Note: Descriptions are shown in the official language in which they were submitted.
10gL~ ~3
This invention relates to the processing of gas streams
containing hydrocarbons and other gases of similar volatility
to remove desired condensable fractions. In particular, the in-
vention is concerned with processing of gas streams such as
natural gas, synthetic gas and refinery gas streams to recover
most of the propane and a major portion of the ethane content
thereof together with substantially all of the heavier hydro-
- carbon content of the gas.
Gas streams COntaining hydrocarbons and other gases
10 of similar volatility which may be processed according to the
present invention include natural gas~ synthetic gas streams
obtained from other hvdrocarhon materials such as coal, crude
:
oil, naphtha, oil shale, tar sands, and lignite. Natural gas
usually has a major proportion of methane and ethane, i.e.,
15 methane and ethane together comprise at least 50 mole percent
of the gas. 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 in accordance
20 with the invention would be, in approximate mole %, 80~ methane,
lOX ethane, 5% propane, 0.5% iso-butane, l.5% normal hutane,
0.25Z iso-pentane, 0.25% normal pentane, 0.5% hexane plus,
with the balance made up of nitrogen and carbon dioxide. Sul-
fur containing gases are also often found in natural gas.
Recent substantlal increases in the market for the
ethane and propane components of natural gas has provided
demand for processes y~eldlng higher recovery levels of these
:' :
- 2 -
?3
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 utilizing the principle of gas
expansion through a mechanical device to produce power while
simultaneously extracting heat from the system. Depending
upon the pressure of the gas source, the richness (ethane and
heavier hydrocarbons content) of the gas and the desired end
products, each of these prior art processes or a combination
thereof may be employed.
The cryogenic expansion type recovery process is now
generally preferred for ethane recovery because it provides
maximum simplicity with ease of start up, operating flexibility,
good efficiency, safety, and good reliability. U.S. Patents
, Nos. 3,360,944, 3,292,380, and 3,292,381 describe relevant pro-
cesses.
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.
; 25 As the gas is cooled, liquids are condensed and are collected
in one or more separators as a high-pressure liquid feed con-
tainlng 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 results in further
cooling of the remaining portion of the liquid. The cooled
stream, comprising a mixture of liquid and vapor, is demethanlzed
in a demethanizer column. The demethanizer is a fractlonating
column in which the expansion-cooled stream is fractionated
to separate residual methane, nitrogen and other volatile gases
as overhead vapor from the deslred products of ethane, propane
and heavier components as bottom products.
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. Additional llquids are condensed as a result
of the further cooling of the stream by expansion. The pres-
sure after the expansion is usually the same pressure at which
- 10 the demethanizer is operated. Liquid thus obtalned is fllso
supplied as a feed to the demethanizer. Typically, the remain-
ing vapor and demethanizer overhead vapor are combined as the
residual methane product gas.
In the ideal operation of such a separation process
the overhead vapors leaving the process will contain substan-
tially all of the methane found in the feed gas to the recovery
plant, and substantially no hydrocarbons equivalent to ethane
or heavièr components. The bottoms fraction leaving the de-
- methanizer will contain substantially all of the heavier com-
- 20 ponents and essentially no methane. In practice, however,
this ldeal sltuatlon ls not nbtalned largely for the reason
that the conventlonal demethanizer is operated largely as a
strlpplng column. The methane product ln the process, there-
fore, typlcally comprlses vapors leavlng the top fractionatlon
stage of the column together with vapors not sub~ected to any
rectification step. Substantial losses of ethane occur because
the vapors remaining from low temperature separatlon steps con-
taln ethane and heavier components whlch could be recovered
lf those vapors could be brought to lower temperAtures or lf
they were brought ln contact with a significant quantlty of
relatively heavy hydrocarbons, for example, C3 and heavier,
--4--
,~
,:
.
:
capable of absorbing the ethane. Overall recovery of ethane
can be further increased by altering the temperature distribu-
tion in the demethanizer column so as to decrease the tempera-
ture at the upper stages of the column by removing heat rom
one or more of the feeds thereto. The present invention pro-
vides the means for achieving either or both of the objectives
that significantly increase the yield of desired products.
In one aspect of this invention there is provided
in a process for separating a feed gas into a volatile residue
gas and relatively less volatile fraction, said feed gas con-
taining hydrocarbons, methane and ethane together comprising
.~ the major portion of said feed gas, wherein said gas is cooled
sufficiently under pressure so as to convert at least some of
said gas to liquid, and,
- (i) any vapor portion under pressure is expanded to a
lower pressure, whereby it partially condenses,
. (ii) ~aid liquid portion under pressure is expanded to
said lower pressure, whereby a part of said liquid portion
- vaporizes to cool the expanded liquid portion; and
(iii) at least the liquid formed upon partial condensation
of said expanded vapor, and the remaining liquid portion of
said expanded liquid portion are supplied to a fractionation
column wherein said relatively less volatile fraction is
separated,
the improvement wherein
(a) at least some of said liquid portion under pressure
is subcooled to a temperature below its bubble point prior to
expansion thereof;
; (b) at least a part of sald subcooled liquid portion
under pressure i6 expanded to said lower pressure; and
_ 5 _
0~;3
(c) at least ~ portion of the liquid remaining in the
. expanded subcooled liquid portion is supplied to said
distillation column as a top liquid feed thereto.
- In another aspect of this invention there is provided
in.a process for separating a feed gas into a volatile residue
gas and a relatively less volatile fraction, said feed gas
containing hydrocarbons, methane and ethane together comprising
. the major portion of said feed gas, wherein said gas under
pressure is cooled sufficiently to form a liquid portion under
10 pressure and a vapor portion under pressure, and
; (i) said vapor portion under pressure is expanded
to a lower pressure, whereby it partially condenses,
.. (ii) said liquid portion under pressure is expanded
.: to said lower pressure, whereby a part of said liquid portion
vaporizes to cool the expanded liquid portion; and
(iii) at least the liquid formed upon partial
condensation of said expanded vapor, and the remaining liquid
portion.of said expanded liquid stream are supplied to a
; fractionation column wherein said relatively less volatile
fraction is separated,
the improvement wherein
(a) at least some of said liquid portion under
pressure is subcooled to a temperature below its bubble point :
prior to expansion thereof,
(b) at least a part of said subcooled liquid portion
: is expanded to said lower pressure, whereby a portion of the
expanded subcooled stream is partially ~aporized to further .
cool said expanded subcooled stream, and
(c) at least a portion of the liquid remaininq
in the expanded subcooled stre~m is supplied to said
distillation column a5 a top liquid feed thereto.
~ - 5(a) ~
QQ3
In a further aspect of this invention there is
pro~ided in a process for separating a feed gas ~nto a
volatile residue gas and a relatively less volatile fraction,
said feed gas containing hydrocarbons, methane and ethane
together comprising the major portion of said feed gas, wherein
. said gas under pressure is cooled sufficiently to form a liquid
; portion under pressure and a vapor portion under pressure, and
. (i) said vapor portion under pressure is expanded
to a lower pressure, whereby it partially condenses;
10 (ii~ said liquid portion under pressure is expanded
to said lower pressure, whereby a part of said liquid portion
vaporizes to cool the expanded liguid portion; and
(iii) at least the liquid formed upon partial
condensation after expansion of the vapor portion in step (i)
and the liquid remaining after expansion of the liquid portion
in step (ii) are supplied to a fractionation column, wherein
- said relatively less volatile fraction is separated,
the improvement wherein
(a) at least some of said liquid portion under
pressure is subcooled to a temperature below its bubble point
prior to expansion thereof,
(b) at least a part of said subcooled liquid portion
is expanded to said lower pressure, whereby it is partially
vaporized to further cool said expanded su~cooled liquid portion,
(c) at least a part of the liquid remaining in the
expanded subcooled liquid portion is supplled to said distilla-
tion column at a first feed position, and
: (d) at least part of the stream resulting from
expansion of said vapor in step (i) is supplied to said
fractionation column at a second feed position, said second
~ - 5(b) ~
,~
~ 1~4~0Q3
feed position being in a lower column position than said
first feed position.
In a still further aspect of this invention there is
provided in a process for separation of a feed gas into a
volatile residue gas and a relatively less volatile fraction,
said feed gas containing hydrocarbons, methane and ethane
together comprising the major portion of said feed gas, wherein
: said feed gas is cooled sufficiently under pressure to condense
it to a liquid, and
(l) at least some of said liquid under pressure is
expanded to a lower pressure and resulting in an expanded ::
stream, whereby part of the liquid vaporizes to cool the
expanded stream, and
(2) at least some of the liquid remaining in said
expanded stream is supplied to a fractionation column wherein
said relatively less volatile fraction is separated,
the improvement wherein said liquid under pressure is
subcooled to a temperature below its bubble point prior to -
expansion thereof by extracting heat therefrom, and
(a) at least a portion of said subcooled liquid
under pressure is expanded to said lower pressure and at
least a part of the expanded liquid is supplied to said frac-
tionation column as the top column liquid feed, and
(b) the heat extracted from said liquid under
pressure is supplied to said column at a mid column position.
