US20140060114A1

US20140060114A1 - Configurations and methods for offshore ngl recovery

Info

Publication number: US20140060114A1
Application number: US14/014,226
Authority: US
Inventors: John Mak
Original assignee: Fluor Technologies Corp
Current assignee: Fluor Technologies Corp
Priority date: 2012-08-30
Filing date: 2013-08-29
Publication date: 2014-03-06
Also published as: KR20150102931A; JP6289471B2; WO2014036322A1; JP2015531851A

Abstract

A natural gas two-column processing plant allows for recovery of at least 95% of C4 and heavier hydrocarbons, and about 60 to 80% of C3 hydrocarbons from a rich feed gas stream in which the first column (absorber) operates at a higher pressure than the second column, with the absorber receiving a compressed gas from the second column, and a turboexpander discharging a two-phase stream to the top of the absorber. Most typically, contemplated configurations and methods operate without the use of external refrigeration.

Description

This application claims priority to U.S. provisional application with the Ser. No. 61/694,949, which was filed on Aug. 30, 2012, and which is incorporated by reference herein.

FIELD OF THE INVENTION

The field of the invention is removal and recovery of natural gas liquids (NGL) from feed gases to meet pipeline hydrocarbon dew point and heating value specifications, especially for offshore applications.

BACKGROUND OF THE INVENTION

Numerous systems and methods are known in the art to recover C2, C3, and heavier components from natural gas, but all or almost all of them are configured for high recovery (i.e., over 90%) of NGL and require use of a turboexpander and deep refrigeration, which are costly and can only be economically justified if there are significant downstream markets. However, this is not case with offshore NGL recovery systems where space is at a premium and economic viability of the installation depends on a relatively small footprint and low operating and capital cost. Therefore, in all or almost all cases, high capital investment and operating costs required for high recovery typically cannot be justified. On the other hand, pipeline operators are required to produce a sales gas to meet the pipeline specification on the hydrocarbon dew point and heating value for safety in transmission. In most cases, recovery of over 95% of the C4 and heavier hydrocarbons from the feed gas may be required, while C3 recovery can be as low as 60%, and C2 recovery is incidental, and can be as low as 30%. In view of the changed demand, the complexity of currently known NGL processing plants that allow over 90% C3 recovery is excessive and often cannot be justified from an economical perspective.
Numerous NGL processing plants with high NGL recovery from a feed gas include cryogenic fractionation and turbo-expansion processes as described in U.S. Pat. No. 4,157,904 to Campbell et al., U.S. Pat. No. 4,251,249 to Gulsby, U.S. Pat. No. 4,617,039 to Buck, U.S. Pat. No. 4,690,702 to Paradowski et al., U.S. Pat. No. 5,275,005 to Campbell et al., U.S. Pat. No. 5,799,507 to Wilkinson et al., and U.S. Pat. No. 5,890,378 to Rambo et al., and U.S. Pat. App. No. 2002/0166336 to Wilkinson et al., and WO 2011/126710 to Johnke et al. These and all other extrinsic materials discussed herein are incorporated by reference in their entirety. Where a definition or use of a term in an incorporated reference is inconsistent or contrary to the definition of that term provided herein, the definition of that term provided herein applies and the definition of that term in the reference does not apply.
While all of these processes can achieve very high NGL recovery, several difficulties still remain. Among other things, the NGL recovery processes use high expansion ratio turboexpanders to produce low levels of refrigeration, which requires recompression of the residue gas. Moreover, when processing a rich gas stream with relatively high levels of C5+ hydrocarbons, additional external refrigeration is often required. Typically, such process configurations are complex and are difficult to operate. For example, Campbell et al. describe in U.S. Pat. No. 6,182,469 a plant in which feed gas is cooled in a heat exchanger using cold residue gas and side reboilers as depicted in Prior Art FIG. 1. The condensed feed gas liquids are then separated in a separator and fed to the demethanizer. Alternatively, as described by Sorensen in U.S. Pat. No. 5,953,935, an absorber may be added upstream of a demethanizer as depicted in Prior Art FIG. 2. In such configurations, the liquids from the feed separator and the absorber bottoms are fed to the demethanizer. To further increase NGL recovery in such configurations, the absorber overhead is cooled and refluxed by chilling with the demethanizer overhead vapor.
In still further known configurations, as described in U.S. Pat. No. 6,244,070 to Lee et al. and U.S. Pat. No. 5,890,377 to Foglietta, the reboiler duties are integrated in feed chilling, and in these configurations, liquids from the intermediate separators are fed to various positions in the downstream demethanizer for NGL recovery. These processes also include various means of providing cooling to the NGL processes. Exemplary known configurations following such schemes are depicted in Prior Art FIGS. 3 and 4. While such complex configurations are suited to achieve high C2 and C3 recovery to over 95%, they tend to be cost prohibitive and not suitable for offshore applications.
Therefore, although various configurations and methods are known to recover NGL from a feed gas, all or almost all of them suffer from one or more disadvantages when dewpointing and moderate C3 recovery is required. Therefore, there is still a need to provide methods and configurations for improved NGL recovery.

