CA1122558A - Cryogenic recovery of liquids from refinery off-gases - Google Patents
Cryogenic recovery of liquids from refinery off-gasesInfo
- Publication number
- CA1122558A CA1122558A CA382,734A CA382734A CA1122558A CA 1122558 A CA1122558 A CA 1122558A CA 382734 A CA382734 A CA 382734A CA 1122558 A CA1122558 A CA 1122558A
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- gas
- expander
- compressor
- vapor
- fraction
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- Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
Abstract
ABSTRACT A vapor fraction containing hydrogen and at least one hydrocarbon selected from the group consisting of C1 to C4 hydrocarbons is separated from a hydrogen-rich refinery off-gas feed to give a liquid product fraction. The refinery off-gas is fed to and compressed in a com-pressor/expander having compressor means and expander means mounted and driven on a common shaft, and then cooled and partially condensed to form a two-phase fluid in a heat exchanger followed by separation of the vapor and liquid product phases of the fluid in a separator or a separator/fractionation column stabilizer unit. The separated vapor phase is transmitted to the compressor/ expander unit wherein the vapor is depressurized and partially condensed, thereby driving the compressor. The partially condensed depressurized vapor fraction from the expander and, optionally, the liquid phase product fraction from the separator are transmitted in separate streams to the heat exchanger for separate thermal contact with the compressed feed gas wherein the partially con-sensed depressurized vapor fraction is fully vaporized and the feed gases are cooled. The fully vaporized fraction and the liquid product fraction are recovered in separate streams.
Description
~ -2-1 BACKGRO~ND OF THE INVENTION
A. TECHNICAL FIELD
This invention relates to the separation of petroleum refinery off-gases. More particularly, it relates to the separation of a light vapor fraction containing hydrogen and one or more C1 to C~ hydro-carbons from hydrogen-rich refinery off-gases to give a liquid product fraction.
B. PRIOR ART
The refining of petroleum, e.g., catalytic reforming of petroleum, is often accompanied by the evolution of significant volumes of off-gases composed predominantly of hydrogen together with substantial quantities of C1 to C4 hydrocarbons (i.e., methane, ethane, propanes`and butanes) and gasoline. Ordinarily, such off-gases are routed to the refinery fuel gas system or simply disposed of by flaring. However, the difference between the value of off-gas components as separated and recovered liquid and their value as refinery fuel is often increased by market conditions to the point where there is economic justification for seeking their separation and recovery. For example, under current market conditions, such a difference can amount to a pretax profit of about 25 cents per~gallon of C4 and heavier components re-covered in the form of gasoline, plus 6 cents per gallon recovered to LPG (liquefied petroleum gas) sales.
Although numerous techniques have been developed for separating gaseous mixtures into their constituents as disclosed, ~or example in U~ S. Pat. No. 2,940,270 issued June 14, 1960 to D. F. Palazzo et al. for GAS SEPARATION;
~. S. Pat. No. 3,026,682 issued March 27, 1962 ~558 1 to D. F. Palazzo et al. for SEPARATION OF HYDROGEN ~ND
~IETHANE;
U. S. Pat. No. 3,119, 677 issued January 2~, 1964 to J. J. Moon et al. for SEPARATION OF GASES;
Uu S. Pat. No. 3,255,596 issued June 14, 1966 to S. G. Greco et al. for PURIFICATION OF HYDROGEN RICH GA5;
U. S. Pat. No. 3,292,380 issued December 20, 1966 to R. W. Bucklin for METHOD AND EQUIPMENT FOR TREATING
- HYDROCARBON GASES FOR PRESSURE REDUCTION AND CONDENSATE
RECOVERY;
, .
U. S. Pat. No. 3,397,138 issued August 13, 1968 to K. H. Bacon for GAS SEPARATION EMPLOYING WORK EXPANSION
OF FEED AND FRACTIONATOR OVERHEAD;
U. S. Pat. No. 3,516,261 issued June ~3, 1970 ~0 to M. L. Hoffman for GAS MIXTURE SEPARATION BY DISTILLA-TION WITH FEED-COLUMN HEAT EXCHANGE AND INTERMEDIATE
PLURAL STAGE WORR EXPANSION OF THE FEED;
U. S. Pat. No. 3,729,944 issued May 1, 1973 to C. S. Kelley et al. for SEPARATION OF GASES;
U. S. Pat. No. 3,996,030 issued December 7, 1976 to E. G. Scheibel for FRACTIONATION OF GASES AT LOW
PRESSURE; and ~ ;-U. S. Pat. No. 4,040,806 issued August 9, 1977 - to K. B. Kennedy for PROCESS FOR P~RIFYING ~YDROCARBO~ GAS
STREAMS, ` a need still exists for a way of economically separating refinery off-gases into a liquid product frac-_ tion, and a light vapor fraction containing hydrogen - ., _ -4-1 and one or more C1 to C4 hydrocarbons beginning at the low end ~nder the aforementioned circumstances.
Accordingly, it is an object of the present invention to provide a means for separating petroleum refinery off-gases, e.g., catalytic re~ormer off-gas, into useful fractions or components.
Another object is to provide a method for separating petroleum refinery off-gases, e.g., catalytic reformer off-gas, into a liquid product fraction and a - fraction containing hydrogen and one or more C1 to C4 hydrocarbons.
These and other objects of the invèntion as well as-a fuller understanding of the advantages thereof can be had by reference to the following description, drawings and claims.
, DISCLOSURE OF INVENTION
The foregoing objects are achieved according to the present invention by the discovery of a cryogenic liquid recovery process for treating hydrogen-rich re-finery off gases such as catalytic reformer off-gas, in which free pressure drop is available in the gas stream to be processed. Some other refinery processes which can also supply feed`gas suitable for this process include hydrotreaters and hydro~esulfurization units processing naphtha boiling between about 100 and about 450F, middle ' 30 distillates boiling between about 350 and about 750~, i heavier distillates boiling between about 650~ and about 1100F,~and residual fuel oil. Accordingly, while the invention is described below primarily in the context of catalytic reformer off-gas feed, it is understood that hydrogen-rich refinery off-gases generally, including those mentioned aboYe, are suitable for use in the present invention as long as there is free pressure drop available ~' -` . , .
.
~5~
1 in the gas stream to be processed.
In one aspect of the in~ention, the process - comprises feeding a hydrogen-rich reformer off-gas to a compressor/expander unit having compressor means and expander means mounted on and driven, either recipro-catingly or, preferably, rotary-wise (centrifugally), by a common shaft; the reformer off-gas is desirably fed at a temperature below about 120~F, desirably between about soD and about 120F and pressure above about 100 psig, desirably between about 140 and about 250 psig, and prefer-' ably at between about 95 and about 105F and between about 150 and 200 psig. The reformer off-gas feed is pressurized in the compressor end of the compressor/
expander unit, with the compressor discharge pressure being - a direct function of the power supplied by the expander on the QppOSite end of the common shaft. Desirably, the off-gas feed is pressurized to a pressure of between about'-180 and about 300 psig, preferably between about 18~ and about 275 psig.- The compressed gas is then transported to and cooled in a heat exchanger wherein the gas is partially condensed to form a two-phase fluid, a step which is desirably carried out so that the gas is cooled to a temperature low enough to achieve the desired amount of condensation, desirably between about -130 and abou , -100F, and prèferably to between about -125 and about ! -115F. The two-phase fluid-obtained from the heat ~ exchanger is transmitted to a separator wherein the liquid ! product phase is removed from the vapor ~hase, the latter ;! 30 containing hydrogen and at least one hydrocarbon selected ! from the group consisting of C1 to C4 h~drocarbons, 1 ' desirably at-least one C1 to C3 hydrocarbon, and j preferably at least one C1 or C2 hydrocarbon. The ! vapor phase is-transmitted'from'~he separator to the turboexpander end of the compressor~expander, in the case ; of a rotary unit, ~Iherein the vapor is depressurizea and _ cool'ed across the turbine blades of the turboexpander and -- . .
, Z~8 _ -6-1 partially condensed thereby, the enthalpy removed from the vapor by the expander supplying power to drive the compres-sor end of the compressor/expander unit. Desirably, the vapor is depressurized and cooled in the turboexpander to between about 20 and about 100 psig and between about -250 and about -140F, preferably between about 55 and - about 65 psig and between about -180 and about -170F.
These conditions typically correspond to the pressure - required for the expanded vapor to ultimately be discharged ; 10 iDto the refinery fuel gas distribution system or the flare. The liquid phase product obtained i~ the separator and the partially condensed and depressurized vapor fraction from the turboexpander are transmitted to the heat exchanger wherein they are brought into therrnal contact with (but Xept physically separate from) the compressed feed gas from the compressor end of the compres-sor/expander unit, wherein the partially condensed and depressurized vapor fraction is fully vaporized and the compressed feed gas is cooled,` after which the fully vaporized fraction and the liquid product fraction are withdrawn from the heat exchanger. Desirably~ the liquid phase product fraction from the separator and the depres-: surized partial condensate from the turboexpander areheated by the feed gas within the heat exchanger to a temperature of about 5F or more below the temperature of the feed gas entering the exchanger, desirably between I about 100 and about 140F, and preferably to between about 130 and about 140F. The partially condensed and depressurized vapor fraction from the turboexpander and 3U the liquid product fraction from the separator are prefer-j ably fed into and withdra~n from the heat exchanger in separate streams.