In a still further aspect of this invention there is
provided in an apparatus for separating a feed gas into a
volatile residue gas and a relatively less volatile fraction,
- said feed ga~ containing hydrocarbons, methane and ethane
together compri~ing the ma~or portion of said feed gas, said
apparatus including
- 51c) ~
~ '
0~3
(i) a first cooling means to receive said feed gas
under pressure and to cool it sufficiently to form a liquid
portion and a vapor portion,
.. (ii) a separation means connected to said first
- cooling means to separate said liquid portion under pressure
and said vapor portion under pressure,
(iii) first expansion means connected to the
separation means to receive said vapor portion under pressure
:: and expand it to a lower pressure, thereby partially condensing
said expanded vapor stream,
(iv) a second expansion means connected to receive
said liquid portion under pressure and to expand said liquid
portion to said lower pressure, thereby to vaporize a portion
of said liquid and to cool the expanded liquid portion, and
(v) a fractionation means connected to said first
and second expansion means to receive at least the liquid
formed from partial condensation of said expanded vapor and the
liquid remaining from expansion of said liquid stream, to
separate said relatively less volatile fraction,
- 20 the improvement comprising
(a) a subcooling means connected intermediate
; said separation means and said second expansion means to
cool said liquid portion under pressure to a temperature below
its bubble point prior to expansion thereof, said subcooling
means being connected to supply at least a portion of said
subcooled liquid to said second expansion means, and
(b) said second expansion means is connected to
supply at least part of the liquid remaining in the expanded
subcooled liquid portion to said distillation column as a
; 30 top liquid feed thereto.
~ - 5(d) ~
' ,I;~
,,~ .
In a still further aspect of this invention there is
provided in an apparatus for geparating a feed gas into a
volatile residue gas and a relatively less volatile fraction~
said feed gas containing hydrocarbons, methane and ethane
together comprising the ma~or portion of said feed gas, said
apparatus including
. (i) a first cooling means to receive said feed
gas under pressure and to cool it sufficiently to form a
liquid portion and a vapor portion,
tii) separation means connected to said cooling
means to separate said liquid portion under pressure and
said vapor portion under pressure, :~
(iii) a first expansion means connected to said
separation means to receive said vapor portion under pressure
and expand it to a lower pressure, thereby partially condensinq
said expanded vapor portion,
(iv) a second expansion means connected to receive
said liquid portion under pressure and to expand said liquid
portion to said lower pressure, thereby to vaporize a portion
of said liquid and to cool the expanded liquid portion, and
(v) a fractionation means connected to said first
and second expansion means to receive at least the liquid
formed from partial condensation of said expanded vapor and
the liquid remaining from expansion of said liquid stream to ;-
: separate said relatively less volatile fraction,
the improvement comprising
(a) a subcooling means connected intermediate
said separation means and said second expansion means to
. cool said liquid portion under pre~sure to a temperature below
30 its bubble point prior to expansion thereof, said second cooling
- 5(e) ~
, . ' ' ' '
' ' , : '
means being connected to supply at least a portion of said
subcooled liquid to said second expansion means,
(b~ said second expansion means being connected
to supply at least a part of the liquid remaining in the
expanded subcooled liquid portiDn to said distillation column
at a first column feed position, and
(c) said first expansion means being connected to
provide at least a part of the expanded vapor stream to
: said fractionation column at a second feed position, said
second feed position being in a lower column position than said
first feed position.
In a still further aspect of this invention there is
provided in an apparatus for separating 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, said
apparatus including
(i) a first cooling means to receive said feed gas
under pressure and to cool it sufficiently to condense it to a
liquid;
; (ii) an expansion means connected to said cooling means
to receive at least a portion of said condensed feed gas and to
expand it to a lower pressure, whereby a portion of the expanded
stream is vaporized to further cool the expanded stream; and
(iii) fractionation means connected to said expansion
means to receive at least the liquid remaining in said expanded
~ stream to separate said relatively less volatile fraction,
the improvement comprising
(a) subcooling means connected intermediate said
first cooling means and said expansion means to receive at
~ - 5(f) ~
1~}~ 3
; least a portion of said condensed gas and to subcool it to
a te~.perature below its bubble point prior t~ expansion thereof
by extracting heat therefrom,
(b) means to supply the heat extracted in said
subcooling means to said fractionation column at a mid column
position, and
(c) said expansion means (ii) being connected to
supply at least a portion of the expanded stream to said
fractionation column as the top feed thereto.
For a better understand~ng of the present lnventlon
reference ls made to the following examples and drawings.
Referring to the drawings:
Figure 1 is a flow diagram of a single-stage cryogenic
expander natural gas processing plant of the prior srt lncor-
porating a set of conditions for a typical rich natural gas
stream;
Figure 2 ls a flow dlagram of a single-stage cryogenic
expander natural gas processing plant of the prlor art incor-
porating a set of conditions for a typical lean natural gas
. 20 stream;
Figure 3 is a flow diagram showing one embodlment of
the present lnventlon whereln the liqulds from the hlgh-pressure
separator are sub-cooled and then combined with the expander
outlet stream to pass to the demethanizer;
Figure 4 is a flow diagram showlng an embodlment of
the pre~ent inYention wherein the lnlet stream is totally con-
densed snd sub-cooled before psqsing to the demethanlzer;
- Flgure 5 ls a flow dlagram showing an embodlment
of the present inventlon whereln the llqulds from ehe hlgh-
pressure separator are sub-cooled and fed to the demethanizer
bo~e the expander outlet stre-m;
- 5(g)
Figure 6 is a flow diagram showing an embodiment
; of the present invention wherein the sub-cooled liquids and
expander outlet stream enter as feeds to the top of the column;
Figure 7 is a flow diagram showing an embodlment of
; 5 the present invention wherein the sub-cooled stream is fed
below the expander stream lnlet;
Figure 8 is a flow diagram showing another embodi-
ment of the present invention wherein a demethanizer side
stream is used to sub-cool the liquids from the high-pressure
separator;
Figure 9 is a flow diagram showing still another
embodiment of the present invention wherein the liquids from
the high-pressure separator are sub-cooled by heat exchange
with vapor product stream;
Figure 10 is a flow diagram showing an embodiment
of the present invention wherein the expander outlet is used
to sub-cool the liquids from the high-pressure separator;
Figure 11 is a flow diagram showing another embodi-
ment of the invention;
Figures 12 and 13 represent further embodiments of the
present invention wherein two parallel expansion steps are em-
ployed; and
Figure 14 is an embodiment of the present invention
wherein a portion of the uncondensed high pressure vapor stream
is condensed by column overhead vapor.
Figures 15A and 15B are graphs OL carbon dioxide
vs. temperature from one embodiment of this invention com-
pared to the prior art.
Q~;3
In the following explanation of the above figures,
tables are provided summarizing flow rates, calculated for
representatlve processing conditlons. In the tables, 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 lnclude all non-
; hydrocarbon components, and are generally larger than the
sum of the stream flow rates for hydrocarbon components. Tem-
peratures indicated are approximate values rounded to the
. 10 nearest degree.
Referring to Figure 1, plant inlet gas from whlch
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 specifications,
or cause icing in the equipment), and which has been dehy-
drated, enters the process at 120F. and 910 psia at stream23. It 1B then divided into two parallel streams and i9
cooled to 45F. by heat exchange with cool residue gas at
5F. in exchanger 10; with product liquids (stream 26) at
20 82F. in exchanger 11; and with demethanizer liquid at 53F.
in demethanizer reboiler 12. From these exchangers, the
streams recombine and enter the gas chil]er, exchanger 13,
where the combined stream is cooled to 10F. with propane
refrigerant at 5F. The cooled stream is again dlvided into
; 25 t~o parallel streams, and further chilled by heat exchange
with cold residue gas (stream 29) at -107DF. in exchanger 14,
- and with demethanizer liquids at -80F. in demethanizer side
reboiler 15. The streams recombine (stream 23a) and enter a
: high-pressure separator 16 at -45F. and 900 psia. The
3~ ctDdented llquld, s~reat 24, ls separated and fed to tùe
-7-
;. ;:
,
demethanizer 19 through expansion valve 30. An expansion engine
may be used in place of the expansion valve 30 if desired.
The cooled gas from the high-pressure separator 16
flows through expander 17 where lt ls work expanded from 900
psia to 290 psia. The work expansion chills the gas to -125F.
Expander 17 is preferably a turbo-expander~ having a compressor
21 mounted on the expander shaft. For convenlence, expander 17
is sometlmes herelnafter referred to as the expanslon means,
In certaln prlor art embodiments, expander 17 ls replaced by
a conventional expansion valve.
Liquid condensed during expanslon ls separated in
low pressure separator 18. The llquld ls fed on level control
through line 25 to the demethanlzer column 19 at the top and
flows from a chimney tray (not shown) as top feed to the column
- 15 19.
- It should be noted that in certain embodiments low
pressure se~arator l8 may be included as part of demethanizer
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 liquid then flows from the chimnev tray as top
feed to the demethanizing section of the column.
As liquid fed to demethanizer 19 flows down the
column, it ls contacted by vapors whlch strip the methane
from the liquid to produce a demethanized liquid product at
the bottom. The heat required to generate stripping vapors
is provided by heat exchangers 12 and 15.
The vapors stripped from the condensed liquid in
demethanizer 19 exit through line 27 to join the cold outlet
gas from separator 18 via llne 28. The comblned vapor stream
''
then flows through line 29 back through heat exchangers 14 and
10. Following these exchangers, the gas flows through compres-
sor 21 driven by expander 17 and directly coupled thereto. Com-
pressor 21 compresses the gas to a discharge pressure of about
305 psia. The gas then enters compressor 22 and is compressed
to a final discharge pressure of son psia.
Inlet and liquid component flow rates, outlet liquid
recoveries and compression requirements for this prior art
process shown in Figure 1 are given in the following table:
TABLE I
Stream Flow Rate Summary - Lb. Moles/Hr.
