NGL Extraction From Natural Gas

Abstract

A process of extraction of natural gas liquids from a hydrocarbon stream is disclosed. The recovered natural gas liquids are flashed to a vapor to provide refrigeration to the process.

Description

CROSS-REFERENCE TO RELATED APPLICATIONS

The subject application claims priority to provisional application No. 61/128,977 filed on May 27, 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 more efficiently and more economically separate ethane, propane, and other hydrocarbon liquids from a hydrocarbon gas stream, i.e., from natural gas or from gases from refineries or petrochemical plants.

BACKGROUND

Natural gas from the wellhead is a mixture of various different gases including methane, ethane, propane, butane, etc. Natural gas liquid (“NGL”) processing plants liquefy and extract these components since they are generally of greater value for purposes other than as a gaseous heating fuel. For example, C₂, C₃ and C₄ hydrocarbons are valuable chemical intermediates and the C₃ and C₄ hydrocarbons are 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. In addition to these NGL components, other components such as ethylene and propylene may be contained in gas streams obtained from refineries or from petrochemical plants that may have a higher value if extracted.

There are several different types of NGL plants including mechanical refrigeration, lean oil and cryogenic turboexpander designs. Of these, the cryogenic turboexpander plant is the current state of the art for deep ethane recovery. In the cryogenic turboexpander plant, the cryogenic refrigeration required in the process is primarily produced by the turboexpander. The turboexpander is typically a major expense in the cryogenic turboexpander design.

There are many known NGL recovery processes that can be designed to process a specific gas composition at specific inlet conditions. However, when the feed gas composition changes, NGL recovery will typically be reduced and potential product revenue can be lost. In general, and specifically in a turboexpander plant, the “turndown” capability, which is the ability to recover a reduced quantity of NGL's from the processed stream, is limited. In many plant designs if the quantity of NGL recovered is desired to be reduced, such as for economic reasons, the efficiency of the plant can be severely decreased. This can occur as a result of equipment such as a turboexpander or a turbine compressor that have limited operational conditions and have diminishing effectiveness below a certain RPM, for example.

A significant cost in NGL recovery processes is related to the refrigeration required to chill the inlet gas. Refrigeration schemes generally include external refrigeration cycles utilizing refrigerants, such as for example ethane or propane. In some applications, mixed refrigerants and cascade refrigeration cycles have been used. An external refrigeration cycle can be a major expense in a NGL recovery process. Refrigeration can also be provided by turboexpansion or work expansion of the compressed gas feed.

In view if the above it would be desirable to have a NGL recovery process that does not require high expense components such as a turboexpander or an external refrigeration cycle. It would also be desirable to have a process with increased efficiency and turndown capability.

SUMMARY

Embodiments of the present invention generally include processes for NGL extraction from a hydrocarbon gas stream wherein the need for a turboexpander is eliminated and the required process refrigeration is produced by expansion of the NGL liquid product. The process may herein be referred to as the SRNE (Self Refrigeration No Expander) process. The liquid NGL is vaporized and passes through a heat exchanger to chill the inlet gas to the necessary cryogenic temperature. The NGL product vapor exiting the heat exchanger is then compressed to a higher pressure and condensed to produce the desired liquid NGL product.

One embodiment of the present invention is a process for the separation of C₂+ or C₃+ hydrocarbons from a hydrocarbon-containing gas feed under pressure. The process includes cooling a first stream of hydrocarbon gas in a first heat exchanger and introducing the cooled first stream into a distillation column at one or more feed trays. The column has a plurality of liquid recovery trays for condensing hydrocarbon liquids in said recovery trays. A vapor hydrocarbon stream is withdrawn from an upper portion of the column having a reduced content of C₂+ or C₃+ hydrocarbons. A liquid NGL hydrocarbon stream is withdrawn from a lower portion of said column having an increased content of C₂+ or C₃+ hydrocarbons. The pressure of the NGL liquid hydrocarbon stream is reduced to vaporize the NGL and generate cold energy and the vaporized NGL stream passes through the first heat exchanger to chill the first stream of hydrocarbon gas.

