CA1200754A - Co.sub.2 removal from high co.sub.2 content hydrocarbon containing streams - Google Patents

Abstract

ABSTRACT OF THE DISCLOSURE
Carbon dioxide is removed from CO2 and hydro-carbon containing gaseous streams. In the instance where the hydrocarbons and CO2 are such that hydrocarbons would condense out during CO2 removal, the gas stream is heated in one or more stages to accomplish hydrocarbon composi-tion control.

Description

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~1----2 REMOVAL FROM HIGH CO ;2 CQNTENT
HYDROC~RBON CONTAINI~aG STREAMS
FIELI) OF THE INVENTION
The invention relates to a process for treating gaseous streams. In one aspect the invention relates to 15 methods for removing carbon dioxide (CO2) from gaseous streams. In another aspect, the invention relates to methods for removing CO2 and/or hydrogen sulfide (H2S) and/or hydrocarbons from gaseous streams.
BACKGROUND ON THE INVENTION
The development of low quality and sour gas res-ervoirs in recent years has required the development of new techniques for low quality gas handling. In addition 9 production of oil and gas by CO2 miscible flooding for enhanced oil recovery can result in sour and low quality , 25 gas streams to be processed. A sour natural gas is a natural gas which contains, in addition to hydrocarbon components, one or more acid gas components. An acid gas component, for example, hydrogen sulfide (H2S) or carbon dioxide (CO2), forms an acidic aqueous solution. Gas 30 sweetening involves almost complete removal of H2S an~
most of the CO2 from sour natural gases. The sweetening is almost always required before the gas can meet sales gas specifications and before the sweet gas can be pro-cessed for production of ethane, propane, butane, and 35 higher hydrocarbon liquid products.
The sour gases encountered today may contain in addition to H2S and CO2 9 carbonyl sulfide, carbon disul-fide, methyl through butyl mercaptans 9 and other volatile ".

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-2-sulfur compounds. Almost complete removal Gf H2S and other volatile sulfur compounds is required to meet rigid sales gas specifications~ CO2 removal may be required, for example, to increase the heating value of the residue 5 sales gas, prevent CO2 frost formation during cryogenic processing, and the like.
FIGURE 1, labeled "Block Diagram of Sour Gas Processing Plant - Background," illustrates the background of the invention in greater detail. A wellstream 11 from 10 a sour gas reservoir is flash separated 12 into a gaseous stream 13 and a liquid stream 14. The liquid stream 14 is stabilized 16 to lower the vapor pressure of the liquid stream, thereby producing a stabilized condensate stream 17 and a vapor f~ction stream 15 which is typically com-15 bined with gaseous stream 13 for gas treatment 18. Gastreatment 18 for a stream from a typical sour gas reser-voir separates an acid gas stream 19 containing predomi-nantly H2S and CO2 which can be further processed in a sulfur plant 20 to produce an elemental sulfllr product 20 stream 21. Gas treatment 18 also typically produces a sweet gas stream 22 which after dehydration and recovery 23 produces a sweet residue gas stream 24, a liquefied petroleum gas (LPG) stream 25, and a natural gasoline liquids (NGL) stream 26. Dotted line 27 indicates gener-25 ally the functional locus of the invention hereindescribed in detail below.
In addition to the routine production of low quality and sour natural gas reservoirs, in recent years reduced petroleum reserves have resulted in development of 30 enhanced oil recovery techniques, such as CO2 miscible flooding, which can result in production of gas streams having a high acid gas content. In the application of CO2 miscible flooding for enhanced oil recovery, the CO2 con-tent of the produced gas increases greatly, after break-35 through, even to levels as high as 98 mol% or higher. Themodification of sour gas treating facilities to process such high, and, in the case of CO2 miscible flooding, variable C02 loading represents a formidable engineering task.

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At low CO2 levels, established technology is and remains attractive; however, at higher levels, ne~ pro-cesses are required. Thus, for example, the MEA (monoe-thanolamine) and DEA (diethanolamine~ processes applied to 5 the gaseous stream ~re economically attractive at CO2 levels -up to, for example, 30 mol/O to 40 mol%. At higher C2 levels, due in part to high energy requirements for regenerating the rich MEA or DEA, other processes become desirable. The invention hereinbelow described is useful 10 for processing high acid gas content natural gas streams and also is adapted for use where the composition of the gas stream, for example, a gas stream derived from a pro~
duction s~ream from a reservoir undergoing CO2 miscible flooding, is highly variable as hereinbelow described in 15 grea~er detail.
SUMMARY OF ~E INVENTION
According to the invention there is provided a process for treating a gaseous stream comprising hydrocar-bons and CO2, the process comprising (a) separating a por~
20 tion of hydrocarbons from the gaseous stream to produce a first stream; (b) separating CO2 from the first stream in a first CO2 removal zone by absorption in a first solution effective for CO2 removal to produce a hydrocarbon enriched second stream; (c) further separating CO2 from ~: 25 the second stream in at least a second CO2 removal zone by absorption in a second solution effective for CO2 removal to produce a further hydrocarbon enriched third stream;
and wherein the gaseous stream comprises CO2 and hydrocar-bons such that in the absence of step (a) hydrocarbon con-30 densation occurs in at least one of steps (b) and (c~.
According to a further aspect of the invention there is provided a process for treating a gaseous stream comprising (a) separating at least a firs~ portion of hydrocarbons from the gaseous stream to produce a first 35 stream having a reduced hydrocarbon content;
(b) separating CO2 from the first stream in a first CO2 removal zone by absorp~ion in a first solution effective for CO2 removal to produce a hydrocarbon enriched second stream; (c) subsequerltly separating a second portion of hydrocarbons from the second stream to produce a ~hird stream having a further reduced hydrocarbon content; and (d) further separa~ing CO2 from the third stream in at 5 least a second CO2 removal zone by absorption in a second solution effective for CO2 removal to produce a further hydrocarbon enriched fourth stream; and wherein the sepa-ration of step (a) is effective for preventing hydrocarbon condensation during CO2 removal in at least the first CO2 10 removal zone~ and the separation of step (c) is effective for preventing hydrocarbon condensation during CO2 removal in said at least a second CO2 removal zone.
According to another aspect of the invention there is provided a process for treating a gaseous stream 15 comprising hydrocarbons and CO2, the process comprising (a) separating a first portion of CO2 from the gaseous stream in a permeation zone by selective permeation of C02 across a differentially permeable membrane to produce a C2 permeate stream and a hydrocarbon enriched first 20 stream; and ~b) further separating CO2 from the hydro-carbon enriched first stream in at least one CO2 absorp-tion zone by absorpLion in a solution effective for CO2 removal to produce a further hydrocarbon enriched process stream A
In a further aspect of the invention, which con-cerns this divi.sional application, there is provided a process for treating a gaseous stream comprisi.ng:
(a) separating a first portion of predomi-nantly higher tc3 and heavier) hydrocarbons from a 30 gaseous stream such that hydrocarbon condensation does not occur from the resulting stream comprising CO2 and at least Cl and C2 hydrocarbons during CO2 removal during Step (b) hereinafter set forth, the gaseous stream comprising CO2 and hydrocarbons including at 35 least Cl and C2 hydrocarbons, and further including C3 - 4a -and heavier hydrocarbons such that hydrocar~on conden-sation occurs during removal o~ C02 unless the gaseous stream is treated prior to C02 removal to remove said first portion of the hydrocarbons; then (b) introducing said resulting stream com-prising C02 and at least C1 and C2 hydrocarbons into a permeation ~one and separating a first portion of C02 from said resulting stream in the permeation zone by selective permeation of C02 across a differentially 1.0 permeable membrane to produce a C02 permeate stream and a hydrocarbon enriched first stream comprising Cl and C2 hydrocarbons; and (c) further separating C02 from the hydro-carbon enriched first stream in at least one C02 15 removal zone by cryogenically fractionating the hydro-; carbon enriched first stream to produce at least a C
stream and a C02 stream.