STREAM METHANE ETHANE PROPANE BUTANES+ TOTAL
23 1100 222 163 130 1647
24 795 202 157 129 1300
16 10 5 1 32
26 3 162 157 130 453
RECOVERIES
Ethane72.9% 29,296 GAL/DAY
Propane96.2% 39,270 GAL/DAY
COMPRESSION HORSEPOWER
Refrigeration 256 BHP
Recompresslon 892 BHP
Total 1148 BHP
In Figure 2 a typical lean natural gas stream 33 ls
processed and cooled uslng a prlor art process simllar to that
~hown in Flgure 1. The lnlet gas stream is cooled to -67F.
at 900 psia t33a) and flows to high-pressure separator 16
where the liquid contained thereln ls separated and fed on level
control through llne 34 and expanslon valve 30 to demethanlzer
: 30 19 ln the mlddle of the col~n.
. , , .:
,
. - :
i -' i
Cold gas from separator 16 flows through expander 17
where because o~ work expansion from sno psia to 250 psia, the
gas is chilled to -153F. The ]iquid condensed during expan-
sion is separated in low pressure separator 18 and is fed on
level control through line 35 to the demethani~er 19 as top
feed to the column.
The data for this case are given in the following
~ table:
,:
TABLE II
(Fig. 2)
Stream Flow Rate Summary - Lb. Moles/Hr.
STREAM METHANE ETHANE PROPANE BUTANES+ TOTAL
33 1447 90 36 43 1647
34 280 42 25 39 391
15 35 133 35 11 4 186
36 2 71 36 43 155
RECOVERIES
Ethane 79.0X17,355 GAL/DAY
Propane 98.2X, 8,935 GAL/DAY
COMPRESSION HORSEPOWER
Refrigeration O BHP
Recompression 1180 BHP
Total 1180 BHP
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 rich gas feed and 79~ for the lean gas feed.
It is recognized that some improvement in yield may result by
adding one or more cooling steps followed by one or more
separation steps, or by altering the temperature of separator
:''
-- ] O--
:
,
. .
16 or the pressure in separator 1.8. Recoveries of ethane and
propane obtained in this manner while possibly improved over
the cases illustrated bv Figure l and Figure 2 are ~signifi- -
cantly less than vields which can be obtained in accordance
with the process of the present invention.
In accordance with the emhodiments o- the irvention
to be described in Examples l to 5, the hYdrccara-~ gas, under
pressure, is cooled sufficiently to form a l.iquid port_on, and
; the liquid portion is expanded to a lower pressure as in the
conventional process. Expansion of the first part n, the
liquid portion vaporizes a portion of it and cools the remain-
ing part which remains as a liquid. Thi.~ expande~ .s ream
usually is supplied to a fractionation col(lmn where it i.s
separated into a top fraction and a hottom fraction. In the
present invention. the foregoing proce.c~s i.s improved by divid-
- ing the remaining part of the liquid portion into fir.st and
second liguid streams. The first ]iquid .stream is diverted
in heat exchange relation with the liquid portion of the feed
stream prior to expansion to pre-cool or sub-cool the liquid
portion prior to expansion. The pre-cooling (or .svnonymou~sly
sub-cooling) of the liquid portion condensed from the feed
gas under pressure prior to expansion reduces the temperature
attained by the aforementioned second liquid stream after ex-
pansion.
The fir.st and .second liquid streams are then sup-
plied to the fractionating column. The second ~i~uid .stream
is supplied to the column at a point higher on the column
than the first liquid stream.
.,
:
Example l
One embodiment of the process of the present inven-
tion is shown ln Figure 3. Prior art processes are used tn
remove sulfur containing compounds, carbon dioxide and to de-
hydrate and cool the inlet gas 23 to -45F., generally as
described in Figure 1, by heat exchangers lO, 11, 12, 13, 14
- and 15. As in Figure 1, the process conditions stated ln
Figure 3, as well as flow rates in Table III below, are for
the case of a rlch feed gas. The cooled and partlally con-
densed gas 23a at -45F. and 900 psia flows to hlgh-pressure
separator 16 where condensed liquid therein is separated.
The cooled gas component of the lnlet stream flows
from high-pressure separator 16 through expander 17 where,
because of work expansion from 900 psia to 290 psia, the gas
is chilled to -125F. As ln the prior art process shown in
Figure 1, expander 17 may have a compressor 21 mounted on the
expander shaft. The expander outlet stream 44 is then com-
bined with the cold stream 43, from valve 42 as it flows to
low pressure separator 50.
' 20 The cooled llquid from high-pressure separator 16,
; stream 24, flows through exchanger 41 where it is sub-cooled
to -132F. by heat exchange with a portion of the cold liquids
from low pressure separator 50, as described below. The sub-
:.
; cooled liquids then undergo expansion and flash vaporization
~ 25 at valve 42 as the pressure is reduced to 290 psia. The cold
stream, 43, from valve 42, then combines with expander outlet
. stream 44, as explained above.
. ' - .
-- ~ .
.
-
A first part of the condensed liquid from separator
50 flows as stream 45a to the top of demethanizer l9 as top
feed to the column. The second part, stream 45b, of liquid
from separator 50 flows through exchanger 41 where it is used
to sub-cool the liquids from hlgh pressure separator 16. From
exchanger 41, the stream flows to demethanizer 19 as feed in
; the middle of the column.
- The vapor stripped from the conden~ed liquid in de-
methani~er 19 leaves through llne 46 to ~oin the cold outlet
vapor 47 from separator 18, and the combined stream flows
~ through llne 48 through the balance of the system.
; Component flow rates, llquid recoverles and compres-
- sion requlrements for this embodlment are glven ln the follow-
lng table:
15TABLE III
tFig. 3)
Stream Flow Rate Summary - Lb. Moles/Hr.
- STREAM METHANE RTHANE PROPANE BUTANES+ TOTAL
23 1100 222 163 130 1647
` 20 24 795 202 157 129 1300
~5 711 201 157 129 1209
49 5 204 161 130 501
RECOVERIES
Ethane 92.1~ 49,647 GAL/DAY
Propane 98.8% 40,333 GAL/DAY
COMPRESSION HORSEPOWER
.. . ..
- Refrlgeratlon 384 BHP
Recompresslon 840 BHP
Total 1224 BHP
-
Example 2
A second embodiment of the process of the present
invention is shown in Figure 4. In this embodiment, inlet
gas 23 is processed and cooled through heat exchangers 10, 11,
12, 13 and 14 and reboiler 15. As in Flgure 1, process condl-
tions and flow rates in the table below are for a rich feed
gas. However, in contrast to Figure l, the feed gas in this
case is cooled to -55F., at which temperature the entire inlet
- gas stream is condensed. The condensed liquid~s then enter ex-
changer 51 wherein they are further cooled to -140F. by
heat exchange with a portion of the cold liquid stream from
low pressure separator 52. From exchanger 5], the cooled in-
- let stream undergoes expansion and flash vaporization through
expansion valve 53. From valve 53, the cold inlet stream en-
ters low pressure separator 52 where the vapor and liquid por-
tions therein are separated.
A first part of the cold liquids 54 from low pres-
sure separator 53, stream 55, enters the demethanizer as top
feed to the column. The second part 56 of the liquid from
separator 52 flows to exchanger 51 where it is used to cool
inlet gas from -55F. to -140F. From exchanger 51, the
stream 56a is fed to demethanizer 19 in the middle of the column.
The data for this case are given in the following
table:
_ABLE IV
(Fig. 4)
; Stream Flow Rate Summarv - Lb. Moles/Hr.
STREAM METHANE ETHANE PROPANE BUTANES+ TOTAL
23 1100 222 163 130 1647
54 1109 221 163 130 1544
57 5 210 162 130 509
.
10~.~0~
; RECOVERIES
Ethane 94.8~ 51,065 GAL/DAY
Propane 99.1~ 40,455 GAL/DAY
CO~PRESSION HORSEPOl~ER
Refrigeration 457 BHP
Recompression 87_ BHP
Total 1328 BHP
Example 3
A third embodiment of the process of the present in-
vention is shown in Figure 5. The prior art process shown in
.~ Figure 2 is used to cool the lnlet gas in the line following
exchanger 14 and reboiler 15 to -67F. at 900 psia, 33a. As
ln Flgure 2, the process conditions given. and flow rates set
forth below in Table V are for a lean feed gas. The gas at -67F.
flows to hlgh-pressure separator 16 where condensed liquld
thereln is separated.
The cooled vapor from separator 16 flows through
expander 17 where because of work expansion from 900 psia to
- 250 psla, it ls chilled to -153F. From expander 17 the chilled
vapor stream flows to demethanizer 19 as its middle feed.
The liquid 34 from separator 16 flows through ex-
- changer 61 where lt ls sub-cooled to -150F. b~ heat exchange
wlth a first part 62 of the cold liquids 63 from low pressure
separator 64. The sub-cooled liquid then undergoes expansion
and flash vaporization at valve 65 as pressure is reduced to
250 psla. The cold stream from expanslon valve 65 flows to
separator 64 where the cold llquid and vapor are separated.
: As mentloned above, the flrst part 62 of the liquld
from separator 64 ls used to sub-cool condensed llquld 34
'
.~,
-15-
"
.06~
from the high pressure separator 16. Stream 62 then flows
to demethanizer 19 as its lowest feed. The second part 66
of the liquid from separator 64 is supplied to demethanizer
19 as the top feed.
