US20090090049A1

US20090090049A1 - Process for producing liqefied natural gas from high co2 natural gas

Info

Publication number: US20090090049A1
Application number: US11/869,488
Authority: US
Inventors: Daniel Chinn; Chung-Nan Nancy Tsai; Paul F. Bryan; Brian M. Frankie; Kaman Ida Chan; Justin I-Ching Pan
Original assignee: Chevron USA Inc
Current assignee: Chevron USA Inc
Priority date: 2007-10-09
Filing date: 2007-10-09
Publication date: 2009-04-09
Also published as: AU2008310984A1; WO2009048869A1

Abstract

A process for producing LNG from high CO₂natural gas. The process includes: separating methane from a hydrocarbon stream containing CO₂to produce a methane-depleted hydrocarbon stream; subjecting the methane-depleted hydrocarbon stream to at least one separation process; and feeding at least one recycle stream from the at least one separation processes into the step for separating methane. The at least one separation process is selected from the group consisting of deethanizing, depropanizing, debutanizing and CO₂separating.

Description

FIELD OF THE INVENTION

The present invention relates to a process for producing liquefied natural gas (LNG) from high-CO₂natural gas. More particularly, the present invention relates to a hybrid distillation process for producing multiple products from high-CO₂natural gas, including LNG, ethane, propane, high-purity CO₂product, and a hydrocarbon condensate stream.

BACKGROUND OF THE INVENTION

Natural gas is a valuable, environmentally-friendly energy source. With gradually decreasing quantities of available or easily-refined crude oil, natural gas has become accepted as a cleaner alternative energy source. Natural gas may be recovered from natural gas reservoirs or as associated gas from crude oil reservoirs. Indeed, natural gas for use in the present process may be recovered from any process which generates light hydrocarbon gases.
Natural gas can be found all over the world. Much of the natural gas reserves found around the world are separate from oil and as new reserves are discovered and processed, growth in the LNG industry will continue. Countries with large natural gas reservoirs include Algeria, Australia, Brunei, Indonesia, Libya, Malaysia, Nigeria, Oman, Qatar, Thailand, and Trinidad and Tobago. Countries that import significant quantities of LNG include China, France, India, Italy, Japan, Malaysia, South Korea, Spain, Taiwan, United Kingdom and United States.
One of the key steps in producing liquefied natural gas (LNG) is the processing of natural gas to remove components such as CO₂, H₂S, H₂O, Hg and aromatics (benzene, toluene, xylene) to ppm levels prior to gas liquefaction. The acid gas components from natural gas (CO₂, H₂S) are normally removed using an aqueous amine process. Amine processes are well known in the art, and typically involve one packed/trayed column for absorption of CO₂and H₂S into the amine solution and a separate packed/trayed column where CO₂and H₂S are stripped (via steam and/or pressure let-down) from the amine solution. Amine units operate only under a narrow range of concentrations and acid gas loadings (at a given CO₂partial pressure in the gas phase) due to corrosion limitations. Because the required amine flowrate is proportional to the amount of CO₂that needs to be removed, amine absorption plants become progressively larger and more expensive with higher CO₂concentrations in the natural gas.
Gas-permeation membranes are a well-known alternative to amine systems in selectively removing CO₂from natural gas. Membranes rely on the pressure driving force of the permeating CO₂, and does not require the use of solvents. Membranes however, have the similar disadvantage of amine systems in that the CO₂is normally recovered at low pressure. Thus, in cases where CO₂must be reinjected, the compression requirements would also be high for membranes. Further, membranes cannot make a perfect separation between CO₂, and hydrocarbons; a small amount of hydrocarbon will always permeate with the CO₂. Thus, the ultimate purity of the CO₂rich product is limited to the order of 97% to perhaps 98%. Membranes are not able to produce a CO₂-rich stream of ultra-high purity, such as on the order of 99.5+%.
Once CO₂is removed from the gas, the CO₂must be captured, processed, sequestered or diverted to some end use. One option currently under study is the capture, compression and re-injection of the into a geologic formation (depleted reservoir, saline aquifer, coal beds, etc.). However, because the CO₂recovered from the overhead of an amine stripping column is slightly above atmospheric pressure, recompression of that CO₂to a state where it could be readily transported/reinjected may be economically impractical.
It is possible to produce LNG from high-CO₂gas using a combination of amine treating (with or without membranes) along with the other related separation processes known in the art (gas dehydration, mercury removal, scrub columns to remove heavy components) and liquefaction processes. Normally, each of these unit operations are conducted in series, with little or no process integration between them. Thus, there remains an opportunity for an improved, integrated process for making LNG, especially in cases where the natural gas contains high levels of CO₂and in situations where it is highly desirable to produce said CO₂at a suitable condition (pressure, purity) for reinjection/geologic sequestration. Further, there is an opportunity for being able to produce LNG with a range of heating values. In cases where the natural gas contains a significant fraction of ethane and propane, there may be an economic incentive to produce a separate, saleable high-purity ethane product instead of leaving most of the ethane into the final LNG product as is typically done in current art LNG production. Having a leaner (i.e., with a lower ethane and propane content) LNG product has several advantages: (1) Many LNG customers actually prefer to have leaner LNG, (2) Excess ethane and propane removed from the LNG may be sold as separate products at higher prices, (3) A greater proportion of the LNG ship's volume is made available for storing liquefied methane, and (4) LNG regasification terminals would not need to install ethane and propane removal units for heating value control.

