Abstract:
A process for the recovery of propane, propylene and heavier hydrocarbon components from a hydrocarbon gas stream is disclosed. The stream is cooled and/or expanded to partially condense it, then separated to provide a first vapor stream. The first vapor stream is directed into a contacting device whereby vapors and liquids are formed. The liquids are directed to a distillation column operating at lower pressure wherein a second vapor stream is separated to recover a product containing the major portion of the C 3  components and heavier hydrocarbon components. The second vapor stream is directed into heat exchange relation with the vapors to cool the second vapor stream and condense at least a part of it, forming a condensed stream. At least a portion of the condensed stream is directed to the contacting device to intimately contact the first vapor stream; the remaining portion (if any) of the condensed stream can be supplied to the distillation column as its top feed. The quantities and temperatures of the feeds to the contacting device and the distillation column are effective to maintain the overhead temperatures of the contacting device and the distillation column at temperatures whereby the major portion of the desired components is recovered.

Description:
BACKGROUND OF THE INVENTION  
         [0001]    This invention relates to a process for the separation of a gas containing hydrocarbons. The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior United States provisional application Serial No. 60/225,260 which was filed on Aug. 15, 2000.  
           [0002]    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.  
           [0003]    The present invention is generally concerned with the recovery of 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, 92.6% methane, 4.7% ethane and other C 2  components, 1.0% propane and other C 3  components, 0.2% iso-butane, 0.2% normal butane, 0.1% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present.  
           [0004]    The historically cyclic fluctuations in the prices of both natural gas and its natural gas liquid (NGL) constituents have reduced the incremental value of propylene, propane, 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 (ethylene, ethane, and heavier hydrocarbons content) of the gas, and the desired end products, each of these processes or a combination thereof may be employed.  
           [0005]    The cryogenic expansion process is now generally preferred for propylene and propane 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; 4,251,249; 4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063; 4,690,702; 4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,568,737; 5,555,748; 5,771,712; 5,799,507; 5,881,569; 5,890,378; and 6,182,468 BI and reissue U.S. Pat. No. 33,408 describe relevant processes.  
           [0006]    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 3 + 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 (deethanizer) column. In the column, the expansion cooled stream(s) is (are) distilled to separate residual methane, C 2  components, nitrogen, and other volatile gases as overhead vapor from the desired C 3  components and heavier hydrocarbon components as bottom liquid product.  
           [0007]    If the feed gas is not totally condensed (typically it is not), 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 slightly below the pressure at which the distillation column is operated. The expanded stream then enters the lower section of an absorption column and is contacted with cold liquids to absorb the C 3  components and heavier components from the vapor portion of the expanded stream. The liquids from the absorption column are then pumped into the deethanizer column at an upper column feed position, perhaps after heating to partially vaporize the stream.  
           [0008]    The overhead distillation stream from the deethanizer passes in heat exchange relation with the residue gas from the absorber column and is cooled, condensing at least a portion of the distillation stream from the deethanizer. The cooled distillation stream then enters the upper section of the absorption column where the cold liquids contained in the stream can contact the vapor portion of the expanded stream as described earlier. Typically, the vapor portion (if any) of the cooled distillation stream and the absorber overhead vapor combine in an upper separator section in the absorber column as residual methane and C 2  component product gas. Alternatively, the cooled distillation stream may be supplied to a separator to provide vapor and liquid streams. The vapor is combined with the absorber column overhead and the liquid is supplied to the absorber column as a top column feed. It may also be advantageous to supply a portion of the cold liquid condensate to the deethanizer tower to serve as reflux.  
           [0009]    The separation that takes place in this process (producing a residue gas leaving the process which contains substantially all of the methane and C 2  components in the feed gas with essentially none of the C 3  components and heavier hydrocarbon components, and a bottoms fraction leaving the deethanizer which contains substantially all of the C 3  components and heavier hydrocarbon components with essentially no methane, C 2  components, or more volatile components) consumes energy for feed gas cooling, for reboiling the deethanizer, for refluxing the deethanizer, and/or for re-compressing the residue gas. The present invention provides a means for achieving this separation while substantially reducing the utility requirements (cooling, reboiling, refluxing, and/or re-compressing) needed for the recovery of the desired products.  