De~ethanizer 19 shows, at 67, an area which repre-
sents trays or packing equivalent to at least one distilla-
tion stage. In this embodiment, an interval of packing or
of trays is provided sufficient to insure that ethane and
higher hydrocarbons contained in the mixture of vapor and
liquid in the feed 68 from expander 17, mixes in the column
with top feed liquids rich in heavier hydrocarbons during
their passage through area 67 of -the demethanizer, and that
the mixing take place under conditions which aid maximum re-
covery of ethane and the higher hydrocarbons. These favor-
able conditions include a top feed that is rich in higher
hydrocarbons, as in stream 66, and column design considera-
tions which prnvide that warmer streams to the column, as
.. .
stream 62a,are spaced sufficiently below the top feed that
in operation, the vapor temperature of the column in the area
adjacent to the top feed will closely approach the temperature
of the top feed stream.
The data for this case are given in the following
table:
TABI.E V
(Fig. 5)
Stream Flow Rate Summary - Lb. Moles/Hr.
STREAM METHANE ETHANE PROPANE BVTANE+ TOTAL
33 1447 90 36 43 1647
61 280 42 25 39 391
63 251 42 25 39 361
69 2 79 36 43 164
-
1~ LV~3
RECOVERIES
Ethane 87.6% 19,240 GAL/DAY
Propane 97.6% 8,883 GAL/DAY
COMPRESSION HORSEPOWER
Refrigeration 0 BHP
Recompression 1181 BHP
Total 1181 BHP
Example 4
A fourth embodiment of the process of this invention
is shown in Figure 6. As in Figure 2, a feed gas 33 is
partially condensed at -67F.and 900 psia ln heat exchangers
10, 11, 12, 14 and 15, and supplied to low pressure separator
16. The process conditions given, and the flow rates in
Table VI below are for a lean feed gas. Vapors from separator
16 are work expanded in expander 17, and supplied to low pres-
sure separator 18 at -153F. and 250 psia, where the liquid
condensed during expansion is separated, and liquid stream 35
is separated.
Liquid stream 34 from separator 16 is sub-cooled in
heat exchanger 71 to -150F., and then expanded through valve
72 to a pressure of 250 psia. A portion of the liquid vaporizes,
thus cooling the remaining part to -158F. Expanded stream
enters separator 73, wherein liquid and vapor are separated.
The cold liquid, stream 74, from separator 73 is combined
with liquid stream 35 from separator 18 to form a combined
stream 75. A first portion 76 of the combined stream is
used to sub-cool condensed liquid 34 in heat exchanger 71
and is then supplied as a feed to demethanizer 19 at a mid-
- column location. The second part 77 of stream 75 ls supplied
to demethanlzer 19 as the ~op feed at -157F.
The data for this case are given in the following
table:
TABLE VI
(Fig. 6)
Stream Flow Rate Summary - Lb. Moles/H_.
STREAM METHANE ETHANE PROPANE BUTANES+ TOTAL
331447 90 36 43 1647
34280 4225 39 391
35133 3511 4 1~6
i~ 10 74251 4225 39 361
782 7436 43 159
RECOVERIES
Ethane 81.6% 17,925 GAL/DAY
Propane 98.0% 8,924 GAI./DAY
COMPRESSION HORSEPOWER
Refrigeration O BHP
Recompression 1182 BHP
'~ Total 1182 BHP
.' :
Example 5
Another embodiment of the invention is shown in
Figure 7. In this case, a cooled feed stream 33a is supplied
as in Figure 5 to high pressure separator 16. Cooled stream
33a is a partially condensed lean feed gas, cooled by means
of a heat exchanger chain as shown in Figure 2.
Cold gas from separator 16 flows to expander 17 where
it is expanded and provides a cold outlet stream at -153F.
to low pressure separator 18. Cold liquid outlet stream 35
from separator 18 is fed on level control to demethanizer 19
as top feed. Cold vapor from separator 18 joins vapors
-18-
stripped from demethanizer 19 and flows to provide heat ex-
change used in cooling feed stream 33a, as in Figure 2 and
then to residue gas compression.
Condensed liquid stream 34 from high pressure separa-
tor 16 ls sub-cooled in heat exchanger 80 and expanded through
an expansion valve 81 to low pressure separator 82. A part of
the stream vaporizes on expansion, thus cooling the remaining
part to -158F. Cold vapor from separator 82 is fed to the
demethanizer 19 at an intermediate level. The cold outlet
liquid stream 79 from separator 82 is divided. A first part,
stream 83, is used to sub-cool stream 34 in heat exchanger 80.
The second part 84 of the liquid stream from separator 82 is
fed to demethanizer column 19 at an intermediate point in the
. column. Alternately, stream 84 may be supplied to the column
at a point just below the top feed, 35, or may be mixed with
the exit stream from expander 17, as shown by the broken lines
85 and 86, respectively.
~Component flow rates, liquid recoveries and composi-
#~ .
tion requirements are given in the following table:
TABLE Vll
(Fig. 7)
Stream Flow Rate Summary - Lb. Moles/Hr.
STREAM METHANE ETHANE PROPANE BUTANES+ TOTAL
33a1447 90 36 43 1647
~5 34 280 42 25 39 391
35 133 35 11 4 186
' 79 251 42 25 39 361
87 2 72 36 43 157
881445 18 0 0 1490
-19-
0(~!3
RECOVERIES
Ethane 79.4~ 17,~40 GALtDAY
Propane 98.2~ 8.935 G.~L/DAY
CO~IPRESSIO~ HORSEPO~ER
Refrigeration O BHP
Recompression 1186 BHP
- Total 1186 BHP
. . .
To sumnarize the foregoing discussion of the first
: five embodiments of our invention, the process feed is par-
10 tially or completelv condensed under pressure by cooling using
product as well as available column side streams and (if neces-
;. sary) external refrigeration. Where the feed gas under pres-
, sure is onlv partially condensed, the remaining vapors are
;~ expanded to provide a cooled and partially condensed expanded
" vapor. The liquid portion obtained bv cooling and refrigera-
tion of the feed gas under pressure is expanded (for instance
by flashing, or by a ~ork engine) whereby a portion of it
vaporizes and the rem2ining part is cooled and used as liquid
feed to a fractiona~or, such as a demethanizer.
Prior to expansion, the liquid portion of the feed
is sub-cooled by bringing it into heat exchange relation with
a portion of the expanded cold ]iquid. This results in t~o
liquid feeds derived from expansion of the liquid portion of
the condensed feed gas, one feed being substantiallv colder
25 than the other.
Advantages of improved recovery may be realized by
utilizing the divided strean in various of process configura-
tions:
-20-
''
0~1l3
(i) By feeding both of the thus-derived liquid
streams directly to the demethanizer tower,
the cooler stream being used as a feed at
a higher point in the column than the hotter
feed. In such a configuration, the colder
liquid stream may be used as all or a portion
of top column feed.
(ii) By combining all or a portion of the expanded
, liquid stream with all or a portion of the work-
expanded vapor stream to form a combined con-
densate, using a portion of the combined con-
densate to sub-cool the liquid portion of the
feed gas and using the remaining condensate as
column top feed.
:
(iii) By using the cold expanded liquid derived
according to the present invention as column
top feed, and feeding liquid or va~or (or both)
from the expanded vapor stream at a column
point below the top, whereby the cold liquid
at the column top will recover absorbable
ethane from the expanded vapors.
In connection with the foregoing, it should be noted
that for clarity in explanation, vapor-liquid separation of
the expanded liquid and vapor streams has been shown external
to the demethanizer. It will be obvious to those of ordinary
skill in the art that such vapor-liquid separation may equally
be accomplished internally of the demethanizer column. Simi-
larly by appropriate selection and control of side stream
eeds and draw-offs, the stream obtained from flash expansior
of the liquid portion of the feed gas can be fed directly
-21-
to the column and internally divided to provide the desired
first portion thereof utilized for sub-cooling the liquid por-
tion of the feed gas. Where the flash expanded liquid is fed
directly to the column at an intermediate column stage, liquid
drawn off from that stage as the source of sub-cooling liquid
will usually contain not only a portion of the liquid from the
- feed but also liquid flowing from the column from higher stages
thereof.
It should also be noted thst as illustrated in the
foregoing examples, the entire condensed liquid stream from
separator 16 is sub-cooled. In some cases, it may prove ad-
vantageous ta treat only a portion of the liquid from separator
in accordance with this invention.
In another embodiment of the present invention sub-
cooling of the condensed process feed under pressure prior
to expansion is accomplished by heat exchange with a liquid
stream available from the demethanizer column. This is illus-
trated in Figure 8 of the present invention.
, . . .
Example 6
Figure 8 is a fragmentary process flow diagram of
one aspect of this embodiment of our invention, and illustrates
a specific case calculated on the assumption of total conden-
sation of feed gas entering the process at a pressure of about
900 psia. Such a total condensation procedure is illustrated,
for example, in Figure 4 when incoming gas at a temperature of
120F. and 910 psia is condensed by heat exchange against
residue gas products, demethanizer column side reboilers, de-
methanizer column reboiler and demethanizer column bottoms
product. For purposes of to~ll condensation, it is usually
-22-
{~
required also to provide supplemental external refrigeration
as illustrated in Fi~ure 4.