The process can further comprise passing a portion of the vapor hydrocarbon stream from an upper portion of said column through the first heat exchanger. A portion of the first stream of hydrocarbon gas can be diverted through a second heat exchanger to provide reboiler duty to the column. The first heat exchanger can provide heat duty to a side reboiler of the column.

The column can be operated with a bottom temperature of 0° F. to 250° F. and a top temperature of −160° F. to 0° F. and at a pressure of 150 psia to 700 psia.

Embodiments of the process does not include a turboexpander, or optionally the process includes a turboexpander that expands at least a portion of the first stream.

The vaporization of at least a portion of the fourth stream can provide at least 50% of the required cooling of the process, optionally at least 75%, optionally 100%.

The process can be capable of operating at a turndown of less than 60% of plant capacity, while at least 50% of the ethane in the first stream is recovered in the fourth stream. Optionally the process is capable of operating from 5% to 100% of plant capacity, while recovering at least 60% of the ethane and at least 90% of C₃+ from the first stream in the fourth stream. Optionally the process is capable of operating from 5% to 100% of plant capacity, while recovering at least 85% of the ethane and at least 95% of C₃+ from the first stream in the fourth stream.

The process can operate with a reduced compression of at least 10% as compared to a comparable turboexpander process.

BRIEF DESCRIPTION OF DRAWINGS

FIG. 1 illustrates a simplified process flow diagram of an embodiment of the present invention.

FIG. 2 illustrates a simplified process flow diagram of an embodiment of the present invention.

FIG. 3 illustrates a simplified process flow diagram of an embodiment of the present invention.

DETAILED DESCRIPTION

One non-limiting embodiment of the design of the present invention is shown in FIG. 1. Inlet gas enters the plant at stream 1 and is routed to an Inlet Dehydration section 20 to remove the majority of any water. The Inlet Dehydration section 20 can comprise equipment and processes known in the industry, such as glycol dehydration systems and/or molecular sieve units, for example.

Depending on the desired NGL recovery, all or a majority of the dehydrated inlet gas is then directed to stream 2, which is the main inlet gas feed into a distillation column, herein referred to as a Demethanizer, T-1. If a higher ethane recovery is required, a portion of the inlet gas can be directed to stream 3 to be used as a reflux stream for the Demethanizer. Depending on the plant operating conditions the reflux stream 3 may need to be compressed in an Inlet Compression section 22, such as in one non-limiting example, if the inlet gas pressure is 550 psig or lower.

Streams 2 and 3 pass through the process exchanger, E-1 and are cooled before entering the Demethanizer. The process exchanger, E-1, can comprise one or more exchangers that can be configured in parallel, in series, or combinations thereof. Control valves CV-1 & CV-2 can decompress and further chill streams 2 and 3 prior to the Demethanizer.

The Demethanizer provides separation of the hydrocarbon components in the inlet gas via boiling point distillation. A vapor overhead stream 4 composed primarily of methane and having a reduced content of NGL's is removed from an upper portion of the Demethanizer. A liquid bottoms stream 5 of extracted NGL's is removed from a lower portion of the Demethanizer. The Demethanizer can be of any suitable design for the separation of hydrocarbons and are generally known to those in the industry. In one non-limiting embodiment the column is operated with a bottom temperature of from 0° F. to 350° F. and a top temperature of from −160° F. to 0° F. and is operated at a pressure of from 150 psia to 700 psia.

All or a portion of the Demethanizer overhead gas can pass through the process exchanger, E-1 before exiting the plant in stream 10. The Demethanizer overhead gas will generally be a lower temperature than the inlet gas and can be used to chill the inlet stream 2 and reflux stream 3. In one embodiment the Demethanizer overhead gas can pre-chill one or both of the inlet streams prior to the inlet streams cooling from heat exchange from stream 5. Residue gas compression 26 may be required to compress the vapor overhead stream 10 into the residue gas pipeline.

