Note: Descriptions are shown in the official language in which they were submitted.
This invention relates to the processing of gas streams contain-
ing hydrocarbons and other gases of similar volatility to remove desired :~
condensable fractloDs. In particular, the invention is concerned with
pr oc e s s m g of ~,a s
.
... ~
~ ~: : ,:
:
: ., -
. ~
,
.
1~48397
streams such as natural gas, synthetic gas and refinery gas
streams to recover nost of the propane and a ma~or portion of
the ethane content thereof, together with sub~tantlally all
of the heavier hydrocarbon content of the gas.
Gas streams containing hydrocarbon~ and other gases
of slmiIar volatility which may be processed according to the
present invéntion include natural gas, synthetic gas streams
obtained from other hydrocarbon materials such as coal, crude
oil, naphtha, oll shale, tar sands, and lignite. The major
hydrocarbon components of natural gas are methane and ethane,
i.e., methane and ethane together comprise at least 50% (molar)
of the gas composition. There may also be lesser amounts of
the relatively heavier hydrocarbons such as propane, butanes,
pentanes, and the like as well as H2, N2, C02, and other gases.
A typical analysis of a natural gas stream to be processed ln
accordance with the invention would be, in approximate mol %,
80% methane, 10% ethane, 5% propane, 0.5% iso-butane, l.5%
normal butane, 0~.25% iso-pentane, 0.25% normal pentane, 0.5%
hexane plus, with the balance made up of nitrogen and carbon
dioxide. Sulfur-containing gases are also often found in natu~
~;
ra1 gas.
Recent substantial increases ln the market for the 1;
ethane and propane components of natural gas has provided de-
mand for proce~sses yielding higher recovery levelis of these
1;~ 25 products. Available processes for separating these materials
include those based upon cooling and refrigeration of gas,
oil absorption, refrigerated oil absorption, and the more
recent cryogenic processes uti]izing the principle of gas
expansion through a mechanical device to produce power while
simultaneously extracting heat from the system. Depending
: -
;
-,'~. ' , .' '~ -: ' : .
r ~
1~8397 `~
upon the pressure of the gas source, the richness (ethane and
heavier hydrocarbons content) of the gas and the desired end
products, each oE these prior art processes or a combinatlon
thereof may be employed. `
The cryogenic expanslon type recovery process is now
generally preferred for ethane recovery because it provides
maximum simplicity with ease of start up, operating flexibillty,
good efficiency, safety, and good reliability. ~.S. Patents
Nos. 3,360,944, 3,292,380, and 3~292,381 describe relevant -
processes. -~
In a typical cryogenic expansion type recovery pro-
cess a feed gas stream under pressure is cooled by heat ex-
change with other streams of the process and/or external sources
of cooling such as a propane compression-refrigeration system.
lS As the gas i8 cooled, liquids flre condensed and are collected
in one or more separators as a high-pressure liquid feed contain-
ing most of the desired C2+ components. The high-pressure
liquid feed is then expanded to a lower pressure. The vaporiza-
tion occurring during expansion of the liquid resu]ts in fur-
ther cooling of the remaining portion of the liquid. The
cooled stream comprising a mixture of liquid and vapor is
demethanized in a demethanizer column. The demethanizer is
a fractionating column in which the expansion-cooled stream
is fractionated to separate residual methane, nitrogen and
other volatile gases as overhead vapor from the desired pro-
ducts of ethane, propane and heavier components as bottom pro-
duct.
If the feed stream is not totally condensed, typically
it is not, the vapor remaining from this partial condensation
is passed through a turbo-expander, or expansion valve, to a
lower pressure at which additional liquids are condensed as a
result of the further cool~ of the stream. The pressure
:. ,: . ., , , ~ ,, ~ :, , ,
: . . . . . . .
: ; :
~48397 ` ~
after the expanslon is usually the same pressure at which the ,~
demethanizer is operated. Llquids thus obtained are also sup~
plied a~ a feed to the demethanizer. Typically, remalning ;'
vapor and the demethanizer overhead vapor are combined as the
residual methane product gas.
In the ideal operation of such a separatlon process
the vapors leaving the process will contaln substantially all '~
of ehe methane found in the feed ~as to the recovery plant,
and substantlally no hydrocarbons equivalent to ethane qr ' '
heavler components. The bottoms fraction leaving the demétha~
nlzer wlll contaln substantially all of the heavier components `~
and essentlally no methane. In practi~e, however, this ldeal
sltuatlon 1~ not obtained for the reason that the conventional
, . ~ !
demethanizer ls operated largely as a stripping column. The ;~
methane product in the process, therefore, typically comprises
vapors leaving the top fr'actionatlon stage of the column to- '~'
gether wlth vapors not sub~ected to any rectlfication step. ;~
Substantia'l losse~ of ethane occur because the vapors discharged
.
from the low temperature separation steps contaln ethane and '~
heavler component6 which could be reco~ered if -those vapors
could be brought to lower te~peratures or if they were brought
in contact wlth a 'slgniflc~nt quantity of relative~y heavv
hydrocarbons, for exa~ple, C3 and heavier, capable of absorbing
the ethane.
As described in applicarlt~s co-pending Canadian patent
application No. 271 j-357 filed February 8, 1977 and havi~ v
Campbell, Wilkinson and Rambo as inventor~, improved
ethane recovery is achievet by cooling the condensed high-pres-
sure liquid prior to expans~on~ Such cooling will reduce the
30 ~emperature of the flash-expsnded liquid feed supplied to the '~
:: .
-4- ~
. ,
F ~
..... . . ..
.
- . . .
,., , , , ,~ . .
~415~397
demethanizer and thus improve ethane recovery. Moreover, as described
in the aforementioned application No. 271,357, by pre-cooling the high
pressure liquid feed, the temperature of the expanded liquid may be suffi-
ciently reduced that it can be used as top col~lmn feed in the demethanizer,
while the expanded vapor is supplied to the demethanizer at a
feed point intermediate the top feed and column bottom. This
variation permits recovery of ethane contained ~in the expanded ;~ -
vapor which would otherwise be lost.
It will be obvious that to supply external refriger-
10 ation at this stage of the process is dif f icult because of theextremely low temperatures encountered. In typical demethanizer
operations ehe expanded liquid and vapor feeds are at temper- ;
atures in the order of -120F. to -190F. Accordingly, cooling
of the condensed hi~ h-pressure liquid stream feed can best be
achieved by heat exchange of the condensed high-pressure liquid `
stream feed with streams derived within the process as described ;~
in the above-identlfled co-pending application No. 271,357.
It will be recogni~ed from the foregoing discussion ;~;
that the high-pressure liquid feed generally contains volatile
20 gases (such as methane)~ as well as ~ases of lower volatility
and that coolin~ ot the high-pressure liquid feed upon expansion
.. . , . . ~ . .
results from v aporizaticn of a portlon of the volatile gases.
