Hydrocarbon gas processing

A process for the recovery of ethane, ethylene, propane, propylene and heavier hydrocarbon components from a hydrocarbon gas stream is disclosed. In recent years, the preferred method of separating a hydrocarbon gas stream generally includes supplying at least a portion of the gas stream to a work expansion device, then directing the work expanded stream to a distillation column at a mid-column feed point below an upper rectification section. The top column feed above the upper rectification section is typically a condensed and subcooled gaseous stream, frequently comprised of a gas stream that would otherwise feed the work expansion device. In the process disclosed, the work expanded stream is further cooled prior to feeding the distillation column at the mid-column feed point, so that a lesser volume of top column feed is required to maintain the column overhead temperature at a temperature whereby the major portion of the desired components is recovered.

BACKGROUND OF THE INVENTION 
This invention relates to a process for the separation of a gas containing 
hydrocarbons. 
Ethylene, ethane, propylene, propane and/or heavier hydrocarbons can be 
recovered from a variety of gases, such as natural gas, refinery gas, and 
synthetic gas streams obtained from other hydrocarbon 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., methane and 
ethane together comprise at least 50 mole percent of the gas. The gas also 
contains relatively lesser amounts of heavier hydrocarbons such as 
propane, butanes, pentanes and the like, as well as hydrogen, nitrogen, 
carbon dioxide and other gases. 
The present invention is generally concerned with the recovery of ethylene, 
ethane, propylene, propane and heavier hydrocarbons from such gas streams. 
A typical analysis of a gas stream to be processed in accordance with this 
invention would be, in approximate mole percent, 85.8% methane, 7.8% 
ethane and other C.sub.2 components, 3.3% propane and other C.sub.3 
components, 0.5% iso-butane, 0.7% normal butane, 0.6% pentanes plus, with 
the balance made up of nitrogen and carbon dioxide. Sulfur containing 
gases are also sometimes present. 
The historically cyclic fluctuations in the prices of both natural gas and 
its natural gas liquid (NGL) constituents have at times reduced the 
incremental value of ethane, ethylene, and heavier components as liquid 
products. This has resulted in a demand for processes that can provide 
more efficient recoveries of these products. Available processes for 
separating these materials include those based upon cooling and 
refrigeration of gas, oil absorption, and refrigerated oil absorption. 
Additionally, cryogenic processes have become popular because of the 
availability of economical equipment that produces power while 
simultaneously expanding and extracting heat from the gas being processed. 
Depending upon the pressure of the gas source, the richness (ethane, 
ethylene, and heavier hydrocarbons content) of the gas, and the desired 
end products, each of these processes or a combination thereof may be 
employed. 
The cryogenic expansion process is now generally preferred for natural gas 
liquids recovery because it provides maximum simplicity with ease of start 
up, operating flexibility, good efficiency, safety, and good reliability. 
U.S. Pat. Nos. 4,157,904, 4,171,964, 4,278,457, 4,519,824, 4,687,499, 
4,854,955, 4,869,740, 4,889,545, 5,275,005, 5,555,748, and 5,568,737, 
reissue U.S. Pat. No. 33,408, co-pending application Ser. No. 08/915,065, 
and co-pending application Ser. No. 60/044,569 describe relevant processes 
(although the description of the present invention in some cases is based 
on different processing conditions than those described in the cited U.S. 
Patents and applications). 
In a typical cryogenic expansion recovery process, a feed gas stream under 
pressure is cooled by heat exchange with other streams of the process 
and/or external sources of refrigeration such as a propane 
compression-refrigeration system. As the gas is cooled, liquids may be 
condensed and collected in one or more separators as high-pressure liquids 
containing some of the desired C.sub.2 + components. Depending on the 
richness of the gas and the amount of liquids formed, the high-pressure 
liquids may be expanded to a lower pressure and fractionated. The 
vaporization occurring during expansion of the liquids results in further 
cooling of the stream. Under some conditions, pre-cooling the high 
pressure liquids prior to the expansion may be desirable in order to 
further lower the temperature resulting from the expansion. The expanded 
stream, comprising a mixture of liquid and vapor, is fractionated in a 
distillation (demethanizer) column. In the column, the expansion cooled 
stream(s) is (are) distilled to separate residual methane, nitrogen, and 
other volatile gases as overhead vapor from the desired C.sub.2 
components, C.sub.3 components, and heavier hydrocarbon components as 
bottom liquid product. 
If the feed gas is not totally condensed (typically it is not), at least a 
portion of the vapor remaining from the partial condensation can be passed 
through a work expansion machine or engine, or an expansion valve, to a 
lower pressure at which additional liquids are condensed as a result of 
further cooling of the stream. The pressure after expansion is essentially 
the same as the pressure at which the distillation column is operated. The 
combined vapor-liquid phases resulting from the expansion are supplied as 
a feed to the column. In recent years, the preferred processes for 
hydrocarbon separation involve feeding this expanded vapor-liquid stream 
at a mid-column feed point, with an upper absorber section providing 
additional rectification of the vapor phase. 
The source of the reflux stream for the upper rectification section is 
typically a portion of the above mentioned vapor remaining after partial 
condensation of the feed gas, but withdrawn prior to work expansion. An 
alternate source for the upper reflux stream may be provided by a recycled 
stream of residue gas supplied under pressure. Regardless of its source, 
this vapor stream is usually cooled to substantial condensation by heat 
exchange with other process streams, e.g., the cold fractionation tower 
overhead. Some or all of the high-pressure liquid resulting from partial 
condensation of the feed gas may be combined with this vapor stream prior 
to cooling. The resulting substantially condensed stream is then expanded 
through an appropriate expansion device, such as an expansion valve, to 
the pressure at which the demethanizer is operated. During expansion, a 
portion of the liquid will usually vaporize, resulting in cooling of the 
total stream. The flash expanded stream is then supplied as top feed to 
the demethanizer. Typically, the vapor portion of the expanded stream and 
the demethanizer overhead vapor combine in an upper separator section in 
the fractionation tower as residual methane product gas. Alternatively, 
the cooled and expanded stream may be supplied to a separator to provide 
vapor and liquid streams, so that thereafter the vapor is combined with 
the tower overhead and the liquid is supplied to the column as a top 
column feed. 