According to another aspect of the inven~ionthere is provided a process for treating a gaseous stream 20 comprising hydrocarbons and CO2, the process comprising (a) separating a first portion of CO2 from the gaseous stream in a permeation zone by selective permeation of CO2 across a differentially permeable membrane to produce a C2 permeate stream and a hydrocarbon enriched first 25 stream; and (b) further separating CO2 from the hydro-carbon enriched first s~ream in at least one CO2 removal zone by cryogenically fractionating the hydrocarbon enriched first stream to produce at Ieas~ a Cl stream and a CO2 stream.
Fur~her aspects of the invention will be described below in the detailed description of the inven-tion and the drawings, in which:

FIGURE 1 represents a block diagram of a sour gas processing plant showing Lhe background of the inven-tion and indicating by dotted line 27 ~he functional site of the invention;
FIGURE 2 represents a block diagram of one embodiment of the invention;
FIGURE 3 represents a block diagram of a second embodimen~ of the invention;
FIGURE 4 represents graphically variations in 10 hydrocarbon composition on an acid gas free basis of well-head separator gas streams with CO2 ~ontent;
FIGURE 5 represents graphically variations in composition of the Butane-Plus (C4 and heavier hydrocar-bons) fraction of a wellhead separator gas stream with CO2 15 content;
FIGURE 6 represents graphically variations in ; H2S content in the acid gas fraction of wellhead separator gas streams with CO2 conten~;
FIGURE 7 represent graphically absorber dewpoint 20 temperatures at 250 psia as a function of the CO2 content of the gaseous stream being processed;
. FIGURF 8 represents graphically permeator dew-point temperatures at 350 psia as a function of the CO2 content of the gaseous stream bein~ processed;
FIGURES 9 and 10 represent graphically absorber dewpoint temperatures of a TEA absorber at 330 psia and a DEA absorber a~ 310 psia operated in a permeation/CO2 absorption process according to the invention.
FIGURE 11 represents schematically one embodi-30 ment of the invention; and FIGURE 12 represents schematically a second embodiment of the invention.
DESCRIPTION OF THE INVENTION
A. CO~ Solution A~sorption Processes With Hydrocarbon Control 1. Introduction According to the invention is a process for CO2 ; removal from a gaseous stream by CO2 solu~ion absorption 7~

in two or more stages with hydrocarbon control before one or more of the stages to prevent operating problems due to hydrocarbon condensation as C02 is removed.
2. The Gaseous Stream The gaseous stream to be treated in accordance with this aspect of the invention comprises hydrocarbons and CO2. The hydrocarbons include C3 and heavier hydro-carbons and ~he gaseous stream is such that, in the absence of control of the hydrocarbon composition of the 10 gaseous stream, hydrocarbon condensation, especially con-densation of C3 and heavier hydrocarbons, can occur during C2 removal by solution absorption. The gaseous stream can also contain H2S and other components.
The gaseous stream can be processed according to 15 the invention to produce (1) a sweet hydrocarbon stream containin~, for example, less than 2 mol/0 C02 and 4 ppm H2S which can be further processed, for example, for NGL
recovery, (2) a CO2 stream of high purity containing, for example, less than about 100 ppm H2S9 and (3) in some 20 instances where H2S is present, an H~S-containing acid gas stream which can be further processed, for example, in a sulfur recovery facility.
The gaseous stream, as indicated, contains C02 and hydrocarbons, especially C3 and heavier hydrocarbons, 25 such that in the absence of hydrocarbon composition con-trol, hydrocarbon condensation can occur during CO2 removal. The gaseous stream can contain, for example, from about 20 mol% CO2 to about 99 mol% CO2, since above about 20 mol% CO2 the CO2 solution absorption process 30 according to the invention becomes increasingly signifi-cant and economic. Further, our studies have indicated that conventional solution processes for CO2 removal, such as alkanolamine solution processes and hot potassium car-bonate solution processes, continue to be energy intensive 35 above about 30 mol% C02 while the energy requirements for the process according ~o the invention drop off above that level. Preferably, then, the gaseous stream comprises in the range from about 30 mol% CO2 to about 95 mol% C02 37~i~

since above about 30 molQ/O CO2 the inventive process looks most competitive compared with conventional processes and above about 95 mol/O CO2, CO2 removal does not appear eco-nomically justified. Most preferably the gaseous stream 5 comprises in the range from about 30 mol% to about 60 mol%
C2 since above about 60 mol% other processes hereinafter described are preferred.
According to a preferred embodiment, the gaseous stream can be from, for example, a wellhead gas separator 10 where the well is producing from a reservoir undergoing enhanced oil recovery using carbon dioxide-miscible flooding. Such a wellhead gas separator stream during the course of production from the reservoir can show widely varying composition. In such a stream, for example, the 15 C~2 content of the wellhead gas can vary from essentially zero percent even to as high as C~8 mol% or higher with a -il rate of CO2 buildup differing for each field, depending on reservoir geometry and heterogeneity, injection schedules, and other factors. In addition, it appears that other 20 components of the produced gas stream will also vary.
Thus, for example, flash calculations indicate that the proportion of C4 and heavier hydrocarbons varies with var-ying amounts of CO2. The proportion of C4 and heavier hydrocarbons shows a marked increase as the CO2 level 25 increases~ for example, during the course of a flood.
Such calculations also indicate 9 for example, that the hydrogen sulfide (H2S) content in the CO2 fraction decreases as the CO2 content of the wellhead gas separator stream increases. These aspects of the gaseous streams 30 will be further discussed below in the detailed descrip-tion of the drawings.

3. Hydrocarbon Composition Control One consequence of the variation in gas composi-tion is that the gas treating facilities for handling such 35 gaseous streams must be capable of handling a wide varia-tion in ~2 content of the produced gas. Further, the variation in heavier hydrocarbon content must also be taken into account in the design and operation of such facilities.

As described in more detail below, in the C02 absorption process according to the invention, CO2 is removed from the gaseous streams in two or more stages or zones utilizing preferably aqueous alkano].amine solutions.
5 The gaseous stream comprising hydrocarbons, CO2, and some-times H2S, is also treated prior to at least one of, and preferably prior to more than one of or even prior to each of, the C02 absorption stages to prevent the presence of heavy (C3 and higher) hydrocarbons from causing operating 10 inefficiencies by hydrocarbon condensat:ion in the absor-bers during CO2 removal.
According to a preferred embodiment, at least a portion of heavy hydrocarbons are removed by chilling prior to a CO2 absorption stage or zone to a ternperature 15 such ~hat, at the operating temperature and pressure of the C0~ absorber, hydrocarbon condensation does not occur during CO2 removal in at least that zone. It is of course possible and in the spirit of the invention to chill the gaseous stream prior to the first or any other CO2 20 absorber such that hydrocarbon condensation does not occur in that absorber or in any subsequent absorber. Prefer-ably, however, since the CO2 removal can occur in two or more stages or zones, the chilling can be accomplished : prior to each CO2 removal stage or zone to a degree effec-25 tive for preventing hydrocarbon condensation during the immediately subsequent CO2 removal zone or stage. Such a C2 absorption process with interstage chilling can resultg for example, in substantial reductions in refri-geration costs since removal of C02 very significantly 30 reduces the volume of gas to be subsequently chilled.
Further, by staging the CO2 removal as described herein, hydrocarbon control by chilling can be accomplished at higher temperatures for a designated stage of CO2 removal than would be required if a single stage of CO2 removal 35 were utilized, thus further reducing refrigeration costs.
To avoid operating probl~ms which can be caused by hydrocarbon condensation in the absorbers, such as, for example, foaming in a ~riethanolamine (TEA) absorber, the process stream is preferably chilled at the absorber operating pressure or somewhat higher pressure to a tem-perature such that the hydrocarbon dewpoint of the gas during C02 removal always is below the operating tempera-5 ture within the absorber. As used herein 3 the hydrocarbondewpoint temperature is that temperature at which hydro-carbon condensation first occurs when a ~as at a given pressure is cooled. As the C02 is removed from the ascending gas in the absorber, the hydrocarbon dewpoint 10 increases forming a hydrocarbon dewpoint temperature curve or profile as a function of C02 removal. It is highly desirable to operate with absorber temperatures well above (greater than) the hydrocarbon dewpoint temperature pro-; file, and therefore it is desired to chill the process 15 stream to a temperature at least 10F less than the hydro-carbon dewpoint temperature of the absorber offgas for the absorber for which hydrocarbon control is desired. Most preferably, the temperature margin between the absorber temperature and the hydrocarbon dew point temperature pro-20 file is greater than 10F, for example, 15F or more. Toaccomplish this, the feed gas to a C02 removal stage is preferably chilled to a temperature such that the hydro-carbon dewpoint tempera~ure of the absorber offgas is, for example, a~ least 10F greater than the absorber top tray 25 operating temperature. It will be appreciated by those skilled in the ar~ that larger temperature margins provide assurance that hydrocarbon condensation will not occur but can re~ult in greater refrigeration costs. Most prefer-ably, therefore, the temperature margin is in the range of 30 abou~ 15F to about 30F.
As indicated, hydrocarbon control is accom-plished according to a preferred embodiment by staged chilling to condense hydrocarbons which would otherwise condense out during C02 removal in the absorber. Other 35 methods, however, for hydrocarbon control in accordance wi~h the invention can also be utilized.