SUMMARY OF THE INVENTION

The inventive subject matter is directed to configurations and methods of recovery of C4 and heavier hydrocarbons, and moderate recovery (up to 90%) of C3 from a gas stream to meet hydrocarbon dew point and heating value specification of a pipeline gas produced from the gas stream.
In one preferred aspect of the inventive subject matter, two columns are operated at different pressures with the first column (absorber) operating at a relatively high pressure of about 550 psig and with the second column (fractionator) operating at about 450 psig. By operating the absorber at relatively high pressure, the compression ratio of the residue gas is reduced, thereby minimizing the overall compression horsepower. With the fractionator operating at about 450 psig, it should be noted that the separation of methane from the ethane and heavier components can be accomplished with less heating requirement due to the favorable relative volatility between components, resulting in a smaller diameter column.
In another preferred aspect of the invention, the vapor stream from the fractionator overhead is advantageously utilized for stripping in the absorber. In one embodiment of this process, the fractionator overhead stream is compressed and the “free” heat of compression is used to efficiently remove the methane components from the NGL from the absorber.
It should also be particularly appreciated that only the liquid portion of the expander discharge is used as reflux to the absorber, which is entirely different from heretofore known configurations and methods, which require the expander discharge to be fed to the mid or the lower section of the absorber, as illustrated in FIGS. 1-4. The expander discharge typically contains about 80% vapor, and by feeding the vapor portion in the top of the absorber, the vapor traffic in the mid- and lower portion of the absorber is significantly reduced, and hence the size of the absorber is smaller. In heretofore known configurations and methods where the expander is discharged to the lower section of the absorber, the column has to be designed to handle the total flow, and not just the liquid flow as presented herein. For example, the size of the absorber in a currently known gas plant is typically 12 ft in diameter for a 1,000 MMscfd feed gas as compared to the absorber size of 10 ft in diameter using configurations and methods presented herein, which significantly reduces space requirement, associated equipment cost and weight, which are of primary importance in an offshore environment.
Additionally, it should be recognized that the second fractionator operates at a lower pressure and temperature, which is not only more efficient in terms of separation, but also allows the use of residue gas compression heat for reboiling the fractionator, thereby eliminating steam requirement or hot oil heating of heretofore known systems and methods.
It is further generally preferred that the fractionator is operated at a pressure between 450 to 550 psig, and that the overhead vapor is compressed to the absorber pressure that is at least 50 psi, and more typically at least 100 psi, and mostly typically at 155 psi higher than the absorber, and that the compressor discharge vapor has a temperature and volume that is sufficient for use as a stripping vapor to the absorber.
Additionally, contemplated methods will also include a step of expanding the vapor phase in a turbo expander and reducing pressure of the liquid phase in a second expansion device before feeding the liquid phase to a feed exchanger. While not limiting to the inventive subject matter, it is typically preferred that the feed gas cooling is performed without use of external refrigeration. In yet another step, the bottom of the absorber is also letdown in pressure via a JT valve providing additional chilling to the feed gas in the feed exchanger.
In another preferred aspect of the inventive subject matter, a processing plant for hydrocarbon dew point control of a natural gas feed gas delivered from a feed gas source (e.g., LNG import terminal, regasification etc.) will include a feed gas exchanger that is fluidly coupled to the feed gas source and configured to cool the feed gas using a liquid phase of the cooled feed gas and an bottom product of an absorber. Contemplated plants will also include a phase separator that is fluidly coupled to the feed gas exchanger and that is configured to separate the cooled feed gas into the liquid phase and a vapor phase. Most typically, the fractionator comprises a top section that is configured to produce a vapor phase that is compressed and used as a stripping gas in the absorber.
Various objects, features, aspects and advantages of the present invention will become apparent from the following detailed description of preferred embodiments of the invention, along with the accompanying drawings.