Referring now to Figure 8, the totally condensed
liquid feed 9O, at a pressure of approximately 900 psia and
a temperature of -55F. passes through heat exchanger 91 where
it is sub-cooled to a temperature of -14nF. The sub-cooled
liquid is then flash expanded through expansion valve 92,
and the expanded product enters separator 93. During flash-
~. ing, a part of the liquid stream vaporizes and cools the re-
: 10 maining liquid to a temperature of about -L46F. The remain-
ing liquid is separated in separator 93 and supplied as stream
94 to the demethanizer column l9 as top feed to the column.
Vapors flashed during the flash expansion step leav-
ing separator 93 as vapor stream 95 are combined with overhead
vapors 96 from demethanizer 19 to form a residue gas stream 101.
As in Figure 4, the residue gas seream ]eaving the demethanizer
column is returned in heat exchange relationship with incoming
feed gas to provide a portion of the cooling required to liquify
the feed gas. Thereafter the residue gas is compressed to
approximately 900 psia and discharged from the process.
The desired liquid product is contained in the de-
methanizer bottoms 100. Before this product leaves the process,
it is heat exchanged with incoming feed to provide inlet gas
cooling as illustrated, for example, in Figure 4.
- 25 To provide sub-cooling of liquid stream 90 in accor-
dance with this embodiment of the present invention, a side
- stream 97 is withdra~n from the demethanizer column 19 and
passed through exchanger 91. The warmed side stream 98 is
then returned to the demethanizer column at a point below the
liquid inlet 94. For the purpose of the embodiment illustrated
in Figure 8, and in the table given below, it was assumed that
demethanizer 19 contained column packing material equivalent
- to one theoretical distillation stage between the side stream
return 98, and the top liquid feed 94.
Inlet and liquid component flow rates, outlet liquid
recovery efficiencies and compression requirements for this
illustration are set forth in the following table:
TABIE VIII
(Fig. 8)
Stream Flow Rate SummarY - Lb. Moles/Hr.
STREAM METHANE ETHANE PROPANE BUTANES+ TOTAL
90 1100 222 163 130 1647
94 1009 221 163 130 1544
100 5 210 162 130 508
101 1095 12 1 0 1139
RECOVERIES
Ethane 94.5X 50,917 GAL/DAY
Propane 99.1% 40,448 GAL/DAY
., COMPRESSION HORSEPOWER
Refrigeration 461 BHP
, Recompression 870 BHP
Total 1331 BHP
It will be recognized that the use of side stream 97
to provide refrigeration to exchanger 91 will result in a side
stream return 98, which is partially vaporized. The column,
therefore, should provide for vapor-liquid contacting means
between liquid inlet 94 and side stream return 98 so that the
warmed vapors rising from side stream return 98 will be cooled
before appearing in column overhead vapors 96.
1~P4~
The fractionation means promotes vapor-liquid con-
tact in this column region to facilitate heat exchange between
the rising vapors and descending liquid. It will be evident
to those skilled in the art that the amount of vapor-liquid
contact thus provided may vary and may be provided by one or
more bubble plates, sieve trays, etc., or by a greater or
lesser amount of packing material.
In preferred embodiments of this invention, the
demethanizer column 19 should provide for sufficient exchange
between liquid inlet 94 and side stream return 98 that the
; vapors rising past the liquid inlet point 94 will have a tem-
perature which did not exceed by more than about 10F.~ the
temperature of the incoming liquid 94.
While the present invention has been described with
- 15 particular reference to an embodiment in which the side stream
withdrawal 97 and return 98 occur at the same point of the de-
methanizer 19, it is not necessary that the side stream return
- 98 correspond to the side stream withdrawal point 97. It
. I ~
may, for example, be advantageous from the standpoint of
`, 20 column efficiency and heat balance to return side stream 98
at a point below where side stream 97 is withdrawn.
- For purposes of illustration, it will also be recog-
nized that while the heat exchanger 9l has been illustrated
as a heat exchanger external to the demethanizer column sup-
plied by side stream withdrawal 97, a fully equivalent result
may be obtained by providing for an internal heat exchanger
within demethanizer 19 in lieu of the external heat exchanger
- 91. In such a case, the internal heat exchanger would be
located so as to correspond to the side stream return point
98.
-25-
,, ~ . : , :............................... -
,'''. ' ' ' ' ' . ' ''',, .
1~4~)Q3
It will also he evident from illustrations o~f other
embodiments of the present invention, such as in Fi~ures 3,
5, 6 and 7, that the Eeed stream 90 need not be chilled to
tlle point of total condensation, nor is it necessary to sub-
cool the entire liquid stream sn. If the feed stream 90 ispartially condensed, as for example in Figure 3, provision
will be made for separation of the partially conden.secl pro-
cess feed. The liquid recnvered from partial condensation
of the feed wil~ be furtller treatcd as illustratc(l in Figure
8. The vapor recovered ma~ ~e ~ork e~pand~d, sucll as ~v a
turbo-expander to produce an expancled and partial]v condensed
vapor stream, the partial condensate that is recovered being
supplied to the demetllanizer column. ,~s is a~parent from
Figures 5, 6 and 7, ~furt~ler variations arc ~os.sible. IDr
example, if the initial feed is onlv partiallv condcnsed, and
a work-expanded vapor is therefore available, all or a ~-art
of ~ork-expanded vapor stream mav, if desircd, be surplied to
the demethanizer column as an intermediate feed, and the sub-
cooled liquid 94 used as a demethanizcr top fæed ~s illustrated
in Figure 8.
Still another embodiment of the present invention
has particular reference to gas separation processes in which
the feed gas under pre.ssure is partiallv condensed to produce
a liquid portion and a vapor portion. The liquid portion is
sub-cooled and expanded to a lo-~er pressure to produce thereby
a cold liquid feed supplied to the fractionation column. The
vapor portion is expanded to the lower pressure resulting in
cooling and partial condensation of the vapor portion. The
refrigeration produced by the expansion of the vapor portion
is emploved to sub-cool the liquified portion of the feed gas
under pressure prior to expansion.
-26-
:
Q3
This embodiment is more spec~fically illustrated in
Figures 9 and 10. Both Figures 9 and 10 represent only
partial flow diagram of an overall gas separation plant. As
indicated on both drawings, the Eragmentary portion illustrated
- 5 is supplied with cooled feed gas. Such cooled feed gas isderived in a conventional manner as shown by the heat exchange
system Figures 1 and 2 involving heat exchangers 10, 11, 12,
13, 14 and 15. These heat exchangers recover refrigeration
values contained in the product and residue gas of the gas
separation plant and incorporate additional external refrigera-
tion to the extent necessary to cno~ the feed gas under pres-
sure to a condition entering the fr"gmentary portion of the
separation process illustrated.
The process conditions described in Figllres 9 and
, 15 10 correspond to the processing of a lean feed gas of the com-
position set forth above in Table ll. The process conditions
in Figures 9 and 10 may be compared with Eigure 2 to illustrate
the present invention. At the inlet conditions in each of
Figures 9 and 10, the cooled lean feed gas 33a i.s at a tem-
;~ 20 perature of -67~F. and a pressure of 900 psia.
:
Example 7
Following is the process of Figure 9, the partially
condensed feed gas 33a derived as described in ~igure 2, com-
prises partially condensed gases containing a liquid portion
and a vapor portion. The partially condensed gas enters a
high pressure separator 16 where liquid and vapors are separated.
Following first the vapors leaving separator 16, the vapors
enter a work expansion engine 17 in which mechanical energy is
extracted from the vapor portion of the high pressure feed.
As that vapor is expanded f.om a pressure of about 900 psia
to a pressure of about 250 psia, the work expansion cools the
-27-
,
.
., ' - ' ' ~ ~ .
:.
OQ3
expanded vapor 113 to a temperature of approximately -153F.
' Expanded and partially condensed vapor 113 is supplied as a
feed to demethanizer 19, wherein the vapors rise and a major
part of C2+ hydrocarbons are absorbed by descending liquid. ~e
methanizer overhead 117 at a temperature of -156F. combined with
vapors 116 from flash vaporization described below to form residue
gas stream 118. The combined cnld residue gas stream 118 then
passes through heat exchanger ll9. The warmed residue gas at
-125F. leaving heat excllanger ~19 then returns to the pre-
liminarv cooling stages as illustrated, for example, in Figure
2, wherein further refrigeration contained in the still cold
r~
residue gas is recovered, and the residue gas is compressed
~; in compressor 21 (see Figure 2) which iq driven bv work expan-
sion engine 17, and then further compressed to a line pressure
of 900 psia by supplementary compressor 22.
Turning to the liquid 34 recovered fro~ separator 16,
liquid 34 passes through heat exchanger 119 in heat exchange
relation with the cold residue gas 118. This results in a
pre-cooling of the liquid portion of the partially condensed
2~ high pressure feed gas. The sub-cooled liquid is then expanded
- through an appropriate expansion device, such as expansion
valve 120, to a pressure of approximately 250 psia. During
expansion a portion of the feed will vaporize, resulting in
cooling of the remaining liquid part. In the process as illus-
trated in Figure 9, the expanded stream leav;ng expansion valve
120 reaches a temperature of -153F. and enters a separator
121. The liquid portion is separated and supplied as stream
115 to the fractionation column 19 as top feed.
It may be noted that by comparison with Figure 2,
,~ 30 the demethani~er feed from eXpansion valve 30 of Figure 2 onlv
achieves a temperature of -134F. Because stream 115 of this
-28-
,
' , ' '
embodiment to the present invention is substantia;ly cooler,
it may be used as top feed to the demethanizer to recover
- ethane in the stream 113. The ethane recovered is withdrawn
in the demethanizer bottoms 125. Demethanizer bottoms 125
are heat exchanged with incoming feed to recover refrigeration
therein as generally illustrated in Figures l and 2.
