Note: Descriptions are shown in the official language in which they were submitted.
Z5S8
~ -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
-3-
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
- .,
S58
_ -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
~' -` . , .
.
2558
~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
.
.
-:
--8--
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
-12-
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.
- 10
- Typical unit feed gas and product compositions
are shown in Table I.
.
15
.
.
.
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- -18-
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. ' .,,
:
.
- : ~
' ~ : '
25S8
_ -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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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.
,. . .' '
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