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_. CO2 Absorption As indicated, according to the invention, a gaseous stream comprising hydrocarbons and CO2 i5 con-tacted in two or more absorber stages or zones with solu-5 tion(s) effective for CO2 removal from the process stream.
Preferably, the gaseous stream from which atleast a first portion of heavy C3 and higher hydrocarbons have been separated is contacted in a first absorber zone with an aqueolls tertiary alkanolamine, for example, 10 trie~hanolamine (TEA)~ methyldiethanolamine (MDEA), and the like. Such ter~iary amines have a relatively lower reactivity than primary or secondary alkanolamines toward acid gas components but are characterized by relatively low energy requirements for regeneration of the rich amine 15 solution; however, as a result of having such lower reac-tivity, such alkanolamines are relatively inefficient for nearly quantitative removal of acid gas components. In accordance with the invention, moreover, ~he absorbers, includin~ the tertiary alkanolamine absorbers, are prefer-20 ably operated such that both H2S and bulk CO2 removaloccurs. Since the H2S is more rapidly removed than the CO2, much of the H2S in the chilled gaseous stream can be removed by the tertiary alkanolamine absorber.
Preferably, the tertiary alkanolamine absorber 25 (first absorber zone) is operated such that the process stream has a CO2 level reduced to a point effective for economical operation of at least one subsequent (second) absorber xone utilizing a primary or secondary alkanola-mine, such as, for example, monoethanolamine (MEA), dieth-30 anolamine (DEA), and the like. Such primary or secondaryalkanolamines are charac~erized by higher energy require-ments for amine regeneration but are efficient or nearly complete removal of acid gas components from the hydro-carbon enriched process s~ream. By arranging the removal 35 Gf the acid gas components so that in a first stage H2S
and bulk CO2 removal occurs using an aqueous alkanolamine having low reactivity and low regeneration energy require-ments followed by a second stage of H2S and CO2 removal ~ 2~

characterized by use of an aqueous alkanolamine having relatively high reactivity and high regeneration energy requirements, a process is provided which accomplishes - nearly complete removal of H2S and C02 while still being 5 relatively energy efficient in comparison to using only a primary or secondary alkanolamine.
The purified hydrocarbon stream from the second stage absorber zone described above can be further pro-cessed, for example, in a natural gas liquids (NGL) 10 recovery plant. The acid gas streams removed in the absorber zone(s) can be further processed in a third selective absorber zone to separate H2S from the C02 to produce a purified C02 stream suitable for pipelining and/or reinjection into the reservoir undergoing enhanced 15 oil recovery. According to a preferred embodiment, the selective absorber zone comprises one or more absorbers using a solution effective for separating H2S from C02, for example, the SELEXOL process developed by Allied Chem-ical Corporation. SELEXOL is a trademark of Allied Chem-20 ical Corporation, U.S. A. Solubility of H2S in theSELEXOL solution (the dimethyl ether of polyethylene glycol) is about nine times the solubility of carbon dioxide, which makes the process useful for selective and nearly quantitative removal of hydrogen sulfide and other 25 sulfur compounds such as, for example, methyl mercaptan and carbonyl sulfide, from the C02 which is present.
Although the SELEXOL process and solution are currently preferred, any chemical absorbent or other process capable of selectivity towards H2S or C02 can be used in accor-30 dance with the invention.B. Combination Staged Processes for CO~ Removal 1. Introduction According to the invention, there are provided combination staged processes for C02 removal from gaseous 3~ streams containing hydrocarbons. According to this aspect of the invention, a first portion of C02 is removed in a permeation zone containing membranes selectively permeable to C02 to produce a hydrocarbon enriched process stream of reduced CO2 content. CO2 is further removed frsm the hydrocarbon enriched process stream by cryogenic fraction-ation and/or by CO2 solution absorption.
At the outset of our investigation into removal 5 of C2 from hydrocarbon gaseous streams containing high levels of CO2, it was not apparent that combination staged : processes for CO2 removal woul~ provide such significant cost and~or energy benefits. As is and was generally known, con~en~ional CO2 removal processes become less 10 costly per unit quantity CO2 removed 35 the mol% CO2 con-tent of the gaseous stream becomes greater. What was no~
immediately apparent, however, was that a hierarchy and staging of CO2 removal processes as described herein would :
provide optimal solutions to the problem of bulk CO2 15 removal from hydrocarbon containing gaseous streams. The C2 removal processes are not combined randomly; rather each process is utilized for CO2 removal in a sequence such that the combined staged processes according to the invention provide highly efficient processes for bulk CO2 j 20 removal.
2. The Gaseous Stream ~;; The gaseous stream to be treated in accordance ' with ~his aspect of the invention comprises hydrocarbons 3 at least C1 and/or C2 hydrocarbons, and carbon dioxide.
25 Any amount of CO2 can be removed in accordance with the invention; preferably, however the gaseous stream will contain above about 30 mol% CO2 since the process according to the invention becomes economically and espe-cially energy competitive above about that level. More 30 preferably the gaseous stream will contain in the range from about 30 mol% CO2 to about 95 mol% CO2 since above about 95 mol% CO2 separation of CO2 and hydrocarbons will not generally be economically justified. Most preferably the gaseous stream will contain in the range from above 35 about 60 mol% CO2 to about 95 mol% CO2 since cryogenic distillative fractionation plants and the CO2 absorption solution processes described hereinabove can be very effi-ciently designed to handle up to about 60 mol% CO2.

Where the gaseous stream does not comprise heavier (C3 and higher) hydrocarbons which can condense out during C02 removal, hydrocarbon control will not gen-erally be necessary. According to a further aspect of the ` 5 invention, however, the gaseous stream can contain such C3 and higher hydrocarbons and in a preferred embodiment the gaseous stream can be from a wellhead gas separator where the well is producing from a reservoir undergoing enhanced oil recovery utilizing C02 miscible flooding as described 10 above. In such cases, hydrocarbon control will be neces-sary as herein described.
3. Hydrocarbon Composition Control As indicated, the invention is directed to pro-cesses for removal of C02, and in some instances J H2S
15 from gaseous streams containing high levels of C02.
According to this aspect of the invention, a gaseous stream containing hydrocarbons and C02 preferably from above about 30 mol% C02 is initially contacted in a per-meation zone containing one or more permeators comprising 20 differentially permeable membranes selectively permeable to C02 relative to hydrocarbons to selectively remove C02 therefrom to a level suitable for a subsequent C02 removal : process.
Where the gaseous stream does not contain C3 and 25 heavier hydrocarbons in an amount which will condense out during C02 removal, control of the hydrocarbon composition will not generally be necessary. Where C3 and heavier ;~ hydrocarbons are present in an amount to condense out~ during C02 removal hydrocarbon control will be necessary.
; 30 Hydrocarbon control can be by any method but preferably will be accomplished by chilling as discussed above with respect to C02 absorption processes. In the case of permeators also, multistage hydrocarbon control is preferred, although single stage control can also be uti-3~ lized. In either instance, the process stream is prefer-ably chilled such that the temperature margin between the dewpoint ~emperature profile of the process s~ream within a permeator and the operating temperature of a respective ;:
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permeator is in the ranges set forth above with respect to C2 absorption processes.