In the ideal operation of such a separation process, the residue gas 
leaving the process will contain substantially all of the methane in the 
feed gas with essentially none of the heavier hydrocarbon components, and 
the bottoms fraction leaving the demethanizer will contain substantially 
all of the heavier components with essentially no methane or more volatile 
components. In practice, however, this ideal situation is not obtained for 
the reason that the top demethanizer liquid feed is not of a sufficient 
quantity to completely absorb the C.sub.2 components and heavier 
hydrocarbon components contained in the rising vapors. This volume of 
liquid (reflux) is typically limited by heat exchange or compression 
(energy) requirements. In such a case, it may be desirable to reduce the 
richness or volume of said rising vapors in the tower, reducing the need 
for additional reflux. The present invention provides a means for 
achieving this objective, resulting in improved C.sub.2 + recoveries for 
equivalent energy requirements, or reduced energy requirements for the 
same C.sub.2 + recoveries. 
In accordance with the present invention, it has been found that C.sub.2 
recoveries can be improved in excess of 5 percent over the prior art. 
Similarly, energy (compression) requirements can be reduced by as much as 
10 percent compared to the prior art while maintaining the same C.sub.2 + 
recovery level. The present invention, although applicable for leaner gas 
streams at lower pressures and warmer temperatures, is particularly 
advantageous when processing richer feed gases at pressures in the range 
of 600 to 1000 psia or higher under conditions requiring column overhead 
temperatures of -110.degree. F. or colder.

In the following explanation of the above figures, tables are provided 
summarizing flow rates calculated for representative process conditions. 
In the tables 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 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. It should also be noted 
that the process design calculations performed for the purpose of 
comparing the processes depicted in the figures are based on the 
assumption of no heat leak from (or to) the surroundings to (or from) the 
process. The quality of commercially available insulating materials makes 
this a very reasonable assumption and one that is typically made by those 
skilled in the art. 
DESCRIPTION OF THE PRIOR ART 
Referring now to FIG. 1, in a simulation of the process according to U.S. 
Pat. No. 4,278,457, inlet gas enters the plant at 120.degree. F. and 900 
psia as stream 31. If the inlet gas contains a concentration of sulfur 
compounds which would prevent the product streams from meeting 
specifications, the sulfur compounds are removed by appropriate 
pretreatment of the feed gas (not illustrated). In addition, the feed 
stream is usually dehydrated to prevent hydrate (ice) formation under 
cryogenic conditions. Solid desiccant has typically been used for this 
purpose. 
The feed stream 31 is cooled in exchanger 10 by heat exchange with cool 
residue gas at -26.degree. F. (stream 36a), bottom liquid product at 
60.degree. F. (stream 37a), demethanizer reboiler liquids at 38.degree. 
F., and demethanizer side reboiler liquids at -22.degree. F. Note that in 
all cases exchanger 10 is representative of either a multitude of 
individual heat exchangers or a single multi-pass heat exchanger, or any 
combination thereof. (The decision as to whether to use more than one heat 
exchanger for the indicated cooling services will depend on a number of 
factors including, but not limited to, inlet gas flow rate, heat exchanger 
size, stream temperatures, etc.) The cooled stream 31a enters separator 11 
at -12.degree. F. and 885 psia where the vapor (stream 32) is separated 
from the condensed liquid (stream 35). 
The vapor (stream 32) from separator 11 is divided into two streams, 33 and 
34. Stream 33, containing about 32 percent of the total vapor, passes 
through heat exchanger 12 in heat exchange relation with the demethanizer 
overhead vapor stream 36 resulting in cooling and substantial condensation 
of the vapor stream. The substantially condensed stream 33a at 
-140.degree. F. is then flash expanded through an appropriate expansion 
device, such as expansion valve 13, to the operating pressure 
(approximately 300 psia) of the fractionation tower 17. During expansion a 
portion of the stream is vaporized, resulting in cooling of the total 
stream. In the process illustrated in FIG. 1, the expanded stream 33b 
leaving expansion valve 13 reaches a temperature of -154.degree. F. and is 
supplied to separator section 17a in the upper region of fractionation 
tower 17. The liquids separated therein become the top feed to 
demethanizing section 17b. 
The remaining 68 percent of the vapor from separator 11 (stream 34) enters 
a work expansion machine 14 in which mechanical energy is extracted from 
this portion of the high pressure feed. The machine 14 expands the vapor 
substantially isentropically from a pressure of about 885 psia to a 
pressure of about 300 psia, with the work expansion cooling the expanded 
stream 34a to a temperature of approximately -91.degree. F. The typical 
commercially available expanders are capable of recovering on the order of 
80-85% of the work theoretically available in an ideal isentropic 
expansion. The work recovered is often used to drive a centrifugal 
compressor (such as item 15), that can be used to re-compress the residue 
gas (stream 36b), for example. The expanded and partially condensed stream 
34a is supplied as feed to the distillation column at an intermediate 
point. 
The separator liquid (stream 35) is likewise expanded to approximately 300 
psia by expansion valve 16, cooling stream 35 to -48.degree. F. (stream 
35a) before it is supplied to the demethanizer in fractionation tower 17 
at a lower mid-column feed point. An alternative routing for the separator 
liquid (stream 35) in accordance with U.S. Pat. No. 4,157,904 is indicated 
by a dashed line whereby at least a portion of the liquid is combined with 
vapor stream 33, thereby constituting a richer (higher C.sub.2 + content) 
top column feed. 
