Economic recovery of ethylene and/or propylene from low concentration feeds

C.sub.2 -C.sub.3 olefins are recovered from a low concentration gaseous feedstream, e.g. a Fischer-Tropsch waste gas. In a first process stage, the weight percent concentration of the olefins is enriched to at least about 25%. One specific enrichment technique includes scrubbing, preferably 2-stage scrubbing, with hydrocarbon scrubbing agents. Another technique comprises condensation of the olefin in regenerators and absorbing same in a purge phase. The resultant enriched stream, in any case, is then subjected to conventional low temperature separation to recover relatively pure individual streams of ethylene and propylene.

BACKGROUND OF THE INVENTION 
This invention relates to a process and apparatus for the recovery of 
low-molecular weight olefins from a gaseous stream containing low 
concentrations of these hydrocarbons and of the corresponding paraffins. 
Low-molecular weight olefins, especially ethylene, constitute important 
starting materials for the chemical industry and are required in large 
quantities. A conventional large-scale industrial process for the 
production of ethylene and propylene comprises the pyrolysis of 
hydrocarbons, using high cracking temperatures and short residence times. 
Such processes usually lead to a product gas containing, in a typical 
cracking operation based on naphtha and gas oil, between 20% and 30% by 
weight of ethylene. This cracked gas, after cooling and separation of 
condensed less volatile materials, is fractionated into its individual 
components in a low temperature gas separation facility. 
Economically feasible operation of such a low-temperature separation 
facility requires that the concentration of the desired low-molecular 
weight olefins in the feed gas be at a concentration of generally at least 
on the order of about 20% by weight, varying according to plant investment 
and operating costs of the particular facility as well as to the market 
price for the final products. 
A lower concentration of the olefins generally results in an economically 
unattractive low-temperature separation, the lower the concentration, the 
poorer the economics. 
In a number of large-scale industrial processes not having a major 
objective of obtaining low-molecular weight olefins, there are 
nevertheless by-product or waste gaseous streams produced containing such 
a low concentration of olefins that the low temperature recovery thereof 
is not economically feasible. Such a gas is, for example, the waste gas 
from a Fischer-Tropsch synthesis oriented primarily to the production of 
hydrocarbons boiling in the gasoline range from a synthesis gas of 
hydrogen and carbon oxides. The waste gas produced in such a process 
contains mostly unreacted synthesis gas along with a minor quantity of 
light hydrocarbons formed during the synthesis, such as methane, ethane, 
ethylene, propane, propylene, and C.sub.4 -hydrocarbons. The ethylene or 
propylene content ranges, for example, at about 2-3 mol-% in a 
conventional Fischer-Tropsch plant. 
Additional information concerning the prior art processes is found in U.S. 
Pat. No. 2,915,881, discussed in greater detail below. 
SUMMARY OF THE INVENTION 
An object of the present invention is to provide an economically feasible 
process for recovering low molecular weight olefins from low concentration 
feedstreams. 
Another object is to provide apparatus for such a process. 
Upon further study of the specification and appended claims, further 
objects and advantages of this invention will become apparent to those 
skilled in the art. 
These objects are attained by providing a process wherein in a first 
process stage, at least one fraction is obtained from the gaseous 
feedstream with an increased concentration of the low-molecular weight 
olefins and of the corresponding paraffins if present; and that the 
low-molecular weight olefins are thereafter recovered from the 
thus-produced fraction(s) in a second process stage by low-temperature 
separation. 
According to the invention, a procedure is thus proposed wherein initially 
an enrichment of all hydrocarbons is effected which lie in the boiling 
range of the olefins to be recovered. The resultant gaseous streams 
enriched in the olefins to be obtained can then be further processed by 
low-temperature separation. By "low molecular weight olefins" is meant 
C.sub.2 - to C.sub.4 -olefins, specifically ethylene and propylene. 