In connection with Figure 9, it should be noted that
for purpnses of heat economy there will be one or more demetha-
nizer side-stream reboilers which exchange heat to cool incom-
ing feed (not shown in Figure 9) as illustrated generally inFigure 2. For purposes of the illustrated process, calcula-
tions appearing in Figure 9 and set forth in the table below,
two such side-stream reboilers have been included, as shown
in Figure 2. The side-stream reboilers are significant to
the overall heat economy of the process. Sub-cooling of the
; liquid stream 34 by residue gas 118 reduces the av~ilable
refrigeration remaining in stream 118 for feed cooling purposes.
However, the increased loading of demethanizer 19 with liquid
stream 115 cooled in accordance with the present invention
provides additional available refrigeration in the side-stream
reboilers. Accordingly the overall heat balance of the pro-
cess remains substantially unaffected.
Inlet and liquid component flow rates, outlet re-
covery efficiencies, and expansion/compression requirements
for the embodiment of this invention as illustrated in Figure
9 are set forth in the following table:
-29-
LQ~3
TABLE IX
(Fig. 9)
Stream Flow Rate Summary - Lb. Moles/Hr.
STREAM METHANE ETHANE PROPANE BUTANES+ TOTAL
5 33a 1447 90 36 43 1647
34 280 42 25 39 391
113 1167 48 11 4 1256
115 251 42 25 39 3~1
116 29 0 0 0 30
10 118 1445 10 1 ~ l483
125 2 8n 35 43 1~4
RECOVERIES
Ethane89.1%19,565 GAL/DAY
Propane97.7~8,894 GAL/DAY
COMPRESSION HORSEPOWER
Refrigeration O BHP
Recompression1177 BHP
Total1177 BHP
The stream flow rate summary set forth in Table IX
corresponds to processing a lean feed gas. For comparison
purposes, reference may be made to Figure 2 and Table II for
the processing of the same feed gas stream without the provi-
sion of pre-cooling of condensed high pressure feed gas liquids
such as in heat exchanger 119.
4 25 The materially improved recoveries of the present
invention result primarily because of the availability of a
substantially colder liquid feed obtained by sub-cooling and
expansion of the liquid condensed from the high pressure feed
gas. By the present invention, this expanded liquid is avail-
able at a temperature sufficiently cold to permit its use as
-30-
- the top feed to demethanizer 19. Because of the significant
propane and C4+ content of this very cold liquld stream, it has
enhanced capability for recovering ethane.
In the foregoing example, sub-cooling of liquid stream
34 with residue gas has been lllustrated using combined residue
gas streams. Other residue gas streams may equally well be used
if they are of sufficient volume relative to stream 34, such as
either of stream 116 or 117 shown in Figure 8, residue gas stream
47 shown in Figure 3, or the residue gas stream 38 shown in
..
Figure 7. As used herein, in any of the claims of this appli-
cation, the term "residue gas" is intended to encompass any oneof the streams or any combination thereof.
; Example ~3
Another illustration of this embodiment of the
present invention is set forth in Figure 10. Following the
process of Figure lO, cooled and partially condensed feed
gas 33a enters high pressure separator 16 at a temperature of
-67F. and`900 psia wherein it is separated into a liquid
portion and a vapor portion. As described, for example,
in Figure 2, this cooled feed gas is obtained by heat exchange
preferably with various process streams to recover the maximum
refrigeration values contained therein, wlth a provision for
addition of supplemental external refrigeration, if required.
Referring first to the vapors recovered from separa-
tor 16, these vapors are work expanded through turbo-expander
17 to a pressure of about 25n psia, and a temperature of -153F.
Ac these conditions the expanded vapor portion of the high pres-
sure feed is partially condensed. The entire expanded and par-
- tially condensed vapors, leavlng turbo-expander 17, then pass
through heat exchanger 13]~ ~herein they are heated to a ~em-
perature of about -137F. and supplied to the demethanizer 19
at a mid-point of the colu~n
-31-
.
1~4~U~
Returning to the liquids leaving separator 16, liquid
stream 34 passes through heat exchanger 131 in a heat exchange
relation with expanded vapor stream from turbo-expander 17.
This results in pre-cooling the liquid 34 from the separator
16 from a temperature of -67F. to a temperature of -148F.
Thereafter, the sub-cooled liquid 34 is expanded through ex-
pansion valve 133 to a temperature of -158F. and enters
separator 134.
In separator 134 vapor evolved as a consequence of
flash expansion is separated from the remaining liquid. The
remaining liquid 135 from the expansion step is supplied to
demethanizer 19 as top feed. Vapor 136 from separator 134 is
taken in combination with demethanizer overhead 137 to form a
combined vapor stream 139 which exits the process. As illus-
: 15 trated in Figure 2, the overhead vapors exiting the process
~ are used to cool and partiallv condense incoming feed gas
.' and are then compressed in compressor 21 driven by turbo-
expander 17 and in supplemental compressor 22 to a line pres-
sure of 900 psia.
Bottom 138 from demethanizer 19 containing the de-
. sired liquid product is also emplo~ed to cool incoming feed
gas and exits the process as a desired product.
Although not specifica]l~ illustrated in Figure lO,
it should be noted that as is customary in demethanizer design,
side-stream reboilers mav be provided as illustrated, for
example, in Figure 2 to control demethanizer operation and
at the same time recover additional refrigeration values use-
ful for pre-cooling the high pressure gas feed. The process
conditions and stream summar~ calculations set forth in Figure
10 and in Table X below are b~sed on the use of two reboilers
as shown in Figure 2.
:
1&4'~
Inlet and liquid component flow rates, outlet l~quld
recovery efficiencies and expansion/compression requlrements
for a process as illustrated in Figure 10, which lncluded de-
; methanizer slde reboilers are given in the following table:
TABLE X
Stream Flow Rates Summary - Lb. Moles/Hr.
STREAM METHANE ETHANE PROPANE BUTANES+ TOTAL
33a 1447 90 36 43 1647
34 280 42 25 39 391
10135 251 42 25 39 361
136 29 0 0 0 30
138 2 79 35 43 161
139 1445 11 1 0 1486
RECOVERIES
Ethane 87.3%19,159 GAL/DAY
Propane 97.6%8,880 GAL/DAY
COMPRESSION HORSEPOWER
; Refrigeration0 BHP
Recompression1180 BHP
Total1180 BHP
The embodiment illustrated in Figure 10 provides
. materially i~proved recovery of ethane values contained in
- the feed gas because of the availability of a cold top feed
: to demethanizer 19, stream 135, provided by pre-cooling stream
: 25 34 prior to flash expansion in accordance with the present
invention for use as column overhead liquid. In this embodi-
ment it will be noted that the entire liguid and vapor stream
leaving expander 17 enters the demethani~er 19 below the liquid
-33-
,' ' '
-
l~Q~3
feed 135. The cold liquid feed 135 contalnlng substantlal
amounts of propane and butane and higher hydrocarbons is
capable of absorbing increased amounts of desirable products
contained in the vapors leaving vapor liquid separator 16.
Example 9
Figure 11 shows another embodiment of this invention.
Plant inlet gas, from which C02, sulfur-containing gases and
moisture has been substantially removed, is cooled by heat
exchange with product streams as shown in Figure 1, and sup-
plied to separator as a cooled, partially condensed feed 23
at 900 psia and -45F. As in Figure 1, process conditions
given, and flow rates in Table XI below are for a rich feed
gas. Cooled gas from separator 16 flows through expander 17,
and the outlet stream thereof, at -125F. and flows to
separator 18. Condensed liquid collected in separator 18
is fed as stream 25 to demethanizer column 19 as top feed.
Vapors from separator 1~3 join the column overhead vapors
from demeth`anizer 19, to form stream 169 which, after heat
exchange and recompression, becomes residue gas.
Condensed liquid from high pressure separator 16 is
fed as stream 24 at -45F. to heat exchanger 16n where it is
i; sub-cooled to -130F. The cooled liquid is then fed through
an expansion valve 161 whereby it is further cooled to -137F.
; and fed to low pressure separator lh2 at 300 psia. Condensed
liquid from separator 162 flows as stream 163 through heat ex-
changer 160 in heat exchange relation with stream 24 and stream
167 where it is warmed to -90F. and fed to separator 164.
Condensed liquid from separator 164 is fed to demethanizer
column 19 as stream 165. The vapor from separator 164 is re-
turned through exchanger 160 ~s stream 167 where it is cooled
to -125F. in exchange with stream 163 and then mixed with
-34-
.
. . , ~.
- / ~
:
vapor stream 166 from separator 162. The combined vapors at
-127F. are fed to the demethanizer column 19 at an inter-
mediate point. Ethane and higher hydrocarbon liquids are
collected as bottoms from demethanizer 19 as stream 168.
Component flow rates, liquid recoveries, and compres-
sion requirements for t'nis embodiment are given in the follow-
ing table:
TABLE XI
:~ (Fig. 11)
Stream Flow Rate Summary - Lb. Moles/Hr.