The demethanizer in fractionation tower 17 is a conventional distillation 
column containing a plurality of vertically spaced trays, one or more 
packed beds, or some combination of trays and packing. As is often the 
case in natural gas processing plants, the fractionation tower may consist 
of two sections. The upper section 17a is a separator wherein the 
partially vaporized top feed is divided into its respective vapor and 
liquid portions, and wherein the vapor rising from the lower distillation 
or demethanizing section 17b is combined with the vapor portion (if any) 
of the top feed to form the cold residue gas distillation stream 36 which 
exits the top of the tower. The lower, demethanizing section 17b contains 
the trays and/or packing and provides the necessary contact between the 
liquids falling downward and the vapors rising upward. The demethanizing 
section also includes reboilers which heat and vaporize a portion of the 
liquids flowing down the column to provide the stripping vapors which flow 
up the column. 
The liquid product (stream 37) exits the bottom of the tower at 53.degree. 
F., based on a typical specification of a methane to ethane ratio of 
0.028:1 on a molar basis in the bottom product. The stream is pumped to 
approximately 805 psia (stream 37a) in pump 18. Stream 37a, now at about 
60.degree. F., is warmed to 115.degree. F. (stream 37b) in exchanger 10 as 
it provides cooling to stream 31. (The discharge pressure of the pump is 
usually set by the ultimate destination of the liquid product. Generally 
the liquid product flows to storage and the pump discharge pressure is set 
so as to prevent any vaporization of stream 37b as it is warmed in 
exchanger 10.) 
The residue gas (stream 36) passes countercurrently to the incoming feed 
gas in: (a) heat exchanger 12 where it is heated to -26.degree. F. (stream 
36a), and (b) heat exchanger 10 where it is heated to 108.degree. F. 
(stream 36b). The residue gas is then re-compressed in two stages. The 
first stage is compressor 15 driven by expansion machine 14. The second 
stage is compressor 19 driven by a supplemental power source which 
compresses the residue gas to 900 psia (stream 36d), sufficient to meet 
line requirements (usually on the order of the inlet pressure). 
A summary of stream flow rates and energy consumption for the process 
illustrated in FIG. 1 is set forth in the following table: 
TABLE I 
______________________________________ 
(FIG. 1) 
Stream Flow Summary-(Lb. Moles/Hr) 
______________________________________ 
Stream 
Methane Ethane Propane 
Butanes+ 
Total 
______________________________________ 
31 23542 2144 898 491 27451 
32 23011 1942 686 217 26222 
35 531 202 212 274 1229 
33 7271 614 217 69 8286 
34 15740 1328 469 148 17936 
36 23490 299 8 0 24129 
37 52 1845 890 491 3322 
______________________________________ 
Recoveries* 
Ethane 86.07% 
Propane 99.09% 
Butanes+ 99.92% 
Horsepower 
Residue Compression 14,195 
______________________________________ 
*(Based on unrounded flow rates) 
The prior art illustrated in FIG. 1 is limited to the ethane recovery shown 
in Table I by the amount of substantially condensed feed gas which can be 
produced to serve as reflux for the upper rectification section of the 
demethanizer. The recovery of C.sub.2 components and heavier hydrocarbon 
components can be improved up to a point either by increasing the amount 
of substantially condensed feed gas supplied as the top feed of the 
demethanizer, or by lowering the temperature of separator 11 to reduce the 
temperature of the work expanded feed gas and thereby reduce the 
temperature and quantity of vapor supplied to the mid-column feed point of 
the demethanizer that must be rectified. Changes of this type can only be 
accomplished by removing more energy from the feed gas, either by adding 
supplemental refrigeration to cool the feed gas further, or by lowering 
the operating pressure of the demethanizer to increase the energy 
recovered by work expansion machine 14. In either case, the utility 
(compression) requirements will increase inordinately while providing only 
marginal increases in C.sub.2 + component recovery levels. 
One way to achieve higher ethane recovery in a case such as this (where the 
recovery is limited by the energy that can be removed from the feed gas) 
without lowering the demethanizer operating pressure is to substantially 
condense a portion of the re-compressed residue gas and recycle it to the 
demethanizer as its top (reflux) feed. In essence, this is a 
compression-refrigeration cycle for the demethanizer using the volatile 
residue gas as the working fluid. FIG. 2 represents such an alternative 
prior art process in accordance with U.S. Pat. No. 4,687,499 that recycles 
a portion of the residue gas product to provide a leaner top feed to the 
demethanizer. The process of FIG. 2 has been applied to the same feed gas 
composition and conditions as described above for FIG. 1. However, in the 
simulation of the process of FIG. 2 a recovery level has been selected 
that is not reasonably achievable with the process of FIG. 1. 
In the simulation of this process, as in the simulation for the process of 
FIG. 1, operating conditions were selected to minimize energy consumption 
for a given recovery level. The feed stream 31 is cooled in exchanger 10 
by heat exchange with cool residue gas at -104.degree. F. (stream 37), 
bottom liquid product at 62.degree. F. (stream 35a), demethanizer reboiler 
liquids at 38.degree. F., and demethanizer side reboiler liquids at 
-34.degree. F. The cooled stream 31a enters separator 11 at -46.degree. F. 
and 885 psia where the vapor (stream 32) is separated from the condensed 
liquid (stream 33). 
The vapor from separator 11 (stream 32) enters a work expansion machine 12 
in which mechanical energy is extracted from this portion of the high 
pressure feed. The machine 12 expands the vapor substantially 
isentropically from a pressure of about 885 psia to the operating pressure 
of the demethanizer of about 323 psia, with the work expansion cooling the 
expanded stream 32a to a temperature of approximately -118.degree. F. The 
expanded and partially condensed stream 32a is supplied as a feed to the 
distillation column at an intermediate point. The separator liquid (stream 
33) is likewise expanded to 323 psia by expansion valve 14, cooling stream 
33 to -93.degree. F. (stream 33a) before it is supplied to the 
demethanizer in fractionation tower 15 at a lower mid-column feed point. 
