 can represent approximately 25% to 30% of the total flow to the Demethanizer 130 while Stream 126 represents approximately 70% to 75% of the Demethanizer feed.

The Demethanizer 130 will typically operate at a pressure from between 150 psig to 550 psig and may have one or more heat sources 131 to provide either a bottoms reboiler and/or one or more side reboilers. The Demethanizer overhead in Stream 142 may have a pressure of from 150 psig to 550 psig and a temperature from −200° F. to 0° F. Stream 142 is then compressed by the compressor 115 end of an Expander-Compressor 114 to the desired pipeline pressure in Stream 144. An optional trim heat exchanger 146 can provide additional heating if required to elevate the temperature of stream 144 to stream 148.

In embodiments of the present invention the compressed overhead stream 144, has either a pressure that is less than 700 psig or a temperature that is warmer than −50° F. at each location in the stream. This can include embodiments that have multiple staged compression of the overhead stream 142.

Additional compression (not shown in FIG. 3) may be used in this process to compress Stream 148, either before or after the optional trim heat exchanger 146, if the expander/compressor 114 does not provide all of the compression required. An embodiment of the present invention provides at least 30% of the required compression from the expander/compressor 114. Alternate embodiments provide at least 40%, at least 50%, at least 60%, at least 70%, at least 80%, at least 90%, and at least 95% of the required compression from the expander/compressor 114. An embodiment of the present invention provides less than 100% of the required compression from the expander/compressor 114 and requires additional compression.

FIG. 3 shows two liquid draws from the Demethanizer 130 passing through exchanger 131, which is used to impart heat to the Demethanizer 130, but alternate embodiments can contain one, or more than two, draws. The reboiler exchanger 131 can comprise one or more exchangers that can be configured in parallel, in series, or combinations thereof.

Optional control valves 150 & 152 can control Streams 124 and 126 prior to the Demethanizer and can control the flow ratios of the two streams.

The Demethanizer 130 provides separation of the hydrocarbon components in the inlet gas via boiling point distillation. A vapor overhead Stream 142 composed primarily of methane and having a reduced content of NGL's is removed from an upper portion of the Demethanizer 130. A liquid bottoms Stream 132 of extracted NGL's is removed from a lower portion of the Demethanizer 132. The Demethanizer 132 can be of any suitable design for the separation of hydrocarbons and are generally known to those in the industry. In one embodiment the Demethanizer 132 is operated with a bottom temperature of about 0° F. to about 250° F. and a top temperature of about −200° F. to about 0° F. and is operated at a pressure of about 150 psig to 550 psig.

There can be several advantages of the above process over the “Condense and Pump to Vaporizer” technologies. It can reduce or eliminate the need for Demethanizer overhead gas recompression other than the expander/compressor, which can reduce capital and utility expenses. It can eliminate the second set of cryogenic pumps and the associated stainless steel surge tank necessary in the Condense and Pump technologies. The overall heat exchanger area can be greatly reduced since the LNG does not have to be vaporized and then completely recondensed. A further benefit is the NGL extraction is located after the LNG vaporization instead of before the LNG vaporization in the Condense and Pump technologies.

FIG. 4 illustrates another embodiment 200 of the present invention. High pressure hydrocarbon gas from the LNG vaporizers (for example at 600 psig to 2500 psig and 0° F. to 80° F.) in Stream 202 flows through a turbo-expander 203 end of an expander-compressor 204 and is expanded/depressured to at or near the Demethanizer pressure (for example 150 psig to 550 psig) in Stream 206. Stream 206 enters the Demethanizer 210 and can represent the majority of the total gas flowing into the Demethanizer 210. In one embodiment Stream 206 can represent from between 50% to 99% of the total gas flowing into the Demethanizer 210. In another embodiment Stream 206 can represent from between 60% to 80% of the total gas flowing into the Demethanizer 210. In embodiments the flow through Stream 202 can range from 100 MMSCF to 500 MMSCF; or from 200 MMSCF to 400 MMSCF; or from 250 MMSCF to 350 MMSCF.