In accordance with the present invention, the tempera~ure drop
obtained upon e~pansi~n of the high-pressure liquld feed csn
be increased bv comblning that f eed with a .process stream hav~
ing a bubble point lo~er than the bubble point of the high~
- 5-
~8397 1l ~:
pressure liquld feed at the pressure to which the high-pressure
f eed is expanded . Prior to expansi~n, the combined stream is
cooled to a temperature below the temperature of the high-pres- , ;
sure liquid feed.
This may be accompllshed by c~olirlg the high-pressure
liquid stream or the gaseous process stream (or both) prior to
combining them; or by cooling the comblned streams if that is ,
more convenient. Upon expanslGn~ the combined stream will achieve ~ I
a lower refrigerated temperature because of the presence of ~n-
hanced quantities of the m~re volatile components which reduces
the bubble point of the combined stream and which vaporize~ at
the lowest pressure to absorb increased quantities of heat of
vaporization.
It will be recognized that in practical situations, the
bubble point temperature of the high-pressure liquid f eed may be
several degrees or more~ above its actual process temperature due
to non-equilibrium conditions arising during the condensation and
separatlon of high-pressure -liquid and vapor feeds. Such a con-
di~ion also arises when the hl~h-pressure liquid feed is cooled
ln accordance wlth ehe inventlon disclosed ln c~-pendlng appli-
cation 271;357. When the bubble point temperature significantly exceeds
~the actual process temperature of the high-pressure liquid feed, the
temperature drop on expansion is less than the temperature drop which would
be obtained by expanding a high-pressure liquid feed at its bubble point. In
accordance with the present invention, such a high-pressure
liquid feed can be combined with a more volatile process stream as described
above, and with moderate further cooling, provide improved proces~ opera-
tions. This i~ because addition of the gaseous process stream to the high-
pressure liquid feed will result in absorption of volatile gases until the actual
bubble point temperature of the high-pressure liquid feed can be reduced to
::
- 6 - ~
~F7 . ,: ~ . ~
`
31~4~397 :
the process temperature of the high-pressure liquid ~eed. Exp~nsiorl of a
liquid of such a reduced bubble point will result in colder refrigerated
temperatures being achieved.
In one aspect of this invention there is provided a process for ~ :
separation o~ a feed gas into a volatile residue gas and a relatively less ~`
volatile fraction, said feed gas containing hydrocarbons, methane and ethane
together comprising a major portion of said feed gas, wherein
(a) said f eed gas under pressure is. cooled suf f iciently ''~
to partially condense said gas forming thereby a liquid portion of .
said feed gas and a vapor feed gas; . , ~'.. '
(b) at least some of the liquid portion is expanded
in an expansion means to a lower pressure whereby a part of said
.: liquid portion vaporizes to cool the expanded li'quid portion to ,,~
a refrigerated temperature; and `:
(c) at least some of the expanded l'iquid port~on' ~s '
subsequently treated ln a fractionation column to separate sald ,~
relatlvely less volatile fraction; ' , . ..
the improvement .comprising ' ~ ,~
. ,. . :, ~- .
(1) combining at least part of l'i;q-uid port~on ~ ..
(a) with a process stream having a bubble point below the
bubble point of said liq~id portlon (a), to form thereby
- combined st~ea~;
(2) supplying said combined stream to s~ld ex-
~: pansion meanC at a temperature whiich is below the bubble .
point of said llquid portion (8); :
(3) thereaf ter expandlng said combine~d strea~ tc> ~ -
sald lower pressure ~ whereby the ref rigerated temper~ture '_ ~ ~
:~.' ~....
achieved in expsnsion ~tep (b) is reduced, ~nd
(4) therenf ~er suppiy~ng ~t leu~t ~o~e of sald
exp~nded comblned 6treasll to ~id frl-ctlon-tls)n eolu~n.
' ~ ' ' `
' ~ `
,:
48397
In another aspect of this invcntion there is provided a,n apparatus ¦
for the separation of a Ieed gas into a volatile residue gas and a relati~ely
less volatile fraction, said feed containing hydrocarbons, methane and ethane
comprising the major portion of said feed gas, said apparatus having
(a) cooling means to cool said gas under pres_ure
sufficiently to partially condense said gas and form thereby
a liquid portion of said gas and 8 vapor ~eed gas; ' j :
~b) expansion means connected to ssid cooling meanq j :~
to receive said partly condensed feed gas and to expand it .. -
10 't a lo~er pressure, ~7hereby lt 18 further cooled; and ~. ,'
(i) means, for combinin~ at leaqe part of the 'liquid
port.ion obtained from said coolin~ meanq ta) with a process ' :~ :
st'ream having a bubble point below ehe bubble point-of said ~ , :
liquid portlon (a) to form thereby a combined stream; ;~
- (il) cooling means for cooling at least one of -said
part of said liquid portion, s&id process stream and said com-
. - . : :
bined stream sufficiently that said combined stream has a tem- ~ .
perature below the bubble point of said liquid portion ta); ~
,.... (iii) ' expansion meanæ connected to receive said com- . -
bined stream at a temperature below the bubble point of said ;
liquid portion (a) and to expand it to said lower pressure; .,~
and . ~ ::
(lv) means connecting sa~d expansion.means to said
fractionation column to recei~e the expanded combined streaD
and to supply at least a portlon o~ it as 8 feed t~o said frac-
t i ona t i on c o l umn . '
`~
~- 7a
~ ~.
F
,... . , ~; ~ ~ , ..... . ... .
. ..
i~41~397 ; ~ ;
For a fuller undersennding of thls invention, refer~
ence may be had to the follnwing drawings in which:
Figure 1 is a flow diagram of a single-stage cryo- ' ~ :
genic expander natural gas processing plant of the prior ar.t .
incorporating a set of conditions for a typical rich natural :~
gas stream;
Figure 2 i~ R flnw diagram of a slngle-stage cry~- .
genic expander natural gas processing plant oE the prior art .
incorporating a set of condition~s for a typical lean natural ::,
10. gas stream; ' :~ :
. Figure 3 is a flow-diagram'from companion applica-
tion, No. 271,357, illu~tnaring,one technique by means
of which high-pressure liquid feed gas r.an he pre-nnoIed prior '
to expansinn; , . . . ~:
Figure 4 is a flow diagram ilIustrating the appl.ica- ., ~ ,,
tion of the present invention to a feed,pre-cnolin~ process as ,.
descrihed in Figure ~; and . , ' ,~
Figure 5 is a fra~mentary fl~w diagram nf the appli- - .
cation of the present inventinn tn a feed pre-conling pro~ess ,
20 wherein the liquid feed i~'pre-cooled~by a flash-expanded ~ , , ,:~
portlon nf said llguld feed. -O . .- :,.
,, ,~
,~
,
: `,~ ,,;;
:
. ' '~
`~
- 7b - j;',- : : .
`' ','``': '~' ' - ,
. 3~ i ' ,, ' '. , ' ' ' ' , , , , ' ' , .
~ 4~3~
F~gures 6A and 6B are gxaphs of carbon d~oxide ~s.
temperature from one embod~ment of th~s invention compared to
the prior art.