Although the aforementioned embodiment utili~es I 35 a single compressor/expander unit, two or more such units, ; preferably connected in series, can be installed depending ; on the overall free pressure drop available and/or the lL~2Z~58 _ -7-1 desired degree of liquid product recovery. In such cases, the feed gas flow to the compressor ends is directly connected from one machine or unit to the next. However, the cold vapor to be expanded across the expander ends of the units should be free of significant amounts of entrained liquid droplets at the inlet of each expander.
Therefore, a separator vessel is advantageously installed ahead of each expander. The liquid recovered in each such separator vessel can then be rejoined with each expander outlet stream via level control valves.
In another aspect of the invention, the two-phase fluid obtained by passing the compressed feed gas through the heat-exhanger is transmitted to the separator portion of a stabilizer comprising said separator and an externally heated, staged (e.g., packed or trayed) fractionation column situated beneath and in communication with the separator, wherein the liquid product phase is recovered from the vapor phase by gravity separation, some constitu-ents of said vapor phase being fractionated from the netliquid product leaving the bottom of the tower, said vapor phase containing hydrogen and at least one hydrocarbon selected from the group consisting o~ C1 to C4 hydro-~ carbons, desirably at least one C1 to C3 hydrocarbon, j 25 and preferably a C1 or C2 hydrocarbon. The liquid product phase discharged from the bottom of the fractiona-tion column of the stabilizer is circulated through a coil within the column whereby heat removed from the liquid I product within the column furnishes supplemental side j 30 reboil heat for the stabilizer. Alternatively, in lieu of j routing the net liquid product through a coil within the j column, a conventional external side reboil heat exchanger can be utilized for this purpose, the details-of the installation of which will be apparent to those skille~ in ! 35 the art. The vapor phase separated in the stabili~er is transmitted to the expander(s3 of the co~pressor~expander _ unit(s) wherein the vapor is depressurized and cooled .
.
-:
1 across the turbine blades of the turboexpander( 5 ) (in the case of a rotary unit) and partially condensed thereby, the enthalpy removed from the vapor by the expander supplying power to arive the compressor. The partially condensed depressurized vapor fraction from the turbo-expander is transmitted to the heat exchanger for thermal contac~ with the pressurized feed gas from the compressor, wherein the partially condensed depressurized vapor fraction is fully vaporiz~d. Finally, the fully vaporized fraction is withdrawn from the heat exchanger and the liquid product phase is withdrawn and recovered from the column.
The two preceding aspects of the invention are particularly applicable where initial capital investment must be minimized and/or where high recovèry of the lighter C1 to C4 hydrocarbon constitue`nts are not major require-ments. In the event that economics or overal- refinery process requirements call for the highest practicable - 20 recovery of the lighter C1 to C4 hydrocarbons, prefer-ably the C2 and` heàvier hydrocarbons or the C3 and heavier hydrocarbons, a further aspect of the invention can be utilized to achieve such higher recovery levels.
In particular, as with the two previously described 2S aspects of the invention, the feed gas is first pres-surized by the compressor ends of one or more turboexpander/
compressor units connected in series as described above.
The resultant compressed feed gas is then cooled to below abo~t 120F in a conventional heat exchanger against cool-ing water or by an air-cooled exchanger. Alternatively, , any part of the cooling required by the compressed gas can be achieved by utilizing the heat available in this stream to supply part or all of the side and/or bottom reboil heat required by the stabilizer fractionation column. The cooled feed gas is then fur~her cooled ana partially condensed to form a two-phase fluid in one _ or more heat exchangers arranged so as ~o achieve the .- . .: ~ .
~ 5~
_g_ I
1 required amount of condensation. Some of the required cooling, if necessary, can be supplied by one or more supplemental external mechanical refrigeration lnits, the implementation of such units being readily ap~arent to those skilled in the art. The two-phase feed stream exiting the last exchanger would then flow to a separator vessel for removal of the liquid phase. The liquid phase -thereby removed can then directly, or indirectly through one of the feed gas heat exchangers, be introduced into an ;10 intermediate stage of the stabilizer column. The vapor ;fraction leaving the high pressure feed- gas separator can then proceed directly to the inlet of the first turbo-expander or be further cooled and partially condensed in one or more heat exc~langers, again follo~ed by another separator vessel, the vapor from the separator flowing to the first turbo-expander inlet and the liquid being charged into an upper stage of the stabilizer column.
In the event of a need for more than one turboexpander, the outlet two-phase stream leaving all but the last turboexpander in series can be separated into liquid and vapor fractions by separator vessels between each expander, ; to insure that no entrained liquid remains in the inlet vapor stream to each machine. Separated liquid fractions can be directly or indirectly introduced at the proper , 25 point to the stabilizer column. The two-phase stream ! exiting the last expander in the series can be discharyed into a separator situated on top of, and in direct commu-nication with, the fractionation section of the stabilizer column, as previously described. The total vapor stream leaving the separator atop the stabilizer, which comprises i the last expander outlet plus the undesired hydrogen and ! lighter C1 to C4 hydrocarbons contained in the column j feed liquid streams, can then be warmed by heat exchange ! with incoming feed gas as discussed previously. l`he ' 35 operating pressure for the stabllizer column, as a minimum~
.
'~ ' . ' ' ~zz55~
--1 o--1 can be adjusted to permit the lean residue gas produc~ to proceed to its final destination, for exa~ple, the refinery fuel gas distribution system. Alternatively, the residue gas product can proceed to outside gas compression facil-ities should a need exist for the hydrogen component, - e.g., in another refinery process. In the event that pressurization of the lean residue gas product is desired, the compressor ends of the expander/compressor units can be utilized for compression of the warmed residue gas product in lieu of feed gas compression as previously described. Also, in a further application of this pro-cess, the power produced by the expanders can be used to drive electric generators, air blowers, dynamometers, or ; other load devices on a common shaft with the expanders.
In a preferred mode of carrying out the afore-mentioned aspects of this invention, the feed gas to the process unit may require one or more types of pretreatment processing. For example, the cryogenic temperatures ; 20 encountered within the basic process may cause some undesirable impurities contained in the feed gas to freeze or form hydrates. Examples of such impurities are water vapor and carbon dioxide, both of which, depending on the process conditions and their concentrations, could free~e within the unit. Thus, if the water content is high I enough to warrant its complete removal, a separate dehydration unit can be installed to process the feed gas prior to its introduction into the facilities described herein. Carbon dioxide can be removed by any of several 1 30 means conventionally utilized for this purpose, such as ! molecular `sieves, amine solution, caustic soda solution, and the like. Where low concentrations of impurities are I present in the feed gas, a freezing point depressant can I be advantageously added to the feed gas prior to its being ' ~5 chilled, in an amount sufficient or as needed to prevent , zss~
-1 1' 1 1 ice formation. Suitable freeæing point depressants incl~de any liquid known to be useful as a feed gas anti-freeze such as a C1 to C3 alcohol, e.g., methanol, ethanol, propanols, or mixtures thereof, these being especially preferred, in view of the fact that the freezing point depressant will substantially remain in the liquid product, - and should therefore be compatible with it and its uses.
BRIEF DESCRIPTION OF THE DRA~INGS
The specific details of the invention and of the best mode known to me for carrying it out can be had by reference to the following description in conjunction with the accompanying drawings wherein:
FIG. 1 is a schematic flow diagram representing a preferred embodiment of the process and apparatus o~
the invention;
FIG. 2 is a schematic flow diagram representing ~0 another preferred embodiment of the invention; and FIG. 3 is a schematic flow diagram representing a third preferred embodiment of the invention.
DESCRIPTION OF PREFERRED EMBODIMENTS
Referring to FIG. 1, which describes the basic process configuration of the i~vention wherein the liquid product fraction of an off-gas feed is recovered in a form suitable for transfer to existing refinery fractionation facilities for stabilization and product separation, the I process flow scheme shown : can be applied to a modern ! catalytic reforming unit which operates at relatively low pressures. Although labeled on the flow diagram as "high pressure rich gas", the flow conditions of reformer off-!35 gas feed stream 1 are usually below about 120F and above ,about 100 psig, desirably between about 90 and about .
.
1~2~558 120DP at between about 140 and about 250 psig, and prefer-ably between about 95D and about 105F at between about 150 and about 200 psig, e.g., 100F at 155 psig. The feed gas 1 is first compressed by the centrifugal compressor 5 end 20 of coMpressor/expander unit 21 to a pressure which is dependent on the power available from the expander, desirably between about 180 and 300 psig, and preerably between about 185 and 275 psig, e.g., about 190 psig. The pressurized feed gas stream 2 from compressor 20 is then - 10 cooled in heat exchanger 22 to a temperature corresponding to the amount of internal or external refrigeration available, desirably between about -130 and about -100F, preferably between about -125 and about -115F, e.g~, -120DF, and partially condensed~ The cold, two-phase 15 stream 4 from heat exchanger 22 proceeds to separator 23 wherein a liquid product phase is recovered from the vapor phase, the latter containing hydrogen and at least one hydrocarbon selected from the group consisting of C1 to C4 hydrocarbons~ The vapor phase stream 5 from separator 20 23 is transmitted to the turboexpander end 24 of expander/
compressor unit ~1 wherein the vapor is depressurized across the turbine blades (not shown) and partially condensed thereby. The work (enthalpy) removed from the ; cold vapor stream 5 by the turboexpander end 24 supplies 25 the power needed to drive the compressor end 20 mounted on the common shaft 25.