Gaseous feedstreams suitable for the process of this invention are 
preferably those containing at least 2 mol-% ethylene and/or 2 mol-% 
propylene. Although an enrichment of these olefins is also possible in 
case of lower concentrations in the gaseous stream, the expenditure for 
the enriching step at lower concentrations in most cases under prevailing 
economic conditions becomes so high that an olefin-producing step is 
economically unjustifiable. Conversely, this invention is most 
advantageous at the present time when the feedstreams contain not more 
than about 12, preferably not more than about 8 mol-% ethylene and/or not 
more than 15, preferably not more than about 10 mol-% propylene. 
To ensure an advantageous operation of the subsequently connected 
low-temperature separation system, it is suitable to proceed with the 
enrichment in the first process stage to such an extent that at least 
about 25 mol-% olefins, preferably more than 35 mol-% olefins are present 
in the concentrate. The proportion of the individual olefins can very in 
this connection, depending on the composition of the gaseous stream. Thus, 
when obtaining olefins from the waste gases of a Fischer-Tropsch 
synthesis, it is advantageous to aim, in the first process stage, for 
ethylene concentrations of between about 10 and 17 mol-%, in order to 
obtain a suitable feed for the second process stage. An even further 
increase in the olefin proportion in the first process stage is desired, 
as long as no intolerably high expense is required for this purpose in the 
first process stage, and since higher olefin concentrations even tend to 
facilitate the separating task in the second process stage. 
When processing the waste gas from a Fischer-Tropsch synthesis according to 
the process of this invention, it is advantageous first to separate carbon 
dioxide from this gas. This component, present in the waste gas in a high 
concentration, e.g. 47 mol-%, can be separated conventionally in a 
preceding process stage. Scrubbing methods or adsorption processes are 
suitable for this purpose, for example. 
Since an individual Fischer-Tropsch synthesis conventionally is based on a 
feedstock of a synthesis gas produced by the gasification of coal or other 
heavy, carbon-containing starting materials, the process of this invention 
constitutes a breakthrough insofar as it provides for the economically 
feasible production of ethylene and/or propylene as by-products from such 
feedstocks. 
The enriching of the olefins in the first process stage can be conducted by 
a scrubbing step in one embodiment of the process of this invention. In 
such a case, it is expedient to utilize scrubbing media which selectively 
absorbs gaseous low mol rate hydrocarbons. It has proven to be especially 
advantageous to use hydrocarbons as the scrubbing media, for example 
hydrocarbons of 4-8 carbon atoms in the molecule. 
In an especially favorable embodiment of this invention, the enrichment of 
the olefins is conducted in a two-stage scrubbing process wherein, in the 
second scrubbing stage, a lower temperature is ambient than in the first 
scrubbing stage, the temperature difference being preferably about 
15.degree. to 35.degree. C., especially 20.degree. to 30.degree. C., and 
wherein differing scrubbing media are employed in the two scrubbing 
stages. The loaded scrubbing media of the individual scrubbing stages are 
regenerated in separate cycles during which step the components scrubbed 
out of the gaseous stream are released and can be introduced into the 
second process stage. 
Although a two-stage process for obtaining ethylene from gases rich in 
C.sub.2 has been known from U.S. Pat. No. 2,915,881, this procedure would 
have considerable disadvantages if applied to the feedstream of the 
invention having a low olefin content. In the patented method, the raw gas 
is fed to two series-connected absorption columns utilizing propylene as 
the scrubbing agent, wherein the temperature in the first scrubbing stage 
is higher than in the second scrubbing stage. The pressures in the two 
columns are about 35 bar. Since a C.sub.2 -rich gas is involved, the 
compression of the entire gas to this pressure has no deleterious effect. 