SUMMARY OF THE INVENTION

The present invention achieves the advantage of a process for producing LNG from high-CO₂natural gas, with the flexibility of producing separated products such as ethane, propane, and a hydrocarbon condensate.
In an aspect of the invention, a process for producing LNG from high-CO₂natural gas includes the steps of: separating methane from a hydrocarbon feed stream containing CO₂to produce a methane-depleted hydrocarbon stream; subjecting the methane-depleted hydrocarbon stream to at least one separation process to produce a hydrocarbon recycle stream; and combining the hydrocarbon recycle stream with the hydrocarbon feed stream prior to separating methane from the hydrocarbon feed stream, wherein the at least one separation process is selected from the group consisting of deethanizing, depropanizing, debutanizing and CO₂separating.
Optionally, in the above process, the step of separating methane includes conducting a full liquid reflux on the separated methane vapor product.
Optionally, in the above process, the step of separating methane includes scrubbing and removing aromatics and heavy hydrocarbons from the hydrocarbon stream containing CO₂.
Optionally, the above process further includes the step of passing the methane to a liquefaction process.
Optionally, the above process further includes the step of passing the methane to a main cryogenic heat exchanger of a liquefaction plant.
Optionally, in the above process, the hydrocarbon recycle stream includes fractionated gas components passed from the at least one separation process.
Optionally, in the above process, the hydrocarbon recycle stream includes a hydrocarbon stream from a front slug catcher.
Optionally, in the above process, the methane contains less than about 100 ppm CO₂and less than about 3 ppm H₂S.
Optionally, in the above process, the step of separating the methane from the hydrocarbon stream is conducted at a pressure in the range of about 38 to about 45 bar, and at a temperature in the range of about −91° C. to about −84° C.
Optionally, in the above process, the step of deethanizing is conducted at a pressure in the range of about 35 to about 44 bar, and at a temperature in the range of about 4° C. at about 35 bar.
Optionally, in the above process, the step of depropanizing is conducted at a pressure in the range of about 17 to about 27 bar, and at a temperature in the range of about −2° C. to about 67° C.
Optionally, in the above process, the step of debutanizing is conducted at a pressure in the range of about 6 to 12 bar, and at a temperature in the range of about 40° C. to about 78° C.
Optionally, in the above process, the step of CO₂separating is conducted at a pressure in the range of about 28 to about 32 bar, and at a temperature in the range of about −6° C. to about −2° C.
Optionally, in the above process, the step of subjecting the methane-depleted hydrocarbon stream to at least one separation process includes blending back at least one principal overhead product with the methane for heating value adjustment.
Optionally, in the above process, the step of subjecting the methane-depleted hydrocarbon stream to at least one separation process includes feeding at least one principal overhead product stream to a fractionation train.
Optionally, in the above process, the step of subjecting the methane-depleted hydrocarbon stream to at least one separation process includes feeding a principal overhead product of ethane and carbon dioxide to an azeotrope separation process.
Optionally, in the above process, the step of subjecting the methane-depleted hydrocarbon stream to at least one separation process includes removing H₂S via adsorption from the methane-depleted hydrocarbon stream.
Optionally, in the above process, the step of removing H₂S is conducted at a pressure in the range of about 17 to 27 bar, and at a temperature in the range of about −2° C. to 67° C.
Optionally, in the above process, the step of subjecting the methane-depleted hydrocarbon stream to at least one separation process includes membrane-separating CO₂from the methane-depleted hydrocarbon stream.
Optionally, in the above process, the step of membrane-separating CO₂is conducted to produce a stream containing about 98 vol % CO₂.