, - ' ' Duè to the work removal from the high-pressure ¦ cold vapor within turboexpander 24, the partially con- -! 30 densed exit stream 7 from the turboexpander is at a temperature of between about -250 and about -100F and pressure of between about 20 and about 100 psig, preferably between about -180 and about -1 70aF at between about 55 and about 65 psig, e.g., -1 75F and 60 psig. The par-35 tially condensed steam 7 from turboexpander 24 proceeds .
--`~
~s~
_ -13-1 back through heat exchanger 22 where it is fully vaporized and heated by the incoming feed gas stream 2 to a temper-ature of about 5 to 10 F. below the temperature of the feed gas entering the exchanger, desirably between about 5 100 and about 140F, and preferably between about 130 and about 140F, e.g. 135F. The thus-vaporized and warm lean gas stream 8 from heat exchanger 22 then passes 'through pressure control station 26 which controls the expander outlet pressure. The low pressure lean gas thus 10 obtained is then routed to the refinery fuel system, the flare or a hydrogen recompressor.
The cold liquid hydrocarbon fraction recovered from separator 23 as stream 9 is transmitted to product 15 pump 27 at the operating temperature of the separator, desirably between about -130 and about -100~F,'and preferably between about -125 and about -115F, e.g., -120F. The high pressure liquid stream 10 discharged from pump 27 procee'ds to heat exchanger 22 wherein it is 20 heated by'compressed feed gas stream 2 to a temperature of between about 100 and about 140F,'a`nd preerably between about 130 and about 140F, e.g. 135F, be~ore being exported as stream 13 to fractionating facilities lnot shown).
! 25 Since reformer off-gas normally contains some water vapor (typically 1S-30 ppm), a freeze point depres-sant or antifreeze stream 3 can, if desired, be'injected into the compressed feed gas stream 2 to prevent ice 30 and~or hydrate formation. Suitable freeze point depres-- sants include C1 to C3 alcohols, e.g., methanol, ethanol, or propanol which will remain in the recovered liquid product phase.
!
In carrying out the process depicted in FI~. 1, ' ~; .
~;
. . .
25~8 _ -14-1 the eY~pander/compressor unit is operated basically on speed control. Thus, when the flow of cold vapor stream 5 from separator 23 is equal to or below the designed throughput, all of such flow is processed through the expander. Should higher flo~ts in stream 5 occur which, if not checked, could cause an overspeed situation - for the expander, the excess flow is automatically by-passed around expander/compressor unit 21 through stream 6. Alternate con'trolling parameters, well-known to those skilled in the art, can easily be adapted to the expander unit, depending on the particular circumstances.
Typical unit feed gas and product compositions are shown in Table I.
Referring now to FIG. 2, which depicts a variant of the present process, the re~erence numerals identical to those in FIG. 1 refer to corresponding elements. In this embodiment, the separator 23 is situated on top of and in communication with a pac~ed fractionation section 29, the entire unit 30 being referred to as a "stabilizer~
Liquid separated from the cold, two-phase stream 4 exiting heat exchanger 22 flows down~ard through the packing in the fractionating column or tower 29 of stabilizer 30.
~5 Lighter components are vaporized from the net liquid product leaving the bottom of fractionation tower 29.
Depending on the lightest component desired in the net ~ ' liguid product phasè, stabilizer 30 operates as a de-methanizer (for ethane recovery to bottoms), as a de-ethanizer (for propane recovery), or a depropanizer (for butane recovery). In the configuration shown in FIG. 2~
the stabllizer 30 is operating as a deethanizer such that all of the et'hane and lower boiling constituents are frac-tionated from the net propane and heavier bottom product.
3~ ' '- ,' ~ - - :
~L~LZZ558 _ -15-1 The recovered liquid product phase stream 10 discharged from product pump 27 is routed to and circulated through coil 31 wound through the packing in fractionation column 29. Heat removed from the liquid 5 product in coil _ furnishes supplemental side reboil heat for stabilizer 30. The main reboil heat to fractionation tower 29 is supplied from an outside source 32, such as a hot refinery process stream, low or high pressure steam, electric heater, and the like.
- Typical unit feed gas and product compositions are shown in Table I.
.
.
.
.
, ~ 25 `I ' ` , .
! 30 1' ' ' . ' . ' ' ' .
~5 ~
.
.. . : : :
.
~8 . ~ ~ W CO
:Z; U ~ r~
C~ o ~ .
C~ ~ .
N ~0 Z .
~, H . ~ ~ O OO N ~ ~1 t:~ O ~:1 1~a~ oo ~ ~ N ~ ~
H H ~ O O ON '1 In 4 H H U) ~ ~ ~ O. CJ~ O -~ ~ ~
~ ~ o N~r o o ~ ~o ~: ~
o ~ H~
~ ~ ~ ~ a~
W ~0 ~ O .~ Q~ ~ ~ o u~
20 ~ a ~ u ~ ~ O O u~ N ~ ~n H O 1;~ S-l O ~
1 P~ Ql OIJ~ NCO N --1 1~l U . ~ N O t` ~ O O
1 1 H .~ r-lCO
25P~ t¢ ~ ~
~I . ~
n ~ I I ~r~ N lSl . . C~
. ' 30 * . ~C o t;~ r~ - ~ N
O tl') O N
~ 35 o q ~
O , 5 U U U ~ ~
U
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~ ' ' 't ~ ~ ' 2S~
; - 1 7 -a ~L.'-)
A. TECHNICAL FIELD
This invention relates to the separation of petroleum refinery off-gases. More particularly, it relates to the separation of a light vapor fraction containing hydrogen and one or more C1 to C~ hydro-carbons from hydrogen-rich refinery off-gases to give a liquid product fraction.
B. PRIOR ART
The refining of petroleum, e.g., catalytic reforming of petroleum, is often accompanied by the evolution of significant volumes of off-gases composed predominantly of hydrogen together with substantial quantities of C1 to C4 hydrocarbons (i.e., methane, ethane, propanes`and butanes) and gasoline. Ordinarily, such off-gases are routed to the refinery fuel gas system or simply disposed of by flaring. However, the difference between the value of off-gas components as separated and recovered liquid and their value as refinery fuel is often increased by market conditions to the point where there is economic justification for seeking their separation and recovery. For example, under current market conditions, such a difference can amount to a pretax profit of about 25 cents per~gallon of C4 and heavier components re-covered in the form of gasoline, plus 6 cents per gallon recovered to LPG (liquefied petroleum gas) sales.
Although numerous techniques have been developed for separating gaseous mixtures into their constituents as disclosed, ~or example in U~ S. Pat. No. 2,940,270 issued June 14, 1960 to D. F. Palazzo et al. for GAS SEPARATION;
~. S. Pat. No. 3,026,682 issued March 27, 1962 ~558 1 to D. F. Palazzo et al. for SEPARATION OF HYDROGEN ~ND
~IETHANE;
U. S. Pat. No. 3,119, 677 issued January 2~, 1964 to J. J. Moon et al. for SEPARATION OF GASES;
Uu S. Pat. No. 3,255,596 issued June 14, 1966 to S. G. Greco et al. for PURIFICATION OF HYDROGEN RICH GA5;
U. S. Pat. No. 3,292,380 issued December 20, 1966 to R. W. Bucklin for METHOD AND EQUIPMENT FOR TREATING
- HYDROCARBON GASES FOR PRESSURE REDUCTION AND CONDENSATE
RECOVERY;
, .
U. S. Pat. No. 3,397,138 issued August 13, 1968 to K. H. Bacon for GAS SEPARATION EMPLOYING WORK EXPANSION
OF FEED AND FRACTIONATOR OVERHEAD;
U. S. Pat. No. 3,516,261 issued June ~3, 1970 ~0 to M. L. Hoffman for GAS MIXTURE SEPARATION BY DISTILLA-TION WITH FEED-COLUMN HEAT EXCHANGE AND INTERMEDIATE
PLURAL STAGE WORR EXPANSION OF THE FEED;
U. S. Pat. No. 3,729,944 issued May 1, 1973 to C. S. Kelley et al. for SEPARATION OF GASES;
U. S. Pat. No. 3,996,030 issued December 7, 1976 to E. G. Scheibel for FRACTIONATION OF GASES AT LOW
PRESSURE; and ~ ;-U. S. Pat. No. 4,040,806 issued August 9, 1977 - to K. B. Kennedy for PROCESS FOR P~RIFYING ~YDROCARBO~ GAS
STREAMS, ` a need still exists for a way of economically separating refinery off-gases into a liquid product frac-_ tion, and a light vapor fraction containing hydrogen - ., _ -4-1 and one or more C1 to C4 hydrocarbons beginning at the low end ~nder the aforementioned circumstances.
Accordingly, it is an object of the present invention to provide a means for separating petroleum refinery off-gases, e.g., catalytic re~ormer off-gas, into useful fractions or components.
Another object is to provide a method for separating petroleum refinery off-gases, e.g., catalytic reformer off-gas, into a liquid product fraction and a - fraction containing hydrogen and one or more C1 to C4 hydrocarbons.
These and other objects of the invèntion as well as-a fuller understanding of the advantages thereof can be had by reference to the following description, drawings and claims.