The minor proportion of these hydrocarbons in the gaseous streams to be 
processed according to this invention would presuppose, if the 
conventional process were utilized, that here, too, the entire gas be 
compressed to 35 bar and cooled to relatively low temperatures, namely 
about -63.degree. C. in the first scrubbing stage and -73.degree. C. in 
the second scrubbing stage. These operations could only be accomplished 
with a very great energy expenditure, and would nowise produce an 
analogous yield of the desired hydrocarbons. Besides, the high proportion 
of inert components at -73.degree. C. and 35 bar would lead at the head of 
the second absorber to a loss of about 50% of the C.sub.3 -hydrocarbons 
contained in the gas; however, these C.sub.3 -hydrocarbons are also 
desirable as a product in such a process. 
A particularly advantageous feature of this invention resides in the use of 
C.sub.4 -or higher hydrocarbons as the scrubbing medium in the first 
scrubbing stage and the use of C.sub.6 - or higher hydrocarbons in the 
second scrubbing stage. The operation of the first scrubbing step with a 
scrubbing medium having a lower boiling point has the advantage of 
requiring a smaller amount of scrubbing medium to scrub out the C.sub.3+ 
-hydrocarbons. This leads to lower initial investment costs and a lower 
refigeration loss due to incomplete heat exchange in the heat exchanger 
between the first scrubbing stage and the associated regenerating column. 
The desired low-molecular weight olefins can be enriched in concentrations 
of up to 15-30 mol-%, without having to compress concomitantly an 
excessive amount of inert compounds. Thus, the hydrocarbons, after 
enrichment, are present in a concentration permitting economically 
attractive processing in a downstream low-temperature separation plant. 
A preferred scrubbing medium in the first scrubbing stage is butane and in 
the second scrubbing stage, hexane. This has the advantage that neither 
the temperatures nor the pressures need to assume extreme values. In this 
case, C.sub.3+ -hydrocarbons are scrubbed out advantageously in the first 
scrubbing stage at about -5.degree. to -25.degree. C., preferably about 
-20.degree. C., and in the second scrubbing stage, C.sub.2 -hydrocarbons 
are scrubbed out at about -30.degree. to -50.degree. C., preferably about 
-40.degree. C. and under pressures of between 10 and 20 bar, preferably 
about 15 bar for the preferred temperatures. Butane boils lower than 
hexane, and therefore the regeneration of the butane can be accomplished 
under a pressure of about 11 bar, and at more importantly sump 
temperatures at which unsaturated hydrocarbons, e.g. diolefins, do not as 
yet polymerize. Thus it is possible to feed a portion of resultant 
concentrate to further separation while already under an initial 
superatmospheric pressure. The refrigeration can be made available at 
moderately low temperatures readily and inexpensively by the propylene 
cycle which is necessary in any case in the low-temperature separation 
stage. 
When using butane and hexane as the scrubbing media, pressures of between 
about 10 and 20 bar are employed in the scrubbing columns. If the waste 
gas to be processed is supplied at a lower pressure, then it is merely 
necessary to replace the scrubbing media by higher homologs. 
According to a further embodiment of the process of this invention, the 
loaded scrubbing medium of the first scrubbing stage is freed from the 
C.sub.3 -hydrocarbons by stripping, and the remaining scrubbing medium is 
withdrawn and, by heating, C.sub.4 - and higher hydrocarbons are vaporized 
therefrom. The resultant gas is used as the stripping gas for removing the 
C.sub.3 -hydrocarbons. For this purpose the separation of the C.sub.3 
-hydrocarbons from the scrubbing medium is preferably conducted in a 
C.sub.3 /C.sub.4 -separating column mounted above a separating column for 
higher hydrocarbons, so that the vapors produced in the last-mentioned 
separating column rise through the liquid in the C.sub.3 /C.sub.4 
-separating column. 