A portion of the high pressure residue gas (stream 40) is withdrawn from 
the main residue flow (stream 34e) to become the top distillation column 
feed. Recycle gas stream 40 passes through heat exchanger 19 in heat 
exchange relation with a portion of the cool residue gas (stream 36) where 
it is cooled to -75.degree. F. (stream 40a). Cooled recycle stream 40a 
then passes through heat exchanger in heat exchange relation with the cold 
demethanizer overhead distillation vapor stream 34 resulting in further 
cooling and substantial condensation of the recycle stream. The further 
cooled stream 40b at -148.degree. F. is then expanded through an 
appropriate expansion device, such as expansion valve 21. As the stream is 
expanded to the demethanizer operating pressure of 323 psia, it is cooled 
to a temperature of approximately -158.degree. F. (stream 40c). The 
expanded stream 40c is supplied to the tower as the top feed. 
The liquid product (stream 35) exits the bottom of tower at 54.degree. F. 
This stream is pumped to approximately 805 psia (stream 35a) in pump 16. 
Stream 35a, now at 62.degree. F., is warmed to 115.degree. F. (stream 35b) 
in exchanger 10 as it provides cooling to stream 31. 
The cold residue gas (stream 34) at a temperature of -153.degree. F. passes 
countercurrently to the recycle gas stream in heat exchanger 20 where it 
is warmed to -104.degree. F. (stream 34a). The warmed residue gas is then 
divided into two portions, streams 36 and 37. One portion, stream 36, 
passes countercurrently to the recycle stream 40 in heat exchanger 19 
where it is heated to 113.degree. F. (stream 36a). The other portion, 
stream 37, passes countercurrently to the incoming feed gas in heat 
exchanger 10 where it is heated to 105.degree. F. (stream 37a). The two 
heated streams then recombine to form the warm residue stream 34b at 
107.degree. F. The recombined warm residue gas is then re-compressed in 
two stages. The first stage is compressor 13 driven by expansion machine 
12. The second stage is compressor 17 driven by a supplemental power 
source which compresses the residue gas to 900 psia (stream 34d). After 
stream 34d is cooled to 120.degree. F. (stream 34e) by heat exchanger 18, 
the recycle stream 40 is withdrawn and the residue gas product (stream 39) 
flows to the sales pipeline. 
A summary of stream flow rates and energy consumption for the process 
illustrated in FIG. 2 is set forth in the following table: 
TABLE II 
______________________________________ 
(FIG. 2) 
Stream Flow Summary-(Lb. Moles/Hr) 
______________________________________ 
Stream 
Methane Ethane Propane 
Butanes+ 
Total 
______________________________________ 
31 23542 2144 898 491 27451 
32 21308 1479 393 90 23605 
33 2234 665 505 401 3846 
40 7687 25 0 0 7810 
34 31171 102 0 0 31670 
39 23484 77 0 0 23860 
35 58 2067 898 491 3591 
______________________________________ 
Recoveries* 
Ethane 96.45% 
Propane 100.00% 
Butanes+ 100.00% 
Horsepower 
Residue compression 18,130 
______________________________________ 
*(Based on unrounded flow rates) 
Comparison of the recovery levels displayed in Tables I and II shows that 
the additional refrigeration in the FIG. 2 process created by recycling a 
portion of the column overhead stream provides a substantial improvement 
in liquids recovery. The FIG. 2 process improves ethane recovery from 
86.07% to 96.45%, propane recovery from 99.09% to 100.00%, and butanes+ 
recovery from 99.92% to 100.00%. However, the horsepower (utility) 
requirement of the FIG. 2 process is substantially higher (by almost 28 
percent) than that of the FIG. 1 process. This means that the liquid 
recovery efficiency of the FIG. 2 process is about 12 percent lower than 
the FIG. 1 process (in terms of ethane recovered per unit of horsepower 
expended). 
DESCRIPTION OF THE INVENTION 
EXAMPLE 1 
FIG. 3 illustrates a flow diagram of a process in accordance with the 
present invention. The feed gas composition and conditions considered in 
the process presented in FIG. 3 are the same as those in FIG. 1. 
Accordingly, the FIG. 3 process can be compared with that of the FIG. 1 
process to illustrate the advantages of the present invention. 
In the simulation of the FIG. 3 process, inlet gas enters at 120.degree. F. 
and a pressure of 900 psia as stream 31. The feed stream 31 is cooled in 
exchanger 10 by heat exchange with cool residue gas at -44.degree. F. 
(stream 36b), bottom liquid product at 57.degree. F. (stream 37a), 
demethanizer reboiler liquids at 34.degree. F., and demethanizer side 
reboiler liquids at -34.degree. F. The cooled stream 31a enters separator 
11 at -25.degree. F. and 885 psia where the vapor (stream 32) is separated 
from the condensed liquid (stream 35). 
The vapor (stream 32) from separator 11 is divided into gaseous first and 
second streams, 33 and 34. Stream 33, containing about 23 percent of the 
total vapor, passes through heat exchanger 12 in heat exchange relation 
with the cool residue gas (stream 36a) where it is cooled to -103.degree. 
F. (stream 33a). The partially cooled stream 33a then passes through heat 
exchanger 13 and is further cooled by heat exchange relation with a 
portion (stream 38) of the -148.degree. F. cold distillation stream 36. 
The resulting substantially condensed stream 33b leaves exchanger 13 at 
-143.degree. F. and is then flash expanded through an appropriate 
expansion device, such as expansion valve 14, to the operating pressure 
(approximately 297 psia) of fractionation tower 19. During expansion a 
portion of the stream is vaporized, resulting in cooling of the total 
stream. In the process illustrated in FIG. 3, the expanded stream 33c 
leaving expansion valve 14 reaches a temperature of -155.degree. F. and is 
supplied to fractionation tower 19 as the top column feed. The vapor 
portion (if any) of stream 33c combines with the vapors rising from the 
top fractionation stage of the column to form distillation stream 36, 
which is withdrawn from an upper region of the tower. 