DESCRIPTION OF THE DRAWINGS

FIG. 1 illustrates an embodiment of the present invention.

FIG. 2 illustrates another embodiment of the present invention.

FIG. 3 illustrates another embodiment of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

So that the above recited features and advantages of the present invention can be understood in detail, a more particular description of the invention, briefly summarized above, may be had by reference to the embodiments thereof that are illustrated in the appended drawings. It is to be noted, however, that the appended drawings illustrate only typical embodiments of this invention and are therefore not to be considered limiting of its scope, for the invention may admit to other equally effective embodiments.
Embodiments describing the process of the present invention are referenced in FIGS. 1 to 3.

First Embodiment

In an embodiment of the invention illustrated in FIG. 1, a hydrocarbon feed stream 101, having a composition as shown in TABLES 1 & 2 (simulated data), is fed to a demethanizer (DeC1) column 150. The DeC1 column 150 may be a packed or trayed-type distillation column equipped with a bottom reboiler, side reboilers, and a condenser, that is designed to process at least two feed streams: a light hydrocarbon feed gas stream and a heavy hydrocarbon liquid solvent stream. The operating pressure of the DeC1 column is in the range of about 38 to about 45 bar. The operating temperature of the overhead condenser is in the range of about −91 to about −84° C.
A hydrocarbon recycle stream 120 from the bottoms of a depropanizer (DeC3) column 160 is also fed into the DeC1 column 150. Essentially, the hydrocarbon recycle stream 120 is fractionated gas components passed from the plurality of separation processes as further described below. The hydrocarbon recycle prevents the CO₂from freezing and acts as a scrubbing agent to remove aromatics and other heavy hydrocarbons from the C1-rich product stream taken overhead.
Although this embodiment shows the hydrocarbon recycle stream 120 being fed from the depropanizer 160, it is also possible to feed a hydrocarbon recycle stream from a front slug catcher into the demethanizer 150. An example of a front slug catcher includes a three phase separator required in the oil and gas industry at an upstream position (typically near a gas wellhead) to separate gas/oil/water.
Methane is taken as the principal overhead product, stream 103, and has the composition as shown in TABLE 1. The separated methane contains less than about 100 ppm CO₂and less than about 3 ppm H₂S. Also, a full liquid reflux (not shown in figure) on the separated methane vapor product is fed back to the demethanizer 150.
The main portion of stream 103 is fed to a liquefaction process. Since the aromatics are reduced to such a low level and the temperature is very cold, the majority of stream 103 may be precooled further and eventually fed to a main cryogenic heat exchanger (MCHE) 170, which is a specially-designed heat exchanger that may be of the spiral-wound type, plate-and-frame type, or any other type known in the LNG art. The purpose of the MCHE is to reduce the temperature of the C1-rich product to a point where it may be readily liquefied, stored, and shipped as LNG. The final steps of the liquefaction process includes nitrogen rejection via endflash or a stripping column. The nitrogen-depleted LNG final product is then pumped to storage and ready to be shipped.
The bottoms product 102 will contain the ethane and heavier hydrocarbon liquids along with most of the CO₂which is fed to a CO₂column 155 (CO₂separating). The bottoms product 102 has a composition as shown in TABLE 1. Two of the components in this stream form an azeotrope system: carbon dioxide and ethane. The CO₂column 155 may be a packed or trayed-type distillation column equipped with a bottom reboiler, side reboilers, and a condenser, that is designed to process at least two feed streams: a hydrocarbon vapor stream and a hydrocarbon liquid stream. The operating pressure of the CO₂column 155 is in the range of about 28 to about 32 bar. The operating temperature of the overhead condenser is in the range of about −6 to about −2° C.