A separate LNG liquid stream 212 is pumped 214 from LNG storage to the Demethanizer in Stream 216 (for example at 150 psig to 550 psig and −225° F. to −260° F.) to act as reflux for the Demethanizer 210. In one embodiment Stream 216 can represent from between 1% to 50% of the total gas flowing into the Demethanizer 210. In another embodiment Stream 216 can represent from between 5% to 40%, or from 5% to 25% of the total gas flowing into the Demethanizer 210. Optional control valves 220 & 222 can regulate Streams 206 and 216 prior to the Demethanizer 210 and control the flow ratios of the two streams. In embodiments the flow through Stream 216 can range from 1 MMSCF to 500 MMSCF; or from 10 MMSCF to 400 MMSCF; or from 20 MMSCF to 300 MMSCF.

The Demethanizer overhead in Stream 224 (150 to 550 psig and −200° F. to 0° F.) is then compressed using the compressor end 205 of the expander/compressor 204, for example to between 800 and 1400 psig, in Stream 226 providing a portion of the power required to compress the gas to pipeline pressure. An embodiment of the present invention provides at least 30% of the required compression from the expander/compressor 204. Alternate embodiments provide at least 40%, at least 50%, at least 60%, at least 70%, at least 80%, at least 90%, or at least 95% of the required compression from the expander/compressor 204. In embodiments of the present invention the compressed overhead stream 226, 230, has either a pressure that is less than 700 psig or a temperature that is warmer than −50° F. at each location in the stream. This can include embodiments that have multiple staged compression of the overhead stream 224. An embodiment of the present invention provides less than 100% of the required compression from the expander/compressor 114 and requires additional compression.

If additional compression is required a non-cryogenic compressor 228 may be used to ensure the compressed gas can reach the desired pipeline pressure of stream 230. The ability to utilize a non-cryogenic compressor rather than a cryogenic compressor can provide advantages over the prior art, such as depicted in FIG. 2. A non-cryogenic compressor can have advantages in cost and operational considerations, for example. In an embodiment the additional compression horsepower requirements can range from 500 HP to 10,000 HP; in alternate embodiments the additional compression horsepower requirements can range from 1,000 HP to 6,000 HP; or from 2,500 HP to 5,000 HP.

The Demethanizer 210 provides separation of the hydrocarbon components in the inlet gas via boiling point distillation. A vapor overhead Stream 224 composed primarily of methane and having a reduced content of NGL's is removed from an upper portion of the Demethanizer 210. A liquid bottoms Stream 232 of extracted NGL's is removed from a lower portion of the Demethanizer 210. The liquid bottoms stream 232 can flow through a surge drum 234 and can be pumped 236 to an outlet stream of NGL product 238. The Demethanizer 210 can be of any suitable design for the separation of hydrocarbons and are generally known to those in the industry. In one embodiment the Demethanizer 210 is operated with a bottom temperature of about 0° F. to about 250° F. and a top temperature of about −200° F. to about 0° F. and is operated at a pressure of about 150 psig to 550 psig.

FIG. 4 shows two liquid draws from the Demethanizer passing through exchanger 240, but alternate embodiments can contain one, or more than two, draws. The reboiler exchanger 240 can comprise one or more exchangers that can be configured in parallel, in series, or combinations thereof.

There can be advantages of the above process over the “Condense and Pump to Vaporizer” process shown in FIG. 2. One or more pieces of major equipment may be eliminated including: (a) 42, the heat exchanger that is typically large and can be expensive; (b) the cryogenic surge tank 64; and (c) cryogenic pumps 68. The present invention can expand the vaporization capacity of the plant by the percentage of feed to the Demethanizer coming from Stream 216 that is used as reflux and then vaporized in the NGL extraction process. The power requirement in the process of the present invention can be comparable to other Condense and Pump technologies, however the secondary compressor 228, if needed, can be non-cryogenic (carbon steel) versus the cryogenic (stainless) compressor required in Condense and Pump technologies. The present invention can be located downstream of the LNG vaporizers whereas the Condense and Pump technologies are located before the LNG vaporizers.