Returning to the gaseous second stream 34, the remaining 77 percent of the 
vapor from separator 11 enters a work expansion machine 15 in which 
mechanical energy is extracted from this portion of the high pressure 
feed. The machine 15 expands the vapor substantially isentropically from a 
pressure of about 885 psia to a pressure of about 300 psia, with the work 
expansion cooling the expanded stream 34a to a temperature of 
approximately -103.degree. F. The expanded and partially condensed stream 
34a then passes through heat exchanger 17 in heat exchange relation with 
the remaining portion (stream 39) of cold distillation stream 36, 
resulting in further cooling of the work expanded stream 34b to a 
temperature of -116.degree. F., which is thereafter supplied as feed to 
the distillation column at a mid-column feed point. 
The condensed liquid (stream 35) from separator 11 is flash expanded 
through an appropriate expansion device, such as expansion valve 18, to 
the operating pressure (approximately 297 psia) of fractionation tower 19, 
cooling stream 35 to a temperature of -66.degree. F. (stream 35a). The 
expanded stream 35a leaving expansion valve 18 is then supplied to 
fractionation tower 19 at a lower mid-column feed point. 
The liquid product (stream 37) exits the bottom of tower 19 at 49.degree. 
F. and is pumped to a pressure of approximately 805 psia (stream 37a) in 
demethanizer bottoms pump 20. The pumped liquid product is then warmed to 
115.degree. F. (stream 37b) in exchanger 10 as it provides cooling to 
inlet gas stream 31. 
The cold distillation stream 36 at -148.degree. F. from the upper section 
of the demethanizer is divided into two portions, streams 38 and 39. 
Stream 38 passes countercurrently to the cooled and partially condensed 
gaseous first stream 33a in heat exchanger 13 where it is warmed to 
-108.degree. F. (stream 38a) as it provides cooling and substantial 
condensation of stream 33a. Similarly, stream 39 passes countercurrently 
to work expanded stream 34a in heat exchanger 17 where it is warmed to 
-108.degree. F. (stream 39a) as it provides further cooling and additional 
condensation of stream 34a. The two partially warmed streams 38a and 39a 
then recombine as stream 36a at a temperature of -108.degree. F. The 
recombined stream passes countercurrently to the gaseous first stream 33 
in heat exchanger 12 where it is heated to -44.degree. F. (stream 36b) as 
it provides cooling and partial condensation of stream 33. 
The cool residue gas stream 36b passes countercurrently to the incoming 
feed gas stream 31 in heat exchanger 10 where it is heated to 108.degree. 
F. (stream 36c) . The warm residue gas stream 36c is then re-compressed in 
two stages. The first stage is compressor 16 driven by expansion machine 
15. The second stage is compressor 21 driven by a supplemental power 
source which compresses the residue gas (stream 36d) to sales line 
pressure of 900 psia (stream 36e). 
A summary of stream flow rates and energy consumption for the process 
illustrated in FIG. 3 is set forth in the following table: 
TABLE III 
______________________________________ 
(FIG. 3) 
Stream Flow Summary-(Lb. Moles/Hr) 
______________________________________ 
Stream 
Methane Ethane Propane 
Butanes+ 
Total 
______________________________________ 
31 23542 2144 898 494 27451 
32 22632 1815 584 161 25551 
35 910 329 314 330 1900 
33 5284 424 136 38 5965 
34 17348 1391 448 123 19586 
36 23489 243 7 0 24050 
37 53 1901 891 491 3401 
______________________________________ 
Recoveries* 
Ethane 88.66% 
Propane 99.28% 
Butanes+ 99.95% 
Horsepower 
Residue Compression 14,195 
______________________________________ 
*(Based on unrounded flow rates) 
Comparison of the recovery levels displayed in Tables I and III shows that 
the present invention in its simplest form improves ethane recovery from 
86.07% to 88.66%, propane recovery from 99.09% to 99.28%, and butanes+ 
recovery from 99.92% to 99.95%. Comparison of Tables I and III further 
shows that the improvement in yields was not simply the result of 
increasing the horsepower (utility) requirements. To the contrary, when 
the present invention is employed as in Example 1, not only do the ethane, 
propane, and butanes+ recoveries increase over those of the prior art 
process, but liquid recovery efficiency also increases by 3.0 percent (in 
terms of ethane recovered per unit of horsepower expended). 
A significant benefit of the process represented by FIG. 3 is an increase 
in the work available from the work expansion machine 15. By further 
cooling the work expanded stream 34a prior to feeding the column (stream 
34b), it is possible to reduce the amount of demethanizer top reflux 
(stream 33) by 28 percent in the FIG. 3 process compared to the FIG. 1 
process. In the process of FIG. 3, a reduction in the flow of stream 33 
corresponds to an increase in the flow of stream 34 feeding work expansion 
machine 15, thereby increasing the heat removed from stream 34 and 
increasing the energy available to booster compressor 16. Such an increase 
in booster compressor horsepower allows a reduction in the fractionation 
tower operating pressure, providing improved product recoveries without 
increasing utility (compression) requirements. 
EXAMPLE 2 
FIG. 4 illustrates a flow diagram in accordance with the preferred 
embodiment of the present invention when applied to the prior art process 
depicted in FIG. 1. In the simulation of the FIG. 4 process, the inlet gas 
cooling and separation scheme is essentially the same as that used in FIG. 
3. The difference lies in the disposition of the work expansion machine 15 
discharge (stream 34a). Rather than cooling the entire mixed-phase stream 
prior to feeding the fractionation tower at a mid-column feed point, it is 
preferable to cool only a portion of the vapor. This further cooled stream 
then feeds the demethanizer at an upper mid-column feed point, above the 
remaining vapor portion and the liquid portion. The feed gas composition 
and conditions considered in the process presented in FIG. 4 are the same 
as those in FIGS. 1 and 3. Accordingly, FIG. 4 can be compared with the 
FIG. 1 process to illustrate the advantages of the present invention, and 
can likewise be compared to the embodiment displayed in FIG. 3. 
In the simulation of the FIG. 4 process, inlet gas enters at 120.degree. F. 
and 900 psia as stream 31 and is cooled in exchanger 10 by heat exchange 
with cool residue gas at -44.degree. F. (stream 40b), bottom liquid 
product at 55.degree. F. (stream 41a), demethanizer reboiler liquids at 
32.degree. F., and demethanizer side reboiler liquids at -31.degree. F. 