4. Permeation for CO~ Topping _ _ As indicated above, in processing a gaseous

5 stream containing hydrocarbons and a high level of CO2 by a CO2 removal process including amine absorbers, the gaseous stream must be compressed and treated to remove hydrocarbons, if necessary, such that as large amounts of C2 are removed, hydrocarbons do not condense out in, for 10 example, the permeation zone. The gaseous stream from which hydrocarbons have been removed, preferably by chilling, to a level effective to prevent hydrocarbon con-densation during subsequent CO2 removal is treated by con-tacting the chilled gaseous stream in the permeation zone lS with one or more permeators containing membranes selec-tively permeable to CO2 to selectively remove CO2 there-from to produce a CO2 stream of high purity and a hydro-carbon enriched stream of reduced CO2 content. The permeators can comprise any membranes selectively perme-20 able to CO2 such as, for example, specially prepared cel-lulose triacetate membranes manufactured by Dow Chemicals, Inc. or polysulfone membranes manufactured by Monsanto Chemical Co. The permeation process is based on the fact ;~ that with a pressure gradient established across such mem-25 branes carbon dioxide permeates through the membrane more rapidly than gaseous hydrocarbon constituents such as methane. The flow rate for each component across the mem-brane is characterized by a permeability rate coefficient defined as the lb mols/hr of flow per square foot of mem-30 brane per psi average partial pressure differential of thecomponent across the membrane. The ratio of the perme-ability coefficient for CO2 to that for a hydrocarbon, for example, methane, represents the selectivity ratio for CO2 relative to that hydrocarbon. Suitable membranes for use 35 in accordance with the invention preferably have selec-tivity ratios greater than about 10 and more preferably greater than about 15, although membranes with selectivity ratios lower than 10 can also be used. It will be appa-~ S4 rent to those skilled in the art that yet higher selectivity ratios are desirable if such selectivity can be accomplished at an adequate permeability rate.
The higher the CO2 content of the gaseous stream 5 being processed, the more effective permeation becomes because high CO2 partial pressure results in high CO2 permeation rate and high purity of the C02 permeate. Con-versely, attempting to produce a hydrocarbon enriched pro-cess gas stream with low CO2 content tends to lower both 10 CO2 permeation rate and permeate purity. By use of mul-tistage permeator UllitS with appropriate interstage recom-pression and hydrocarbon control, for example, by chilling, it is possible at some cost to achieve any desired degree of separation.
In accordance with the invention, however, per-meators comprising membranes selectively permeable to CO2 are utilized to top off CO2 from the gaseous stream being processed. For the instance where, after the permeation zone, the gaseous stream is subjected to cryogenic frac-20 tionation discussed in greater detail below the CO2 level is preferably reduced to an efficient operating range for cryogenic fractionators, for example, to a CO2 level of about 30 mol% to about 60 mol% CO2. Other levels of CO2 removal by the permeators can, of course, also be utilized 25 in accordance with the invention.
Alternatively, after the permeation zone the gaseous process stream can be treated by contacting the stream with an aqueous alkanolamine stream under condi-tions effective for bulk removal of carbon dioxide from 30 the stream.
5. Permeation/Solution Absorption According to this aspect of the invention, a gaseous stream comprising hydrocarbons and broadly above about 30 mol% CO2 can utilize, as a suitable subsequent 35 C2 removal process after the permeation zone, chemical solution absorption, for example, by aqueous alkanolamines as described in greater detail below. Conventional alka-nolamine processes can be efficiently used up to about 30 mol% CO2 in the case oE primary or secondary alkanolamines, for example, MEA and/or DE~. A combined absorption process as described herein can be used effi-ciently up to about 60 mol% CO2 or even higher. By 5 installing a membrane permeator unit ahead of, for example, an alkanolamine solution absorber, to top out CO2, and thereby reduce the CO2 load on the absorber system, a significant reduction in cost and/or energy requirements for CO2 removal can be effected for feeds 10 with higher CO2 content. In this situation, as described in greater detail below, the feed treatment necessary to prevent liquid hydrocarbon dropout in the permeation zone can also be effective for conditioning the process stream to prevent hydrocarbon condensation or dropout in the 15 absorption zone.

6. Permeation/Cr~ogenic Fractionation Where the gaseous stream contains above about 30 mol% CO2 an alternative suitable subsequent CO2 removal process is a cryogenic distillative fractionation process 20 as hereinafter described in greater detail. In this ins-tance the cryogenic fractionation process is preferably designed to handle gaseous feeds containing, for example, from about 30 mol% to 60 mol% CO2, although the invention can also be practiced with cryogenic fractionation plants 25 for other CO2 ranges. When the CO2 content of ~hè gaseous stream exceeds the design level for the cryogenic frac-tionation process, a permeation zone comprising one or more permeators utilizing membranes selectively permeable to CO2 relative to hydrocarbons is utilized to produce a 30 relatively high purity CO2 permeate (for example, about 95 mol% minimum CO2 content where the CO~ stream is to be utilized for miscible Elooding for enhanced oil recovery) and to produce a hydrocarbon rich process stream with a reduced CO2 content which can be processed in the cryo-35 genic fractionation system.
The combination of permeation ancl cryogenicfractionation as herein described has certain advantages.
The energy consumption and capital investment costs for the cryogenic fractionation system, especially refrigeration costs, increase with CO2 content. On the other hand, the permeation process becomes more energy and investment efficient at higher CO2 levels due in part to 5 the high CO2 permeation rate and high purity of the CO2 permeate at high partial pressures of CO2. The permeation process can thus serve as an effective CO2 topping method for the cryogenic fractionation process and can thereby extend the range of CO2 content in the gaseous feeds which 10 can be economically processed. The combination of the permeation and cryogenic processes also is advantageous in that the drying and feed preparation process appropriate for protection of the permeation membrane(s) can also be used for operation of the cryogenic fractionation process 15 so that the combination of permeation and cryogenic frac-tionation integrates in fact very nicely.
Thus, in cryogenic fractionation t.he feed must be further chilled and compressed, for example, ~o a tem-perature preferably in the range of about -15F to about 2Q -30F prior to entering, for example, a cryod~methanizer column in which Cl hydrocarbon (methane) is separated from C2 and other gaseous and other components of the process stream. Such chilling of course results in liquefaction of most of the CO2, ethane 9 and heavier hydrocarbons in 25 the feed. Consequently, ~he chilling of the gaseo~s stream to the permeation zone to prevent hydrocarbon con-densation in the permeation zone can be integrated effi-ciently with the chilling required for the cryogenic frac-tionation process.
Generally, in design and operation of the com-bined or integrated permeation/cryogenic fractiona~ion process, the operating pressure for the permeator and the degree of CO2 removal ("topping") in the permeator(s) is preferably selected to minimize energy requirements and/or 35 overall CO2 removal costs.
DETAILED DESCRIPTION OF THE DRAWINGS AND EXAMPLES
Referring now to the drawings and in particular to FIGURE 1, dotted line 27 indicates generally the func-'`