BRIEF DESCRIPTION OF THE DRAWINGS

Prior Art FIG. 1 is a schematic of one known configuration for NGL recovery in which feed gas is cooled in a heat exchanger using cold residue gas and side reboilers.

Prior Art FIG. 2 is a schematic of another known configuration for NGL recovery in which an absorber/fractionator column is positioned upstream of a demethanizer.

Prior Art FIG. 3 is a schematic of yet another known configuration for NGL recovery in which reboiler and feed gas compression are integrated in feed chilling.

Prior Art FIG. 4 is a schematic of a further known configuration for NGL recovery in which reboiler and compressed residue gas recycle are integrated in feed chilling.

FIG. 5 is a schematic of an exemplary configuration for NGL recovery according to the inventive subject matter.

FIG. 6 is a table listing calculated compositions of gas streams in the exemplary NGL recovery plant of FIG. 5.

DETAILED DESCRIPTION

The inventor has discovered various configurations and methods of NGL recovery in which capital and operating cost can be significantly reduced, and especially in offshore applications, where a rich feed gas is processed and where C4+ recovery with moderate C2 and C3 recovery is required. Among other advantages, contemplated configurations and methods significantly reduce complexity and cost by reducing the number of equipment services, by elimination of external refrigeration and external heating while lowering residue gas compression requirements.
In particularly preferred configurations and methods, the feed gas (typically a natural gas comprising C1, C2, C3, and C4, and heavier components) is cooled at relatively high pressure to thereby effect partial condensation. The vapor and liquid phases are then separated, with the liquid phase being expanded to a lower pressure to so provide cooling to the feed gas. After reduction in pressure, the liquid phase is fed to the lower section of a fractionation column, while the vapor phase is expanded via a turboexpander and fed into the top section of a first fractionator (absorber). As the absorber is operated at relatively high pressures (typically 550 to 650 psig), residue gas recompression requirements are significantly reduced.
One exemplary plant configuration is depicted in FIG. 5, in which wet feed gas 1 at a pressure of about 1,000 psig and a temperature of about 100° F., having a typical composition as shown in the table of FIG. 6, is dried in a molecular sieve drier 51, forming stream 2. The so dried gas stream 2 is cooled to a temperature of about −65° F. in exchanger 52, forming stream 3, utilizing the refrigeration content from residue gas stream 9 and liquid streams 6 and 11. The so chilled gas stream 3 is then separated in phase separator 53 into a liquid portion, stream 5, and a vapor portion, stream 4.
The liquid portion 5 is letdown in pressure via JT valve 54 to a pressure of about 475 psig, chilled to about −106° F. forming stream 6, which is heated in exchanger 52 to about 70° F. prior to entering as stream 7 to the lower section of fractionator 59. The vapor portion 4 is expanded via the turboexpander 55 to about 550 psig at about −109° F. to form stream 8, which is fed to the top of absorber 70. As used herein, the term “about” in conjunction with a numeral refers to a range of that numeral starting from 20% below the absolute of the numeral to 20% above the absolute of the numeral, inclusive. For example, the term “about −150° F.” refers to a range of −120° F. to −180° F., and the term “about 1500 psig” refers to a range of 1200 psig to 1800 psig. Moreover, and unless the context dictates the contrary, all ranges set forth herein should be interpreted as being inclusive of their endpoints, and open-ended ranges should be interpreted to include commercially practical values. Similarly, all lists of values should be considered as inclusive of intermediate values unless the context indicates the contrary. It is noted that in some embodiments, the energy of expansion of gas within the turboexpander 55 may used to drive compressor 56 or other device to recover expansion energy. In some embodiments, the energy of expansion in turboexpander 55 can be used to also drive compressor 57.
It is contemplated that the vapor portion 4 is being expanded via the turboexpander 55 in such a way that partially condenses vapor portion 4 to produce a two-phase stream 8 comprising a vapor phase and a liquid phase. In some embodiments, at least 5 vol. % of stream 8 is in the vapor phase. In some embodiments, at least 10 vol. % of stream 8 is in the vapor phase. In yet some embodiments, at least 20 vol % of stream 8 is in the vapor phase. In yet some other embodiments, at least 30 vol % of stream 8 is in the vapor phase. In yet some embodiments, at least 40 vol % of stream 8 is in the vapor phase. In yet some embodiments, at least 60 vol % of stream 8 is in the vapor phase. In yet some embodiments, at least 80 vol % of stream 8 is in the vapor phase. The remainder of the expanded stream is in the liquid phase to serve as a reflux stream. Therefore, in some embodiments, at least 5 vol % of stream 8 is in a liquid phase. In yet some embodiments, at least 20 vol % of stream 8 is in a liquid phase. In yet some other embodiments, at least 30 vol % of stream 8 is in a liquid phase. In yet some embodiments, at least 40 vol % of stream 8 is in a liquid phase. In yet some embodiments, at least 60 vol % of stream 8 is in a liquid phase. In yet some embodiments, at least 80 vol % of stream 8 is in a liquid phase.