The cooled stream 31a enters separator 11 at -23.degree. F. and 885 psia 
where the vapor (stream 32) is separated from the condensed liquid (stream 
35). 
The vapor from separator 11 (stream 32) is divided into gaseous first and 
second streams, 33 and 34. Stream 33, containing about 23 percent of the 
total vapor, passes through heat exchanger 12 in heat exchange relation 
with the cool residue gas (stream 40a) where it is cooled to -103.degree. 
F. (stream 33a). The partially cooled stream 33a then passes through 
exchanger 13 and is further cooled by heat exchange relation with a 
portion (stream 42) of the -152.degree. F. cold distillation stream 40. 
The resulting substantially condensed stream 33b leaves exchanger 13 at 
-147.degree. F. and is then flash expanded through an appropriate 
expansion device, such as expansion valve 14, to the operating pressure 
(approximately 294 psia) of fractionation tower 21. The expanded stream 
33c leaving expansion valve 14 reaches a temperature of -156.degree. F. 
and is supplied to fractionation tower 21 as the top column feed. 
Returning to the gaseous second stream 34, the remaining 77 percent of the 
vapor from separator 11 enters a work expansion machine 15 in which 
mechanical energy is extracted from this portion of the high pressure 
feed. The machine 15 expands the vapor substantially isentropically from a 
pressure of about 885 psia to a pressure of about 297 psia, with the work 
expansion cooling the expanded stream 34a to a temperature of 
approximately -103.degree. F. The expanded and partially condensed stream 
34a then enters separator 17 where the vapor (stream 36) is separated from 
the condensed liquid (stream 39). 
The vapor (stream 36) from separator 17 is divided into two streams, 37 and 
38. Stream 37, containing about 53 percent of the total vapor, passes 
through heat exchanger 18 in heat exchange relation with the remaining 
portion (stream 43) of the cold distillation stream 40, resulting in 
cooling and partial condensation of stream 37a, which is thereafter 
supplied to the distillation column at an upper mid-column feed point. 
The remaining 47 percent of the vapor from separator 17 (stream 38) is fed 
to the distillation column at a second upper mid-column feed point. The 
condensed liquid (stream 39) from separator 17 is typically pumped to 
overcome hydrostatic head by pump 19 prior to feeding the distillation 
column at a third upper mid-column feed point. The decision as to whether 
to include a pump for the liquid from separator 17 will depend on the 
acceptable back-pressure on the expansion machine. If a pump is not 
included, the discharge pressure of the expansion machine must be 
increased by an amount equal to the hydrostatic head of the column of 
liquid (stream 39), less the pressure drop of the first vapor portion 
(stream 37) through exchanger 18. Depending on the particular conditions, 
an alternate solution may consist of mixing the second vapor portion 
(stream 38) with the condensed liquids (stream 39), thereby reducing the 
hydrostatic head. 
The condensed liquid (stream 35) from separator 11 is flash expanded to the 
operating pressure (approximately 294 psia) of fractionation tower 21 
through an appropriate expansion device, such as expansion valve 20, 
cooling stream 35 to a temperature of -64.degree. F. (stream 35a). The 
expanded stream 35a leaving expansion valve 20 is then supplied to 
fractionation tower 21 at a lower mid-column feed point. 
The liquid product (stream 41) exits the bottom of tower 21 at 47.degree. 
F. and is pumped to a pressure of approximately 805 psia (stream 41a) in 
demethanizer bottoms pump 22. The pumped liquid product is then warmed to 
115.degree. F. (stream 41b) in exchanger 10 as it provides cooling to 
inlet gas stream 31. 
The cold distillation stream 40 at -152.degree. F. from the upper section 
of the demethanizer is divided into two portions, streams 42 and 43. 
Stream 42 passes countercurrently to the cooled and partially condensed 
gaseous first stream 33a in heat exchanger 13 where it is warmed to 
-108.degree. F. (stream 42a) as it provides cooling and substantial 
condensation of stream 33a. Similarly, stream 43 passes countercurrently 
to the first portion of vapor from separator 17 (stream 37) in heat 
exchanger 18 where it is warmed to -108.degree. F. (stream 43a) as it 
provides cooling and partial condensation of stream 37. The two partially 
warmed streams 42a and 43a then recombine as stream 40a at a temperature 
of -108.degree. F. The recombined stream passes countercurrently to the 
gaseous first stream 33 in heat exchanger 12 where it is heated to 
-44.degree. F. (stream 40b) as it provides cooling and partial 
condensation of stream 33. 
The cool residue gas stream 40b passes countercurrently to the incoming 
feed gas stream 31 in heat exchanger 10 where it is heated to 108.degree. 
F. (stream 40c). The warm residue gas stream 40c is then re-compressed in 
two stages. The first stage is compressor 16 driven by expansion machine 
15. The second stage is compressor 23 driven by a supplemental power 
source which compresses the residue gas (stream 40d) to sales line 
pressure of 900 psia (stream 40e). 
A summary of stream flow rates and energy consumption for the process 
illustrated in FIG. 4 is set forth in the following table: 
TABLE IV 
______________________________________ 
(FIG. 4) 
Stream Flow Summary-(Lb. Moles/Hr) 
______________________________________ 
Stream 
Methane Ethane Propane 
Butanes+ 
Total 
______________________________________ 
31 23542 2144 898 491 27451 
32 22675 1829 594 166 25624 
35 867 315 304 325 1827 
33 5177 418 136 38 5850 
34 17498 1411 458 128 19774 
36 16896 905 89 5 18154 
39 602 506 369 123 1620 
37 8935 479 47 3 9600 
38 7961 426 42 2 8554 
40 23487 160 6 0 23962 
41 55 1984 892 491 3489 
______________________________________ 
Recoveries* 
Ethane 92.53% 
Propane 99.40% 
Butanes+ 99.95% 
Horsepower 
Residue Compression 14,195 
______________________________________ 
*(Based on unrounded flow rates) 
A comparison of Tables III and IV shows that preferred embodiment of the 
present invention (FIG. 4) is capable of achieving significantly higher 
product recoveries than the simpler embodiment of FIG. 3. In addition, 
comparison of Tables I and IV that, compared to the prior art, the present 
invention improves ethane recovery from 86.07% to 92.53%, propane recovery 
99.09% to 99.40%, and butanes+ recovery from 99.92% to 99.95%. Comparison 
of Tables I and IV further shows that the improvement in yields was 
achieved using equivalent horsepower (utility) requirements. When the 
present invention is employed as in Example 2, not only do the product 
recoveries increase over those of the prior art process, but liquid 
recovery efficiency also increases by 7.5 percent (in terms of ethane 
recovered per unit of horsepower expended). 