7~i~- 1 8 -tional site of the process according to the invention in a sour gas processing plant. Other uses of the invention can of course be made and it is not intended ~y the illus-tration of FIGURE 1 to limit the invention thereto but to 5 illustrate a preferred application of the invention.
Other portions of FIGURE l have been described above relating to the Background of the Invention and will not be repeated here.
Referring now to FIGURE 2, FIGURE 2 represents 10 schematically a block diagram of the process according to the invention using chemical solution absorption for CO2 removal either separately or as part of a combination pro-cess after a permea~ion zone. Line 13 represents a gaseous stream containing carbon dioxide and hydrocarbons 15 as described above, from, for example, a wellhead gas sep-arator. After compression to the operating pressure of the permeation zone 30, stream 13 is provided to first treater 28 (TR1) in which the stream 13 is, for example, chilled. In an alternative embodiment, stream 13' can be 20 provided directly to the CO2 first absorption zone 35 as shown after treating, for example, by chilling in an optional second treater (TR2) 32 to remove hydrocarbons via stream 47 which could otherwise condense out during C2 removal. Stream 13 is chilled in first treater 28 to 25 a level effective for preventing hydrocarbon condensation during subsequent removal of carbon dioxide in permeation zone 30 and preferably at least in first absorption zone 35. Thus, for example, if subsequent removal of CO2 is to be from a first level in stream 13 to a second level after 30 first absorption zone 35~ treater 28 chills the stream 13 to a temperature effective to remove hydrocarbons from the process stream which can otherwise condense, for ~xample, in the permeation zone 30 and/or in the first absorption zone 35 as carbon dioxide is removed to the second level.
35 It is, of course, also possible and in accordance with the invention to chill stream 13 in first treater 28 to a tem-perature effective for preven~ing hydrocarbon condensation in the permeation zone 30 and thereafter operating first absorption zone 35 at a temperature which would require additional treating, for example, by chilling and separa-tion of hydrocarbons in optional second treater 32 with hydrocarbons removed via stream 47 indicated by dotted 5 lines following permeation zone 30.
The chilled and compressed stream 29 from which a first portion of predominantly heavier (C3 and higher) hydrocarbons have been removed is introduced into a per-meation zone 30 in which at least a portion of the CO2 in 10 the chilled compressed stream 29 is removed by selective permeation to produce high puri~y CO~ stream 31 and a hydrocarbon enriched process stream 34 of reduced CO2 con-tent. Process stream 34, or in a alternative embodiment, chilled stream 7 from second treater 32, is preferably 15 provided to the first absorption zone 35. Alternatively, if necessary, stream 34 can be provided to a second treater to remove at least a second portion of hydrocar-bons, especially C3 and heavier hydrocarbons, which could otherwise condense in first absorber zone 35 and then 20 introduced into first absorber zone 35. First absorber zone 35 preferably utilizes an aqueous tertiary alkanola-mine solution having a relatively low reactivity for carbon dioxide and hydrogen sulfide and relatively low energy requirements for regeneration. Suitable aqueous 25 alkanolamines include, for example, triethanolamine (TEA~
and methyldiethanolamine (MDEA). Preferably, TEA is used, for example, in an aqueous solution, comprising preferably about 35 to about 40 wt% TEA and preferably the TEA
absorber is operated in the range of about 160F to about 30 180F to minimize viscosity effects and because operation at this temperature range also integrates efficiently with permeation to reduce refrigeration duty as described in more detail below, for example, by eliminating or reducing interstage treating. Carbon dioxide and hydrogen sulfide 35 are removed from stream 34 in first absorber zone 35 to produce a hydrocarbon enriched process stream 37 of reduced CO2 content and an acid gas stream 36 containing C2 and H2S. The level to which CO2 and H2S are removed in first absorption zone 35 is a function, among other thi.ngs, of total absorption energy requirements. Typi-cally, it is preferred in the first absorber zone that the C2 be reduced from a first level to a second level, typi-5 cally about 10 mol% G02 to about 20 mol% C02, where thesecond level is such that the second absorber zone by ~ti-lizing an aqueous alkanolamine solution characterized by relatively high reactivity toward C02 and H2S and rela-tively higher energy requirements for regeneration of the 10 alkanolamine can be operated for substantially complete removal of acid gas components.
Stream 37 can be provided to third treater 38 (TR3) to remove a third portion o hydrocarbons, espe-cially C3 and heavier hydrocarbons, in stream 37 which 15 could o~herwise condense in second absorber zone 40. The treated stream 39 from third treater 38 is provided to second absorber zone 40 utilizing, for example, an aqueous secondary alkanolamine, such as diethanolamine (DEA) or a primary alkanolamine s-uch as monoethanolamine (MEA).
20 Preferably, DEA is used in an aqueous alkanolamine solu-tion comprising, for example, about 30 to about 40 wt%
DEA. The second absorber zone 40 further removes C02 and H2S from process stream 39 to produce a sweet gas stream 43 having less than, for example, about 2 mol% C0~, and 25 less than about 4 ppm mol% H2S, if H2S is present, and an ~ acid gas stream 41.
: In the illustrated embodiment acid gas stream 31 from permeation zone 30, acid gas stream 36 from the first absorption (absorber) zon~ 35 and acid gas stream 41 from : 30 second absorption (absorber) zone 40 are combined in stream 42 and provided to a selective sweetening zone 45 ~ in which C02 is separated from H2S to produce a sweet ; (i.e., less, for example, ~han about 100 ppm H2S) C02 stream 44, which can be pipelined and/or reinjected for 35 C0~ and an H2S containing stream 46 which can be, for example, sent to a sulfur plant ~see FIGURE 1). Hydro-carbon streams 46, 47, and 48 can be combined as appro-;; priate, for example, in stream 49 which can be further processed for NGL recovery.