In the following explanation of the above figures,
tables are provided summarizing flow rates calculated for
representative process conditions. In the table appearing
herein, the values for flow rates (in pound moles per hour)
have been rounded to the nearest whole number, for convenience.
The total stream flow rates shown in the tables include all
non-hydrocarbon components and hence are generally larger
than the sum of the stream flow rates for the hydrocarbon
components. Temperatures indicated are approximate values,
rounded to the nearest degree.
~;
Referring to Figure 1, for a fuller description of
a typlcal conventional ethane recovery process, plant inlet
gas from which carbon dioxide and sulfur compounds have been
- removed (if the concentration of these compounds in the plant
inlet gas would cause the product stream not to meet speci~
fications, or cause icing in the equipment), and which has
been dehydrated enters the process at 120F and 910 psia as -
stream 23. It is divided into two parallel streams and cool~d
to 45F by heat exchange with cool residue gas at 5F in ex-
changer 10; with product liquids Istream 26) at 82F. in ex~
changer 11; and with demethanizer liquid at 53F. in demethani~
zer reboiler 12. From these exchangers, the streams recombine
and enter the gas chiller, exchanger 13, where the combined ;~
stream is cooled to 10F. With propane refrigerant at 5F. The
cooled stream propane is again di~ided into two parallel streams
and further chilled by heat exchangei~ith cold residue gas
(stream 291 at ~1~7~. in exchanger 14, and with demethanizer
-8- ~ `
; .
,~ ~ - , ' :
` - / ^`
~4~3~7
liquids at -80F. in demethanizer side reboiler 15. The strea~s
are reco~bined, as stream 23a, and enter a high-pressure separ-
ator 16 at -45F. and 900 psia. The condensed liquid (stream
24) is separated and fed to the demethanizer 19 through expan-
sion valve 30. An expansion engine may be used in place of the
expansion valve 30 if desired.
The cooled gas from the high pre~ssure separator 16
flows through expander 17 where it is work expanded from 900
psia to 290 psia. The work expansion chills the gas to -L25~F.
Expander 17 is preferably a turbo-expander~ having a compressor
21 mounted on the expander shaft. ~or convenience, expander ;
17 is sometimes hereinafter referred to as the expansion means.
In certain prior art eml~odiments, expander 17 is replaced bv
a conventional expansion valve.
Liquid condensed during expansion is separated in
low pressure separator 18. The liquid is fed on level control
through line 25 to the demethanizer column 19 at the top and
flows from a chimney tray (not shown) as top feed Co the
column 19.
It should be noted that in certain embodiments low
pressure separator l8 may be included as part of clemethanizer
19~ occupying the top section of the column. In this case,
the expander outlet stream enters above a chimney tray at
the bottom of the separator section, located at the top of
the column. The 1iquid then flows from the chimne~ trav as
top feed to the demethanizing sectic~n of the column.
As liquid fed to demethanizer 19 flows down the ;
column, it is contacted by vapors which strip the methane
from the liquid to produce a demethanized liquid product at
_9_
... .... .. . , , . . . . -
:: .
~4~397 ~ ~;
the bottom. The heat required to generate stripping vapors ~
is provided ~y heat exchangers 12 and 15. . ::
The vapors stripped from the condensed llquid in
demethanizer 19 exit thro~lgh line 27 to join the cold o~ltlet ;
gas from separator 18 via line 28. The comblned vapor stream
then flo~,s through ].ine 29 back through heat exchangers 14
and 10. ~ollowing these exchanaers, the gas flo-~ throll~h :
compressor 21 driven b~ expander 17 and directly coupled
thereto. Compres~sor 21 compresses the.gas to a ~l~charge
pressure ~f about 305 psia. The gas then enters a compressor:
.
22 and is compressed to a final discharge pressllre of 900 p5ia.
Inlet and liquid component flow rates, outlet liquid
recoveries and oompression re~uirements for this prior art ~ : :
:
process shown in Figure L are given in tlle following table~
15 : TABLE I
: (Fig. 1) :
Stream low Rate_Summary - I.b. ~1 0 1 ~' 9 / H .
STREAM METHANE ET~HA.~E PROP~NE~ BUTANES+ TOTAL : :~
23 1100 222 163 13() 1647:
24 795 202 157 129 1300 ~`
16 ~ 10 S 1 32 ::
: 26 3 162 157 130 453
REC0VERIES
Ethane 72.9~29,296 CAL/DAY ~ ~:
~:, ~ ..
Propaneqn.,?~.39,270 ~Al./DAY
CO~PRESSION HORSEPOWER
--~
~ ~ .
: Refrigeration256 BHP ~:
~: :,
Recompre~sion 892 BHP
. . _
: .:
Total 1148 B~lP
'~."
: ,' ,
. ,
-- I O--
39~f
In Figure 2 a typical lean natural gas stream is
processed and cooled using a prior art process si~ilar to
that shown in Figure 1. The inlet gas stream 33 is cooled to
-69F. and flows to high press~re separator 16 as stream 33a
where the liquid contained therein is separated and fed on
level control through line 34 and expansion valve 30 to ~le-
methanizer 19 in the middle of the column. !
Cold gas from separator 16 flows through expander 17where because of work expansion from 900 psia eo 225 psia, the
gas is chilled to -160F. The llquid condensed during expan-
sion is separated in low ~ressure separator ]R and is~fed on
level control through line 35 to the demethanizer 19 as top ~;
feed to the column.
The data ~or this case are oiven in the following
table:
TABLE II
(Fig. 2)
Stream Flow Rate SI~mmary - Lb. Mo1es/ _
STREAM METHANE ETHANE PROPANE BUTANES+ TOT~L
~ 20 33 1447 90 36 43 1647
`~ 34 280 42 25 39 `~ 391
35 ~133 35 11 4 186 '~
; 36 2 71 36 43 ]55
RECOVERIES
~; 25 Ethane 79.0%17,355 GAL/DAY
Propane 98.2%8,935 GAL/DAY
COMPRESSION HORSEPOI~ER
Refrigeration 0 BPH
Recompression 11~0 BHP
30Total 118'j RHP
: ' '
,,' , :
~ :. . . ~ . :. . .. , . , . .
:... . . : . . .
.:: . , : : , , ,
. . . . . . . . .