To provide the necessary process refrigeration, the liquid NGL product in stream 5 is flashed, such as through a static expansion valve, CV-3, to a lower pressure then routed to the process exchanger E-1 to chill the inlet gas streams. The NGL product vapor leaving the process exchanger in stream 6 is then compressed and condensed to produce the desired NGL liquid product stream. The NGL liquid product stream is then pumped out of the plant via stream 7, the flow regulated by valve CV-5. A portion of the NGL liquid product stream can be diverted from stream 7 to stream 5 to provide additional refrigeration capability if needed. In this embodiment, not shown, a portion of the NGL liquid product can cycle as needed based on refrigeration needs and operational considerations.

As shown in FIG. 1, a portion of reflux stream 3 can be directed to stream 8 to be used as a heat source for the Demethanizer and utilized as reboiler duty. In an alternate embodiment (not shown) a portion of feed stream 2 can be directed to stream 8 to be used as a heat source for the Demethanizer and utilized as reboiler duty rather than a portion of reflux stream 3. In another alternate embodiment (not shown) a portion of either or both of feed stream 2 and reflux stream 3 can be directed to stream 8 to be used as a heat source for the Demethanizer and utilized as reboiler duty. Control valve CV-4 can regulate the flow through stream 8 to provide thermal energy to one or more reboilers of the Demethanizer. Heat exchanger E-2 can exchange heat duty from stream 8 to warm the fluids within the lower portion of the Demethanizer to strip residual methane from the NGL liquid product stream. Stream 8 will generally be separated from reflux stream 3 after any compression 22 that may be required for reflux stream 3, as the compression will warm reflux stream 3. FIG. 1 shows two liquid draws from the Demethanizer passing through exchanger E-2, but alternate embodiments can contain one, or more than two, draws. The reboiler exchanger E-2 can comprise one or more exchangers that can be configured in parallel, in series, or combinations thereof.

A side stream draw from the Demethanizer is shown as stream 9, which is warmed by cross exchange with the process exchanger E-1, acting as a reboiler, to provide heat to the middle section of the Demethanizer and assist in the separation of NGL from methane. One or more side stream draws can be incorporated into the present invention, but are not required. In one embodiment of the present invention the design includes an upper side stream reboiler utilizing heat from the process exchanger E-1, a middle reboiler utilizing heat from a portion of the inlet feed, and a lower reboiler utilizing heat from a portion of the inlet feed, where the three reboilers provide a temperature profile within the Demethanizer that is preferred as compared to a design having only two reboilers.

Simulation data for three examples of Demethanizer design, each utilizing a total reboiler duty of 59 and providing a bottom reboiler out temperature of 63° F. is shown in Table 1. The first design having two reboilers as shown in FIG. 2 gives a middle reboiler out temperature of −16° F. and provides a temperature difference of 79° F. between the reboiler outlets. The second design having three reboilers as shown in FIG. 1 gives a middle reboiler out temperature of 40° F. and a top reboiler out temperature of −13° F., providing a temperature difference of 76° F. between the top and bottom reboiler outlets. The third design also has three reboilers but a different duty profile, giving a middle reboiler out temperature of 40° F. and a top reboiler out temperature of −54° F. The modified three reboiler design provides a more defined temperature profile within the Demethanizer and a greater temperature difference of 117° F. between the top and bottom reboiler outlets.

	TABLE 1
	Two		Three
	Reboilers	Three	Reboilers
	Only	Reboilers	(Modified)
			T ° F.			T ° F.			T ° F.
	Duty	T ° F. (in)	(out)	Duty	T ° F. (in)	(out)	Duty	T ° F. (in)	(out)
Top	na	na	na	22	−66	−13	27	−108	−54
Reboiler
Middle	35	−66	−16	22	25	40	16	25	40
Reboiler
Bottom	24	46	63	15	48	63	16	48	63
Reboiler

In one embodiment a portion of the Demethanizer overhead gas stream 10 can be used as an optional reflux stream 11 to the Demethanizer T-1 as shown in FIG. 3. The optional reflux stream 1 can be cooled in exchanger E-1, its flow controlled by control valve CV-6, and can enter the upper portion of the Demethanizer. If the optional reflux stream 11 is fed to the Demethanizer T-1, the reflux stream 3 will enter at a location below the optional reflux stream 11.