, DISCLOSURE OF INVENTION
The foregoing objects are achieved according to the present invention by the discovery of a cryogenic liquid recovery process for treating hydrogen-rich re-finery off gases such as catalytic reformer off-gas, in which free pressure drop is available in the gas stream to be processed. Some other refinery processes which can also supply feed`gas suitable for this process include hydrotreaters and hydro~esulfurization units processing naphtha boiling between about 100 and about 450F, middle ' 30 distillates boiling between about 350 and about 750~, i heavier distillates boiling between about 650~ and about 1100F,~and residual fuel oil. Accordingly, while the invention is described below primarily in the context of catalytic reformer off-gas feed, it is understood that hydrogen-rich refinery off-gases generally, including those mentioned aboYe, are suitable for use in the present invention as long as there is free pressure drop available ~' -` . , .
.
~5~
1 in the gas stream to be processed.
In one aspect of the in~ention, the process - comprises feeding a hydrogen-rich reformer off-gas to a compressor/expander unit having compressor means and expander means mounted on and driven, either recipro-catingly or, preferably, rotary-wise (centrifugally), by a common shaft; the reformer off-gas is desirably fed at a temperature below about 120~F, desirably between about soD and about 120F and pressure above about 100 psig, desirably between about 140 and about 250 psig, and prefer-' ably at between about 95 and about 105F and between about 150 and 200 psig. The reformer off-gas feed is pressurized in the compressor end of the compressor/
expander unit, with the compressor discharge pressure being - a direct function of the power supplied by the expander on the QppOSite end of the common shaft. Desirably, the off-gas feed is pressurized to a pressure of between about'-180 and about 300 psig, preferably between about 18~ and about 275 psig.- The compressed gas is then transported to and cooled in a heat exchanger wherein the gas is partially condensed to form a two-phase fluid, a step which is desirably carried out so that the gas is cooled to a temperature low enough to achieve the desired amount of condensation, desirably between about -130 and abou , -100F, and prèferably to between about -125 and about ! -115F. The two-phase fluid-obtained from the heat ~ exchanger is transmitted to a separator wherein the liquid ! product phase is removed from the vapor ~hase, the latter ;! 30 containing hydrogen and at least one hydrocarbon selected ! from the group consisting of C1 to C4 h~drocarbons, 1 ' desirably at-least one C1 to C3 hydrocarbon, and j preferably at least one C1 or C2 hydrocarbon. The ! vapor phase is-transmitted'from'~he separator to the turboexpander end of the compressor~expander, in the case ; of a rotary unit, ~Iherein the vapor is depressurizea and _ cool'ed across the turbine blades of the turboexpander and -- . .
, Z~8 _ -6-1 partially condensed thereby, the enthalpy removed from the vapor by the expander supplying power to drive the compres-sor end of the compressor/expander unit. Desirably, the vapor is depressurized and cooled in the turboexpander to between about 20 and about 100 psig and between about -250 and about -140F, preferably between about 55 and - about 65 psig and between about -180 and about -170F.
These conditions typically correspond to the pressure - required for the expanded vapor to ultimately be discharged ; 10 iDto the refinery fuel gas distribution system or the flare. The liquid phase product obtained i~ the separator and the partially condensed and depressurized vapor fraction from the turboexpander are transmitted to the heat exchanger wherein they are brought into therrnal contact with (but Xept physically separate from) the compressed feed gas from the compressor end of the compres-sor/expander unit, wherein the partially condensed and depressurized vapor fraction is fully vaporized and the compressed feed gas is cooled,` after which the fully vaporized fraction and the liquid product fraction are withdrawn from the heat exchanger. Desirably~ the liquid phase product fraction from the separator and the depres-: surized partial condensate from the turboexpander areheated by the feed gas within the heat exchanger to a temperature of about 5F or more below the temperature of the feed gas entering the exchanger, desirably between I about 100 and about 140F, and preferably to between about 130 and about 140F. The partially condensed and depressurized vapor fraction from the turboexpander and 3U the liquid product fraction from the separator are prefer-j ably fed into and withdra~n from the heat exchanger in separate streams.
Although the aforementioned embodiment utili~es I 35 a single compressor/expander unit, two or more such units, ; preferably connected in series, can be installed depending ; on the overall free pressure drop available and/or the lL~2Z~58 _ -7-1 desired degree of liquid product recovery. In such cases, the feed gas flow to the compressor ends is directly connected from one machine or unit to the next. However, the cold vapor to be expanded across the expander ends of the units should be free of significant amounts of entrained liquid droplets at the inlet of each expander.
Therefore, a separator vessel is advantageously installed ahead of each expander. The liquid recovered in each such separator vessel can then be rejoined with each expander outlet stream via level control valves.
In another aspect of the invention, the two-phase fluid obtained by passing the compressed feed gas through the heat-exhanger is transmitted to the separator portion of a stabilizer comprising said separator and an externally heated, staged (e.g., packed or trayed) fractionation column situated beneath and in communication with the separator, wherein the liquid product phase is recovered from the vapor phase by gravity separation, some constitu-ents of said vapor phase being fractionated from the netliquid product leaving the bottom of the tower, said vapor phase containing hydrogen and at least one hydrocarbon selected from the group consisting o~ C1 to C4 hydro-~ carbons, desirably at least one C1 to C3 hydrocarbon, j 25 and preferably a C1 or C2 hydrocarbon. The liquid product phase discharged from the bottom of the fractiona-tion column of the stabilizer is circulated through a coil within the column whereby heat removed from the liquid I product within the column furnishes supplemental side j 30 reboil heat for the stabilizer. Alternatively, in lieu of j routing the net liquid product through a coil within the j column, a conventional external side reboil heat exchanger can be utilized for this purpose, the details-of the installation of which will be apparent to those skille~ in ! 35 the art. The vapor phase separated in the stabili~er is transmitted to the expander(s3 of the co~pressor~expander _ unit(s) wherein the vapor is depressurized and cooled .
.
-:
1 across the turbine blades of the turboexpander( 5 ) (in the case of a rotary unit) and partially condensed thereby, the enthalpy removed from the vapor by the expander supplying power to arive the compressor. The partially condensed depressurized vapor fraction from the turbo-expander is transmitted to the heat exchanger for thermal contac~ with the pressurized feed gas from the compressor, wherein the partially condensed depressurized vapor fraction is fully vaporiz~d. Finally, the fully vaporized fraction is withdrawn from the heat exchanger and the liquid product phase is withdrawn and recovered from the column.
The two preceding aspects of the invention are particularly applicable where initial capital investment must be minimized and/or where high recovèry of the lighter C1 to C4 hydrocarbon constitue`nts are not major require-ments. In the event that economics or overal- refinery process requirements call for the highest practicable - 20 recovery of the lighter C1 to C4 hydrocarbons, prefer-ably the C2 and` heàvier hydrocarbons or the C3 and heavier hydrocarbons, a further aspect of the invention can be utilized to achieve such higher recovery levels.
In particular, as with the two previously described 2S aspects of the invention, the feed gas is first pres-surized by the compressor ends of one or more turboexpander/
compressor units connected in series as described above.
The resultant compressed feed gas is then cooled to below abo~t 120F in a conventional heat exchanger against cool-ing water or by an air-cooled exchanger. Alternatively, , any part of the cooling required by the compressed gas can be achieved by utilizing the heat available in this stream to supply part or all of the side and/or bottom reboil heat required by the stabilizer fractionation column. The cooled feed gas is then fur~her cooled ana partially condensed to form a two-phase fluid in one _ or more heat exchangers arranged so as ~o achieve the .- . .: ~ .
~ 5~
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1 required amount of condensation. Some of the required cooling, if necessary, can be supplied by one or more supplemental external mechanical refrigeration lnits, the implementation of such units being readily ap~arent to those skilled in the art. The two-phase feed stream exiting the last exchanger would then flow to a separator vessel for removal of the liquid phase. The liquid phase -thereby removed can then directly, or indirectly through one of the feed gas heat exchangers, be introduced into an ;10 intermediate stage of the stabilizer column. The vapor ;fraction leaving the high pressure feed- gas separator can then proceed directly to the inlet of the first turbo-expander or be further cooled and partially condensed in one or more heat exc~langers, again follo~ed by another separator vessel, the vapor from the separator flowing to the first turbo-expander inlet and the liquid being charged into an upper stage of the stabilizer column.
In the event of a need for more than one turboexpander, the outlet two-phase stream leaving all but the last turboexpander in series can be separated into liquid and vapor fractions by separator vessels between each expander, ; to insure that no entrained liquid remains in the inlet vapor stream to each machine. Separated liquid fractions can be directly or indirectly introduced at the proper , 25 point to the stabilizer column. The two-phase stream ! exiting the last expander in the series can be discharyed into a separator situated on top of, and in direct commu-nication with, the fractionation section of the stabilizer column, as previously described. The total vapor stream leaving the separator atop the stabilizer, which comprises i the last expander outlet plus the undesired hydrogen and ! lighter C1 to C4 hydrocarbons contained in the column j feed liquid streams, can then be warmed by heat exchange ! with incoming feed gas as discussed previously. l`he ' 35 operating pressure for the stabllizer column, as a minimum~
.
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--1 o--1 can be adjusted to permit the lean residue gas produc~ to proceed to its final destination, for exa~ple, the refinery fuel gas distribution system. Alternatively, the residue gas product can proceed to outside gas compression facil-ities should a need exist for the hydrogen component, - e.g., in another refinery process. In the event that pressurization of the lean residue gas product is desired, the compressor ends of the expander/compressor units can be utilized for compression of the warmed residue gas product in lieu of feed gas compression as previously described. Also, in a further application of this pro-cess, the power produced by the expanders can be used to drive electric generators, air blowers, dynamometers, or ; other load devices on a common shaft with the expanders.