A hydrocarbon recycle stream 121 from the bottoms of the DeC3 column 160 is also fed into the CO₂column 155. The hydrocarbon recycle breaks the azeotrope formed by the carbon dioxide and ethane.
Carbon dioxide is taken as the principal overhead product, stream 105, and has the composition as shown in TABLE 1. Since stream 105 is a high purity CO₂stream, it is suitable for geologic reinjection or enhanced oil recovery (EOR).
The bottoms product 104 is fed to the DeC3 column 160. The bottoms product 104 has a composition as shown in TABLE 1. The DeC3 column 160 may be a packed or trayed-type distillation column equipped with a reboiler and condenser. The operating pressure of the DeC3 column 160 is in the range of about 17 to about 27 bar. The operating temperature of the overhead condenser is in the range of about −2 to about 29° C.
Ethane and propane are taken as the principal overhead product, stream 107, and has the composition as shown in TABLE 1.
Stream 107 is fed to an H₂S separator 180. The H₂S separator may be a fixed bed adsorber that may be regenerative or non-regenerative, or any other process known in the art for selective-removal of H₂S from hydrocarbon streams. An H₂S-rich stream 123 may have several alternative destinations, depending on the design basis. For example, stream 123 may be fed to a Claus plant for further conversion to elemental sulfur, burned in a thermal or catalytic oxidizer, blended with the CO₂-rich product for reinjection, or reinjected separately. The operating pressure of the H₂S separator is in the range of about 17 to 27 bar, and the operating temperature is in the range of about −2 to 29° C.
The principal overhead product output of the H₂S separator 180, stream 113, is fed to a small fractionation train 175 to recover sufficient ethane and propane for sales and refrigerant makeup in a liquefaction process. A portion of the ethane and propane, stream 116, from the small fractionation train 175 may be blended back with stream 103 for heating value adjustment. For example, in cases where it is desirable for the LNG product to have a higher heating value greater than that of pure methane (about 1000 BTU/SCF), additional ethane, propane, and/or butane may be added.
For situations where C2, C3, and C4 products are neither required for LNG heating value adjustment nor for refrigerant makeup, each of the components may be exported from the facility as separate, saleable products as shown in stream 114, 115, and 125.
The bottoms product 106 is fed to the hydrocarbon recycle streams 120 and 121 via stream 119, and to a debutanizer (DeC4) column 165, via stream 108. The bottoms product is output as stream 110. The DeC4 column 165 may be a packed or trayed-type distillation column equipped with a reboiler and condenser. The operating pressure of the DeC4 column 165 is in the range of about 6 to about 12 bar, and the operating temperature is in the range of about 40° C. to about 78° C. It is noted that the DeC4 column is only necessary if a separate, C4-rich product stream is desired. If the DeC4 column is omitted, a greater amount of C4 will be present in the hydrocarbon recycle streams 120, 121 and condensate export stream 108.
Butane is taken as the principal overhead product via stream 109.
Stream 109 is fed to an H₂S separator 185. The H₂S separator 185 may be a fixed bed adsorber that may be regenerative or non-regenerative, or the separator may be any other process known in the art for selective-removal of H₂S from hydrocarbon streams. An H₂S-rich stream 111 may have several alternate destinations, depending on the design basis. For example, stream 111 may be fed to a Claus plant for further conversion to elemental sulfur, burned in a thermal or catalytic oxidizer, blended with the CO₂-rich product for reinjection, or reinjected separately. The operating pressure of the H₂S separator 185 is about 10 bar, and the operating temperature is up to 80° C.
A portion (stream 117) of the output of the H₂S separator 185, stream 112, is blended back with stream 103 for heating value adjustment, and the remainder, stream 125, is fed to sales and refrigerant makeup of a liquefaction process.