There can be several advantages of the present invention over the current state of the art turboexpander plants. There can be significant capital and maintenance savings produced by reducing or eliminating the turboexpander, the turboexpander inlet separator, and the turboexpander auxiliary systems.

The overall plant reliability can be improved by eliminating the turboexpander, which is a rotating device operating at high speeds, for example between 10,000-20,000 rpm, and can be subject to failure.

Embodiments of the present invention can produce the necessary cryogenic refrigeration significantly more efficiently than the turboexpander design. This is particularly true when the inlet gas pressure is at 550 psig or lower. Embodiments of the present invention can provide 100% of the required cooling. Alternate embodiments can provide at least 90% of the required cooling, optionally at least 80%, optionally at least 75%, optionally at least 50%.

In some low inlet pressure applications, the present invention can reduce the required compression horsepower by at least 10%, optionally at least 25%, and optionally at least 50%. By utilizing the refrigeration from the NGL expansion and reducing or eliminating the need for an external refrigeration system the initial capital expense and the ongoing fuel gas and plant maintenance expenses can be significantly reduced.

The plant operating range can be improved with the present invention as compared to a conventional turboexpander plant. Most turboexpander plants can only efficiently turn down to approximately 60% of the maximum inlet flow due to limitations in the expander ability to turn down. Embodiments of the present design, being not limited by an expander, can efficiently turn down to as low as 5% of the maximum inlet flow. Simulation data for an expander plant and a plant of the present invention is shown in Table 2 and Table 3 that illustrate the ability of the two designs to reduce plant processing capabilities, which can be referred to as “turn down” of the plant. The expander plant exhibits increased fuel usage as the throughput is reduced from 100% to 60%, and is not able to operate at levels of 40% or less. The design of an embodiment of the present invention, shown in Table 3, maintains fuel efficiency and recoveries at reduced throughput down to 10% of design capacity. The fuel efficiency is shown as Fuel Usage (BTU/Gal) in Tables 2 & 3, which is the BTU value of the fuel usage of the process per gallon of NGL recovered by the process.

	TABLE 2
	Expander Plant Design Capacity (%)
	100%	80%	60%	40%	20%	10%
Recovery %
Ethane	88%	88%	88%	na	na	na
Propane	98%	98%	98%	na	na	na
C4+	100%	100%	100%	na	na	na
Fuel Usage	6600	6798	7062	na	na	na
(BTU/Gal)

	TABLE 3
	SRNE Plant Design Capacity (%)
	100%	80%	60%	40%	20%	10%
Recovery %
Ethane	88%	88%	88%	88%	88%	88%
Propane	98%	98%	98%	98%	98%	98%
C4+	100%	100%	100%	100%	100%	100%
Fuel Usage	6400	6400	6400	6400	6400	6400
(BTU/Gal)

Embodiments of the present invention include a plant design capable of operating at throughputs of from 5% to 100% of plant design while achieving ethane recovery of at least 50%, optionally at least 60%, optionally at least 80%. Alternate embodiments can include an optional turboexpander located on the inlet feed stream that can provide cooling while operating from 100% down to 60% of plant design capacity while the turboexpander can be bypassed to enable the plant to continue to operate at throughputs below 60% of plant design capacity. Another embodiment is a process of the current design that reduces the net compression of the process by at least 10% as compared to a comparable turboexpander plant.

Use of the term “optionally” with respect to any element of a claim is intended to mean that the subject element is required, or alternatively, is not required. Both alternatives are intended to be within the scope of the claim. Use of broader terms such as comprises, includes, having, etc. should be understood to provide substrate for narrower terms such as consisting of, consisting essentially of, comprised substantially of, etc.

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 hydrocarbon-containing gas feed under pressure, comprising:

providing a first stream of hydrocarbon gas;

cooling at least a portion of the first stream of hydrocarbon gas in a first heat exchanger to form a second stream;

introducing the second stream into a distillation column;

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

withdrawing a fourth stream of liquid natural gas liquid (NGL) hydrocarbon from as compared to the first stream;

reducing the pressure of the fourth stream to vaporize at least a portion of the fourth stream to form a fifth stream;

passing the fifth stream through the first heat exchanger to cool the first stream of hydrocarbon gas.