           [0010]    In accordance with the present invention, it has been found that C 3  recoveries in excess of 99 percent can be maintained while providing essentially complete rejection of C 2  components to the residue gas stream. In addition, the present invention makes possible essentially 100 percent separation of C 2  components and lighter components from the C 3  components and heavier hydrocarbon components at reduced energy requirements. The present invention, although applicable at lower pressures and warmer temperatures, is particularly advantageous when processing feed gases in the range of 400 to 1500 psia [2,758 to 10,342 kPa(a)] or higher under conditions requiring column overhead temperatures of −50° F. [−46° C.] or colder. 
       
    
    
       [0011]    For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings:  
         [0012]    [0012]FIG. 1 is a flow diagram of a prior art cryogenic natural gas processing plant in accordance with U.S. Pat. No. 5,771,712;  
         [0013]    [0013]FIG. 2 is a flow diagram of a prior art cryogenic natural gas processing plant of an alternative system in accordance with U.S. Pat. No. 5,771,712;  
         [0014]    [0014]FIG. 3 is a flow diagram of a natural gas processing plant in accordance with the present invention;  
         [0015]    [0015]FIG. 4 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream;  
         [0016]    [0016]FIG. 5 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream;  
         [0017]    [0017]FIG. 6 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream;  
         [0018]    [0018]FIG. 7 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream;  
         [0019]    [0019]FIG. 8 is a flow diagram illustrating an alternative means of application of the present invention to a natural gas stream;  
         [0020]    [0020]FIG. 9 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream;  
         [0021]    [0021]FIG. 10 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream;  
         [0022]    [0022]FIG. 11 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream;  
         [0023]    [0023]FIG. 12 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream;  
         [0024]    [0024]FIG. 13 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream;  
         [0025]    [0025]FIG. 14 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream;  
         [0026]    [0026]FIG. 15 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream;  
         [0027]    [0027]FIG. 16 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream;  
         [0028]    [0028]FIG. 17 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream; and  
         [0029]    [0029]FIG. 18 is a flow diagram illustrating an alternative means of application of the present invention to a hydrocarbon gas stream. 
     
    
       [0030]    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 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.  
         [0031]    For convenience, process parameters are reported in both the traditional British units and in the units of the International System of Units (SI). The molar flow rates given in the tables may be interpreted as either pound moles per hour or kilogram moles per hour. The energy consumptions reported as horsepower (HP) and thousand British Thermal Units per hour (MBTU/hour) correspond to the stated molar flow rates in pound moles per hour. The energy consumptions reported as kilowatts (kW) correspond to the stated molar flow rates in kilogram moles per hour.  
       DESCRIPTION OF THE PRIOR ART  
       [0032]    Referring now to FIG. 1, in a simulation of prior art according to U.S. Pat. No. 5,771,712, inlet gas enters the plant at 80° F. [27° C.] and 1215 psia [8,377 kPa(a)] 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.  
         [0033]    The feed stream  31  is cooled in exchanger  10  by heat exchange with cool residue gas at −76° F. [−60° C.] (stream  34   a ). (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.). For the conditions stated, the feed stream pressure is above the cricondenbar, so no liquid will condense as the stream is cooled. Hence, the cooled stream  31   a  (a dense-phase fluid at these conditions) is supplied directly to work expansion machine  13  at −14° F. [−26° C.]. (The cricondenbar is the maximum pressure at which a vapor phase can exist in a multi-phase fluid. At pressures below the cricondenbar, a separator or scrubber would typically be used to separate any condensed liquid contained in stream  31   a  from the vapor so that only the vapor is supplied to work expansion machine  13 .)  