As with the process of FIG. 3, a significant portion of the benefit 
achieved by the preferred embodiment of FIG. 4 is derived from a reduction 
in top column reflux, corresponding to an increase in booster compressor 
horsepower and lower fractionation tower operating pressure. However, in 
the preferred embodiment, only that portion of the vapor leaving work 
expansion machine 15 (stream 37) that can be effectively cooled using the 
available portion of cold distillation tower overhead (stream 43) is 
routed through exchanger 18. Upon being supplied to the tower at the first 
upper mid-column feed point, the condensed liquids resulting from cooling 
of this first vapor portion of the work expanded stream then act as 
additional reflux to the vapors rising up the distillation column. In this 
manner, the vapor requiring rectification by the top reflux stream is not 
only reduced in quantity, but is also leaner in composition. The result is 
increased C.sub.2 + component recoveries for the FIG. 4 process while 
using essentially the same amount of top reflux as the FIG. 3 embodiment 
of the present invention. 
EXAMPLE 3 
FIG. 5 illustrates a flow diagram in accordance with the preferred 
embodiment of the present invention when applied to the prior art process 
depicted in FIG. 2. In the simulation of the FIG. 5 process, the inlet gas 
cooling and separation scheme is essentially the same as that used in FIG. 
2. The difference lies in the disposition of the work expansion machine 12 
discharge (stream 32a). In accordance with the preferred embodiment of the 
present invention, expanded and partially condensed stream 32a enters 
separator 14 where the vapor (stream 33) is separated from the condensed 
liquid (stream 36). 
The vapor (stream 33) from separator 14 is divided into two streams, 34 and 
35. Stream 34, containing about 58 percent of the total vapor, passes 
through heat exchanger 15 in heat exchange relation with a portion (stream 
41) of cold distillation stream 38, resulting in further cooling and 
partial condensation (stream 34a). The further cooled and partially 
condensed stream 34a is then supplied to the distillation column at an 
upper mid-column feed point. 
The remaining 42 percent of the vapor from separator 14 (stream 35) is fed 
to the distillation column at a second upper mid-column feed point. The 
condensed liquid (stream 36) from separator 14 is pumped in pump 16 prior 
to feeding the distillation column at a third upper mid-column feed point. 
As in the process of FIG. 2, the source of the top column feed (reflux) is 
compressed warm residue gas that has been cooled, substantially condensed, 
and flashed to the operating pressure of the demethanizer. 
A summary of stream flow rates and energy consumption for the process 
illustrated in FIG. 5 is set forth in the following table: 
TABLE V 
______________________________________ 
(FIG. 5) 
Stream Flow Summary-(Lb. Moles/Hr) 
______________________________________ 
Stream 
Methane Ethane Propane 
Butanes+ 
Total 
______________________________________ 
31 23542 2144 898 491 27451 
32 21822 1593 447 107 24314 
37 1720 551 451 384 3137 
33 20712 870 63 3 21961 
36 1110 723 384 104 2353 
34 12013 505 37 2 12737 
35 8699 365 26 1 9224 
47 5762 19 0 0 5856 
38 29246 95 0 0 29724 
46 23484 76 0 0 23868 
39 58 2068 898 491 3583 
Recoveries* 
Ethane 96.45% 
Propane 100.00% 
Butanes+ 100.00% 
Horsepower 
Residue Compression 16,395 
______________________________________ 
*(Based on unrounded flow rates) 
Contrary to the comparisons of Tables I, III, and IV, the stimulation of 
FIG. 5 has been presented such that the recovery levels are substantially 
constant, so that the benefit of the present invention manifests itself as 
a reduction in utility (horsepower) requirements. Comparison of Tables II 
and V shows that the prior art process of FIG. 2 essentially matches the 
C.sub.2 + recovery levels of the present invention. However, the FIG. 2 
process does so at the expense of greatly increased horsepower (utility) 
consumption. The present invention achieves the same recovery levels using 
only 90 percent of the external power required by the FIG. 2 prior art 
process. 
The reduction in compression horsepower for the FIG. 5 process stems from 
the top reflux feed for the FIG. 5 process (stream 47 in Table V) being 25 
percent less than the top reflux feed for the FIG. 2 process (stream 40 in 
Table II). The cooling and partial condensation of a portion of the vapor 
phase leaving the work expansion machine results in additional liquids 
being supplied to the fractionation tower at the upper mid-column feed 
position, whereupon these liquids act like reflux on the vapor rising up 
the tower. This in turn reduces the amount of vapor to be rectified by the 
top reflux stream (and also creates a leaner vapor composition), allowing 
the corresponding reduction in the amount of top reflux feed required to 
achieve the desired C.sub.2 + component recovery level. 
Other Embodiments 
One skilled in the art will recognize that the present invention gains much 
of its benefit by providing additional cooling to the mid-column feed or 
feeds, which typically comprise the majority of the vapor requiring 
rectification in the column. With this additional cooling, less reflux 
must be supplied to the upper section of the column, thereby reducing 
utility requirements for a given product recovery level, or improving 
product recovery levels for a given utility consumption, or some 
combination thereof. Therefore, the present invention is generally 
applicable to any process dependent on an upper reflux section consisting 
of any number of feed streams produced by substantially condensing a 
portion of the feed gas or residue gas and supplying the resulting reflux 
stream(s) to the column above the feed point(s) of the majority of the 
vapor to be rectified. 