2. The process of claim 1, further comprising:

passing at least a portion of the third stream through the first heat exchanger to form a sixth stream.

3. The process of claim 1, further comprising:

diverting a portion of the first stream through a second heat exchanger to provide reboiler duty to the distillation column.

4. The process of claim 1, wherein the first heat exchanger provides heat duty to a side reboiler of the distillation column.

5. The process of claim 1, wherein the distillation column is operated with a bottom temperature of 0° F. to 350° F. and a top temperature of −160° F. to 0° F.

6. The process of claim 1, wherein the distillation column is operated at a pressure of 150 psia to 700 psia.

7. The process of claim 1, wherein a portion of the sixth stream is diverted to an upper region of the distillation column as a reflux stream.

8. The process of claim 1, wherein the process does not include a turboexpander.

9. The process of claim 1, wherein the process includes a turboexpander that expands at least a portion of the first stream.

10. The process of claim 1, wherein the vaporization of at least a portion of the fourth stream provides at least 50% of the required cooling of the process.

11. The process of claim 1, wherein the vaporization of at least a portion of the fourth stream provides at least 75% of the required cooling of the process.

12. The process of claim 1, wherein the vaporization of at least a portion of the fourth stream provides 100% of the required cooling of the process.

13. The process of claim 1, wherein the process is capable of operating at a turndown of less than 60% of plant capacity, while at least 50% of the ethane in the first stream is recovered in the fourth stream.

14. The process of claim 1, wherein the process is capable of operating at a turndown of 10% of plant capacity, while at least 50% of the ethane in the first stream is recovered in the fourth stream.

15. The process of claim 1, wherein the process is capable of operating from 5% to 100% of plant capacity, while recovering at least 60% of the ethane and at least 90% of C3+ from the first stream in the fourth stream.

16. The process of claim 1, wherein the process is capable of operating from 10% to 100% of plant capacity, while recovering at least 85% of the ethane and at least 95% of C3+ from the first stream in the fourth stream.

17. The process of claim 1, wherein the process operates with a reduced compression of at least 10% as compared to a comparable turboexpander process.

18. A process for the separation of natural gas liquids from a hydrocarbon-containing gas feed comprising:

providing a first stream of hydrocarbon gas;

cooling at least a portion of the first stream of hydrocarbon gas in a first heat exchanger to form a second stream;

introducing the second stream into a distillation column;

diverting a portion of the first stream through a second heat exchanger to provide reboiler duty to the distillation column;

wherein the distillation column includes a middle reboiler stream and a lower reboiler stream that are warmed by heat exchange with the second heat exchanger to provide heat duty to the distillation column;

wherein the distillation column includes a side stream that is warmed by heat exchange with the first heat exchanger to provide heat duty to the distillation column;

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

passing at least a portion of the third stream through the first heat exchanger to form a sixth stream;

withdrawing a fourth stream of liquid natural gas liquid (NGL) hydrocarbon from a lower portion of said column having an increased content of C2+ or C3+ hydrocarbons as compared to the first stream;

reducing the pressure of the fourth stream to vaporize at least a portion of the fourth stream to form a fifth stream;

passing the fifth stream through the first heat exchanger to cool the first stream of hydrocarbon gas;

wherein the process does not include a turboexpander and the distillation column is operated with a bottom temperature of 0° F. to 350° F., a top temperature of −160° F. to 0° F. and a pressure of 150 psia to 700 psia;

wherein the vaporization of at least a portion of the fourth stream provides at least 75% of the required cooling of the process; and

wherein the process is capable of operating from 5% to 100% of plant capacity, while recovering at least 85% of the ethane and at least 95% of C3+ from the first stream in the fourth stream.

19. The process of claim 18, wherein a portion of the sixth stream is diverted to an upper region of the distillation column as a reflux stream.

20. The process of claim 18, wherein the process operates with a reduced compression of at least 10% as compared to a comparable turboexpander process.