In another embodiment of the process of this invention, the enrichment of 
the olefins in the first process stage is effected with the use of 
regenerators. In such a method, several cyclically operated regenerators 
are used. The regenerators are traversed by the gaseous stream during the 
first or loading phase, cooling said gaseous stream thereby. The 
components condensed therefrom during the cooling step settle on the 
surface of the regenerator packing while uncondensed components are 
withdrawn. In a subsequent purge phase, the condensed components are then 
revaporized by a stream of purge gas and driven out. Conventionally, a 
cooling phase follows such a purge phase, during which the regenerator is 
cooled down to such low temperatures that condensation of the components 
to be separated takes place in the subsequent, new loading phase. Such a 
regenerator process is preferably conducted with the use of at least three 
cyclically interchangeable regenerators. 
In the recovery of low-molecular weight olefins according to this invention 
from a gaseous stream consisting essentially of components having a lower 
boiling point than the olefins, for example hydrogen, carbon monoxide, 
nitrogen, and methane, the olefins and higher-boiling components of the 
gaseous stream are condensed while flowing through regenerators and are 
thus retained. In such a mode of operation, it is advantageous to proceed 
with the cooling of the gaseous stream in the regenerators to such a point 
that the olefins are essentially completely condensed, so that in a 
subsequent purge phase the purge stream can remove the olefins at an 
enriched concentration. It proved to be advantageous to cool the gaseous 
stream to such a degree that C.sub.2 -hydrocarbons are condensed to an 
extent of at least 95%, preferably 99% or more. A substantial further 
cooling is not only economically unattractive owing to the increased 
demand for refrigeration resulting in an only minor rise in the olefin 
yield, but also leads in many cases to a reduction in the olefin 
concentration in the condensate, since the lower-boiling components, e.g., 
methane, are increasingly condensed. 
To obtain the olefin-rich fraction during a purge phase, it is advantageous 
to employ as the purge gas a partial stream of the gaseous stream not 
condensed during a loading phase. Since the purge step takes place 
advantageously under a lower pressure than the loading procedure, only a 
small amount of scavenging (purge) gas is generally required, so as to 
mitigate dilution of the olefin-rich condensed fraction with lower-boiling 
components. 
While the loading of the regenerators is conducted in a pressure range 
between about 5 and 15 bar, preferably between 8 and 10 bar, the pressure 
during a purge phase is suitably between 0.1 and 2.0 bar, for example 
approximately at atmospheric pressure. The purge operation conducted under 
such low pressures, preferably 0.2 to 1.3 bar, has the advantage that the 
amount of scavenging gas is reduced and thus the dilution of the 
components to be recovered is decreased; conversely, extremely low 
subatomspheric pressures entail high costs. 
In an advantageous embodiment of the process of this invention, the 
refrigeration requirement for operating the regenerators is covered at 
least in part by engine expansion of the uncondensed, cold gaseous stream 
or of a partial stream thereof, In a preferred embodiment of the 
technique, the refrigeration demand is supplied by engine expansion of the 
purge gas stream, branched off from the uncondensed gaseous stream.

DETAILED DESCRIPTION OF THE DRAWINGS 
Via conduit 1 of the embodiment shown in FIG. 1 4,461.5 kmol/h of waste gas 
from a Fischer-Tropsch synthesis, from which the carbon dioxide, initially 
present in an amount of 47 mol-% in the waste gas, had been conventionally 
removed in a scrubbing column, not shown, is introduced under a pressure 
of 15 bar and at a temperature of about 30.degree. C. into a heat 
exchanger 2. The gas is not compressed before entering the scrubbing 
stage, since it already has the desired pressure of 15 bar from the 
Fischer-Tropsch synthesis. In heat exchanger 2, the gas is cooled to about 
5.degree. C. countercurrently to scrubbed gas and then passes via conduit 
3 into the bottom of the first scrubbing stage 4. The waste gas has the 
composition set forth in Table 1. 