In a preferred mode of carrying out the afore-mentioned aspects of this invention, the feed gas to the process unit may require one or more types of pretreatment processing. For example, the cryogenic temperatures ; 20 encountered within the basic process may cause some undesirable impurities contained in the feed gas to freeze or form hydrates. Examples of such impurities are water vapor and carbon dioxide, both of which, depending on the process conditions and their concentrations, could free~e within the unit. Thus, if the water content is high I enough to warrant its complete removal, a separate dehydration unit can be installed to process the feed gas prior to its introduction into the facilities described herein. Carbon dioxide can be removed by any of several 1 30 means conventionally utilized for this purpose, such as ! molecular `sieves, amine solution, caustic soda solution, and the like. Where low concentrations of impurities are I present in the feed gas, a freezing point depressant can I be advantageously added to the feed gas prior to its being ' ~5 chilled, in an amount sufficient or as needed to prevent , zss~
-1 1' 1 1 ice formation. Suitable freeæing point depressants incl~de any liquid known to be useful as a feed gas anti-freeze such as a C1 to C3 alcohol, e.g., methanol, ethanol, propanols, or mixtures thereof, these being especially preferred, in view of the fact that the freezing point depressant will substantially remain in the liquid product, - and should therefore be compatible with it and its uses.
BRIEF DESCRIPTION OF THE DRA~INGS
The specific details of the invention and of the best mode known to me for carrying it out can be had by reference to the following description in conjunction with the accompanying drawings wherein:
FIG. 1 is a schematic flow diagram representing a preferred embodiment of the process and apparatus o~
the invention;
FIG. 2 is a schematic flow diagram representing ~0 another preferred embodiment of the invention; and FIG. 3 is a schematic flow diagram representing a third preferred embodiment of the invention.
DESCRIPTION OF PREFERRED EMBODIMENTS
Referring to FIG. 1, which describes the basic process configuration of the i~vention wherein the liquid product fraction of an off-gas feed is recovered in a form suitable for transfer to existing refinery fractionation facilities for stabilization and product separation, the I process flow scheme shown : can be applied to a modern ! catalytic reforming unit which operates at relatively low pressures. Although labeled on the flow diagram as "high pressure rich gas", the flow conditions of reformer off-!35 gas feed stream 1 are usually below about 120F and above ,about 100 psig, desirably between about 90 and about .
.
1~2~558 120DP at between about 140 and about 250 psig, and prefer-ably between about 95D and about 105F at between about 150 and about 200 psig, e.g., 100F at 155 psig. The feed gas 1 is first compressed by the centrifugal compressor 5 end 20 of coMpressor/expander unit 21 to a pressure which is dependent on the power available from the expander, desirably between about 180 and 300 psig, and preerably between about 185 and 275 psig, e.g., about 190 psig. The pressurized feed gas stream 2 from compressor 20 is then - 10 cooled in heat exchanger 22 to a temperature corresponding to the amount of internal or external refrigeration available, desirably between about -130 and about -100F, preferably between about -125 and about -115F, e.g~, -120DF, and partially condensed~ The cold, two-phase 15 stream 4 from heat exchanger 22 proceeds to separator 23 wherein a liquid product phase is recovered from the vapor phase, the latter containing hydrogen and at least one hydrocarbon selected from the group consisting of C1 to C4 hydrocarbons~ The vapor phase stream 5 from separator 20 23 is transmitted to the turboexpander end 24 of expander/
compressor unit ~1 wherein the vapor is depressurized across the turbine blades (not shown) and partially condensed thereby. The work (enthalpy) removed from the ; cold vapor stream 5 by the turboexpander end 24 supplies 25 the power needed to drive the compressor end 20 mounted on the common shaft 25.
, - ' ' Duè to the work removal from the high-pressure ¦ cold vapor within turboexpander 24, the partially con- -! 30 densed exit stream 7 from the turboexpander is at a temperature of between about -250 and about -100F and pressure of between about 20 and about 100 psig, preferably between about -180 and about -1 70aF at between about 55 and about 65 psig, e.g., -1 75F and 60 psig. The par-35 tially condensed steam 7 from turboexpander 24 proceeds .
--`~
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_ -13-1 back through heat exchanger 22 where it is fully vaporized and heated by the incoming feed gas stream 2 to a temper-ature of about 5 to 10 F. below the temperature of the feed gas entering the exchanger, desirably between about 5 100 and about 140F, and preferably between about 130 and about 140F, e.g. 135F. The thus-vaporized and warm lean gas stream 8 from heat exchanger 22 then passes 'through pressure control station 26 which controls the expander outlet pressure. The low pressure lean gas thus 10 obtained is then routed to the refinery fuel system, the flare or a hydrogen recompressor.
The cold liquid hydrocarbon fraction recovered from separator 23 as stream 9 is transmitted to product 15 pump 27 at the operating temperature of the separator, desirably between about -130 and about -100~F,'and preferably between about -125 and about -115F, e.g., -120F. The high pressure liquid stream 10 discharged from pump 27 procee'ds to heat exchanger 22 wherein it is 20 heated by'compressed feed gas stream 2 to a temperature of between about 100 and about 140F,'a`nd preerably between about 130 and about 140F, e.g. 135F, be~ore being exported as stream 13 to fractionating facilities lnot shown).
! 25 Since reformer off-gas normally contains some water vapor (typically 1S-30 ppm), a freeze point depres-sant or antifreeze stream 3 can, if desired, be'injected into the compressed feed gas stream 2 to prevent ice 30 and~or hydrate formation. Suitable freeze point depres-- sants include C1 to C3 alcohols, e.g., methanol, ethanol, or propanol which will remain in the recovered liquid product phase.
!
In carrying out the process depicted in FI~. 1, ' ~; .
~;
. . .
25~8 _ -14-1 the eY~pander/compressor unit is operated basically on speed control. Thus, when the flow of cold vapor stream 5 from separator 23 is equal to or below the designed throughput, all of such flow is processed through the expander. Should higher flo~ts in stream 5 occur which, if not checked, could cause an overspeed situation - for the expander, the excess flow is automatically by-passed around expander/compressor unit 21 through stream 6. Alternate con'trolling parameters, well-known to those skilled in the art, can easily be adapted to the expander unit, depending on the particular circumstances.
Typical unit feed gas and product compositions are shown in Table I.
Referring now to FIG. 2, which depicts a variant of the present process, the re~erence numerals identical to those in FIG. 1 refer to corresponding elements. In this embodiment, the separator 23 is situated on top of and in communication with a pac~ed fractionation section 29, the entire unit 30 being referred to as a "stabilizer~
Liquid separated from the cold, two-phase stream 4 exiting heat exchanger 22 flows down~ard through the packing in the fractionating column or tower 29 of stabilizer 30.
~5 Lighter components are vaporized from the net liquid product leaving the bottom of fractionation tower 29.
Depending on the lightest component desired in the net ~ ' liguid product phasè, stabilizer 30 operates as a de-methanizer (for ethane recovery to bottoms), as a de-ethanizer (for propane recovery), or a depropanizer (for butane recovery). In the configuration shown in FIG. 2~
the stabllizer 30 is operating as a deethanizer such that all of the et'hane and lower boiling constituents are frac-tionated from the net propane and heavier bottom product.
3~ ' '- ,' ~ - - :
~L~LZZ558 _ -15-1 The recovered liquid product phase stream 10 discharged from product pump 27 is routed to and circulated through coil 31 wound through the packing in fractionation column 29. Heat removed from the liquid 5 product in coil _ furnishes supplemental side reboil heat for stabilizer 30. The main reboil heat to fractionation tower 29 is supplied from an outside source 32, such as a hot refinery process stream, low or high pressure steam, electric heater, and the like.
- Typical unit feed gas and product compositions are shown in Table I.
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~ZZS5~ ' 1 Referring now to FIG. 3, an alternate embodi-ment for higher recovery is depicted, in which reEerence numerals id~ntical to those in FIGS. 1 and 2 re~er to corresponding elements. In this embodiment, the feed gas 1 can comprise off-gas from a modern low pressure catalytic reforming unit, reformate stabilizer distillate vapor, and the off-gas from a hydrodesulfurization unit upstream from the reformer. Again, as indicated previously, the operatinq conditions indicated are typical of those encountered in a modern low pressure catalytic reforming unit, but the process is, in fact, adaptable to older catalytic reforming units, operated at 300 psig to 550 psig, or any hydrogen-rich refinery or petrochemical off-gas stream where free pressure drops would produce higher thermodynamic efficiencyt i.e., higher recovery and/or less required external refrigeration~
.
Combined feed gas 1 from a-molecular sieve dehydration unit (not shown) at 100F and 185 psig pro-ceeds initially to the first stage compressar 20, which ison the same shaft 50 as the second stage expander 51, ; wherein the feed gas is pressurized and heated to about 223 psig and 131F. ~rom there, the ~eed gas Elows-to the second stage compressor 52, which is on the same shat 53 as the irst stage expander 24, wherein the feed gas is further pressurized and heated to 265 psig and 167F.
Pressurized feed gas 2 leaving the second stage compressor 52 is then cooled to 115F by an air-cooled ex~hanger 54.