STREAM METHANE ETHANE PROPANE BUTANES+ TOTAL
23a 1100 222163 130 1647
24 795 202157 129 1300
16 10 5 1 32
15163 711 201157 129 1209
165 184 169153 128 636
168 4 17516n 130 470
169 1097 47 3 O 1177
RECOVERIES
20Ethane 78.9~ 42,511 GALtDAY
Propane 98.2%40,090 GAI./DAY
COMPRESSION HORSEPOIJE_
Refrigeration 440 BHP
Recompression 815 BHP
Total 1255 BHP
Example 10
In still another example of the present invention,
it may be advantageous to provide two separate expansion valves
for expansion of the sub-cooled high-pressure liquid conden-
sate. This modification may be better understood by reference
-35-
- `
to Flgure 12, which may be compared wlth Flgure 5. Referring
flrst to Figure 5, lt will be noted that the high-pressure
llquid condensate from separator 16 (stream 34~ is sub-cooled
in heat exchanger 61, expanded in expansion valve 65 and
separated into a liquid and vapor stream in separator 64.
In accordance with the modification of the present
lnvention in this Example 10', liquid stream 34 is sub-cooled
in exchanger 150. The sub-cooled liquids from exchanger 150
are divlded prior to expansion into two portions (streams 170
and 171). Stream 170 expands through expansion valve 172 and
achieves an expanded temperature of -158F. The expansion pro-
ducts are used to sub-cool llquid stream 34 in heat exchanger
150 and then supplied to the demethanizer 19 as feed stream
170a. Portion 171 of the sub-cooled hlgh-pressure liquid is
expanded ln expansion valve 173, again reaching an expansion
temperature of -158F. and supplied to the demethanizer 19
as top feed 171a.
Both feed streams 17na and 171a are vapor-liquid
mixtures and, accordingly, the demethanizer column wlll be
provlded at the feed points with approprlate column lnternals
(such as chlmney trays or the like~ which wlll effect vapor-
llquld separation of the feeds. In the embodiment illustrated
in the present example, there will normally also be one or
more distlllation trays, sleve travs, or an appropriate column
packing between the feed points of streams 170a and 171a to
provide for vapor-liquld contact between the liqulds falling
through the column from feed 171a and vapors rlsing through
the column from feed stream 170a.
0~3
It will be recognized that Figure 12 is a fragmentary
: flow diagram illustrating only the treatment of high-pressure
liquid stream 34 from separator 16. As will be apparent from
- a comparison with Figure 5, appropriate provisions also will
be made for expansion of the vapors leaving separator 16 and
. supplying those vapors to the demethanizer as an appropriate
feed stream. Provisions also will be made for cooling and
partial condensation of high-pressure feed gas initially en-
tering the process at a pressure of 910 psia and a temperature
of 120F. by heat exchange with residue gas demethanizer side
reboilers and demethanizer bottoms liquid (none of these heat
exchangers being shown in fragmentary Figure 12).
Provision for two expansion devices as 172 and 173
of this example and feeding the expanded products directly to
the demethanizer column 19 provides improved mechanical
simplicity since it eliminates the need for vapor-liquid
separators if the separation is done external to the column
and eliminates piping for side stream wlthdrawal and return
as in Example 6 (Figure 8). The performance of this
modification can be seen from the following process flow stream
summary (the flow stream conditions belng comparable to Figure
, 5):
TABLE XII
(Fig. 12)
Str_am Flow Rate Summary - Lb. Moles!Hr.
STREAM METHANE ETHANE PROPANE B~TANES+ TOTAL
- 34 280 43 25 39 391
: 170 140 21 12 19 195
171 140 21 13 20 196
'
:; -3~-
:: 1~*(~3
The component recovery of C2+ fraction for this
illustration should be increased relative to the component
recoveries of Figure 2 above, and the horsepower requirements
, should be reduced.
Example 11
The following is another example of the use of two
separate expansion valves for the expansion of the sub-cooled
liquid condensate and may be understood by reference to Figure
13.
Figure 13 is a fragmentary flow process diagram for
; 10 the separation of cooled and partially condensed high-pressure
gas 174 supplied to separator 16 at a temperature of -55F.
and a pressure of 900 psia. Prior art processes similar to
those shown in Figures 1 and 2 are used to cool the inlet
gas to -55F. These include provision for heat exchange with
residue gas, external refrigeration (if needed), demethanizer
bottoms and one or more demethanizer side reboilers as illus-
trated in Figures 1 and 2 but not shown in the fragmentary
drawing Figure 13.
The process flow conditions indicated in Figure 13
differ from those set forth in Figures 1 and 2 since the
assumed composition employed for purposes of inlet feed gas
174 of Figure 13 was intermediate in composition between the
rich and lean gases on which Figures 1 and 2 are based. For
purposes of the calculations, two demethanizer side boilers
(not shown) were assumed, as in Figures 1 and 2.
Referring to Figure 13, the cooled vapor from separa-
tor 16 is divided into two portions. The first portion 176
flows throùgh expander 17 where, because of work expansion from
900 to 290 psia, it is coolet to about -133F. From expander
17 the chilled vapor flows to demethanizer 19 as ies middle
feed. The second vapor portio~ 177 is combined with a portion
-3B-
179 of the sub-cooled liquid from heat exchanger 184 as ex-
plained below.
The cooled liquid 175 from separator 16 flows through
exchanger 184 where it is sub-cooled to -130~F. by heat ex-
change with the cold stream from expansion valve 182. Thesub-cooled liquid is then divided into two portions. The
first portion 178 flows through expansion valve 182 where it
undergoes expansion and flash vaporization as the pressure is
reduced from about 900 to 250 psia. The cold stream from
expansion valve 182 then flows through exchanger 184 where it
is used to sub-cool the liquids from separator 16. From ex-
changer 184, the stream flows to demethanizer 19 as its lowest
feed.
The second liquid portion 179 from exchanger 184,
still at high pressure, is combined with portion 177 of the
vapor stream from separator 16. The combined stream then
flows through heat exchanger 185 where it is sub-cooled to
approximate,ly -140F. by heat exchange with cold vapor stream
180. The sub-cooled stream then enters expansion valve 183
where it undergoes expansion and flash vaporization as the
pressure is reduced from 895 psia to 250 psia. From expansion
valve 183, the cold stream proceeds to demethanizer 19 as its
top feed.
Inlet and liquid component flow rates, outlet re-
covery efficiencies, and expansion/compression requirementsfor the embodiment of this invention as illustrated in Figure
13 are given in the following table:
, ,
_~9_
1~4~V/~!3
.:
TABLE XIII
(Fig. 13)
Stream Flow Rate Summary - Lb. Moles/Hr.
STREAM METHANE ETHANE PROPANE BUTANES+ TOTAL
5 1741304 162 80 54 1647
175486 109 66 51 723
176723 47 12 2 817
177 95 6 2 1 107
178243 54 33 26 361
10 179243 55 33 25 362
- 1801301 14 1 0 1362
181 3 148 79 54 285
P~ECOVERIES
Ethane 91.47% 36,036 CAL/DAY
15 Propane 98.38% 19,732 OAL/DAY
HORSEPOWER REQUIREMENTS
Refrigeration 130 BHP
Recompression 987 BHP
Total 1117 BHP
..
; 20 Examp]e 12
Another embodiment of the present invention is
illustrated by Figure 14. Figure 14 is a frag~entary flow
diagram showing the treatment of a partially-condensed feed
gas 33a entering high pressure separator 16 at -67F. and
- 25 900 psia. The feed gas is partially condensed by heat ex-
change with residue gas, ethane product and demethanizer
- liquids as shown in Figure 2. As in Figure 2, the process
conditions set forth, as the flow rates in Table 14 below,
are for a lean feed gas.
_ 4 () _
.' ' . ' . ' ' .
,
1~4~ 3
Following the process of Figure 14, the liquid
stream 34 from separator 16 is sub-cooled through heat ex-
changer 190 in heat exchange relatlon with a portion 196 of
the overhead vapor stream 200 from the demethanizer 19 re-
sulting in sub-cooling of the liquid stream 34. The sub-
cooled stream is then expanded through an appropriate expan-
sion device, such as expansion valve 192, to a pressure oE
approximately 250 psia. During expansion a portion of the
feed will vaporize. resulting in cooling of the remaining
part. In the process illustrated in Figure 14, the expanded
liquids leaving expansion valve 192 reach a temperature of
about -158F., and are supplied to the demethanizer column l9
as an intermediate feed 34a.
The vapor from separator 16 is split into streams
193 and 194 as it leaves the top of the separator. The first
portion, stream 193, flows through exchanger 195 where it is
chilled to about -159F. by heat exchange with another portion
197 of demethanizer overhead vapors 200. From exchanger 195
the chilled vapor portion flows through expansion valve 198,
where it undergoes expansion and flash vaporization as the
pressure is reduced to about 250 psia. The flash expansion
further cools the stream to about -169F. From expansion
valve 198 the stream flows to demethanizer 19 as top feed to
the column.
The second portion, 194, of vapor from ~eparator
16 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 expanded vapor in s~ream 194 to a temperature of
approximately -153F. The expanded and partially condensed
1~4~O~L~3
vapor is supplied as feed to demethanizer 19 at an intermediate
~- point on the column, below the sub-cooled liquid stream feed
34a.
It may be noted that by comparison with Figure 3
the stream leaving expander 17 and entering the demethanizer
column achieves a temperature of about -153F. As a result
of splitting the vapor stream from separator 16, and cooling
one of the streams prior to expansion, a colder demethanizer
; top feed can be realized.
The stream flow rates, component recoveries, and
expansion/compression requirement for the process illustrated
in Figure 14 are given in the following table:
TABLE XlV
(Fig. 14)
Stream Flow Rate Summary - Lb. Moles/Hr.
_ _ _ _ _ _ _ _ .