l~C~075 Referring now to FIGURE 3, FIGURF 3 represents a block diagram of a process according to the invention using cryogenic fractionation for CO2 removal after per-meation. Line 53 represents a gaseous stream containing 5 carbon dioxide and hydrocarbons as described above, from, for example, a wellhead gas separator. Stream 53 is pro-vided to treater 54 in which compressed stream 53 is, for example, chilled to remove a first portion of hydrocar-bons, especially C3 and heavier hydrocarbons, in stream 62 10 ~o prevent hydrocarbon condensation in the permeation zone 56 as CO2 is removed. Chilled compressed stream 55 from treater 54 is provided to permeation zone 56 comprising one or more permeators with interstage recompression and chilling, if necessary or appropriate, and containing mem-15 branes selectively permeable to CO2 to remove a portion ofthe CO2 therein as high purity sour CO2 stream 63 and Lo produce a hydrocarbon enriched process stream 57 of reduced CO2 content. Stream 57 is provided to cryogenic fractionation system 5S to produce sweet gas stream 61 20 containing C1 (methane) and some lighter hydrocarbons and less, for example, than 2 mol% CO2 and/or 4 ppm H2S, stream 61 containing sweet CO2 containing, for example, less than about 100 ppm H2S, and a sour natural gas liquid stream 59.
Referring now to ~IGURE 4, FIGURE 4 illustrates graphically the composition and variation in hydrocarbon content (reported on an acid gas free basis) expected in the flash gas from a field separator operating, for example, at 100F and 30 psia, for varying mol% CO2 in the 30 flash gas. The curves are based upon calcula~ions based on flashing mixtures of a West Texas reservoir oil of known composition using a modified Redlich-Kwong Equation of State (see, for example, K. C. Chao and R. L. Robinson, Jr., Equations of S~ate in Engineering and Research, ~ 35 Ch. 21, "Applications of a Generaliæed Equation of Sta~e to Petroleum ~eservoir Fluids," Adv. Chem. Ser. 182, Am.
Chem. Soc., Washington, D.C., (1979); and Turek, et al., Phase Equilibria in Carbon Dioxide-Multicomponent Hydro-carbon Systems: Experimelltal Data and an Improved Prediction Techni~ue, Print of paper presented at 55th Annual Fall Technical Conference and Exhibition of the Society of Petroleum Engineers, Dallas, Texas, September, 5 1980).
FIGURE 4 illustrates that, for example, as ~he C2 content of the well separator flash gas increases~ for example, after CO2 breakthrough during recovery from a CO2 - miscible flooding recovery project, the hydrocarbon frac-10 tion comprising C4 and heavier hydrocarbons, designated "Butane Plus" in FIGURE 4, increases rapidly with an increase in C02 content especially above about 60 mol/0 C02 in the flash gas. Conversely, the methane frac~ion falls off increasingly rapidly especially above about 60 mol%
15 CO2. Nitrogen, ethane, and propane functions also decline at higher C02 levels. Accordingly, the process according to the invention for removal of heavy hydrocarbons becomes increasingly important at higher CO2 levels such as above : about 60 mol% CO2.
FIGURE 5 (based upon flash calculations using a modified Redlich Kwong Equation of State as set forth above in reference to FIGURE 4) represents graphically variations in the composition of the Butane Plus fraction (C4 and heavier hydrocarbons3 from a wellhead separator 25 gas stream as a function of CO2 content. As indicated above for FIGURE 4, the Butane Plus fraction increases rapidly at high C02 levels, for example, above about 60 mol% C02. FIGURE 5 indicates that the increase in Butane-Plus fraction is largely a function of an increase 30 in C6 (Hexane) and heavier ("Heptane Plus") hydrocarbons.
As indicated above, the content of C3 and heavier hydro-carbons must be adjus~ed to prevent hydrocarbon condensate formation during removal of C02. FIGURE 5 indicates that ;: in the illustrated case, the adjustment involves removal 35 of hydrocarbons which contain an increasing proportion of ; C6 and heavier hydrocarbons.
FIGURE 6 (based on flash calculations using a modified Redlich-Kwong Equation of State as set forth ~q~5~
above in reference to FIGURE 4) represents graphically variations in H2S content the acid gas fraction in the wellhead separator gas stream as a function of CO2 con-tent. As indicated in FIGURE 6, the H2S content of the 5 acid gas fraction decreases as the GO2 level increases, and in the case illustrated actually falls below the 100 ppm level when CO2 content reaches abou~ 95 mol%.
FIGURE 7 (based on a modified Redlich-Kwong Equation of State as set forth above in reference to 10 FIGURE 4) represents graphically absorber hydrocarbon dew-point temperatures at 250 psia as a function of CO2 con-tent remaining in the process stream in, for example, a TEA absorber operating without a permeation zone preceding the absorber. In accordance with the invention, it has 15 been discovered that as CO2 is removed, for example, from the vapor ascending in a first absorber zone using an aqueous TEA solution for CO2 removal, the hydrocarbon dew-point temperature of the vapor steadily increases, and heavy liquid hydrocarbons can condense out if the dewpoint 20 temperature exceeds the absorber operating temperature.
This is illustrated in FIGURE 7 by curves which illustrate ~; ~ dewpoint behavior in the absorber feed stream having a certain mol% CO2 and which have been chilled to a tempera-ture indicated and hydrocarbons condensed and removed.
25 Thus, the legend "90 mol% CO2 feed, 60F chiller" indi-cates a feed stream to the absorber containing 90 mol% CO2 which has been chilled to the temperature indicated (60F) at or about 250 psia pressure. To avoid operating prob-lems associated with the presence of liquid hydrocarbons 30 in, for example, a T~A first absorber system, especially foaming and hydrocarbon loss in the CO2 offgas from the stripper, it is highly desirable to operate with a temper-ature profile above the hydrocarbon dewpoint profile of the ascending vapor, i.e., the process stream,at all 35 points in, for example, ~he TEA absorber.
FIGURE 8 (based on a modified Redlich-Kwong equation of state as set forth above with reference to FIGURE 4~ represents graphically permeator hydrocarbon dew point temperatures at 350 psia. The curves illustrate hydrocarbon dewpoint behavior in a permeator feed stream havin~ a certain mol% CO~ and which have been chilled to a temperature indicated and hydrocarbons condensed and 5 removed, Thus, the legend "95 mol% CO2, 40F chiller"
indicates a feedstream to the permea~or containing 95 mol%
C2 which has been chilled to 40F at or about 350 psia.
As CO2 is removed by the permeator, the hydrocarbon dew-point temperature steadily increases and heavy (especially 10 C3 and heavier) hydrocarbons can condense Oll$ if the dew-point temperature exceeds the operating temperature of the permeator.
As indicated above, the temperature margin after chilling and hydrocarbon removal between the operating 15 temperature of the permeator and the hydrocarbon dewpoint profile of the gas stream from which CO2 has been removed is at least lO~F and preferably in the range of abou~ 15F
to about 30F although of course even wider margins can be used.
FIGURES 9 and 10 (based on a modified Redlich-Kwong Equation of State as set forth above in reference to FIGURE 4) represent graphically TEA absorber and DEA
absorber hydrocarbon dewpoint behavior, respectively, where the CO2 removal system comprises a permeation zone 25 at 350 psia, a second stage comprising a TEA absorber at 330 psia, and a third stage comprising a DEA absorber at ; 310 psia. Legends are as described above.
EXAMPLE
In accordance with the invention, as indicated, 30 the gaseous stream is compressed and chilled to remove heavy hydrocarbons to prevent condensation or dropout of liquid hydrocarbons in the CO2 removal process. The chilling temperature is set so that no condensation of hydrocarbons can occur as CO2 is removed from the gaseous 35 stream, at least during the immediately subsequent C02 removal stage. This is illustrated by FIGURES 8, 9 and 10, which show the variation in hydrocarbon dewpoint of several feedstreams at 350 psia in a permeator, and at )7 5 330 psia in a TEA absorber, and at 310 psia in a DEA
absorber, respectively, as CO2 is removed.
For example, consider the curve designated "80 mol% CO2 feed, 60F chiller" in FIGURE 8. This repre-5 sents a feed ~as with 80 mol% CO2 which has been chilledto 60F at 355 psia with condensatiorl and removal of heavy hydrocarbons. From FIGURE 8, it can be seen that as the mol~/O CO2 is reduced, for example by CO2 removal from the stream in a permeator operating at 350 psia, the dew point 10 increases from 60F at 80 mol% CO2 to about 95F at 60 mol% CO2 remaining in the feed gas. To avoid condensa-tion of hydrocarbons on the permeator membranes which can cause a reduction in capacity and/or deterioration in the membrane itself, it is desirable to operate with a temper-15 ature profile within the permeator well above the hydro-carbon dewpoint temperature profile of the vapor during permeation. With, for example, the permeator main~ained ~ at about 105F at 350 psia, ~here would be a 10F margin ;~ above the dewpoint currently considered t~ be a minimum 20 operating margin, in this case when 60 mol% CO2 process gas is produced. Further cooling of the process gas to further remove heavy hydrocarbons may be required if, as discussed in greater detail below, subsequent removal of ;; carbon dioxide is to be accomplished by further stages of 25 permeation or chemical absorption solution processes.
Where the subsequent CO2 removal process after the permeator is by chemical absorption using, for examp]e, aqueous alkanolamine solutions, the gaseous feed stream may be chilled either before or after the permea-30 tion zone to a level effective for preventing hydrocarboncondensation or dropout during at least the immediately following CO2 ~bsorption stage.
For example, consider the curve designated "60 Mol% CO2 Feed - 90F Chiller" in FIGURE 9. This curve 35 represents an absorber feed gas with 60 mol% CO2 content which has been chilled to 90F at 255 psia with condensa-tion and removal of heavy hydrocarbons. FIGURE 9 shows that as CO2 is removed from this gas in, for example, a ~,-TEA absorber operating at 250 psia, the dewpoint increases, and, for example, at 20 mol% CO2 the dewpoint is about 132F at 250 psia. To avoid operating problems associated with the presence of liquid hydrocarbons which, 5 for example, in the TEA system, would cause foaming, it is highly desirable to operate with a temperature profile well above the dewpoint of the vapor. With, for example, the TEA absorber top tray maintained at 145F at 250 psia, there would be a greater than 10F margin above the dew-10 point in this case when 20% CO2 absorber offgas is pro-duced. Note that if the TEA absorber is operated at a preferred temperature above 160F 3 such as in the range of 160F to 180F, no chiller is needed after the permeator to lower the dewpoint temperature.
Further cooling of the TEA absorber offgas can be required to prevent hydrocarbon dropout in, for example, a subsequent DEA absorber, which can be operated, ~ for example, to produce an offgas with less than, for ; example, 2 mol% CO2 with a top tray temperature normally 20 about 125F. For a DEA absorber operating at 125~F at 310 psia to reduce CO2 content of feed gas to that stage, Figure 10 shows that for the illustrated stream chilling ; to 100F at about 315 psia is seen to be appropriate to pro~ide about a 10F temperature margin for hydrocarbon 25 control.
Referring now to FIGURE 11, ~IGURE 11 represents an integration in accordance with the invention of mem-brane permeators and cryogenic fractionation for CO2 removal. In the discussion of FIGURE 11 herein exemplary 30 calculated material balance, compositions, and operating parameters are set forth to illustrate practicing the invention; however, it will be clear to those skilled in the art that many other embodiments in accordance with the invention can be u~ilized.
In the illustrated embodiment of FIGURE 11 a feed stream 101 after drying is chilled in chiller 102 to produce chilled stream 103 to separator 104. For illus-tration purposes, stream 101 has the following composi-tion:

~ ~(3~5~

Mols/hr Mol/O

C2 6~48 80 Cl (meth~r~e ) 7 50 C2(ethane) 310 C3(propane) 253 C4~(Butane Plus) 139 Total ~060 Separator 104 at about 300 psia and 0F sepa-rates a liquid fraction stream 109 which is pumped 110 to cryodemethanizer column 111. Vapor s~ream 105 from sepa-rator 104 is superheated in heat exchanger 106 in heat : 15 exchange relationship with hydrocarbon enriched process stream 113 from permeator 108 as hereinafter described to a temperature effective to prevent hydrocarbon condensa-tion during passage of stream 107 through permeator 108.
~: As indicated ab~ve, an advantage of integrating ~: 20 membrane and cryogenic processes for CO2 removal in accor-dance with the invention is tha~ chilling of the feed :; stream 101 is necessary prior to entering cryodeme~hanizer : column 11~ so that treating of feed 101, for example, by chiller 102, to prevent hydrocarbon dropout or condensa-25 tion in the permeator can be accomplished, in part, simul-; taneously. Removal of a significant proportion of the CO2 in the permeator results in a further substantial reduc-tion in the refrigeration duties in chilling the feed to the cryodemethanizer and CO2. Similarly, the hydrocarbon 30 enriched s~ream 113 can be cooled in heat exchange rela-tionship with the vapor stream 105 to reheat the vapor stream to the permeation unit operating temperature.
Superheated stream 107, for example 9 at about.
50F is provided to p~rmeator 108 which in the illustrated embodiment produces a high purity CO2 permeate stream 112 at 50 psi.a and a hydrocarbon enriched process stream 113 of reduced CO2 con~ent a~ 50F.

The composition of permeate stream 112 is, for example, as follows:
Mols/hr Mol%
: N2 17 C2 5379 97.4 c4~ 2 Total 5524;
while the composition of hydrocarbon enriched process stream 113 is as follows:
Mols/hr Mol%

C2 825 38.3 Cl 657 c3 202 C4~ 65 Total 2155.
Process stream 113 is air cooled in exchanger 114 and fur-ther chilled in exchangers 106 and 124 and introduced into 25 cryodemethanizer column 111 at a temperature in the range of -15F to -30F preferably at an intermediate level.
Liquid stream 109 from separator 104, having ~he following exemplary composition Mo 1 s/hr C~+ 72 Total 381, ;

i ~ ~C~ 5 is introduced aL an intermedia~.e level to cryodemethanizer column 111.
Demethaniæer column 111 can be a conventional bubble or valve tray ~ype; alternatively, a packed column 5 can be used. Lean oil stream 112 is introduced into the top of cryodemethanizer column 111 to facilitate reduction of the CO2 content of the methane and li~hter overhead gas to about 2 mol% or less without CO2 freezeup. O~erhead vapors (distillate) stream 113 from ~he top of column 111 10 are conducted through line 114 to chiller 115 where the vapors are partially condensed and the condensed liquid from separator 116 is returned ~o column 111 through line 117.
The residue gas stream from separator 116 has 15 the following exemplary eomposition Mol/hr Mol%

Cl 658 C2 12 1.5 Total 797.
Liquid stream 119, which contains the bulk of the CO2 plus C2 and heavier hydrocarbons, from the bottom of column 111 passes through reboiler 120 via line 121 to ~; maintain the temperature at the bottom of column 111 at a 25 desired level and from cryodemethanizer column 111 to C02 column 1~8 preferably entering a~ an intermediate level.
Lean oil, preferably C4-C5 hydrocarbons, from line 123 via line 126 is introduced with into the top of C2 column 127 to facilitate separation of carbon dioxide 30 from ethane. In addition, the lean oil enhances the vola-tility of C02 relative to H2S and column 127 thereby pro~
duces a sweet high purity carbon dioxide product stream 12~. Ethane and o~her hydrocarbons together wi~h lean oil exit from the bottom of C02 column 127 in stream 129 and 35 are introduced into a depropanizer column 130 at an inter-mediate level. In the depropanizer butane and heavier hydrocarbons are separated as bo~tom stream 131. The VapGr s~ream 133 from the top of depropanizer 130 is pref-3LZ~ 75~a erably subjected to convlentional amine scrubbing in amine unit 134 and the acid gases CO2 and H~S sent to sulfur recovery ~not shown) by line 136. The sweet C2-C3 stream 135 from the amine unit 134 can be sold as an "E-P mix" or 5 further fractionated.
From the foregoing it will be appare~t to those skilled in the art that an efficient integrated membrane permeation-cryogenic fractionation process for removal of C2 from a process gas stre~m containing high levels of 10 CO2 has been described.
Referring now to FIGURE 12, FIGURE 12 represents schematically an integration in accordance with the inven-tion of membrane permeators and chemical absorption for C2 removal. In the discussion of FIGURE 12 herein exem-lS plary material balances, compositions and operating param-eters are set forth to illustrate practicing the inven-tion; however, it will be clear to those skilled in the art that many other embodiments in accordance with the invention can be utilized.
In the illustrated embodiment of FICURE 12 a ~; gaseous feed stream 151 having a composition, for example, as set forth below Mols/hr Mol%

C2 6432 80.0 S ~2 Cl 480 C4 2~4 C5-~ 278 Total 8015 at 335 psia is chilled in exchanger 155 and chiller 157 operated at 355 psia and 60F to remove heavy hydrocarbons 35 in sufficient amount such that ~he vapor passing through the permeator unit has a hydrocarbon dewpoint above ~he operating temperature of the permeator. Chilled com-pressed stream 156 is sent to chiller flash drum 158. A

material across ~he flash drum 158 is, for example, as ~: follows:
Stream 159 Stream 160 Liquid Vapor Mols/hr Mol/O Mols/hr Mol/O

C2 151 31.8 6281 ~3.3 ~1 4 476 : 10 C2 13 273 C5~ ~97 81 Total475 74~0 Hydrocarbon liquid stream 159 from chiller flash drum 158 is sent by line 166 to stabilizer column 164 ; where H2S, CO2 and propane and lighter hydrocarbons are taken overhead in stream 170, chilled in chiller 172, and liquids refluxed via line 171 to the top of stabilizer 20 column 164. C3 and lighter hydrocarbons together with H2S
and CO2 are combined with vapor stream 160 from chiller flash drum 158 and combined stream 161 is superheated in exchanger 155 to 110F at 350 psia and sent to the per-` meator unit 162.
; 25 The butane and heavier bottom streams from sta-bilizer 164 is taken, for example, to a natural gas liquids train via line 167. A portion of the bottom is used for temperature control of the stabilizer column 164 via line 169 and exchanger 168.
Permeator unit 162 separates combined stream 161 into a high purity CO2 stream 164. The material balance across permeator 162 is, for example, as follows:

Stream 161 Stre~m 163 Stream 164 Feed Process Permeate : Mo].s/hr ol%~ols/hr Mol% Mols/hr Mol/O

5 C2 6442 81.21803 55.8 4639 98.7 H~S 12 6 6 Cl ~7 ~5] 36 C3 336 332 ~;

C5~ 130 129 Total7935 3234 4701 Hydrocarbon enriched process stream 163 is introduced into TEA absorber 174, having a bottom pressure 15 of 335 psia and contacted with 40 wt% lean TF.A from TEA
Flash Tower 180 via line 183 by pump 182 to the top of TEA
absorber 174.
In the illustrated embodiment of FIGURE 12~ the perm~ator 162 is designed to yield a hydrocarbon ~nriched 20 process stream lS3 with about 56 mol% CO2 to be sent to TEA absorber 174. The TEA absorber 174 is designed to remove CO2 from its feed to yield an overhead vapor stream 175 containing about 20 mol% CO2 at 160F and 330 psia.
Overhead vapor stream 175 is sent to DEA unit 25 177 after interstage chilling in chiller 152 ~o 100F and hydrocarbon separa~ion in separator 153 to prevent hydro-carbon condensa~ion in the DEA unit 177. The material balance across the DEA treater comprising chiller 152 and ; separator 153 is as follows:
3~

` 35 ' f~ 7~:

Stream 175Separator 153 Separator 153 Feed Vapor Liquid ~ Mol% M~ r Mol%Mols/hr Mol/O
N~ 12 12 ~:5 C2 304 19.0 294 20.7 10 5.3 :H2S 2 2 Cl 43$ 429 7 ; C2 287 269 18 C3 331 2~3 48 10 C4 152 ~7 55 C5~ ~0 31 49 Total 1604 1417 187 DEA unit 177 is functional for removal of mosL
all of the remaining CO2 to produce a lean sweet hydro-15 carbon stream 178 which can be further processed for N&Lrecovery and an acid gas stream 179 which can be sent to selective stripper 188 for selective separation of CO2 H2S ~
TEA flash tower 180 having a bottom pressure of 20 about 30 psia strips CO2 and H2S from rich TEA stream 176 introduced into the top of TEA s~ripper 180 to produce a sour gas overhead stream 184 and lean TEA stream 181 which is recirculated to TEA absorber 174 by pump 182 and line ~ 1$3.
: 25 Sour gas overhead stream 184 can be air cooled : 185 and liquids returned to the top of TEA flash tower 180 ; via line 18$ to produce sour CO2 stream 187.
When H2S is present, selective sweetening of sour CO2 s~ream 187 may be required, for example, using a selective sweetening process such as, for example, SELEXOL
process 188, available from Allied Chemicals, Inc., ~o produce sweet CO2 stream 189, having, for example, less than 100 ppm H2S, suitable for pipelining and/or reinjec-tion for CO2 miscible flooding and an H2S stream 190 which can be sent, for example, to a Claus type sulfur recovery unit (not shown).
Although the invention has been described as required in terms of exemplary and preferred embodiments, 0~ 54 -3~-the invention is not. limited thereto but by the claims hereto appended.

,, '. ~
. 30 . ~

Claims (14)

WHAT IS CLAIMED IS:

1. Process for treating a gaseous stream com-prising:
(a) separating a first portion of predomi-nantly higher (C3 and heavier) hydrocarbons from a gaseous stream such that hydrocarbon condensation does not occur from the resulting stream comprising CO2 and at least C1 and C2 hydrocarbons during CO2 removal during Step (b) hereinafter set forth, the gaseous stream comprising CO2 and hydrocarbons including at least C1 and C2 hydrocarbons, and further including C3 and heavier hydrocarbons such that hydrocarbon conden-sation occurs during removal of CO2 unless the gaseous stream is treated prior to CO2 removal to remove said first portion of the hydrocarbons; then (b) introducing said resulting stream com-prising CO2 and at least C1 and C2 hydrocarbons into a permeation zone and separating a first portion of CO2 from said resulting stream in the permeation zone by selective permeation of CO2 across a differentially permeable membrane to produce a CO2 permeate stream and a hydrocarbon enriched first stream comprising C1 and C2 hydrocarbons; and (c) further separating CO2 from the hydro-carbon enriched first stream in at least one CO2 removal zone by cryogenically fractionating the hydro-carbon enriched first stream to produce at least a C1 stream and a CO2 stream.

2. Process as in Claim 1 wherein the gaseous stream comprises in excess of about 30 mol% CO2.

3. Process as in Claim 2 wherein the separating of Step (a) of Claim 1 is by chilling the gaseous stream to a temperature such that the hydrocarbon dewpoint temperature profile during permeation is less than the temperature of the hydro-carbon enriched first stream produced in Step (b).

4. Process as in Claim 3 wherein the gaseous stream is chilled to a tempera-ture such that the hydrocarbon dewpoint temperature profile is at least 10°F less than the temperature of the thus produced hydrocarbon enriched first stream.

5. Process as in Claim 2 comprising:
chilling the gaseous stream to separate said portion of hydrocarbons from the gaseous stream to pro-duce a chilled gaseous second stream;
heating the thus chilled gaseous second stream to a temperature effective for maintaining the hydrocarbon dewpoint temperature profile during CO2 removal in a permeation zone such that hydrocarbon con-densation does not occur during CO2 removal in said permeation zone;
separating a first portion of CO2 from the thus heated gaseous second stream in a permeation zone by selective permeation of CO2 across a differentially permeable membrane to produce a CO2 permeate stream and a hydrocarbon enriched first stream; and chilling the hydrocarbon enriched first stream to a temperature effective for cryodemethan-izing.

6. Process as in Claim 5 wherein the chilled gaseous second stream is heated in heat exchange relation with the hydrocarbon enriched second stream.

7. Process as in Claim 1 wherein the gaseous stream comprises in the range of about 30 mol% CO2 to about 95 mol% CO2.

8. Process as in Claim 1 wherein the gaseous stream comprises in the range of about 30 mol% CO2 to about 60 mol% CO2.

9. Process as in Claim 1 wherein the gaseous stream is a gaseous stream from a CO2 miscible flood produced reservoir and is character-ized by containing greater than about 30 mol% CO2 and by a variable C3 and heavier hydrocarbon composition.

10. Process as in Claim 1 wherein the gaseous stream is further characterized by a generally increasing C4 and heavier hydrocarbon fraction as a function of increased CO2 content in the gaseous stream from the CO2 miscible flood produced reservoir.

11. Process as in Claim 1 comprising:
introducing the hydrocarbon enriched first stream into a cryodemethanizer column and separating C1 hydrocarbon from CO2 and other components of the hydro-carbon enriched first stream.

12. Process as in Claim 11 further comprising:
introducing a stream comprising hydrocarbons removed from the gaseous stream during Step (a) of Claim 1 at an intermediate level into the cryodemethan-izer column.

13. Process as in Claim 11 comprising:
chilling the gaseous stream to separate said first portion of hydrocarbons from the gaseous stream to produce a chilled gaseous second stream;
heating the thus chilled gaseous second stream to a temperature effective for maintaining the hydrocarbon dewpoint temperature profile during CO2 removal in the permeation zone such that hydrocarbon condensation does not occur during CO2 removal in the permeation zone;
separating a first portion of CO2 from the thus heated gaseous second stream in the permeation zone by selective permeation of CO2 across a differen-tially permeable membrane to produce a CO2 permeate stream and a hydrocarbon enriched first stream;
chilling the hydrocarbon enriched first stream to a temperature effective for cryodemethan-izing; and cryogenically fractionating the hydrocarbon enriched first stream to produce a C1 stream and a CO2 stream.

14. Process as in Claim 13 further comprising:
introducing a stream comprising hydrocarbons removed from the gaseous stream during Step (a) of Claim 1 at an intermediate level into the cryodemethan-izer column.

Respectfully submitted,