31L04839~ : ~:
In the prior art cases discussed with respect to
Figure 1 and Figure 2 above, recoveries of ethane are 73%
for the case of the r~ch gas feed and 79X for the lean gas
feed. It is recopnized that some improvement in vield may ',
result by adding one or more cooling steps follo~ed by one
or more separation steps, or by altering the temperature of ¦
separator 16 nr the pressure in separator 18. Recoveries of
ethane and propane obtain,ed in this manner, while possibly
improved over the cases illustrated by Figure l and Fi~ure
2, are significant]y less than yields which can be obtained,
in accordance with the process of the present inven~inn. ~y
way of illustrat;on th,e process conditions ~f Figure 7 can ~e
altered by re~ucin~ column pres~ure t~ 2~5 psia. At the lower ~ ,
- pressure ethane and propane recoveries are somewhat increàsed ~
(to 82.96X and 98.6h%, respectively): however, the lower ~, I
operatin~ pressure requlres a substantial increase in the - '
horsepower requirements of the process to 1353 BHP. ~ - ,: :`
Figure 3 shows one means, as described in the above-
identified application,No 271,357~filed Februar~ 8, 1977,for pre-cool-
ing the high-pressure liquld feed. In the process of Figur~ 3~ ¦'~;~' -~''
the partially condensed feed gas 33a at -67F. and 900 psio is ,
obtained as described in Fi,gure 2. The~feed gas was a,ssumed to
- be a lean feed gas of the composition,of stream 33 in Figure 2. ,~
The partially condensed gas 33a enters a high-pressure separator
16 where liquid and vapors are separated. ' j,'
Followln~ flrst the vapors 113 leaving separator 16,
the vapors enter a work expansion eng~ne'17 in which mechani~
' cal ener~y ls extracted fsom the vapor portion of the high '''~
'pressure feed. As the vapor ls expanded ~rom a pressure of
;'~
- 1 2
,
' ' ~ ' ' ' "' . ' ., ,.: '. ' .' '
,', - ', " ,~' ~"' ' ' , , , ' '
,' , ". '' ' ,'' ": ' ' " ' ', ." ' ' ' .' ,,
3397
about 900 psia to a pressure of about 250 psia, work expan
sion cools the expanded vapor 113 to a temperature of approxi-
mately -153F. The expanded and partially condensed vapor
113 is supplicd as a feed to demethanizer 19, wherein the
vapor portion rises and forms part of demethani7er overhead
117. Demethanizer overhead 117 at a temperature of -156F.
comhines with vapor.s 116 from fla~sh vaporization described
below to form residue gas stream llR. The combined, cold
residue gas stream 118 then passes throu2h heat exchanger 119.
10 The warmed residue gas at -125F. leavlng heat exchanger 119
then returns to the preliminary cooling stages a~ illustrated,
for example, in Figure 2~ wherein further refrigeration con-
tained in the still col.d vapor ~stream is recovered, and the
vapor is compressed, via com~re~Sor 2~ (see Figure 2) which is
driven bv work expansion engine 17, nnd then further compressed
to a line prcs~sure of 90n p~sia bv supplementarv compres.sor 22.
Turning to the liquid 34 recovered from separator
16, liquid 34 pa.sse~s through heat eYchan~er 119 in heat ex-
change re1ation with the co]d re.sidue gas ]lR. This re.sults .
20 in a pre-cooling of the liquid portion of the partially con- -
densed high pressure feed gas. The .sub-cooled liquid is then
expanded through an appropriate expansion cdevice. .such as ex-
pansion valve l2n, to a pressure of approYimatelv 250 p~sia.
During expansion a portion of the feed wi]i vaporize, result-
ing in coo1ing of the remaining part. In the process as
illustrated in Figure 3, the expanded ~stream leaving expan.sion
valve 120 reaches a temperature of -158F. and enter~s a separa-
tor. The liqui-d portion is separated and supplied as stream
115 to the fractionation column 19 as top feed. It may be
-13-
,: : :, - .
~ : , . . - - , , ~ ........................ :. :
: ... , .. : ., . . .. :
- .~
~483~7
noted that by comparison with Figure 2, the expanded liquids
through line 34 entering the demethanizer column only achieve
a temperature of -134F. Because stream 115 of Figure 3 is
substantiallv cooler, it may be used as top feed to the de-
methanizer to recover ethane in the stream 113. The ethanerecovered is withdrawn in the demethanizer bottoms 125. De-
methanizer bottoms 125 are heat exchanged with incoming feed
to recover refrigeration therein as generally illustrated
in Figures 1 and 2.
In connection with Figure 3~ it should be noted
that for purposes of heat economy there will be one or more
demethanizer reboilers which exchange heat to cool incoming
feed (not shown in Figure 3) as illustrated generallv in
Figures 1 and 2. For purposes of the illustrated process
calculations appearin2 in Figure 3 and set forth in the table
below, two such reboilers have been included as shown in
Figure 2. The rehoilers are significant to the overall heat
economy of the process. Sub-cooling of the liquid stream 34
by overhead vapors 118 reduces the available refrigeration
remaining in stream 118 for feed cooling purposes. However,
the increased loading of demethanizer 19 with liquid stream
' 115 cooled in accordance with Figure 3 provides additional
available refrigeration in the reboilers. Accordingly, the
overall heat balance of the process remains substantiallv un-
affected.
Inlet and liquid component flow rates, outlet re-
covery efficiencies, and expansion/compression requirements
; for the process illustrated in Figure 3 are set forth in the
following table:
.
-14-
~)4~ 7 ~ ~
TABLE III
(Fig. 3)
Stream Flow Rflte Summarv - Lb, ~loles/Hr.
STREA~ METHANE F.THANE rR~PANE BUTANES+ TOTAL
5 33a1447 90 36 43 1647
34 280 42 25 3~ 391
1131167 48 11 ~ 1256
115 ~51 42 25 39 361 ,
116 29 0 0 0 3 n
101181445 10 1 0 14~3
125 2 80 35 43 164
REC~VERIES
Ethane89.1~19,656 6AI./D~
Propane 97.7% 8,~9~ G~L~Y
15 ~COMPRESSI0~1 HORSEr~~FR
Refrigeration n BHP
Recompression ll77 BHP
Total ll77 B~P
For purposes nf further comparison with tile present
invention in the exampl~s set forth below, a second hase ca~se
was cAlculated following the flow plan of Figure 3 ~nd emploving
the same lean ~eed gas. In the modified fl--w plan. the Feed gas
,
to the process .~t 120F. and 9ln psia was coo]ed to -67F. in
the feed pre-coolers (fnr example, exeilangers lO, ll, 12, 14
and 15 of Figure 2) ra~her than -67F. and the c(-lumn was operated
at slightly lower pressure, i.e., 240 ~sia rather than 250 psia.
The result was a slight increase ~ln recbverv of ethane and propane, ~-
together with an increase in horsepower requirements for the pro-
cess. A summary of the modiEied flow conditions and flow rates
for the alternate base case is set forth in Table I~' below:
~ '
~'`';'i '`~
~15-
~` `
` - ~
1~8397 ~
TABLE_IV
(Fig. 3)
Streflm Flow Rate Summary - Lb. Moles/Hr.
_ _ . _ __ _ _ _ __ _ ___ _ _ _
STREAM METHANE ETHANE PROPANE BUTANES+ l'OTAL CONDITIONS
533a1447 90 3643 1647 68F.; 900 psia
34 308 44 2639 424 -68F.; 900 psia
34a 308 44 2639 424 -153F.; 9no psia
34b 308 44 2639 424 -161QF.; 240 psia
1131139 46 10 4 1223 -68F.; 900 psia
10113a1139 46 10 4 1223 -153F.; 240 psia
l15 278 44 2639 297 -l61F.; 240 psia ;
1181446 R 1 0 1479 -160F.; 740 psia
118a1446 8 1 0 1479 -125F.; 740 psia
125 1 87 35~,3 l~ 3~~.