Part of the cooled feed gas 56 is then routed to the 30 stabilizer side rèboiler 58 where it is further cooled to about 30F. The remaining portion 60 of the feed gas, ! amounting to about 70% of the total, is cooled -in heat -exchanger 62 to about -10F by thermal contact with the cooler residue gas product 64. ~he total feed gas is then 35 recombined at 66 and flows through a refrigerant chiller .
.
l~ZZ~ii58 _ -19-1 exchanger 68 which further cools the feed stream to about -40F. The cooling achieved in this exchanger is obtained by an external conventional mechanical refrigeration unit 70. Upon leaving the chiller 68, the cold, partially condensed, two-phase feed stream 72 flows into a separator vessel 74 operating at about -40F and 245 psig. Liquid collected in separator 74 is removed through a level control valve 76 and introduced as stream 78 to stabilizer column 30. Vapor 80 from separator 74 is routed through another heat exchanger 82 where it is further cooled and partially condensed by the cooler residue gas product 64 before flowing as stream 84 into the first expander inlet separator 86, operating at about -119F and 240 psig.
.- Liquid 88 removed by separator 86 then flows via level 15 control 90 to an upper feed stage in fractionation column _ of stabilizer 30. The vapor 85 from separator 86 flows_ to first expander 24. The inlet nozzle-vanes (not shown) on this machine are actuated by a pressure controller (not shown) located remotely in an upstream unit. The cold, 20 partially condensed stream 92 leaving the first stage expander 24 tHen flows into second expander inlet sepa-.-rator 94, opërating at about -155F and 115 psig. Liquid recovered by separator 94 is removed by level controller 96 and joins liquid stream 88 flowing from first expander 25 inlet separator 86 to stabilizer 30~ Vapor 9~ from second expander inlet separator 94 flows into second expander 51.
The inlet nozzle vanes (not sh~wn) on this machine are actuated to control the pressure in second expander inlet separator 94. Both expanders 24 and 51 are also.equipped 30 with bypass control~valves (not shown), who-se operation has been discussed previously. The cold, 1QW pressure, two-phase second expander effluent stream 100, operating at about -188F and 55 psig flows into the separator - section 23 situated above the ~ractionation section 29 of 35 stabilizeFcolumn 30. ' .,, :
.
- : ~
' ~ : ' _ -20-1 Second expander outlet vapor plus fractionated vap~rs rising from the top stage o~ stabilizer column 30 form the residue gas stream 64. This cold vapor stream 64, operating at about -162F and 55 psig, flows back through the previously discussed exchangers 82 and 62, which warm the stream to about 105F. The residue gas stream 64 then flows through a pressure control station 102, and subsequently to the refinery fuel gas distribution system (not shown). The pressure control - 10 station 102 assures that the pressure within this unit - is maintained at a constant value sufficient to permit residue stream gas 64 to flow freely into the fuel system.
Part o~ the reboil heat required for the stabi-lizer column 29, which in this case is operated as a de-ethanizer, is obtained from side reboiler 58, extract-ing heat from incoming feed gas stream 56. The remaining reboil heat for the bottom reboiler can be obtained from any outside heat source 104. The liquid product stream 106 leaving the bottom of stabilizer column 29 i5 pumped to external fractionation facilities (not shown), where it is separated into propane, butanes, and gasoline products.
Typical feed gas, residue gas, and liquid product compositions for this high recovery mode of operation are presented in Table 2.
. Pressure levels maintained withln the cryogenic recovery unit of the apparatus of the invention can vary widely from one application to another, depending on the feed gas supply pressure ~100 psig to 2,500 psig) and ! the residue gas pressure requirement at its destination ' 35 ~ psig to 1,500 psiy).
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~ZZS5~ ' 1 Referring now to FIG. 3, an alternate embodi-ment for higher recovery is depicted, in which reEerence numerals id~ntical to those in FIGS. 1 and 2 re~er to corresponding elements. In this embodiment, the feed gas 1 can comprise off-gas from a modern low pressure catalytic reforming unit, reformate stabilizer distillate vapor, and the off-gas from a hydrodesulfurization unit upstream from the reformer. Again, as indicated previously, the operatinq conditions indicated are typical of those encountered in a modern low pressure catalytic reforming unit, but the process is, in fact, adaptable to older catalytic reforming units, operated at 300 psig to 550 psig, or any hydrogen-rich refinery or petrochemical off-gas stream where free pressure drops would produce higher thermodynamic efficiencyt i.e., higher recovery and/or less required external refrigeration~
.
Combined feed gas 1 from a-molecular sieve dehydration unit (not shown) at 100F and 185 psig pro-ceeds initially to the first stage compressar 20, which ison the same shaft 50 as the second stage expander 51, ; wherein the feed gas is pressurized and heated to about 223 psig and 131F. ~rom there, the ~eed gas Elows-to the second stage compressor 52, which is on the same shat 53 as the irst stage expander 24, wherein the feed gas is further pressurized and heated to 265 psig and 167F.
Pressurized feed gas 2 leaving the second stage compressor 52 is then cooled to 115F by an air-cooled ex~hanger 54.
Part of the cooled feed gas 56 is then routed to the 30 stabilizer side rèboiler 58 where it is further cooled to about 30F. The remaining portion 60 of the feed gas, ! amounting to about 70% of the total, is cooled -in heat -exchanger 62 to about -10F by thermal contact with the cooler residue gas product 64. ~he total feed gas is then 35 recombined at 66 and flows through a refrigerant chiller .
.
l~ZZ~ii58 _ -19-1 exchanger 68 which further cools the feed stream to about -40F. The cooling achieved in this exchanger is obtained by an external conventional mechanical refrigeration unit 70. Upon leaving the chiller 68, the cold, partially condensed, two-phase feed stream 72 flows into a separator vessel 74 operating at about -40F and 245 psig. Liquid collected in separator 74 is removed through a level control valve 76 and introduced as stream 78 to stabilizer column 30. Vapor 80 from separator 74 is routed through another heat exchanger 82 where it is further cooled and partially condensed by the cooler residue gas product 64 before flowing as stream 84 into the first expander inlet separator 86, operating at about -119F and 240 psig.
.- Liquid 88 removed by separator 86 then flows via level 15 control 90 to an upper feed stage in fractionation column _ of stabilizer 30. The vapor 85 from separator 86 flows_ to first expander 24. The inlet nozzle-vanes (not shown) on this machine are actuated by a pressure controller (not shown) located remotely in an upstream unit. The cold, 20 partially condensed stream 92 leaving the first stage expander 24 tHen flows into second expander inlet sepa-.-rator 94, opërating at about -155F and 115 psig. Liquid recovered by separator 94 is removed by level controller 96 and joins liquid stream 88 flowing from first expander 25 inlet separator 86 to stabilizer 30~ Vapor 9~ from second expander inlet separator 94 flows into second expander 51.
The inlet nozzle vanes (not sh~wn) on this machine are actuated to control the pressure in second expander inlet separator 94. Both expanders 24 and 51 are also.equipped 30 with bypass control~valves (not shown), who-se operation has been discussed previously. The cold, 1QW pressure, two-phase second expander effluent stream 100, operating at about -188F and 55 psig flows into the separator - section 23 situated above the ~ractionation section 29 of 35 stabilizeFcolumn 30. ' .,, :
.
- : ~
' ~ : ' _ -20-1 Second expander outlet vapor plus fractionated vap~rs rising from the top stage o~ stabilizer column 30 form the residue gas stream 64. This cold vapor stream 64, operating at about -162F and 55 psig, flows back through the previously discussed exchangers 82 and 62, which warm the stream to about 105F. The residue gas stream 64 then flows through a pressure control station 102, and subsequently to the refinery fuel gas distribution system (not shown). The pressure control - 10 station 102 assures that the pressure within this unit - is maintained at a constant value sufficient to permit residue stream gas 64 to flow freely into the fuel system.
Part o~ the reboil heat required for the stabi-lizer column 29, which in this case is operated as a de-ethanizer, is obtained from side reboiler 58, extract-ing heat from incoming feed gas stream 56. The remaining reboil heat for the bottom reboiler can be obtained from any outside heat source 104. The liquid product stream 106 leaving the bottom of stabilizer column 29 i5 pumped to external fractionation facilities (not shown), where it is separated into propane, butanes, and gasoline products.
Typical feed gas, residue gas, and liquid product compositions for this high recovery mode of operation are presented in Table 2.
. Pressure levels maintained withln the cryogenic recovery unit of the apparatus of the invention can vary widely from one application to another, depending on the feed gas supply pressure ~100 psig to 2,500 psig) and ! the residue gas pressure requirement at its destination ' 35 ~ psig to 1,500 psiy).
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W ~ u~ a~ u~ ~ co ,1 o~ ~ 1~ 0 g ~ ~ a~ r,l , P~ ~ ~ ~ ~ ~ ~ ~ ~ o~ a~ O , ~ ~ 3:: ~1 ~ ~ o ~ ~ ~ ~ ~ o ~ ... ... ....,11 ,~
O ~ o o ~D
~; ~ ~1~
,~ . . ~ J ~ :
3 ~ ~
ow ~
. ~ a) u, a) - ~, . z l ~ ~ ~ :n ~ ~) v V ~D a O
o 3:: ~ v :r: v ,~ m * ~ ~ -o t~
.
. - . : :... :
~lZ2S5~
_ -22-1 The foregoing embodiments are presented for the purpose of illustrating, without limitation, the process and apparatus of the present invention. It is understood, of course, that changes and variations therein can be made 5 without departing from the scope of the invention which isdefined in the claims.