In the scrubbing column 4, the gas is scrubbed countercurrently to 
downflowing liquid butane at 15 bar, said liquid having entered the column 
at -20.degree. C. The major portion of the C.sub.3+ -hydrocarbons 
contained in the gas is thereby removed from the gaseous feedstream (F-T 
waste stream). The gas exiting from the head of the scrubbing column 4, 
the composition of this gas being derivable from Table 1, has a 
temperature of about -15.degree. C. and is conducted via conduit 5 into a 
heat exchanger 6. In the latter heat exchanger, the gas is again cooled 
counter-currently to scrubbed gas and then passes via conduit 7 into the 
bottom of the second scrubbing column 8, this column operating at about 14 
bar due to minor pressure drops between this column and the scrubbing 
column 4. In the scrubbing column 8, C.sub.2 -hydrocarbons and higher 
hydrocarbons still present in the gas are scrubbed out by means of 
entering hexane at -40.degree. C. in a countercurrent operation. The 
scrubbed gas having the composition set forth in Table 1 exits from the 
scrubbing column 8 with a temperature of about -40.degree. C. and is 
removed via conduit 9 and via the heat exchangers 6 and 2. 
As can be seen from Table 1, all C.sub.2 - and higher hydrocarbons 
contained in the feedstream are completely scrubbed out except for minute 
amounts of C.sub.2 -hydrocarbons. The thus-scrubbed gas still contains 
small amounts of hexane due to saturation from the scrubbing medium. 
From the C.sub.4 -scrubbing stage of the first scrubbing column 4, the 
scrubbing medium, loaded with C.sub.3+ -hydrocarbons, is withdrawn at the 
sump via conduit 10 and fed into a heat exchanger 11. In the latter, the 
loaded butane is warmed countercurrently to regenerated butane and 
conducted via conduit 12 into the upper part of a separating colum 13, 
said column having a bottom part and an upper part which are separated 
from each other with regard to the liquid flow. Both parts are supplied 
with bubble trays for intimately contacting the downflowing liquid with 
the raising vapours in the column. 
In the upper part of column 13, C.sub.3 -hydrocarbons and all lower-boiling 
components are separated overhead. The scrubbing medium is discharged from 
the bottom zone of the upper part of the column via conduit 20. A portion 
of the scrubbing medium is fed into a heat exchanger 22 via pump 21 and 
conduit 20 and, after cooling to about -15.degree. C., in heat exchanger 
22, reintroduced into column 4 as the scrubbing medium. The other portion 
is introduced via conduit 14 into the lower portion of column 13; from the 
sump of this column, there are withdrawn C.sub.4+ -hydrocarbons. The 
overhead of the bottom part of this column comprising gaseous C.sub.4 
-hydrocarbons is fed as heating vapor to the upper part of column 13. 
The gas, enriched with C.sub.3 -hydrocarbons leaves the separating column 
13 overhead at about 30.degree. C. and is conducted via conduit 15 to a 
heat exchanger 16 wherein residues of higher-boiling hydrocarbons are 
condensed and separated in phase separator 17. The condensates pass via 
conduit 18 as reflux liquid back into the separating column 13 whereas the 
gas rich in C.sub.3 is conducted at a pressure of about 12 bar via conduit 
19 into a low-temperature separation system, not shown. The gas rich in 
C.sub.3 has the composition inducated in Table 1. 
The scrubbing medium of the second scrubbing stage, for exmple hexane, is 
regenerated in the same way. The loaded hexane at a temperature of about 
-30.degree. C. leaves the sump of the scrubbing column 8 via conduit 23 
and passes into a heat exchanger 24 wherein it is warmed countercurrently 
to regenerated hexane, whereupon it is conducted via conduit 25 after 
throttling in throttling means (not shown) to a pressure of about 5 bar 
into the middle of a separating column 26. In the latter, the C.sub.2 - 
and any remaining residues of higher hydrocarbons are heated in the 
reboiler to about 120.degree. C., vaporized therein, and withdrawn as 
overhead at about 30.degree. C. via conduit 27 and then conducted into a 
heat exchanger 28. In this latter heat exchanger, the C.sub.3 - and 
C.sub.4 -hydrocarbons are condensed. These condensates are separated in a 
separator 29 and recycled into the separating column 26 via conduit 30. 