As stated earlier, the purpose of the turbo-expander in the processes of the present invention is to provide recompression horsepower and not to provide process refrigeration as described in FIG. 1 above depicting a prior art cryogenic expander plant for NGL extraction from natural gas not associated with an LNG vaporization terminal. In the prior art plant in FIG. 1, the turbo-expander is located after the inlet heat exchange and the gas inlet temperature to the turbo-expander is cooled to between −10° F. to −100° F. In an embodiment of the present invention, the turbo-expander gas inlet is before any heat exchanger and is warmer, generally ranging from 0° F. to 100° F., than the process shown in FIG. 1.

In addition, in the prior art process in FIG. 1, the residue gas flowing to the expander booster compressor is after a heat exchange typically ranges from −10° F. to 80° F. In contrast, in the processes of the present invention, the residue gas flowing to the compressor end of the expander-compressor comes directly from the Demethanizer overhead and thus is significantly colder, generally ranging from 0° F. to −200° F. The effect of this placement difference of the expander-compressor in the process is that the expander-compressor in the prior art plant in FIG. 1 only provides 10% to 25% of the necessary recompression horsepower; whereas, the expander-compressor in the processes of the present invention can provide from 30% to 100% of the necessary recompression horsepower to reach pipeline pressure.

The prior art process shown in FIG. 1 has an inlet stream to the expander 11, Stream 10, which is at a lower temperature than the inlet stream to the compressor 13, Stream 26. For example the temperature of the inlet stream 10 to the expander can be from 0° F. to 180° F. less than the temperature of the inlet stream 26 to the compressor.

In contrast to the prior art process shown in FIG. 1, in the processes of the present invention the inlet stream to the expander, Stream 112 in FIG. 3 and Stream 202 in FIG. 4, are at a higher temperature than the inlet stream to the compressor, Stream 142 in FIG. 3 and Stream 224 in FIG. 4. For example the temperature of the inlet stream to the expander can be from 0° F. to 300° F. higher than the temperature of the inlet stream to the compressor.

In one embodiment of the present invention the temperature of the inlet stream to the expander can be higher than the temperature of the inlet stream to the compressor. In alternate embodiments of the present invention the temperature of the inlet stream to the expander can be higher than the temperature the inlet stream to the compressor in an amount greater than 50° F., greater than 100° F., greater than 150° F., greater than 175° F., or greater than 200° F.

The present invention does not include the production of a LNG product but involves the extraction of NGL product from a LNG stream.

As used herein the term “LNG” shall mean natural gas that is in a liquid state.

As used herein the term “NGL” shall mean natural gas liquids that can include ethane, propane, butanes and natural gasolines.

Depending on the context, all references herein to the “invention” may in some cases refer to certain specific embodiments only. In other cases it may refer to subject matter recited in one or more, but not necessarily all, of the claims. While the foregoing is directed to embodiments, versions and examples of the present invention, which are included to enable a person of ordinary skill in the art to make and use the inventions when the information in this patent is combined with available information and technology, the inventions are not limited to only these particular embodiments, versions and examples. Other and further embodiments, versions and examples of the invention may be devised without departing from the basic scope thereof and the scope thereof is determined by the claims that follow.

Claims

1. A process for the separation of C2+ or C3+ hydrocarbons from a liquefied natural gas (LNG) stream, comprising:

providing a first stream of liquid LNG having a pressure of at least 600 psig;

passing the first stream through a first heat exchanger to vaporize the LNG and form a second stream;

expanding the second stream in an turbo-expander to form a third stream having reduced pressure and temperature as compared to the second stream;

introducing a first portion of the third stream to a distillation column at one or more feed trays;

passing a second portion of the third stream through the first heat exchanger to be cooled by heat exchange with the first stream to form a fourth stream;

introducing the fourth stream to the distillation column at a point higher than the first portion of the third stream inlet;

withdrawing a fifth stream of vapor hydrocarbon from an upper portion of the distillation column having a reduced content of C2+ or C3+ hydrocarbons as compared to the first stream;

compressing the fifth stream in a turbo-compressor coupled to and powered by the turbo-expander;

withdrawing a sixth stream of liquid hydrocarbon from a lower portion of the distillation column having an increased content of C2+ or C3+ hydrocarbons as compared to the first stream.

2. The process of claim 1, wherein said distillation column is operated with a bottom temperature of about 0° F. to about 250° F. and a top temperature of about −200° F. to about 0° F.