         [0034]    The work expansion machine  13  extracts mechanical energy from the high pressure feed by expanding the stream substantially isentropically from a pressure of about 1210 psia [8,343 kPa(a)] to a pressure of about 435 psia [2,999 kPa(a)] (the operating pressure of separator/absorber tower  15 ), with the work expansion cooling the expanded stream  31   b  to a temperature of approximately −104° F. [−76° C.]. The expanded and partially condensed stream  31   b  is supplied to absorbing section  15   b  in a lower region of separator/absorber tower  15 . The liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section and the combined liquid stream  35  exits the bottom of separator/absorber tower  15  at −106° F. [−77° C.]. The vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C 3  components and heavier components.  
         [0035]    The separator/absorber tower  15  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 separator/absorber tower may consist of two sections. The upper section  15   a  is a separator wherein any vapor contained in the top feed is separated from its corresponding liquid portion, and wherein the vapor rising from the lower distillation or absorbing section  15   b  is combined with the vapor portion (if any) of the top feed to form the cold distillation stream  34  which exits the top of the tower. The lower, absorbing section  15   b  contains the trays and/or packing and provides the necessary contact between the liquids falling downward and the vapors rising upward to condense and absorb the C 3  components and heavier components.  
         [0036]    The combined liquid stream  35  from the bottom of the separator/absorber tower  15  is supplied to deethanizer  17  by pump  16 , entering at a mid-column feed point at −105° F. [−76° C.] as stream  35   a  to be stripped of its methane and C 2  components. The deethanizer in tower  17 , operating at about 450 psia [3,103 kPa(a)], is also a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The deethanizer tower may also consist of two sections: an upper section  17   a  wherein any vapor contained in the top feed is separated from its corresponding liquid portion, and wherein the vapor rising from the lower distillation or deethanizing section  17   b  is combined with the vapor portion (if any) of the top feed to form distillation stream  36  which exits the top of the tower; and a lower, deethanizing section  17   b  that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The deethanizing section  17   b  also includes a reboiler  18  which heats and vaporizes a portion of the liquid at the bottom of the column to provide the stripping vapors which flow up the column to strip the liquid product, stream  37 , of methane and C 2  components. A typical specification for the bottom liquid product is to have an ethane to propane ratio of 0.02:1 on a molar basis. The liquid product stream  37  exits the bottom of the deethanizer at 207° F. [97° C.] and is cooled to 110° F. [43° C.] (stream  37   a  ) in heat exchanger  19  before flowing to storage.  
         [0037]    The operating pressure in deethanizer  17  is maintained slightly above the operating pressure of separator/absorber tower  15 . This allows the deethanizer overhead vapor (stream  36 ) to pressure flow through heat exchanger  20  and thence into the upper section of separator/absorber tower  15 . In heat exchanger  20 , the deethanizer overhead at −36° F. [−38° C.] is directed in heat exchange relation with the overhead (stream  34 ) from separator/absorber tower  15 , cooling the stream to −107° F. [−77° C.] (stream  36   a ) and partially condensing it. The partially condensed stream is then supplied to the separator section in separator/absorber tower  15  where the condensed liquid is separated from the uncondensed vapor. The uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream  34  leaving the upper region of separator/absorber tower  15 . The condensed liquid is divided into two portions. One portion, stream  39 , is routed to the lower absorbing section of separator/absorber tower  15  as the cold liquid that contacts the vapors rising upward through the absorbing section. The other portion, stream  38 , is supplied to deethanizer  17  as reflux by pump  21 , with reflux stream  38   a  flowing to a top feed point on deethanizer  17  at −107° F. [−77° C.].  