15 RECOVERIES
Ethane90.66~ 19,907 ~AI./~AY
Propane98.08~ 8~928 GAI.!DA-'
COMPRESSION HORSEPOWER
Refrigeration 0 BHP
20Recompression 125~ BHP
Total l258 BHP
; '
Exam~le l
The present invention is il]ustrated by the follow-
ing example which should be considered in conjunction with
Figure 4. Figure 4 is a fragmentary flow diagram wherein a
lean feed gas 33a at 900 psia is cooled to -67F. and supplied
to separator 16. The feed gas is cooled and partially con-
densed by heat exchange with various process streams (these
heat exchangers not being s~own), includin~ side reboilers :;
on the demethanizer column 19 (side reboilers not shown),
-16-
, , , , , ; ~ ~ ! : ' '; ~; ; ;
1a3 4~339~ -
heat exchange with demethanizer bottoms and product gas as
described in Figure 2. If necessary, as indicated in Figures
1 and 2, supplementary external refrigeration may also be pro-
vided. The process conditions described in Figure 4 and the
f]ow rates set forth in Table V below, correspond to the
process of a lean feed gas of the composition set forth in
Table II and Figu~e 2.
Following the process of ~igure 4. the partially
condensed gas 33a containing a liquid portion and a vapor
portion, enters high pressure separator 16 where the liquid
portion is separated. The liquid from separator 16 (~qtream
34) is combined with a portion of the vapor from separator
16 (stream 169). The combined stream then passes through
heat exchanger 154 in heat exchange relation with the overhead
vapor stream 158 from the demethanizer resulting in coo]ing and
condensa~ion of the combined stream. The cooled stream at -152F.
is then expanded through an appro~riate expansion device, such
as expansidn valve 155, to a pressure of approximately 250
psia. Durin~ expansion a portion of the feed will vaporize,
resulting in cooling of the remaining part. In the process
illustrated in Figure 4, the expanded stream l57 leaving expan-
sion valve 155 reaches a temperature of -162F., and is sup-
plied to the fractionarion column 19 a.q top feed.
The remaining vapor from separator 16 (stream 170)
enters a work expansion engine 17 in which mechanical energy
is extracted from this portion of the high pressure feed. As
that vapor is expanded from a pressure of about 900 psia to ~
a pressure of about 250 psia, the work expansion cools the ~;s~:
expanded vapor 153 to a temperature of approximately -153F.
The expanded and partially condensed vapor l53 is supplied
as feed to demethanizer 19 at an intermediate point.
-17-
", .
, . , . . . . . :
397
It may he noted that by comparison with the Eirst
base case of Figure 3 the liquid 115 of said Figure 3 enter-
ing the demethanizer column achieves a temperature of about
-158F. To achieve a lower temperatur~ of -161F. at the
column top in the alternate hase case~ a reduced column
pressure was necessarv The reduced column ~ressur.~ increased
hvrsepower requirement, but only slightl~ improved ~ield. In
Figure 4, as a result of combining the liquid 34 from separator
16 with a portion o~ the high pressure Eeed vapor 169 prior to
sub-cooling in heat exchanger 154, the co]der demethanizer top
feed of -162F. can be realized without lowering the demethanizer
pressure.
Inlet and liquid component flow rates, outlet re-
coveries, and expansion/compression requlrements for the process
of Example 1 are set forth in the followin~ Table V.
TABLE V
(Fig. 4)
Stream F].ow Rate Summary - Lb. MolesiHr.
.
STREAM METHANE ETHANE PROPANE BUTANES+ T0TAL
33a 1447 90 36 43 1647
34280 42 25 39 391
,, . . :
157444 48 27 40 ~67
1581445 7 1 0 i476
`
1592 83 35 ~3 171 ~:
169164 6 2 1 176 -
1701003 42 9 3 1080
RECOVF,RIES
Ethane 92.2X 20,261 GAL/DAY ~ .
Propane 98.3% 8,949 GAL/DAY
~`:~ '`' '
~ , . .
-18
~ 8397
COMPRESSION HORSEPOWER
Refrigeration O ~HP
Recompression 1221 ~IIP
Total 1221 BHP
Comparison of the ethane and propane recoveries as
between Tables III and V shows that in the absence of enrLch-
ing the liquids from separator l6 ethane recovery i~s 89.1%
and propane recovery is 97.7%. Enrichment o~ the separator
liquids in accordance with Example 1 (see Figure 4) increases
ethane and propane recoveries to 92.2% and 98.3%, re~spectively.
Comparison of Tables IV and V further shows that
the improvement in yields obtained in the present process
was not simply the resu1t of increasing the.horsepower re-
quirements. To the contrary, Table IV shows that even when
the process conditions of the base case were a~tered to supply
the demethanizer at a lower pressure, thus increasing horse-
power requirements of the base case to 1.258 horsepower, ethane
and propan`e recovery increased only to 90.~6% and 98.087.,
respectively. When the present invention was employed, as
in Example l, ethane and propane rec.overies increased over
those set forth in the alternate hase case, evcn though some- ~ :
what less horsepower was actually required.
From a preferred design standpoint in the practice of ;~
this invention, particularly for leaner gases, all of the liquid
from separator 16 will be combined with some portion of the vapor
from separator 16. The combined stream will then be cooled and :~
expanded as described. The amount of vapor employed in the com-
bined stream will be sufficient that the combined stream will
provide the cooling duty and temperature needed to control the
top temperature of the demethanizer. The liquids from separator
16, when added to the vapor forming the top column feed, increase
397
the surface tension of the feed at column condltion, therebg
minimizing the formation of small liquid particles whlch are ;
difficult to separate from the top vapor stream.
For richer gases~ where there is more liquid from
separator 16 than required to maintain the column top condition,
it may be more economical from a design standpoint to d~vide
the liquid from separator 16, and to expand a ~ortion directly
into the tower, or possibly after some sub-cooling, This may
make possible savings in heat exchange requirements and higher ~;
10 recovery.
As set forth in the above-mentioned application ¦
No. 271,357there-are a variety of modified nOw }
plans characteri~ed by æub-cooling of some or all of the liquid
feed obtained from separator 16 to which the present invention
is applicable. T~o or more of these techniques may be used
concurrently. Among these flow plans are the following:
1. Uncondensed vapors leavin~ separator 16 may be ;` ~;
expanded such as in a work expansinn engine to produce a cnld
partially c,ondensed liquld and gas. The liquids are separated
and supplied to the demethanizer column. All o~ a portion
of the liquid thus separated may be used as a sour~e of re-
fri~eration for sub-coolin~ liquid condensace 34 recovered
in separator 16 AlternDtively, all or a portion of the
cntire expanded vapor stream may be used. Additlonally, side
demethanlzer reboilers may be used to provi~e sub-cooling
of condensate 34 from separator 16. In Accordance wlth the
present invention. flow plans can be modlfied by eomhin$ng
liquid condensate from separator 16 wlth ~ portion of the vapor
~.
from that separator prior eo sub-cooling and flashlng of the
llquid condensate.