TABLE 1

Summary of Major Streams Only (Compositions in mol %)

Stream No.
(FIG. 1)	C₁	C₂	C₃	C₄	C₄+	CO₂	N₂	H₂S

101	77.8	3.2	0.9	0.3	0.1	14.7	3.0	0.003
102	0.1	10.8	5.8	15.9	18.2	49.2	0
103	96.3	0	0	0	0	0	3.7
104	0	7.6	9.1	38.3	45	0	0
105	0.2	0.4	0.6	0	0	98.8	0
106	0	0	8	42.1	49.9	0	0
107	0	76.4	18.9	4.6	0.01	0.04	0	0.07
108	0	0	8	42.1	49.9	0	0
120	0	0	8	42.1	49.9	0	0
121	0	0	8	42.1	49.9	0	0

TABLE 2

Summary of Major Streams Only

Stream No.	Flowrate	Pressure	Temp
(FIG. 1)	(kgmole/hr)	(bar)	(° C.)

101	59,770	41	−67
102	17,870	39	22
103	48,250	38	−91
104	24,860	33	141
105	8,890	33	−2
106	22,380	17	117
107	2,477	17	−3
108	153	43	18
120	6,350	42	−62
121	15,880	43	18

Second Embodiment

In another embodiment of the invention as illustrated in FIG. 2, a hydrocarbon feed stream 201, having a composition as shown in TABLES 3 & 4 (simulated data), is fed to a DeC1 column 250. The DeC1 column 250 is similar to that described above. The operating pressure of the DeC1 column 250 is in the range of about 38 to about 45 bar. The operating temperature of the overhead condenser is in the range of about −91 to about −84° C.
A hydrocarbon recycle stream 228 from the bottoms of a DeC3 column 260 is also fed into the DeC1 column 250. The hydrocarbon recycle prevents the CO₂from freezing and acts as a scrubbing agent to remove aromatics and other heavy hydrocarbons from the C1-rich product stream taken overhead.
Methane is taken as the principal overhead product, stream 203, and has the composition as shown in TABLE 3. The separated methane contains less than about 100 ppm CO₂and less than about 3 ppm H₂S. Also, a full liquid reflux on the separated methane is fed back to the demethanizer 250.
The main portion of stream 203 is fed to a liquefaction process. Since the aromatics are reduced to such a low level and the temperature is very cold, the majority of stream 203 may be fed to a MCHE 270. The MCHE 270 is similar to that described above.
The bottoms product 202 is fed to a deethanizer (DeC2) column 255. The bottoms product 202 has a composition as shown in TABLE 3. The DeC2 column 255 may be a packed or trayed-type distillation column equipped with a reboiler and condenser. The operating pressure of the DeC2 column 255 is in the range of about 35 to about 44 bar. The operating temperature of the overhead condenser is in the range of about −4° C. (at 35 bar).
Ethane and carbon dioxide are the main components of the overhead product, stream 205, and has the composition as shown in TABLE 3.
The overhead product stream 205 is mixed with stream 216 and is fed to a CO₂membrane 290, via stream 213, at a pressure of about 34 bar. In the membrane unit 290, the gases pass over a semi-permeable membrane through which the carbon dioxide passes much more readily than ethane. The surface area of the membrane available and residence time are controlled so that a stream containing up to about 98 vol % CO₂, at a pressure of about 2 bar, is produced. It is well known in the art that the operating parameters of gas-separation membranes (e.g., membrane area, feed and downstream pressure, temperature, and degree of staging) may be varied to yield any desired combination of (CO₂) product purity and (C2) recovery.
The gas exiting the membrane system rich in ethane, stream 215, is fed to a Azeo column 295 at a pressure of about 25 bar. The overhead product stream 216 from the Azeo column 295, substantially the binary azeotrope of ethane and carbon dioxide, is returned to the entrance of the membrane unit 290. A bottoms product 217, which is substantially ethane, is blended back with stream 203, via stream 218, for LNG heating value adjustment, fed to sales (stream 219), and/or supplied to the refrigerant makeup of a liquefaction process.
A bottoms product 204 is fed to a DeC3 column 260 (depropanizer). The bottoms product 204 has a composition as shown in TABLE 3. The DeC3 column 260 is similar to that described above. The operating pressure of the DeC3 column 260 is in the range of about 17 to about 27 bar. The operating temperature of the overhead condenser is in the range of about −2 to about 67° C.
Propane is taken as the principal overhead product, stream 207, and has the composition as shown in TABLE 3.
Stream 207 is fed to an H₂S separator 280. The H₂S separator 280 is similar to that described above. An H₂S stream 220 is Fed to a Claus plant to produce sulfur if the sulfur (tonne/day) is sufficiently high, or disposed of using alternatives known in the art such as reinjection as a separate stream, reinjection as a mixed stream with the CO₂(stream 214), or burned in a thermal or catalytic oxidizer. The operating pressure of the H₂S separator 280 is in the range of about 17 to about 27 bar. The operating temperature of separator 280 is in the range of about −2 to about 67° C.
A portion of the output of the H₂S separator 280, stream 221, is blended back with stream 203, via stream 223, for heating value adjustment. The remainder of the output is fed to sales (stream 222) and refrigerant makeup in a liquefaction process.
A bottoms product 206 is fed to the hydrocarbon recycle stream 228, and to a DeC4 column 265, via stream 208. The bottoms product 206 has a composition as shown in TABLE 3. The DeC4 column 265 is similar to that described above. The operating pressure of the DeC4 column 265 is in the range of about 6 to about 12 bar. The operating temperature of column 265 is in the range of about 40° C. to about 78° C. The bottoms product is output as stream 210.
Butane is taken as the principal overhead product, stream 209, and has the composition as shown in TABLE 3.
Stream 209 is fed to an H₂S separator 285. The H₂S separator 285 is similar to that described above. An H₂S stream 224 is fed to a Claus plant to produce sulfur, or any of the other alternatives for managing H₂S known in the art as described previously. The operating pressure of the H₂S separator is in the range of about 11 bar, and the operating temperature is up to about 78° C.
A portion of the output of the H₂S separator 285, stream 225, is blended back with stream 203, via stream 227, for heating value adjustment, and the remainder of the stream is fed to sales (stream 226) and refrigerant makeup of a liquefaction process.