In accordance with this invention, the cooling of at least a portion of the 
work expanded stream may be accomplished in several ways. In the process 
of FIG. 3 the entire work expanded stream is cooled, while in the 
processes of FIGS. 4 and 5 only a portion of the vapor phase of the work 
expanded stream is cooled. However, this cooling as described in the 
present invention may be applied to any portion of the work expanded 
stream, such as the entire vapor stream, the condensed liquid stream, or 
any combination of the vapor and liquid. 
This cooling of at least a portion of the mid-column feed(s) may be 
effectively carried out in any number of alternate process configurations. 
One such alternate method of cooling is presented in FIG. 6. In the 
process of FIG. 6, a portion of the demethanizer liquid is withdrawn from 
the distillation column as stream 36. After being pumped in liquid 
circulation pump 18, stream 36a is subcooled in exchanger 19 in heat 
exchange relation with a portion (stream 40) of the cold distillation 
stream 37. Subcooled stream 36b is returned to the column at a mid-column 
feed point above the point at which it was withdrawn. In this manner, low 
temperature cooling duty available in the demethanizer overhead vapor 
stream is transferred indirectly (via the side liquid stream) to the vapor 
phase of the work expanded stream rising up the fractionation tower. 
Although the process configuration of FIG. 6 appears significantly 
different from the processes of FIGS. 3, 4, and 5, the benefits observed 
arise from substantially identical process conditions. One skilled in the 
art will recognize that any such method of removing a sufficient amount of 
duty from the majority of the column vapors will provide similar results. 
Other examples include, but are not limited to, cooling and return of a 
distillation column vapor side draw and cooling of the column vapors by 
use of process equipment such as an internal heat exchanger or 
dephlegmator. The selection of a method for providing cooling will depend 
on a number of factors including, but not limited to, inlet gas 
composition and conditions, plant size, equipment availability and cost, 
etc. 
In addition to providing cooling by any number of different process methods 
as described above, it will be recognized that this cooling can be 
provided by a number of different sources. In the processes of FIGS. 3 
through 6, the cooling has been provided by a portion of the distillation 
column overhead vapor stream. However, any such stream at a temperature 
colder than the process stream being cooled may be utilized. For instance, 
as shown in flow diagram FIG. 7, a portion of the fractionation tower 
liquids (stream 42) could be withdrawn from substantially the same region 
of the distillation column as the work expanded feed. After cooling at 
least a portion (stream 37) of the work expanded stream, the partially 
vaporized liquids (stream 42a) could then be returned to the distillation 
column at a point below one or more of the work expanded feed(s). Other 
potential sources of cooling include, but are not limited to, flashed high 
pressure separator liquids and mechanical refrigeration systems. The 
selection of a source of cooling will depend on a number of factors 
including, but not limited to, inlet gas composition and conditions, plant 
size, heat exchanger size, potential cooling source temperature, etc. 
One skilled in the art will also recognize that any combination of the 
above cooling sources or methods of cooling may be employed in combination 
to achieve the desired benefit. For example, a first portion of the vapor 
in the work expanded stream may be cooled by distillation column overhead 
vapor, while a second portion may be simultaneously cooled by 
fractionation column liquids. A second scenario could involve partial 
cooling of a mid-column feed with one source of cooling, with further 
cooling being provided by a second source of cooling. 
In accordance with this invention, the use of external refrigeration to 
supplement the cooling available to the inlet gas from other process 
streams may be employed, particularly in the case of an inlet gas richer 
than that used in Example 1. The use and distribution of demethanizer 
liquids for process heat exchange, and the particular arrangement of heat 
exchangers for inlet gas cooling must be evaluated for each particular 
application, as well as the choice of process streams for specific heat 
exchange services. 
The high pressure liquid in FIGS. 3, 4, 6, and 7 (stream 35) need not be 
expanded and fed to the lower mid-column feed point on the distillation 
column. Alternatively, this liquid stream (or a portion thereof) may be 
combined with the portion of the separator vapor (stream 33) flowing to 
heat exchanger 12. (This is shown by the dashed line in FIG. 3.) The 
liquid stream may also be used for inlet gas cooling or other heat 
exchange service before or after the expansion step prior to flowing to 
the demethanizer. 
It will also be recognized that the relative amount of feed found in each 
branch of the column feed streams will depend on several factors, 
including gas pressure, feed gas composition, the amount of heat which can 
economically be extracted from the feed and the quantity of horsepower 
available. More feed to the top of the column may increase recovery while 
decreasing power recovered from the expansion machine thereby increasing 
the recompression horsepower requirements. Increasing feed lower in the 
column reduces the horsepower consumption but may also reduce product 
recovery. The mid-column feed positions depicted in FIGS. 3 through 7 are 
the preferred feed locations for the process operating conditions 
described. However, the relative locations of the mid-column feeds may 
vary depending on inlet composition or other factors such as desired 
recovery levels and amount of liquid formed during inlet gas cooling. 
Moreover, two or more of the feed streams, or portions thereof, may be 
combined depending on the relative temperatures and quantities of 
individual streams, and the combined stream then fed to a mid-column feed 
position. FIGS. 3 through 7 are the preferred embodiments for the 
compositions and pressure conditions shown. Although individual stream 
expansion is depicted in particular expansion devices, alternative 
expansion means may be employed where appropriate. For example, conditions 
may warrant work expansion of the substantially condensed portion of the 
feed stream (33b in FIGS. 3, 4, and 6, and 33a in FIG. 7) or the 
substantially condensed recycle stream (47c in FIG. 5). 
While there have been described what are believed to be preferred 
embodiments of the invention, those skilled in the art will recognize that 
other and further modifications may be made thereto, e.g. to adapt the 
invention to various conditions, types of feed, or other requirements, 
without departing from the spirit of the present invention as defined by 
the following claims.