STREAM METHANE ETHANE PROPANE BUTANES+ TOTAL
33a 1447 90 36 43 1647
34 280 42 25 39 391
193 102 4 1 0 110
20194 1065 44 ln 4 1146
196 956 4 0 0 976
197 489 2 0 0 500
199 2 84 36 43 171
200 1445 6 0 0 1476
25RECO ERIES
Ethane 92.8%20,377 GAL/DAY
Propane 99.5%9,057 GAL/DAY
HORSEPOWER REO~UIREMENTS
Refrigeration 0 BHP
30Recompression 1224 BHP
Total 1224 eHP
'
-42-
,
:
; In light of the foregoing disclosures, still other
variations of the process of the present invention will be
evident:
1. As already noted it may, in appropriate cases,
be desirable to sub-cool only a portion of the high-pressure
condensed liquid feed prior to expansion.
2. As explained in the same applicant~ Canadian
patent appllcation No.-271,343 of Campbell and Wilkinson,
filed February 8, 1977, it may be desirable to
combine the sub-cooled, high-pressure liquid feed (either
before or after sub-cooling) with a process stream containing
substantial quantlties of volatile fractions capable of re-
ducing the bubble point of the higll-pressure s~b-cooled liquid
feed (for example, as illustrated in Figure 13 of this appli-
cation).
3. The enhanced refrlgeration obtained in the flash,-
sub-cooled liquid in accordance with the present invention
may, in appropriate cases, be advantageouslv employed bv di-
.~ recting all or a part of the sub-cooled liquid into heat ex-
change relation wlth other process streams. Bv way of illus-
tration, flash-expanded, sub-cooled liquid mav be employed
to partially cool or condense all or a portion of the high-
pressure vapors obtained from the partially-condensed feed
stream before or after e~pansion of the vapor stream.
- 25 4. ~'arlations in the methods of sub-cooline may
be employed; and in this respect, two or more of the sub-
cooling techniques described ln the examples above mav be
emploved in combination.
-~3-
. .
1~4~
5. Process flow plans and examples of the present
invention have been described for convenience using shell and
tube heat exchangers. In cryogenic operations, it is usually
preferred to use specially designed heat exchangers such as
plate-fin heat exchangers. Such special heat exchangers have
improved heat transfer characteristics which may permit closer
temperature approaches in the heat exchangers, lower cost, and
also permit flow arrangements to accommodate heat exchange of
several streams concurrently as illustrated (for example) in
10 Figure 11 (exchanger 160).
To summarize the foregoing, for a given demethaniæer
pressure and expansion ratio in prior art processing as illustrated
for example in Figures 1 and 2, the liquid recoverable is sub-
ject to practical limitations, and often the desired recovery is
i:.J ,
greater than the recovery which can be practicably obtained in
single stage gas separation plants within the available limits
of pressure and expansion. To increase recovery, greater ex-
pansion ratios must be used. However, the increased expansion
ratios increase the horsepower requirements of the process at
increasing rates and thus economics limit the recovery normally
obtained in single stage gas recovery processes.
The limitations on single stage gas recovery units have
led to the use of processes having more than one separator stage
for condensed liquid vapor prior to éxpansion for the same ex-
pansion ratio and demethanizer pressure. Two stage operation
- may provide in the order of two to ten percentage points improve-
- ment in ethane recovery. However, this increase is also limited
and further increases cannot be obtained without prohibitive
increases in horsepower requirements.
Surprisingly, we have found in the present invention
that substantial increases in ethane recovery can be obtained in
-44-
.V~3
single stage operation without increasing expansion ratios. In-
deed, as may be seen for example in comparing the ethane recovery
for a typical lean gas plant of the prior art (e.g., Figure 2)
with the ethane recovery of processes in accordance with the
present invention such as Figures 5, 6, 9, and 10, signiflcant
- improvements in ethane recovery can be obtained without material-
ly increasing the horsepower requirements of the process.
Similar improvements in ethane yield can be obtained
when processing a rich gas stream, as can be seen by comparing
Figures 3, 4, and 8 (Examples 1, 2, and 6) with Flgure 1. In
~ Figure 1 (a typical prior art process for such gas) ethane re-
,~ covery was 72.9~, while in Figures 3, 4, and 8, ethane recovery
from the same gas when treated in accordance with the present
invention was between 92.1~, and 94.8%, depending on process flow
~- 15 plan. The ho.sepower requirement in Figures 3, 4, and 8 was
,. ~
between 1224 and 1331.
;~ ~he processes of Figures 3, 4, and 8 required more
horsepower,thar, the prior art process of Figure 1 for treating
the same gas. The increased horsepower requirement resulted be-
cause the increased recovery was withdrawn as a condensed liquid.
Employing the same increased horsepower to ~he prior process of
Figure 1 will not provide comparable improvements in yield.
This can be seen by considering the flow plan of Figure l where
the demethanizer is operated at a lower pressure, e.g., 250 psia
instead of 290 psia. Reducing column pressure to 250 psia in
the process of Figure 1 only increases ethane recovery to 77.1%.
At the same time horsepower required increased to 1315 BHP at
the lower column pressure.
The choice of a particular flow plan based on the
present invention will depend upon the composition of the gas to
-45-
1~h~ 3
be treated. This may be seen. for example from Figure 7
(Example 5). Where the flow plan of Figure 7 was employed to
process a lean feed gas, ethane recovery was 79.4~, and process
horsepower was 1186 BHP. When the same gas processed in accord-
ance with Figure 2, ethane recovery was 79.1~ and process horse-
power was 1180 BHP. By contrast when processing a rich gas in
the process of Figure 7, ethane recovery is 88.4~ and process
horsepower is 1195. This contrasts with processing a r~ch gas
following Figure 1 where ethane recovery is 72.9 to 77.1%, and
the horsepower required is 1148 to 1315 depending on column
pressure.
! The increased recovery of the present invention wil]
in some circumstances require increased process horsepower
(such as for recompression in compressor 22 of Figure 1 or feed
gas cooling as in exchanger 13 of Figure 1) to provide the nec-
essary cooling and refrigeration to condense the additional gases
withdrawn from the rrocess as a liquid. In our invention, this
additional required duty can usually he supplied in a manner
requiring significantly less additional horsepower thaT- would he
required in a prior art process such as Figure 1 to increase the
e;hane recovery level to the same level.
As is well known, natural gas streams usually contain
carbon dioxide in substantial amounts. The presence of carbon
dioxide in the demethanizer can lead to icing of the column in-
ternals under cryogenic conditions. Even when the feed containsless than 1~ carbon dioxide, it fractionates in the demethanizer,
and can build up to concentrations of 5% to 10% or more. At
such cqncentrations carbon dioxide can freeze out, depending
on temperature, pressure, whether the carbon dioxide is in the
liquid or vapor phase, and the solubility of carbon dioxide in
the liquid phase.
.
--46-
In the present invention, it has been found that when
- the vapor from the high-pressure separator is expanded and sup-
plied 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 as a mid column feed, therefor,
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 readily be
seen by plotting carbon dioxide concentration and temperature
for various trays of the demethanizer when practicing the present
invention and when following the prior art. A chart thus con-
structed for processing the gas as described above in Example 8
(see Figure 10 and Table X), and containing 0.72% carbon dioxide,
can be compared with a similar chart constructed for the process
of Figure 2 (prior art), applied to the same gas; (see Figures
15A and 15B). These charts also include equilibria for vapor-
solid and liquid-solid conditions. The equilibrium data given
- 20 in Figures 15A and 15B are for the methane-carbon dioxide system.
These data are generally considered representative for the methane
- ane ethane systems. If the C02 concentration at a particular
temperature in the column is at or above the equilibrium line for that
temperature, icing can be expected. For practical design purposes,
the engineer will usually require a margin of safety, i.e., the
actual concentration should be less than the "icing" concentration
by a suitable safety factor.
--q7-
,,
As is evident, when following the prior art process
of Figure 2 (per Figure 15A~ the vapor conditions at point A
touches the line representing solid-vapor phase equilibria. By
contrast, in Figure 15B, neither the vapor nor liquid conditions
reach or exceed their related equilibrium conditions.
It should be noted in connection with the foregoing
that when designing demethanizer columns for use in the pre~ent
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 the process 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 concen-
tration shifting to the right on the plots of Figures 15A and
15B. Depending on the particular application, the result can
be an objectionably high concentration of carbon dioxide near
; the top of the column. For such a circumstance, it may be neces-
sary to accept a somewhat lower ethane recovery to avoid column
;~ icing, or to pre-treat the feed gas to reduce carbon dioxide
- 20 levels to the point where they can be tolerated in the demethanizer.
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. This will tend to increase both the
temperature of the expanded vapor stream as well as the amount of
it. If this can be done within the limitations of the process
heat balance, icing may be avoided without losing ethane recovery.
-48-
-
OQ3
~.
In connection with the foregoing description of our
invention, it should be noted in some embodiments the feed to
the top of the demethanizer is wholly or partially a liquified
portion of vapors from the high-pressure separator (see, for
instance, Figures 13 and 14) which is flash expanded to the
demethanizer pressure. In some cases it may be advantageous
to provide for auto-cooling of this stream. This may be ac-
complished by dividing the liquified high-pressure vapor into
two streams either before or after expansion. (If the vapor
is divided before expansion, both parts are expanded). There-
after, one of the two divided streams after expansion is directed
into a heat exchange relation to the high-pressure vapor prior
to expansion.
. .
-49-
,