TABLE 3

Summary of Major Streams Only (Compositions in mol %)

Stream
No.
(FIG. 2)	C₁	C₂	C₃	C₄	C₄+	CO₂	N₂	H₂S	BTEX

201	77.9	3.3	0.9	0.3	0.2	14.8	2.5	0.003	0.06
202	0.01	14	4.1	7.7	9.7	62.7	0	0.01	1.8
203	96.9	0	0.01	0.04	0.01	0	3.1	0	0
204	0	0.01	17.7	33	41.5	0	0	0.01	7.8
205	0.01	18.2	0.03	0	0	81.8	0	0.01	0
206	0	0	2	39.3	49.4	0	0	0	9.3
207	0	0.03	0.99	0.5	0	0.02	0	0.05	0
208	0	0	2	39.3	49.4	0	0	0	9.3
209	0	0	4.9	95.1	0.04	0	0	0	0
210	0	0	0	0.07	84	0	0	0	15.9
213	0.2	23.3	0.02	0	0	76.5	0	0.01	0
214	0.01	1.6	0	0	0	98.4	0	0.01	0
215	0.4	49.8	0.04	0	0	49.7	0	0.01	0
216	0.6	33.2	0	0	0	66.3	0	0	0
217	0	99.8	0.2	0	0	0.01	0	0.02	0
228	0	0	2	39.3	49.4	0	0	0	9.3

TABLE 4

Summary of Major Streams Only

Stream No.	Flowrate	Pressure	Temp
(FIG. 2)	(kgmole/hr)	(bar)	(° C.)

201	58,550	45	−49
202	13,760	45	19
203	47,090	44.5	−84
204	3,209	35	178
205	10,560	34.5	−4
206	2,693	25	175
207	516	24.5	67
208	393	25	175
209	162	10.5	78
210	231	11	171
213	16,000	34.5	0.2
214	8,801	2.1	0.2
215	7,200	34.5	0.2
216	5,401	24	−19
217	1,799	24.5	0.9
228	2,300	45	−32