The gas rich in C.sub.2, the composition of which can be derived from 
Table 1, is conducted from separator 29 via conduit 31 into a 
low-temperature separation system, not illustrated. 
Part of the condensate, especially the C.sub.4 -hydrocarbons, can be 
branched off from conduit 30 and passed to the same level as conduit 12 
and fed via conduit 32 into the separating column 13. 
Part of the thus-regenerated hexane is passed to the reboiler at the sump 
of the separating column 26, and the other part is withdrawn via conduit 
33 and cooled in heat exchanger 24 countercurrently to loaded hexane. Via 
pump 34, the regenerated hexane is finally passed into another heat 
exchanger 35 wherein it is cooled to about -40.degree. C. and introduced 
at the head of the scrubbing column 8 into the scrubbing stage. 
As can be seen from Table 1, a C.sub.2 -enriched gas is obtained by the use 
of the process of this invention, whith about 60% C.sub.2 -hydrocarbons, 
and a C.sub.3 -enriched gas is furthermore obtained with about 60% C.sub.3 
-hydrocarbons. The high olefin proportion in these fractions can be 
separated in the usual way from the remaining components in 
low-temperature separation plants, and can be obtained as the desired 
product of the process. 
In the embodiment illustrated in FIG. 2, the starting material is the same 
amount of a gaseous stream having the same composition as in the 
aforedescribed embodiment. The waste gas of the Fischer-Tropsch synthesis, 
from which carbon dioxide has again been removed in a preceding process 
stage, not illustrated, is conducted via conduit 36 into a first of three 
parallel-connected regnerators 37, 38 and 39. The regenerators are in 
communication with one another via a conduit system, not shown in the 
figure, in such a way that by operating various valves a cyclic 
interchangeability of the operating conditions of the individual 
regenerators is made possible. 
The feedstream fed via conduit 36 is under a pressure of about 10 bar and 
is introduced into the regenerator 37, the packing of which had been 
cooled in a preceding process stage by a specific temperature difference. 
During cooling of the gas to about -150.degree. C., all C.sub.2 - and 
higher hydrocarbons are condensed and settle on the regenerator packing. 
The gas exiting from the cold regenerator end via conduit 40 and having 
the composition set forth in Table 2, is conducted through heat exchanger 
41 through conduit 42, and through heat exchanger 43, the gas being cooled 
by about a further 3.degree. C. during this procedure. The gas is then 
passed via conduit 44 into a phase separator 45 wherein additional 
condensate, formed during the cooling step, is separated. This condensate 
is withdrawn via the conduit 47, equipped with a valve 46, and after being 
warmed in heat exchangers 43 and 41, is fed into regenerator 38. The 
gaseous phase from separator 45 is introduced via conduit 48 to the heat 
exchanger 41 and thereafter subdivided into two partial streams. The 
largest part of this gaseous stream is introduced via conduit 49 into 
regenerator 39 wherein it is warmed and then leaves the system discussed 
herein via conduit 50 at a pressure of about 5 bar. A smaller portion of 
the gas warmed in heat exchanger 41 from conduit 48 is conducted via 
conduit 51 to an expansion turbine 52 and is work expanded therein with 
production of refrigeration. The thus-cooled, partially expanded gas 
passes via conduit 53 again into heat exchanger 41 and transfers its cold 
content to the gaseous stream in conduit 40. After reheating, further 
expansion is effected in expansion turbine 54 to obtain additional cold 
for covering the refrigeration demand of the process. The gas, expanded to 
about 1.1 bar, is conducted via conduit 55 to conduit 47 and is combined 
at 56 with the expanded condensate from separator 45. After being reheated 
in heat exchanger 41, this stream is conducted through regenerator 38 and 
serves to absorb the condensates precipitated in a preceding cycle. The 
gas withdrawn via conduit 57 during such a purge cycle is the fraction 
enriched with olefins which, after compression in a compressor 58, is 
conducted via conduit 59 to a low-temperature separation system. The 
composition of this fraction can be seen from Table 2. 