         [0038]    The distillation stream leaving the top of separator/absorber tower  15  at −112° F. [−80° C.] is the cold residue gas stream  34 . The residue gas stream passes countercurrently to deethanizer overhead stream  36  in heat exchanger  20  and is warmed to −76° F. [−60° C.] (stream  34   a ) as it provides cooling and partial condensation of the deethanizer overhead stream. The residue gas is further warmed to 54° F. [12° C.] (stream  34   b ) as it passes countercurrently to the incoming feed stream in heat exchanger  10 . The residue gas is then re-compressed in two stages. The first stage is compressor  14  driven by expansion machine  13 . The second stage is compressor  22  driven by a supplemental power source which compresses the residue gas (stream  34   d ) to sales line pressure. After cooling in discharge cooler  23 , the residue gas product (stream  34   e ) flows to the sales gas pipeline at 110° F. [43° C.] and 1215 psia [8,377 kPa(a)].  
         [0039]    A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 1 is set forth in the following table:  
                                                           TABLE I                           (FIG. 1)       Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]            Stream   Methane   Ethane   Propane   Butanes+   Total               31   81,340   4,128   878   439   87,840       35    5,023   2,108   880   439    8,558       36    7,473   3,432    6    0   11,080       39    3,674   2,011    4    0    5,782       38    2,449   1,341    2    0    3,854       34   81,340   4,111    3    0   86,508       37      0     18   876   439    1,333            Recoveries*                   Propane    99.70%       Butanes+   100.00%       Power       Residue Gas Compression   37,593 HP   [61,802 kW]       Utility Heat       Deethanizer Reboiler   53,478 MBTU/Hr   [34,544 kW]                          
 
         [0040]    In the prior art illustrated in FIG. 1, the refrigeration generated by work expansion machine  13  is not used efficiently in the process. This is evidenced by the relatively cool residue gas (stream  34   b ) temperature of 54° F. [12° C.] leaving heat exchanger  10  (compared to the temperature of the inlet stream  31 ), and by the fact that the process cooling available in the cold liquid (stream  35 ) leaving the bottom of separator/absorber  15  is not needed to provide a portion of the inlet gas cooling in heat exchanger  10 . Ordinarily, this would indicate that less expansion ratio is needed across work expansion machine  13  to maintain the desired C 3  component recovery efficiency, and that the operating pressure of separator/absorber  15  could be raised to reduce the external power requirements in compressor  22 .  
         [0041]    However, since the operating pressure of deethanizer  17  must of necessity be maintained somewhat higher than that of separator/absorber  15  so that its overhead stream  36  can pressure flow through heat exchanger  20  and into the separator section of separator/absorber  15 , reducing the expansion ratio across work expansion machine  13  also means raising the operating pressure of deethanizer  17 . Unfortunately, this is not advisable in this instance because of the detrimental effect on distillation performance in deethanizer  17  that would result from the higher operating pressure. This effect is manifested by poor mass transfer in deethanizer  17  due to the phase behavior of its vapor and liquid streams. Of particular concern are the physical properties that affect the vapor liquid separation efficiency, namely the liquid surface tension and the differential in the densities of the two phases. As a result, the operating pressure of deethanizer  17  should not be raised above the value shown in FIG. 1, so there is no means available to reduce the power consumption of compressor  22  using the prior art process.  
         [0042]    [0042]FIG. 2 represents an alternative application of the prior art process in accordance with U.S. Pat. No. 5,771,712. The process of FIG. 2 has been applied to the same feed gas composition as described above for FIG. 1, but in this simulation of the process the inlet gas (stream  31 ) enters the plant at 80° F. [27° C.] and 580 psia [3,999 kPa(a)]. The feed stream  31  is cooled in exchanger  10  by heat exchange with cool residue gas at −95° .F [−71° C.] (stream  34   a ), with separator liquids at −92° F. [−69° C.] (stream  33   a ), and with separator/absorber liquids at −107° F. [−77° C] (stream  35   a ). At this operating pressure the feed stream is below the cricondenbar, so the cooled stream  31   a  enters separator  11  at −77° F. [−60° C.] and 570 psia [3,930 kPa(a)] where the vapor (stream  32 ) is separated from the condensed liquid (stream  33 ).  