Third Embodiment

In another embodiment of the invention as illustrated in FIG. 3, a hydrocarbon feed stream 301, having a composition as shown in TABLES 5 & 6 (simulated data), is fed to a DeC1 column 350.
The DeC1 column 350 is similar to that described above. The operating pressure of the DeC1 column 350 is in the range of about 38 to about 45 bar. The operating temperature of the overhead condenser is in the range of about −91 to about −84° C.
A hydrocarbon recycle stream 312 from the bottoms of a DeC3 column 360 is also fed into the DeC1 column 350. The hydrocarbon recycle prevents the CO₂from freezing and acts as a scrubbing agent to remove aromatics and other heavy hydrocarbons from the C1-rich product stream taken overhead. Also, a full liquid reflux on the separated methane is fed back to the demethanizer 350.
Since the aromatics are reduced to such a low level and the temperature is very cold, the majority of stream 303 may be fed to a MCHE 370. The MCHE is similar to that described above.
In stream 303, the separated methane contains less than about 100 ppm CO₂and less than about 3 ppm H₂S.
A bottoms product 302 is fed to a DeC2 column 355 at a pressure of about −40 bar. The bottoms product 302 has a composition as shown in TABLE 5.
The DeC2 column 355 is similar to that described above. The operating pressure of the DeC2 column 355 is in the range of about 35 to about 44 bar, and the operating temperature of the overhead condenser is in the range of about 4° C. (at 35, bar).
Ethane and carbon dioxide are major components in the principal overhead product, stream 305, and has the composition as shown in TABLE 5.
The overhead product stream 305 is fed to an Azeo column 375 at a pressure of about 25 bar. The overhead product 318 from the Azeo column 375, which is substantially a binary azeotrope of ethane and carbon dioxide, is mixed with stream 321 and is fed to the entrance of a CO₂removal membrane 395 via stream 322. A bottoms product 319, which is substantially carbon dioxide, is sent to a re-injection process. In the membrane unit 395, the gases pass over a semi-permeable membrane through which the carbon dioxide passes much more readily than ethane. The surface area of the membrane available and residence time are controlled so that a stream containing about 98 vol % CO₂, at a pressure of about 2 bar, is produced. The membrane unit 395 is similar to that described above. The permeate output (stream 323) from the CO₂membrane 395 is fed to a compressor 380 and output as a vapor stream 320. The vapor stream 320 is recycled back to the Azeo column 375.
The gas (stream 324) exiting the membrane unit 395 rich in ethane is fed to an ethane recovery (C2 Rec) column 396 at a pressure of about 24 bar. The overhead product 321 from this C2 Rec column 396, substantially the binary azeotrope of ethane and carbon dioxide, is returned to the entrance of the membrane unit 395. A bottoms product 325, which is ethane, is blended back with stream 303 for heating value adjustment, fed to sales, and/or refrigerant makeup of a liquefaction process.
A bottoms product 304 is fed to a DeC3 column 360 at a pressure of about 25 bar. The bottoms product 304 has a composition as shown in TABLE 5. The DeC3 column 360 is similar to that described above. The operating pressure of the DeC3 column 360 is in the range of about 17 to about 27 bar. The operating temperature of the overhead condenser is in the range of about −2 to about 67° C.
Propane is taken as the principal overhead product, stream 307, and has the composition as shown in TABLE 5.
Stream 307 is fed to an H₂S separator 385. The H₂S separator 385 is similar to that describe above. An H₂S stream 327 is fed to a Claus plant to produce sulfur, or any of the other H₂S mitigation processes as described above. The operating pressure of the H₂S separator 385 is in the range of about 17 to about 27 bar. The operating temperature of the overhead condenser is in the range of about −2 to about 67° C.
A portion of the output of the H₂S separator 385, stream 326, is blended back with stream 303 for heating value adjustment. The remainder of the output is fed to sales and/or refrigerant makeup in a liquefaction process.
A bottoms product 306 is fed to the hydrocarbon recycle stream 312, and to a DeC4 column 365, via stream 309, at a pressure of about 11 bar. The bottoms product 306 has a composition as shown in TABLE 5.
The DeC4 column 365 is similar to that described above. The operating pressure of the DeC4 column 365 is in the range of about 6 to about 12 bar, and the operating temperature of the overhead condenser is in the range of about 40° C. to about 78° C. The bottoms product is output as stream 310.
Butane taken as the principal overhead product, stream 311, and has the composition as shown in TABLE 5.
A portion of the stream 311, is blended back with stream 303 for heating value adjustment, and the remainder of stream is fed to sales and/or refrigerant makeup of a liquefaction process.

TABLE 5

Summary of Major Streams Only (mol %)

Stream
No.
(FIG. 3)	C₁	C₂	C₃	C₄	C₄+	CO₂	N₂	H₂S	BTEX

301	77.9	3.3	0.9	0.3	0.2	14.8	2.5	0.003	0.06
302	0.01	14	4.1	7.7	9.7	62.7	0	0.01	1.8
303	96.9	0	0.01	0.04	0.01	0.01	3.1	0	0
304	0	0.01	17.7	33	41.5	0	0	0.01	7.8
305	0.01	81.8	0.03	0	0	18.2	0	0.01	0
306	0	0	2	39.3	49.4	0	0	0	9.3
307	0	0.03	99.5	0.4	0	0.02	0	0.05	0
309	0	0	2	39.3	49.4	0	0	0	9.3
310	0	0	0	0.07	84.2	0	0	0	15.9
311	0	0	4.9	95.1	0.04	0	0	0	0
312	0	0	2	39.3	49.4	0	0	0	9.3
318	0.09	26.1	0	0	0	73.9	0	0	0
319	0	0.1	0.03	0	0	99.8	0	0.02	0
320	0.1	2	0	0	0	97.9	0	0	0
321	2.7	34.5	0	0	0	62.8	0	0	0
322	1.5	30.4	0	0	0	68.2	0	0	0
323	0.1	2	0	0	0	97.9	0	0	0
324	2.2	46.8	0	0	0	51.0	0	0	0
325	0	99.9	0	0	0	0.01	0	0	0

TABLE 6

Summary of Major Streams Only

Stream No.	Flowrate	Pressure	Temp
(FIG. 3)	(kgmole/hr)	(bar)	(° C.)