The aforedescribed process with three regenerators is especially 
advantageous with an intake pressure of the compressor 58 of 1.1 bar, 
since the required amount of purge gas corresponds precisely to the 
quantity of gas which must be expanded in turbines 52, 54 for covering the 
refrigeration losses. 
In a modification of the described regenerator process, the recovery of the 
olefin-enriched fraction condensed in a regenerator can be effected by 
purging at a lower pressure, for example at 0.1 bar. Such a modification 
has the advantage that the condensed hydrocarbons are driven out by a 
smaller quantity of scavenging gas and thus are obtained in a more 
concentrated form. Due to the lower requirement of purge gas, though, it 
is impossible to cover the refrigeration losses of such a method by 
expansion of the purge gas. It is therefore advantageous in such a case 
either to resort to external refrigeration or to select a process with 
four regenerators wherein the remainder of the waste gas stream expanded 
to cover the refrigeration requirement is introduced into the fourth 
regenerators. 
The low-temperature separation system not illustrated in detail in the 
drawing is conventional and is described in the following references 
incorporated herein: Information Leaflets "C.sub.2 H.sub.4 -Plant for 
production of ethylene, propylene, acetylene, butadiene, gasoline and 
aromatics" and "C.sub.2 H.sub.4, C.sub.3 H.sub.6 -Plant for the production 
of ethylene and propylene," published 1978 by Linde A. G. Further relevant 
references are German Pat. No. 2509689. U.S. Pat. No. 4,218,229 and 
copending U.S. patent application Ser. No. 082,452, filed Oct. 9, 1979. 
TABLE 1 
__________________________________________________________________________ 
Gas After Gas After Gas After C.sub.3 Enriched 
C.sub.2 Enriched 
CO.sub.2 Removal 
C.sub.4 Scrubbing Step 
C.sub.6 Scrubbing Step 
Gas Gas C.sub.4+ 
__________________________________________________________________________ 
H.sub.2 
56.24 65.31 72.49 2.45 6.75 -- 
N.sub.2 + CO 
9.28 10.68 11.56 1.12 4.06 -- 
CH.sub.4 
14.66 16.34 15.81 5.42 25.37 -- 
C.sub.2 H.sub.4 
3.81 3.04 0.01 10.26 34.51 -- 
C.sub.2 H.sub.6 
3.13 2.50 0.01 8.42 28.35 -- 
C.sub.3 
7.14 0.08 -- 59.94 0.96 -- 
C.sub.4 
3.22 2.05 -- 12.39 -- 42.05 
C.sub.5+ 
2.52 -- 0.12 -- -- 57.95 
__________________________________________________________________________ 
(Data in Mol%) 
TABLE 2 
______________________________________ 
Gas After Waste Gas From 
Concentrate 
CO.sub.2 Enrichment From Enrichment 
Removal (Conduit 50) (Conduit 57) 
______________________________________ 
H.sub.2 56.24 73.18 27.09 
N.sub.2 + CO 
9.28 11.72 5.09 
CH.sub.4 
14.66 15.04 14.02 
C.sub.2 H.sub.4 
3.81 0.06 10.27 
C.sub.2 H.sub.6 
3.13 -- 8.48 
C.sub.3 H.sub.6 
5.83 -- 15.88 
C.sub.3 H.sub.8 
1.31 -- 3.56 
C.sub.4 3.22 -- 8.75 
C.sub.5+ 
2.52 -- 6.86 
______________________________________ 
(Data in Mol%) 
The preceding examples can be repeated with similar success by substituting 
the generically or specifically described reactants and/or operating 
conditions of this invention for those used in the preceding examples. 
From the foregoing description, one skilled in the art can easily ascertain 
the essential characteristics of this invention, and without departing 
from the spirit and scope thereof, can make various changes and 
modifications of the invention to adapt it to various usages and 
conditions.