3. The process of claim 1, wherein said column is operated at a pressure of about 150 psig to about 550 psig.

4. The process of claim 1, wherein the fourth stream provides from 1% to 50% of the total feed to the distillation column.

5. The process of claim 1, wherein the fourth stream provides from 20% to 40% of the total feed to the distillation column.

6. The process of claim 1, wherein the expansion of the second stream in the turbo-expander to form the third stream generates work and said work is utilized to compress the fifth stream in the turbo-compressor.

7. The process of claim 1, wherein the turbo-compressor provides at least 30% of the required compression for the fifth stream to reach a pipeline pressure.

8. The process of claim 1, wherein the turbo-compressor provides at least 50% of the required compression for the fifth stream to reach a pipeline pressure.

9. The process of claim 1, wherein the turbo-compressor provides less than 100% of the required compression for the fifth stream to reach a pipeline pressure.

10. The process of claim 1, wherein the second stream entering the turbo-expander is at a higher temperature than the fifth stream entering the turbo-compressor.

11. The process of claim 1, wherein the temperature of the second stream entering the turbo-expander is at least 50° F. higher temperature than the fifth stream entering the turbo-compressor.

12. The process of claim 1, wherein the temperature of the second stream entering the turbo-expander is at least 100° F. higher temperature than the fifth stream entering the turbo-compressor.

13. The process of claim 1, wherein the fifth stream before and after the turbo-compressor has either a pressure of at least 700 psig, or a temperature greater than −50° F., or both.

14. A process for the separation of C2+ or C3+ hydrocarbons from a liquefied natural gas (LNG) stream, comprising:

vaporizing a first LNG stream to provide a first stream of vaporized natural gas having a pressure of at least 600 psig;

expanding the first stream in an turbo-expander to form a second stream having reduced pressure and temperature as compared to the first stream;

introducing the second stream to a distillation column at one or more feed trays;

providing a third stream of liquid LNG having a pressure of at least 150 psig;

introducing the third stream to the distillation column at a point higher than the second stream inlet;

withdrawing a fourth stream of vapor hydrocarbon from an upper portion of the distillation column having a reduced content of C2+ or C3+ hydrocarbons as compared to the first stream;

compressing the fourth stream in a turbo-compressor coupled to and powered by the turbo-expander;

withdrawing a fifth stream of liquid NGL hydrocarbon from a lower portion of said distillation column having an increased content of C2+ or C3+ hydrocarbons as compared to the first stream.

15. The process of claim 14, wherein said distillation column is operated with a bottom temperature of about 0° F. to about 250° F. and a top temperature of about −200° F. to about 0° F.

16. The process of claim 14, wherein said column is operated at a pressure of about 150 psig to about 550 psig.

17. The process of claim 14, wherein the second stream provides from 1% to 50% of the total feed to the distillation column.

18. The process of claim 14, wherein the second stream provides from 20% to 40% of the total feed to the distillation column.

19. The process of claim 14, wherein the expansion of the first stream in the turbo-expander to form the second stream generates work and said work is utilized to compress the fourth stream in the turbo-compressor.

20. The process of claim 14, wherein the turbo-compressor provides at least 30% of the required compression for the fourth stream to reach a pipeline pressure.

21. The process of claim 14, wherein the turbo-compressor provides at least 50% of the required compression for the fourth stream to reach a pipeline pressure.

22. The process of claim 14, wherein the turbo-compressor provides less than 100% of the required compression for the fourth stream to reach a pipeline pressure.

23. The process of claim 14, wherein the first stream entering the turbo-expander is at a higher temperature than the fourth stream entering the turbo-compressor.

24. The process of claim 14, wherein the temperature of the first stream entering the turbo-expander is at least 50° F. higher temperature than the fourth stream entering the turbo-compressor.

25. The process of claim 14, wherein the temperature of the first stream entering the turbo-expander is at least 100° F. higher temperature than the fourth stream entering the turbo-compressor.

26. The process of claim 14, wherein the fourth stream before and after the turbo-compressor has either a pressure of at least 700 psig, or a temperature greater than −50° F., or both.