: .
-~0~
:: . ~:. . : , . : , .
~-...... , . .. . : . . : .- . '
3~7 :-
2. Liquid condensate from separator 16 may he di- ;
rected through a suh-cooling heat exchanger flnd thence to an
expansion valve wherein it is expanded from line pressure
(e.g., 900 psia as in Figures 1-4) to demethanizer col~mn
operating pressure. This will result Ln a vapor-liquid mix-
ture which can he separated elther in a separate low pres-
sure separator or maV be ~ed directly to the demethanizer
column with co]umn internals de.signed to effect the necessary
vapor-liquid separation. The flashing results in furtber
cooling of the feed to the column. A portion of the further
cooled liquid thereby obtained is employed as the coolanr in
heat exchange with the high pressure condensate from separator
16 and then supplied to the demethanizer column as a secolld
feed at an intermediate point in the cnlumn. In accordance with
the present inventiol~, such a process can be improved by en-
riching the liquid condensate leaving separator 16 with a portion
of the vapors from that separator prior to sub-cooling and flash~
ing of the liquid condensate. ~ ~'x
3. The uncondensed vapor leaving separator 16 may
be expanded such as in a work expansion engine from a high
pressure (e.g., 900 psia as in Figures 1-4) to the operating
pressure of the demethanizer and the entire cooled gas-liquid
mixture resulting from expansion may be used to sub-conl the ;~
condensate recovered in separator 16. The sub-cooled con-
25 densate is thereafter flashed and is supplied to the de- -
methanizer as a feed. This embodiment may be improved by
enriching the liquid condensate recovered from separator 16
with a portion of the vapors leaving that separator prior
to sub-cooling and flashing of the liquid condensate.
-21-
., ~ , . . . .
-,:. ,- . , :. ,,, :,
::, - : :.: : ",
: : :- , . ', : : ,, : ,
, : , .. , ,, ., . :: . . ~ .. :
::
:
~4~33~7 ~
In lieu of or in addltion to the foregoing additional
external refrigeration may be provided if increased yields are
required; however, one of the advantages of the invention de-
scribed in said application No. 27l,357 i9 that wherethe condensate ~ ~ : '' `.":
is sub-cooled, i~proved yields frequently may be-obtained with-
out the necessity of increasing process horsepower requirements.
Still another embodiment of the present invention
i~ set forth in the following example, which should read in
~conjunction with related Figure 5: -
Example 2
Figure S is a fragmentary flow process diagram for - `~
recovery of ethane and heavier-components from a hydrocarbon ~ ;
feed gas containing methane, ethane and heavier hydrocarbons. `` `
As illustrated in Figure 5, a partially condensed;high-pressure
~eed gas 174 is provided to a separator 16 at -55F.~ and 900 ;
psia. Cooling of the feed gas to -55F. may be provided as ;
shown for example, ln Figures 1 and 2 by heat exchange to the
feed gas with residue methane gas and other process streams such
as demethani=er side reboil~ers and bottom streams (these heat ;
exchangers not being shown), and, if necessary, with appropriate
external refrigeration. For purposes of calculations on whicb
this example is bssed, two demethanizer reboilers have been
assumed. However, in contrast to Figures l~and 2. the process
calculations (e.g., temperature, pressure and flow rates h~ve ;~;
been based on sn sssumed gaseous feed intermediate in composl~
tion between the lean and rich case gases set forth in Figure~ ~ -~
l and 2 and Sable I and II.
~ ' ~
:, ! .
22 ~ !
339i7
As ind~cated in accompanying Figure 5, the liquid
and vapor portions of the partially condensed feed 174 are
separated in separator 16. The vapor from separator 16 is
divided into two portions. The first portion 176 flows
through expander 17 where, because of work expansion from 900
to 290 psia, it is cooled to about -133F. From expander 17
the c~illed vapor flows to demethanizer 19 as its middle feed. --
The second vapor portion 177 is combined with a portion of
the sub-cooled liquid from heat exchanger 184 as it flows to ~-
heat exchanger 185.
The liquid 175 from separator 16 flows through ex-
changer 184 where it is sub-cooled to -130F. by heat exchange ~
with the cold stream from expansion valve 182. The sub- ~ ~ `
cooled liquid is then divided into two portions. The first
portion 178 flows through expansion valve 182 where it under~
goes expansion and flash vaporization as the pressure is re-
duced from about 900 to 250 psia. The cold stream from expan-
sion valve 182 then flows through exchanger 184 where it is -
used to sub-cool the liquids from separator 16. From exchanger ~ ;
184, the strearn flows to demethanizer 19 as its lowest feed
at -67~
The remaining liquid port1on 179 from exchanger 184,
still at high pressure, combines with a po~tion 177 of the ~ ;
vapor stream from separator 16. The combined stream then flows
through heat exchanger 185 where it is cooled to approxi- ;~
mately -140F. by heat exchange~ with column overhsad stream
180. At this temperature, the combined strearn is substantially
-.
condensed. The condensed stream then enters expansion valve
183 where it undergoes expansion and flash Yaporization as
the pressure is reduced from 895 psia to 2S0 psia. From ex
pansion valve 183 the cold stream proceeds to demethanizer
19 as its top feed.
:: :
-23-
. -- -- - . : ..... : : ,, :
.,: ;.:: :. , :
83~7
, .
Inlet and liquid component flow rates~ outlet re-
covery efficiencies, and expansion/compression requirements ~ ~-
for the embodiment of this invention as illustrated in Figure
5 are given in the following table~
TABLE V :~
(Fig. 5)
Stream Flow Rate Summary - Lb. Moles/Hr. , :
STREAM METHANE ETHANE PROPANE BUTANES~TOTAL
174 1304 162 80 54 1647 ;.
175 486 109 66 51: 723 `~
176 723 47 12 2 817
177 95 6 2 1 107
178 ~43 54 33` 26 361
179 243 55 33 25 362
.. ~i .
: 180 1301 14 1 0 1362
181 3 148 79 :54 285
. s ~
~ECOVERIES
Ethane 91.47% 36,039 GAL/DAY
,- ,
: Propane 98.38% 19,732 GAL/DAY ; .~:~
:20
HORSEPOWER REQUIREMENTS -
Refrigeration 130 BHP ; `~
: ~-': : '
: Recompression987 BHP
~: Total1117 BHP
~ .,. : ,
'
~ ",~
-2
'
397
It is noted that in addition to the procedure out~
lined in Figure 5 for handling the cooled liquid 175 from
separator 16, other alternate procedures may he used in
some situations to advantage. One alternate procedure involves
carrying a portion o~ the cooled liquid 179 directly from the
separator through another expansion valve dlrectly into the
demethanizer column l9 at intermecliate level.
In a second alternate proceclure, the liquid 17~
from separator 16 can be sub-cooled by residue gas instead of
auto refrigeration as illustrated in Figure S. In this alter-
nate, high pressure condensate may be cooled in two successive
heat exchangers, each employing residue gas as one refrigerant.