INDUSTRIAL APPLICABILITY
-The present invention provides a cryogenic liquid recovery process which has particular application to petroleum refinery off-gases where free pressure drop is available in the gas stream to be processed, for example, in catalytic reforming units or various types of 15 hydrotreating units where significant volumes of off-gas would otherwise be routinely depressurized to the refinery ; fuel gas system or flared~
Although the refinery off-gas streams are composed 20 predominately of hydrogen, they also contain signifi-cant quantities of ethane, propane, butànes, and gasoline;
the economic justification for the process is derivèd from the difference in value of these prQducts as recovered liquid over their refinery fuel value. Although local 25 marketing conditions woula determine the economics of recovery for ethane and propane, the recovery of butanes.
- and heavier components is justified for most refineriès.
,. . .' ' , . . .
.', - : ' ' . Il - ' ~ -1 3~ .
(~ 0 3 . ~ td ~ C~ ~ . .` ~It)o~
~ ~ P~; a O ~ O ~0 O Or c,~ u~ ~ O s~
W ~ u~ a~ u~ ~ co ,1 o~ ~ 1~ 0 g ~ ~ a~ r,l , P~ ~ ~ ~ ~ ~ ~ ~ ~ o~ a~ O , ~ ~ 3:: ~1 ~ ~ o ~ ~ ~ ~ ~ o ~ ... ... ....,11 ,~
O ~ o o ~D
~; ~ ~1~
,~ . . ~ J ~ :
3 ~ ~
ow ~
. ~ a) u, a) - ~, . z l ~ ~ ~ :n ~ ~) v V ~D a O
o 3:: ~ v :r: v ,~ m * ~ ~ -o t~
.
. - . : :... :
~lZ2S5~
_ -22-1 The foregoing embodiments are presented for the purpose of illustrating, without limitation, the process and apparatus of the present invention. It is understood, of course, that changes and variations therein can be made 5 without departing from the scope of the invention which isdefined in the claims.
INDUSTRIAL APPLICABILITY
-The present invention provides a cryogenic liquid recovery process which has particular application to petroleum refinery off-gases where free pressure drop is available in the gas stream to be processed, for example, in catalytic reforming units or various types of 15 hydrotreating units where significant volumes of off-gas would otherwise be routinely depressurized to the refinery ; fuel gas system or flared~
Although the refinery off-gas streams are composed 20 predominately of hydrogen, they also contain signifi-cant quantities of ethane, propane, butànes, and gasoline;
the economic justification for the process is derivèd from the difference in value of these prQducts as recovered liquid over their refinery fuel value. Although local 25 marketing conditions woula determine the economics of recovery for ethane and propane, the recovery of butanes.
- and heavier components is justified for most refineriès.
,. . .' ' , . . .
.', - : ' ' . Il - ' ~ -1 3~ .
Claims (12)
1. A process for cryogenic recovery of liquid from refinery off-gas, comprising:
(a) feeding the off-gas to a compressor/expander having compressor means and expander means mounted and driven on a common shaft;
(b) compressing the off-gas feed in the compressor of the compressor/expander;
(c) cooling the compressed gas obtained from step (b) in a heat exchanger wherein the gas is partially condensed to form a two-phase fluid;
(d) transmitting the two-phase fluid obtained in step (c) to a separator wherein the liquid product phase is recovered from the vapor phase, said vapor phase containing hydrogen and at least one hydrocarbon selected from the group consisting of C1 to C4 hydro-carbons;
(e) transmitting the vapor phase separated in step (d) to the expander of the compressor/expander wherein the vapor is depressurized and cooled and par-tially condensed thereby, the enthalpy removed from the vapor by the expander supplying power to drive the compressor;
(f) transmitting the liquid phase product separated in step (d) and the partially condensed depressurized vapor fraction obtained in step (e) to the heat exchanger for separate thermal contact with the compressed feed gas obtained in step (b) wherein the partially condensed depressurized vapor fraction is fully vaporized and the feed gas is cooled; and (g) withdrawing the fully vaporized fraction and the liquid product fraction from the heat exchanger in separate streams.
(a) feeding the off-gas to a compressor/expander having compressor means and expander means mounted and driven on a common shaft;
(b) compressing the off-gas feed in the compressor of the compressor/expander;
(c) cooling the compressed gas obtained from step (b) in a heat exchanger wherein the gas is partially condensed to form a two-phase fluid;
(d) transmitting the two-phase fluid obtained in step (c) to a separator wherein the liquid product phase is recovered from the vapor phase, said vapor phase containing hydrogen and at least one hydrocarbon selected from the group consisting of C1 to C4 hydro-carbons;
(e) transmitting the vapor phase separated in step (d) to the expander of the compressor/expander wherein the vapor is depressurized and cooled and par-tially condensed thereby, the enthalpy removed from the vapor by the expander supplying power to drive the compressor;
(f) transmitting the liquid phase product separated in step (d) and the partially condensed depressurized vapor fraction obtained in step (e) to the heat exchanger for separate thermal contact with the compressed feed gas obtained in step (b) wherein the partially condensed depressurized vapor fraction is fully vaporized and the feed gas is cooled; and (g) withdrawing the fully vaporized fraction and the liquid product fraction from the heat exchanger in separate streams.
2. A process according to claim 1 wherein:
the off-gas feed is catalytic reformer off-gas;
the compressor/expander is of the rotary type comprising a centrifugal compressor means and a turbo-expander means mounted and driven-on a common shaft;
the reformer off-gas is fed to the compressor/
expander in step (a) at a temperature of between about 90°
and about 120°F and a pressure of between about 140 and about 170 psig;
the reformer off-gas feed is compressed in step (b) to a pressure of between about 180 and about 210 psig:
the compressed reformer off-gas is cooled in step (c) to a temperature of between about -130°
and about -100°F;
the vapor phase recovered in step (d) contains at least one hydrocarbon selected from the group con-sisting of C1 to C3 hydrocarbons;
the vapor phase recovered in step (d) is depres-surized and cooled in the turboexpander in step (e) to a pressure of between about 50 and about 80 psig and temper-ature of between about -190° and about -160°F; and the liquid phase product fraction separated in step (d) and the partially condensed depressurized vapor fraction are heated to a temperature of between about 100° and about 140°F by the feed gas in the heat exchanger in step (f).
the off-gas feed is catalytic reformer off-gas;
the compressor/expander is of the rotary type comprising a centrifugal compressor means and a turbo-expander means mounted and driven-on a common shaft;
the reformer off-gas is fed to the compressor/
expander in step (a) at a temperature of between about 90°
and about 120°F and a pressure of between about 140 and about 170 psig;
the reformer off-gas feed is compressed in step (b) to a pressure of between about 180 and about 210 psig:
the compressed reformer off-gas is cooled in step (c) to a temperature of between about -130°
and about -100°F;
the vapor phase recovered in step (d) contains at least one hydrocarbon selected from the group con-sisting of C1 to C3 hydrocarbons;
the vapor phase recovered in step (d) is depres-surized and cooled in the turboexpander in step (e) to a pressure of between about 50 and about 80 psig and temper-ature of between about -190° and about -160°F; and the liquid phase product fraction separated in step (d) and the partially condensed depressurized vapor fraction are heated to a temperature of between about 100° and about 140°F by the feed gas in the heat exchanger in step (f).
3. A process according to claim 2 wherein:
the compressor/expander comprises a plurality of compressor/expander units connected and operated in series;
the reformer off-gas is fed to the compressor/
expander unit in step (a) at a temperature of between about 95° and about 105°F and a pressure of between about 150 and about 160 psig;
the reformer off-gas feed is compressed in step (b) to a pressure of between about 185 and about 195 psig;
the compressed reformer off-gas is cooled in step (c) to a temperature of between about -125° and about -115°F;
the vapor phase recovered in step (d) contains at least one hydrocarbon selected from the group con-sisting of C1 and C2 hydrocarbons;
the vapor phase separated in step (d) is de-pressurized and cooled in the turboexpander in step (e) to a pressure of between about 55 and 65 psig and temperature of between about -180° and about -170°F;
the liquid phase product fraction separated in step (d) and the partially condensed depressurized vapor fraction are heated to a temperature of between about 130° and about 140°F by the feed gas in the heat exchanger in step (f).
the compressor/expander comprises a plurality of compressor/expander units connected and operated in series;
the reformer off-gas is fed to the compressor/
expander unit in step (a) at a temperature of between about 95° and about 105°F and a pressure of between about 150 and about 160 psig;
the reformer off-gas feed is compressed in step (b) to a pressure of between about 185 and about 195 psig;
the compressed reformer off-gas is cooled in step (c) to a temperature of between about -125° and about -115°F;
the vapor phase recovered in step (d) contains at least one hydrocarbon selected from the group con-sisting of C1 and C2 hydrocarbons;
the vapor phase separated in step (d) is de-pressurized and cooled in the turboexpander in step (e) to a pressure of between about 55 and 65 psig and temperature of between about -180° and about -170°F;
the liquid phase product fraction separated in step (d) and the partially condensed depressurized vapor fraction are heated to a temperature of between about 130° and about 140°F by the feed gas in the heat exchanger in step (f).
4. A process according to claim 1, 2 or 3 wherein a portion of the liquid product fraction withdrawn from the heat exchanger is recycled to the compressed reformer off-gas feed obtained in step (b) in an amount sufficient to dissolve frozen solids.