The operating pressure of the absorber 70 is in the range of about 550 to about 650 psig or higher, and the top section temperature is about −100° F., and the bottom section is about −15° F. It should be noted that only the liquid portion from the expander discharge is used as the reflux and the vapor portion forms part of the residue gas. The absorber is stripped with hot compressor discharge stream 16 from the fractionator column 59.
In some embodiments, the overhead gas stream 9 comprises the residue gas from the absorber 70 and at least some of the vapor portion of stream 8. In some embodiments, overhead gas stream 9 has a methane content of about 95 mol %. Overhead gas stream 9 that comes out of absorber 70 has a low temperature (at about −100° F.), and the refrigeration content of the overhead gas stream 9 is used to chill natural gas feed 2. The absorber bottom stream 10 is letdown in pressure to about 450 psig and chilled to −14° F., forming stream 11, and the refrigerant content is used to chill the feed gas in exchanger 52 to form stream 21. The heated gas is flashed to the top of the fractionator column 59.
After being used to chill natural gas feed 2, the warmed gas stream 17 that comes out of the heat exchanger 52 is being compressed by compressor 56, and turns into compressed gas stream 18. In some embodiments, gas stream 18 is further compressed by compressor 57 to form compressed gas stream 19, which is used for reboiling products from the fractionator 59 in reboiler 62. So cooled residue gas stream 15 is then fed to air cooler 58 prior to leaving the plant as residue gas stream 20 (e.g., as pipeline gas). It should be recognized that such configuration does not require external heating or fuel gas heater while producing on spec product which is advantageous for offshore operation and eliminating noxioius or otherwise undesireable emissions.
Fractionator 59 uses reboiler 62 to maintain the methane content in the bottom liquid stream 12 to preferably no more than 2 mol % or as required to meet the vapor pressure specification of the NGL product. Because of the relatively low operating pressure in the fractionator, the reboiler can use the low temperature compression heat from the residue gas compressor discharge stream 19 for reboiling fractionator bottom product 13, eliminating external heating requirement. The fractionator 59 is configured to produce a fractionator overhead product 14 that is passed to compressor 63. As mentioned above, the compressed stream 16 is then passed into the bottom section of the absorber 70, while a portion of the bottom product leaves as C2+ NGL product stream 12.
With respect to the feed gas it is generally contemplated that suitable feed gases will include C1, C2 and C3+, and may further comprise N2 and CO2. Consequently, it should be appreciated that the nature of the feed gas may vary considerably, and all feed gases in plants are considered suitable feed gases no long as they comprise C1 and C3 components, and more typically C1 to C5 and heavier components, and most typically C1 to C6 and heavier components. Therefore, particularly preferred feed gases include natural gas (e.g., after regasification from LNG, after CO2 removal where produced from a gas well), refinery gas, and synthetic gas streams obtained from other hydrocarbon materials such as coal, crude oil, naphtha, oil shale, tar sands, and lignite. Suitable gases may also contain relatively lesser amounts of heavier hydrocarbons such as propane, butanes, pentanes and the like, as well as hydrogen, nitrogen, carbon dioxide and other gases. Depending on the particular feed gas, the pressure of the feed gas may vary. However, it is generally preferred that the feed gas has a pressure between about 700 psig to about 1400 psig, and more typically between about 900 psig to about 1200 psig.