After passage through the first exchanger, partly cooled high :
; ~ pressure condensate is divided into two parts. The first part
~ is expanded through an expansion valve and supplied to the de-
methanizer column 19 as an intermedia~te feed. The second part
of the partly cooled condensate continues through the second
exchanger where it is further cooled and then comblned with
vapor from separator 16. The combined stream is then further
. 20 cooled and expanded whereafter lt is sopplied to the column 19
as top feed. Alternatively, vapor from separator 16 could be
~: , : ,
added to the second part of the partly cooled liquid stream be-
fore entering the second exchanger, thus eliminating the need ~;
for subsequent cooling of the combined streams.
. . v
In still another modification of the present inven~
tion, the flash-expanded stream, such as stream 186 of Figure
~ :
. ,` ~
- -25- ~ -
:
~ 8397 ~
5, may be dlrected into heat exchsnge relation with the work
expanded vapor stream 187, thus cooling stream 187 and warming
: ~
stream 186. If stream 187 i9 thus cooled sufficiently, it may ~
be advantageous to employ stream 187 as the top feed to the de- ~ ~;
methanizer and stream 186 as an ineermediate feed since, as is
evident from the process flow plan of Figure 5, stream 186 is
,
richer in heavier components, i.e., C2~, and stream 187 con-
tains more lighter components, e.g., methane and uncondensed .
gases. ;~
Other alternate procedures for obtaining the cooled `
liquid 175 are described in aforementioned application
No. 271,357 These alternate procedures may be used in
various combinations, when appropriate. Also, these various
schemes may be used in place of or ln conjunction with cooling
provl-ded by residue gas to the enriched stream, prior to its use
as top feed to the column 19
These alternate procedures are particularly useful when, ~
i :: :
because of the richness of the feed to the p-rocess, the cooling
capacity of the overhead gas stream 180 ls insufficient to cool ¦ ~ i
the entire volume of liquid recovered through line-175 to the
I desired low temperature.
.:. ~ :
- As is well known, natural gaa streams usually contain
carbon dioxide, so~etimes in substantial amounts, The presence
of carbon dloxide in th~ demethanizer can lead to icing of the I~
column internals under cryopenic conditions. Even when feed
feed gas contains less than lX carbon dio~ide, it fractio~a~es
in the demethan$zer and can build up to concentrat$Dns of as much j~
.`'~
.
-26- .~
: ~ ~
-:- :
,: .,: : : ,
~0~3397 ~
as 5% to 107, or more. At such high concentrations, carbon di-
oxide can freeze out depending on temperatures, pressure,
whether the carbon dioxide is in the liquid or vapor phase,
and the liquid phase solubility.
In the present invention it has been foùnd that when
the vapor from the high-pressure separator is expanded and
supplied to the demethanizer below the top column feed position
the problem of carbon dioxide icing can be substantially miti-
gated. The high-pressure separator gas typically contains a
large amount of methane relative to the amount of ethane and
carbon dioxide. When supplied at a mid column feed position,
therefore, the high-pressure separator gas tends to dilute the
carbon dioxide concentration, and to prevent it from increasing
to icing levels.
The advantage of the present invention can be readily
seen by plotting carbon dioxide concentration and temperature
for various trays of the demethanizer. To illustrate the pre-
paration of such a chart the flow process illustrated in Figure
4 was applied to the treatment of a feed gas of the following
composition~
Feed Gas Composition
Methane 93.82
Ethane 3.16
Propane 1.06
; 25 Butane + .80
C2 .59
N2 57 ;
The principal operating conditions for the process ;
were the following~
-27-
3~
:
Pressure in high-pressure separator 16 895 psia
Temperature of high-pressure separator 16 0F
% of feed condensed in separator 16 .44%
% of gas from separator 16 to expander 17 60%
S Temperature of combined stream to expansion
valve 155 -120F
Temperature of expansion valve outlet -147F
Column overhead temperature -145F
Temperature of gas from expander 17 -83F
Pressure in demethanizer 360 psia
% ethane recovery B7.33~ ~ ~
% propane recovery 97.05% ~ -
Horsepower~
Recompression 3194 BHP
Refrigeration 0
: Total : ~ 3194 BHP
For base-case purposes the same feed gas was treated
; also in accordance with the process ~of Figure 2. However, :;
for more efficient utilization of available heat duty, the feed
pre-cool~ing exchangers prior to the high-pressure separator
were slightly rearranged. The principal operating parameters
were the following~
::
Pressure in high-pressure separator 16 895 psia ;~
Temperature in high-pressure separator 16 -70F
% of feed gas condensed in separator 16 2.95%
:: Temperature of expanded gases leaving
- expander 17 -136F `~
;:~ Temperature of expanded liquid leaving
flash valve 30 -116F
Temperature of demethanizer overhead vapor-134F
~; Pressure of demethanizer 360 psia :
' -28-
, ~ :
, " " ., ., . .. . .. . . , .. . . ..... , ...... . . . ~ . . , , . . . .. ~ . :
... . , , : :
3397
% ethane recovery 60.92%
% propane recovery 90,58%
Horsepower:
Recompression 3074 BHP
Refrigeration O_BHP
Total 3074 BHP
Plots were made for each of these cases of C2 con-
centration as a function of temperature in the demethanizer,
as shown in Figure 6A and 6B. Also shown on these figures are
the liquid-solid and vapor-solid equilibria. The equilibrium
data given in Figures 6A and 6B are for the methane-carbon di- ~ -
oxide system. These data are considered generally representa-
tive for the methane and ethane systems. If the C02 concen-
tration at a particular temperature in the column is at or above
the equilibrium line at that temperature, icing can be expected.
For practical design purposes, the engineer usually requires a
margin of safety, i.e., the actuaI concentratlon should be less
- than the "lcing" concentration by a suitable safety factor.
factor.
as in Figure 6A the carbon dioxide concentration ln the demeth-
anizer rose well above the tolerable level. Such a gas could
not be used in a conventlonal process, therefore, without pre- ;
treating it to remove a substantial amount of the carbon
dioxide. By contrast, when the expanded vapors are employed as
a mid column feed in accordance with the present invention, the
C2 concentration is reduced in the demethanizer to a point well
below the "icing" level.
!,
'~ ;'''''''~'
-29- ~
~,' . ~.,
, j .. . . . . . . . . . . . . . .
-: . . . . .
~IDgL~35~ :
It should be noted in connection with the foregoing
that when designing demethanizer columns for use in the present
invention, the designer will routinely verify that icing in
the column will not occur. Even when vapor is fed at a mid-
column position it is possible that icing may occur if theprocess is designed for the highest possible ethane recovery.
Such designs normally call for the coldest practical temperature
at the top of the column. This will result in the carbon dioxide
concentration shifting to the right on the plots of Figures 6A
and 6B. Depending on the particular applicstion, the result can
be an objectionably high concentration of carbon dioxide near
the top of the column. For such circumstance, it may be neces-
sary to accept a somewhat lower ethsne recovery to avoid column
icing, or to pre-treat the feed gas to reduce carbon dioxide
levels to the point where they can be tolerated in the demethan-
æer. In the alternative, it may be possible to avoid icing
in such a circumstance by other modifications in the process
, .~
conditions. For instance, it may be possible to operate the
high-pressure separator at a higher temperature, to increase
20 the relative amount of gas from the high-pressure separator `';
which is expanded through expander 17, or to expand a part of
the vapors from the high-pressure separator through an isen- `
thalpic expansion valve. If such alterations can be made with~
in the limitations of the process heat balance, icing may be
avoided without losing ethane recovery.