301	58,550	45	−49
302	13,760	45	19
303	47,090	44.5	−84
304	3,209	35	178
305	10,560	34.5	−4
306	2,693	25	175
307	516	24.5	67
309	393	11	40
310	231	11	171
311	162	10.5	78
312	2,300	45	−32
318	7,779	24.6	−17
319	8,632	25	−11
320	5,856	25	40
321	8,208	24.7	−18
322	15,990	24.7	−18
323	5,846	24.2	−4
324	10,140	24.2	−4
325	1,898	24.3	0

Claims

1) A process for producing LNG from high-CO₂natural gas, comprising the steps of:

separating methane from a hydrocarbon feed stream containing CO₂to produce a methane-depleted hydrocarbon stream;

subjecting the methane-depleted hydrocarbon stream to at least one separation process to produce a hydrocarbon recycle stream; and

combining the hydrocarbon recycle stream with the hydrocarbon feed stream prior to separating methane from the hydrocarbon feed stream,

wherein the at least one separation process is selected from the group consisting of deethanizing, depropanizing, debutanizing and CO₂separating.

2) The process according to claim 1, wherein the step of separating methane comprises:

conducting a full liquid reflux on the separated methane vapor product.

3) The process according to claim 1, wherein the step of separating methane comprises:

scrubbing and removing aromatics and heavy hydrocarbons from the hydrocarbon stream containing CO₂.

4) The process according to claim 1, further comprising the step of:

passing the methane to a liquefaction process.

5) The process according to claim 1, further comprising the step of:

passing the methane to a main cryogenic heat exchanger of a liquefaction plant.

6) The process according to claim 1, wherein the hydrocarbon recycle stream comprises fractionated gas components passed from the at least one separation process.

7) The process according to claim 1, wherein the hydrocarbon recycle stream comprises a hydrocarbon stream from a front slug catcher.

8) The process according to claim 1, wherein the methane contains less than about 100 ppm CO₂and less than about 3 ppm H₂S.

9) The process according to claim 1, wherein the step of separating the methane from the hydrocarbon stream is conducted at a pressure in the range of about 38 to about 45 bar, and at a temperature in the range of about −91° C. to about −84° C.

10) The process according to claim 1, wherein the deethanizing is conducted at a pressure in the range of about 35 to about 44 bar, and at a temperature in the range of about 4, C at about 35° C.

11) The process according to claim 1, wherein the depropanizing is conducted at a pressure in the range of about 17 to about 27 bar, and at a temperature in the range of about −2° C. to about 67° C.

12) The process according to claim 1, wherein the debutanizing is conducted at a pressure in the range of about 6 to about 12 bar, and at a temperature in the range of about 40° C. to about 78° C.

13) The process according to claim 1, wherein the step of CO₂separating is conducted at a pressure in the range of about 28 to about 32 bar, and at a temperature in the range of about −6° C. to about −2° C.

14) The process according to claim 1, wherein the step of subjecting the methane-depleted hydrocarbon stream to at least one separation process comprises blending back at least one principal overhead product with the methane for heating value adjustment.

15) The process according to claim 1, wherein the step of subjecting the methane-depleted hydrocarbon stream to at least one separation process comprises feeding at least one principal overhead product stream to a fractionation train.

16) The process according to claim 1, wherein the step of subjecting the methane-depleted hydrocarbon stream to at least one separation process comprises feeding a principal overhead product of ethane and carbon dioxide to an azeotrope separation process.

17) The process according to claim 1, wherein the step of subjecting the methane-depleted hydrocarbon stream to at least one separation process comprises removing H₂S via adsorption from the methane-depleted hydrocarbon stream.

18) The process according to claim 17, wherein the step of removing H₂S is conducted at a pressure in the range of about 1.7 to 27 bar, and at a temperature in the range of about −2° C. to 67° C.

19) The process according to claim 1, wherein the step of subjecting the methane-depleted hydrocarbon stream to at least one separation process comprises membrane-separating CO₂from the methane-depleted hydrocarbon stream.

20) The process according to claim 19, wherein the step of membrane-separating CO₂is conducted to produce a stream containing about 98 vol % CO₂.