With respect to most suitable applications, contemplated configurations and methods use a single fractionator to recover at least 95% of the C4 and heavier hydrocarbons, and 60% to 80% of the C3 component, and 20% to 50% C2 component, without the use of external refrigeration. Therefore, it should be noted that feed gas cooling and/or cooling of the vapor product are performed without use of external refrigeration (e.g., at least 90% of refrigeration requirements are produced from expansion of process streams). It should also be recognized that white a single column configuration can also be used with two separate columns stacked on top of each other, with functions corresponding to the absorber and fractionator are also deemed suitable for use herein. It is still further contemplated that the dryer, separator, fractionator, heat exchanger, JT-valves, residue gas compressor, and turboexpander used in present configurations and methods are conventional devices well known to the skilled artisan.
Among other advantages of contemplated configurations, it should be particularly recognized that the phase separator produces a C5+ enriched liquid and a C5+ depleted vapor from a feed gas. Thus, so produced C5 enriched liquids may advantageously be fractionated in the lower section of the fractionator to meet the product liquid specification. Additionally, it should be recognized that by using a feed cooler and feed phase separator, and further by cooling of the vapors from the feed cooler and separation of the cooled vapors in the separator (to form a C5+ enriched liquid and a C5+ depleted vapor) most, if not all of the heavier components are removed from the feed gas. Consequently, the composition of the material flowing through the cold section is substantially stabilized as processing of heavy components in the feed gas in the upper section of the fractionator can be eliminated. Therefore, the heat duties, the turbo expander, and the fractionator will operate at the most efficient points. Thus, contemplated configurations and processes allow handling of a rich feed gas composition, thereby eliminating the complexity of a refrigeration unit of most prior arts. Viewed from another perspective, contemplated processes maintain constant operating conditions for the NGL recovery plant by removal of the C5+ components in the feed gas.
According to previously performed calculations (data not shown), contemplated configurations will achieve at least 60%, and more typically 78% propane recovery, and at least 85%, and more typically 95% butane recovery (see FIG. 6). Further contemplations, configurations, and methods suitable for use herein are described in U.S. Pat. Nos. 6,601,406, 6,837,7070, 7,051,552, 7,051,552 and 7,377,127, all of which are incorporated by reference herein. It is contemplated that a natural gas processing plant does not have to include all of the features described above to achieve efficiency in NGL recovery. Thus, a natural processing plant may include only a subset of the features described above. In some of these embodiments, the natural processing, plant may also include additional features that are not disclosed herein.
For example, a natural gas processing plant of some embodiments may include a turboexpander and an absorber. The turboexpander is configured to reduce pressure of a vapor stream to generate a two-phase stream having a liquid phase and a vapor phase. The absorber is configured to receive the two phase stream in a position such as to allow use of the liquid phase as a reflux. The absorber is further configured to produce an absorber overhead product and an absorber bottom product. Preferably but not necessarily, the vapor stream that enters into the turboexpander comprises natural gas feed that is cooled by a heat exchanger. In some of these embodiments, the absorber overhead product is led back into the heat exchanger in which refrigeration content of the absorber overhead product is used to chill the natural gas stream.
Preferably but not necessarily, after being used to chill the natural gas feed, the absorber overhead product is being compressed and used to reboil content within a fractionator. Moreover, the vapor stream that enters into the turboexpander comprises natural gas feed that is cooled by a heat exchanger. In some of these embodiments, the absorber bottom product is recycled back into the heat exchanger in which refrigeration content of the absorber bottom product is used to chill the natural gas stream.
It should be apparent to those skilled in the art that many more modifications besides those already described are possible without departing from the inventive concepts herein. The inventive subject matter, therefore, is not to be restricted except in the spirit of the appended claims. Moreover, in interpreting both the specification 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 the referenced elements, components, or steps may be present, or utilized, or combined with other elements, components, or steps that are not expressly referenced. Where the specification claims refers to at least one of something selected from the group consisting of A, B, C . . . and N, the text should be interpreted as requiring only one element from the group, not A plus N, or B plus N, etc.