. ~
In connection with the process descrihed above, it
should be noted that in some instances the feed to the top of
the demethanizer is a liquid which is expanded from a high
pressure to the pressure of the demethanizer (see for example,
Figures 4 and 5). In such cases it may be desirable to auto-
~v ~30-
" .. , - . . ....... . . . . . : :, . ,. ., -. . . -~ .
,... .. . . ....... ..... . .
~. .. ... . . .. . . . .
1~8397
cool the top column liquid feed. This is accomplished by
divlding the top column liquid feed into two streams either
before or after expansion. (Both streams are expanded if the
top column feed is divided before expansion). One of the two
expanded streams thus obtained is directed into a heat exchange
relation with the top column feed prior to expansion.
.-
~. .
' ~
:, ' .
-31-
~48397
SUPPLEMENTAR~ DISCLOSURE ;~
Figure 7 is a process flow plan illustrating the `~
importance of wor~ expanding at least part of the high-
pressure vapor.
Figure 8 is a carbon dioxide-temperature diagram-
comparing the processes of Figures 4 and 7.
In carrying out the present invention, it is ~;
important that at least part of the high-pressure vapor remain-
:
ing after cooling and partial condensation of the feed be
expanded in a work expansion engine to the demethanizer and ~ -~
supplied as a mid-column feed. There are two reasons for
this:
-, . . .
~ 1) Extraction of work energy from the high-pressure
vapor stream by expansion in a worX engine provides a signifi-
cant amount of refrigeration to the process. If work is not
extracted from this stream, it is necessary to supply external
~` ' refrigeration and, because of the low temperatures required,
;~ ~providing that refrigeration may become uneconomic. Additional~
ly,-recompress~.on requirements are increased, since if the
high pressure vapor is work expanded to cool it the energy
extracted is available to supply some of the recompression
requirements of the process.
~ 2~ The vapor supplied to the mid-column feed position
serves to dilute the carbon dioxide present in the liquids
supplied to the top of the column. If the carbon dioxide i8 .
not diluted, it will tend to accumulate in the upper stages
of the column and cause CO2 icing.
~ It should be noted that where rich gases are pro~
cessed, the liquifaction temperature may be sufficiently
high that total condensation can be practical, as shown in ~ ; :
- SD-32 ~
- ~ ~ .
:
~48397
the ab~ve-mentioned application No 271,357. ~ :
The importance of utilizing at least a portion of ~ : :~
the high-pressure vapor stream in a work expansion engine
may be seen by the following illustrative case, in which the
feed gas of Example 3 is processed. In the illustrative ~.
case, all of the high-pressure vapor is recombined with the
high-pressure liquid condensat`ion prior to flash expansion ;
of the latter to the fractionation co~umn.
; In explaining this illustrative case, reference will
10 be made to Figure 7. As shown in Figure 7, incoming feed ~ :~
is cooled by heat exchange with product liquid (exchanger 191~, ~
demethanizer reboiler (exchangar 192) and partially warmed ;; ~ :
residue gas texchanger 193). The feed is further cooled by ~
external propane refrigeration to -14F. (exchanger 194). - ~;
Additional cooling is extracted from residue gas (exchanger 196). ;
In this manner, the major portion of the incoming feed gas
is cooled and supplied to separator 197 at -79F. and 895 psia.
Liquid ~rom separator 197 is further cooled in heat exchanger- -
198, and then recombined wi~h vapor therefrom. Separate
cooling of ~he liquid permits advantageous design of the
Iiquid-liquid heat exchanger (see, for example, the discussion
of this in United States patent 3,874,184 to ~arper et al.). .
The recombined stream is further cooled in exchanger l99to
-94F., flash expanded to the demethanizer pressure of 250 psia
in flash valve 200 and supplie~ as the top column feed to
- -demethanizer 19 at -145F.
;~
Inlet and liquid component flow rates, outlet - ~ :
recovery and expansion/compression requirements for the pro~ess
of Figure 7 is shown in the ~ollow~ng Table VIII: ~
~''.' " :': ~`
- SD-33 - ~ .
. '' ~ ;'
,, , . . ,j . . . ,., : :
~83~7 ~ i
TABLE VI~
(Figl 7)
.... .. .. . , ,.~ .
STREAM FLOW PLAN SUMMARY - LB. MOLESfHR, ~-
STREAM METHANE ETHANE PROPANE BUTANE+ TO~L
33 6181 208 71 123 6588
201 3819 =~4=-~-- 14 6 3972
202 2362 124 57 117 2616
203 6177 25 3 1 6263
204 4 183 68 122 325 ~`
RECOVERIES
Ethane 87.6% 20,243 GAL/D~Y
Propane g5.8%8,982 GAL/DAY
HORSEPOWER REQUIREMENTS
Recompression 3792 BHP
Refrigeration 261 BHP
Total 4053 BHP
As can be seen by comparing the foregoing with
: .: : --
Example 3, to achieve essentially the same recovery of ethane `~
a great increase in horsepower is re~uired. The increase~in `-i ;
~20 hoxsepower arises not only because of the unavailability of
expansion work through expansion of a portion of the high
pressure vapor stream, but also because external refrigeration
was required to achieve the temperature level needed to obtain
the desired ethane recovery.
It is also important to note that by expanding a
portion of the high-pressure vapor and providing it as a mid~
colum~ feed, the carbon dioxide level in the column is
reduced, and column icing conditions are thereby avoided. '~`
This is best seen by constructing carbon-dioxide-temperature ~ ~;
diagrams in the ~ame manner as Figures 6A and 6B were
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constructed. ~hen following the process of the prior art,
serious carbon dioxide icing problems are encountered in
both`liquid phase ~line 205 of Fig. 8) and vapor phase ~line
206 of Fig. 8). ~owever, when the process of the present
invention is used, the carbon dioxide icing is avoided, see
Figure 6B.
In the practice of the present invention, it will
be recognized that the amount of vapor which is work expanded ~
and supplied to the mid-column feed position will depend upo~ ~ -
the amount of refrigeration which can be economically
extracted from lt balanced against the advantage of a reduced
column overhead temperature which can be obtained by using
that same gas to enrich the high-pressure liquid which is
flash expanded to supply the top column feed. Selection of
the amount of vapor work expanded and supplied to a mid-
column feed position will also take into consideration the
amount of ~apor which must be supplied to the mid-column
position in order to avoid carbon dioxide icing. Ag a
general rule of thumb, we have found that for best results
at least about 25% of the gas should be work expanded and
supplied to the mid-column feed position and, for lean gases,
about 50% or more of the gas should be work expanded.
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