5. A process according to claim 1, 2 or 3 wherein a water freezing point depressant selected from the group consisting of C1 to C3 alcohols is added to the reformer off-gas feed obtained in step (b) in an amount sufficient to prevent ice formation, said de-pressant remaining substantially in the liquid product fraction.
6. A process according to claim 1, 2 or 3 wherein a portion of the vapor phase flow to the expander of the compressor/expander in step (e) is adapted to be bypassed to the heat exchanger to prevent overspeed in the compressor/expander.
7. A process for cryogenic recovery of liquid from refinery off-gas, comprising:
(a) feeding the off-gas to a compressor/expander haying compressor means and expander means mounted and driven on a common shaft;
(b) compressing the off-gas feed in the compressor of the compressor/expander;
(c) cooling the compressed gas obtained from step (b) in a heat exchanger wherein the gas is partially condensed to form a two-phase fluid;
(d) transmitting the two-phase fluid obtained in step (c) to the separator portion of a stabilizer comprising said separator and an externally heated packed fractionation column situated beneath and in communication with the separator, wherein the liquid product phase is recovered from the vapor phase by gravity separation, the constituents of said vapor phase being fractionated from the net liquid product leaving the bottom of the tower, said vapor phase containing hydrogen and at least one hydrocarbon selected from the group consisting of C1 to C4 hydrocarbons;
(e) circulating the liquid product phase discharged from the bottom of the fractionation column of the stabi-lizer through a coil within the column whereby heat removed from the liquid product within the column furnished supple-mental side reboil heat for the stabilizer;
(f) transmitting the vapor phase separated in step (d) to the expander of the compressor/expander wherein the vapor is depressurized and cooled and par-tially condensed thereby, the enthalpy removed from the vapor by the expander supplying power to drive the com-pressor;
(g) transmitting the partially condensed depres-surized vapor fraction obtained in step (f) to the heat exchanger for thermal contact with the compressed feed gas obtained in step (b) wherein the partially condensed depressurized vapor fraction is fully vaporized;
(h) withdrawing the fully vaporized fraction from the heat exchanger; and (i) withdrawing the liquid product phase from the column.
(a) feeding the off-gas to a compressor/expander haying compressor means and expander means mounted and driven on a common shaft;
(b) compressing the off-gas feed in the compressor of the compressor/expander;
(c) cooling the compressed gas obtained from step (b) in a heat exchanger wherein the gas is partially condensed to form a two-phase fluid;
(d) transmitting the two-phase fluid obtained in step (c) to the separator portion of a stabilizer comprising said separator and an externally heated packed fractionation column situated beneath and in communication with the separator, wherein the liquid product phase is recovered from the vapor phase by gravity separation, the constituents of said vapor phase being fractionated from the net liquid product leaving the bottom of the tower, said vapor phase containing hydrogen and at least one hydrocarbon selected from the group consisting of C1 to C4 hydrocarbons;
(e) circulating the liquid product phase discharged from the bottom of the fractionation column of the stabi-lizer through a coil within the column whereby heat removed from the liquid product within the column furnished supple-mental side reboil heat for the stabilizer;
(f) transmitting the vapor phase separated in step (d) to the expander of the compressor/expander wherein the vapor is depressurized and cooled and par-tially condensed thereby, the enthalpy removed from the vapor by the expander supplying power to drive the com-pressor;
(g) transmitting the partially condensed depres-surized vapor fraction obtained in step (f) to the heat exchanger for thermal contact with the compressed feed gas obtained in step (b) wherein the partially condensed depressurized vapor fraction is fully vaporized;
(h) withdrawing the fully vaporized fraction from the heat exchanger; and (i) withdrawing the liquid product phase from the column.
8. A process according to claim 7 wherein:
the off-gas feed is catalytic reformer off-gas;
the compressor/expander is of the rotary type comprising centrifugal compressor means and turboexpander means mounted and driven on a common shaft;
the reformer off-gas is fed to the compressor/
expander in step (a) at a temperature of between about 90°
and about 120°F and a pressure of between about 140 and about 170 psig;
the reformer off-gas feed is compressed in step (b) to a pressure of between about 180 and about 210 psig;
the compressed reformer off-gas is cooled in step (c) to a temperature of between about -130° and about -100°F;
the vapor phase recovered in step (d) contains at least one hydrocarbon selected from the group con-sisting of C1 to C3 hydrocabons;
the vapor phase separated in step (d) is depres-surized and cooled in the turboexpander in step (f) to a pressure of between about 50 and 80 psig and temperature of between about -190° and about -160°F; and the partially condensed depressurized vapor fraction is heated to a temperature of between about 100°
and about 140°F by the feed gas in the heat exchanger in step (g).
the off-gas feed is catalytic reformer off-gas;
the compressor/expander is of the rotary type comprising centrifugal compressor means and turboexpander means mounted and driven on a common shaft;
the reformer off-gas is fed to the compressor/
expander in step (a) at a temperature of between about 90°
and about 120°F and a pressure of between about 140 and about 170 psig;
the reformer off-gas feed is compressed in step (b) to a pressure of between about 180 and about 210 psig;
the compressed reformer off-gas is cooled in step (c) to a temperature of between about -130° and about -100°F;
the vapor phase recovered in step (d) contains at least one hydrocarbon selected from the group con-sisting of C1 to C3 hydrocabons;
the vapor phase separated in step (d) is depres-surized and cooled in the turboexpander in step (f) to a pressure of between about 50 and 80 psig and temperature of between about -190° and about -160°F; and the partially condensed depressurized vapor fraction is heated to a temperature of between about 100°
and about 140°F by the feed gas in the heat exchanger in step (g).
9. A process according to claim 8 wherein:
the compressor/expander comprises a plurality of compressor/expander units connected and operated in series;
the reformer off-gas is fed to the compressor/
expander unit in step (a) at a temperature of between about 95° and about 105°F and a pressure of between about 150 and about 160 psig;
the reformer off-gas feed is compressed in step (b) to a pressure of between about 185 and about 195 psig;
the compressed reformer off-gas is cooled in step (c) to a temperature of between about -125 and about -115°F;
the vapor phase recovered in step (d) contains at least one hydrocarbon selected from the group con-sisting of C1 and C2 hydrocarbons;
the vapor phase recovered in step (d) is depres-surized and cooled in the turboexpander in step (f) to a pressure of between about 55 and 65 psig and temperature of between about -180° and about -170°F; and the partially condensed depressurized vapor fraction-is heated to a temperature of between about 100°
and about 140°F by the feed gas in the heat exchanger in step (g).
the compressor/expander comprises a plurality of compressor/expander units connected and operated in series;
the reformer off-gas is fed to the compressor/
expander unit in step (a) at a temperature of between about 95° and about 105°F and a pressure of between about 150 and about 160 psig;
the reformer off-gas feed is compressed in step (b) to a pressure of between about 185 and about 195 psig;
the compressed reformer off-gas is cooled in step (c) to a temperature of between about -125 and about -115°F;
the vapor phase recovered in step (d) contains at least one hydrocarbon selected from the group con-sisting of C1 and C2 hydrocarbons;
the vapor phase recovered in step (d) is depres-surized and cooled in the turboexpander in step (f) to a pressure of between about 55 and 65 psig and temperature of between about -180° and about -170°F; and the partially condensed depressurized vapor fraction-is heated to a temperature of between about 100°
and about 140°F by the feed gas in the heat exchanger in step (g).
10. A process according to claim 7, 8 or 9 wherein a portion of the liquid product fraction withdrawn from the column in step (i) is recycled to the compressed reformer off-gas feed obtained in step (b) in an amount sufficient to dissolve frozen solids.
11. A process according to claim 7, 8 or 9 wherein a water freezing point depressant selected from the group consisting of C1 to C3 alcohols is added to the reformer off-gas feed obtained in step (b) in an amount sufficient to prevent ice formation, said depressant remaining substantially in the liquid product fraction.
12. A process according to claim 7, 8 or 9 wherein a portion of the vapor phase flow to the expander of the compressor/expander in step (e) is adapted to be bypassed to the heat exchanger to prevent overspeed in the compressor/expander.
Priority Applications (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CA382,734A CA1122558A (en) | 1979-04-04 | 1981-07-28 | Cryogenic recovery of liquids from refinery off-gases |
Applications Claiming Priority (4)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| US26,809 | 1979-04-04 | ||
| US06/026,809 US4272270A (en) | 1979-04-04 | 1979-04-04 | Cryogenic recovery of liquid hydrocarbons from hydrogen-rich |
| CA348,107A CA1108522A (en) | 1979-04-04 | 1980-03-21 | Cryogenic recovery of liquids from refinery off-gases |
| CA382,734A CA1122558A (en) | 1979-04-04 | 1981-07-28 | Cryogenic recovery of liquids from refinery off-gases |
Publications (1)
| Publication Number | Publication Date |
|---|---|
| CA1122558A true CA1122558A (en) | 1982-04-27 |
Family
ID=27166623
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CA382,734A Expired CA1122558A (en) | 1979-04-04 | 1981-07-28 | Cryogenic recovery of liquids from refinery off-gases |
Country Status (1)
| Country | Link |
|---|---|
| CA (1) | CA1122558A (en) |
-
1981
- 1981-07-28 CA CA382,734A patent/CA1122558A/en not_active Expired
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