Claims

What is claimed is:

1. A method of processing a natural gas stream, comprising:

cooling the natural gas stream and separating the cooled natural gas stream into a vapor portion and a liquid portion;

using a turboexpander to reduce pressure of the vapor portion to thereby generate a two-phase stream having a liquid phase and a vapor phase;

feeding the two phase stream to an absorber such that the liquid phase is a reflux in an absorber that produces an absorber overhead product and an absorber bottom product;

reducing the bottom product in pressure and feeding the bottom product after pressure reduction into a fractionator that produces a fractionator bottom product and a fractionator overhead product; and

compressing the fractionator overhead product and using the compressed fractionator overhead product as a stripping gas in the absorber.

2. The method of claim 1, further comprising a step of using refrigeration content of the absorber overhead product for the step of cooling the natural gas stream.

3. The method of claim 1 or claim 2, further comprising a step of using refrigeration content of the liquid portion and the bottom product after pressure reduction for the step of cooling the natural gas stream.

4. The method of claim 1 further comprising a step of compressing the absorber overhead product, and using heat content of the compressed absorber overhead for reboiling the fractionator.

5. A natural gas processing plant, comprising:

a heat exchanger configured to cool a natural gas stream, and a phase separator that is configured to receive and separate the cooled natural gas stream into a vapor portion and a liquid portion;

a turboexpander coupled to an absorber and phase separator, and configured to reduce pressure of the vapor portion to thereby generate a two-phase stream having a liquid phase and a vapor phase;

wherein the absorber is configured to receive the two phase stream in a position such as to allow use of the liquid phase as a reflux, and wherein the absorber is further configured to produce an absorber overhead product and an absorber bottom product;

a pressure reduction device that is fluidly coupled to the absorber and configured to reduce the bottom product in pressure;

a fractionator that is configured to receive the bottom product after pressure reduction and that is further configured to produce a fractionator bottom product and a fractionator overhead product;

a compressor fluidly coupled between the fractionator and the absorber, wherein the compressor is configured to receive and compress the fractionator overhead product and to provide the compressed fractionator overhead product as a stripping gas in the absorber.

6. The plant of claim 5, wherein the heat exchanger is configured such as to allow use of refrigeration content of the absorber overhead product to cool the natural gas stream.

7. The plant of claim 6 or claim 7, wherein the heat exchanger is further configured such as to allow use of refrigeration content of the liquid portion and the bottom product after pressure reduction to cool the natural gas stream.

8. The plant of claim 5 further comprising a residue gas compressor that is configured to compress the absorber overhead product, and wherein the fractionator uses a reboiler that is configured to use heat content of the compressed absorber overhead for reboiling the fractionator.

9. A method of operating an absorber in a natural gas processing plant, comprising a step of feeding a two-phase stream having a liquid phase and a vapor phase into a top portion of the absorber such that the liquid phase operates as a reflux stream, wherein the two-phase stream is produced by expansion of a vapor portion of a natural gas stream.

10. The method of claim 9 wherein the expansion of the vapor portion of the feed gas is performed using a turboexpander.

11. The method of claim 9 wherein the natural gas stream is a rich natural gas stream with at least 3% C3+ and wherein the natural gas stream has a pressure of at least 1000 psig.

12. A method of operating a fractionator reboiler of a natural gas processing plant having an absorber, and a fractionator, comprising:

using the absorber to produce an absorber overhead product, and compressing the absorber overhead product to a delivery pressure;

using heat content from the compressed absorber overhead product in the fractionator reboiler.

13. The method of claim 12, further comprising a step of heating the absorber overhead product in a heat exchanger using heat from a natural gas feed stream prior to the step of compressing the absorber overhead product.

14. A natural gas processing plant, comprising:

turboexpander configured to reduce pressure of a vapor stream to generate a two-phase stream having a liquid phase and a vapor phase, wherein the two-phase stream is produced by expansion of a vapor portion of a natural gas stream; and

an absorber coupled to the turboexpander, and configured to receive the two phase stream in a position such as to allow use of the liquid phase as a reflux, and wherein the absorber is further configured to produce an absorber overhead product and an absorber bottom product.

15. The natural gas processing plant of claim 14, further comprising:

a heat exchanger configured to use the absorber overhead product to cool the natural gas stream, and a phase separator that is configured to receive and separate the cooled natural gas stream into the vapor portion and a liquid portion.

16. The natural gas processing plant of claim 14, wherein the natural gas stream is a rich natural gas stream with at least 3% C3+ and wherein the natural gas stream has a pressure of at least 1000 psig.

17. The natural gas processing plant of claim 14, wherein the absorber overhead product is compressed and the compressed overhead product is used for reboiling a fractionator.

18. The natural gas processing plant of claim 17, wherein a fractionator overhead product from the fractionators is compressed, and the compressed fractionators overhead product is fed into the absorber as a stripping gas.

19. A natural gas processing plant, comprising:

an absorber configured to produce an absorber overhead product and compress the absorber overhead product to a delivery pressure; and

fractionator configured to use heat content from the compressed absorber overhead product to reboil content within the fractionator.

20. The natural gas processing plant of claim 19, wherein the absorber overhead product is heated in a heat exchanger using heat front a natural gas feed stream prior to entering into the absorber.

21. A method of processing a natural gas stream, comprising:

cooling the natural gas stream and separating the cooled natural gas stream into a vapor portion and a liquid portion; and

feeding the cooled vapor portion to an absorber that produces an absorber overhead product and an absorber bottom product, wherein refrigeration content of the absorber overhead product is used for the step of cooling the natural gas stream.

22. The method of claim 21, wherein the step of cooling the natural gas stream comprises using a turboexpander to reduce pressure of the vapor portion.

23. A natural gas processing plant, comprising:

a heat exchanger configured to coot a natural gas stream, and a phase separator that is configured to receive and separate the cooled natural gas stream into a vapor portion and a liquid portion;

an absorber configured to receive the vapor portion and produce an absorber overhead product, wherein refrigeration content of the absorber overhead product is used for cooling the natural gas stream in the heat exchanger.

24. The natural gas processing plant of claim 23, further comprising a turboexpander configured to reduce pressure of the vapor portion before the vapor portion reaching the absorber.