         [0043]    The vapor (stream  32 ) from separator  11  enters work expansion machine  13  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  13  expands the vapor substantially isentropically from a pressure of about 570 psia [3,930 kPa(a)] to a pressure of about 380 psia [2,620 kPa(a)] (the operating pressure of separator/absorber  15 ), with the work expansion cooling the expanded stream  32   a  to a temperature of approximately −107° F. [−77° C.]. The expanded and partially condensed stream  32   a  enters the lower section of separator/absorber  15 . The liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section and the combined liquid stream  35  exits the bottom of separator/absorber  15  at −108° F. [−78° C.]. The vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C 3  components and heavier components.  
         [0044]    The combined liquid stream  35  from the bottom of the separator/absorber  15  is routed to heat exchanger  10  by pump  16  where it (stream  35   a ) is heated as it provides cooling of the incoming feed gas as described earlier. The combined liquid stream is heated to −85° F. [−65° C.], partially vaporizing stream  35   b  before it is supplied as a mid-column feed to deethanizer  17 . The separator liquid (stream  33 ) is flash expanded to slightly above the 395 psia [2,723 kPa(a)] operating pressure of deethanizer  17  by expansion valve  12 , cooling stream  33  to −92° F. [−69° C.] (stream  33   a ) before it provides cooling to the incoming feed gas as described earlier. Stream  33   b , now at 65° F. [18° C.], then enters deethanizer  17  at a lower mid-column feed point. In the deethanizer, streams  35   b  and  33   b  are stripped of their methane and C 2  components. The resulting liquid product stream  37  exits the bottom of the deethanizer at 195° F. [91° C.] and is cooled to 110° F. [43° C.] (stream  37   a ) in heat exchanger  19  before flowing to storage.  
         [0045]    The operating pressure in deethanizer  17  is maintained slightly above the operating pressure of separator/absorber  15 . This allows the deethanizer overhead vapor (stream  36 ) to pressure flow through heat exchanger  20  and thence into the upper section of separator/absorber  15 . In heat exchanger  20 , the deethanizer overhead at −29° F. [−34° C.] is directed in heat exchange relation with the overhead (stream  34 ) from separator/absorber  15 , cooling the stream to −108° F. [−78° C.] (stream  36   a ) and partially condensing it. The partially condensed stream is then supplied to the separator section in separator/absorber tower  15  where the condensed liquid is separated from the uncondensed vapor. The uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream  34  leaving the upper region of separator/absorber  15 . The condensed liquid is divided into two portions. One portion, stream  39 , is routed to the lower absorbing section of separator/absorber  15  as the cold liquid that contacts the vapors rising upward through the absorbing section. The other portion, stream  38 , is supplied to deethanizer  17  as reflux by pump  21 , with reflux stream  38   a  flowing to a top feed point on deethanizer  17  at −108° F. [−78° C.].  
         [0046]    The distillation stream leaving the top of separator/absorber  15  at −113° F. [−81° C.] is the cold residue gas stream  34 . The residue gas stream passes countercurrently to deethanizer overhead stream  36  in heat exchanger  20  and is warmed to −95° F. [−71° C.] (stream  34   a ) as it provides cooling and partial condensation of the deethanizer overhead stream. The residue gas is further warmed to 75° F. [24° C.] (stream  34   b ) as it passes countercurrently to the incoming feed gas in heat exchanger  10 . The residue gas is then re-compressed in two stages. The first stage is compressor  14  driven by expansion machine  13 . The second stage is compressor  22  driven by a supplemental power source which compresses the residue gas (stream  34   d ) to sales line pressure. After cooling in discharge cooler  23 , the residue gas product (stream  34   e ) flows to the sales gas pipeline at 110° F. [43° C.] and 613 psia [4,226 kPa(a)].  
         [0047]    A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 2 is set forth in the table below:  
                                                           TABLE II                           (FIG. 2)       Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]            Stream   Methane   Ethane   Propane   Butanes+   Total               31   81,340   4,128   878   439   87,840       32   80,218   3,702   573   126   85,651       33    1,122     427   306   313    2,189       35    1,876     964   547   126    3,556       36    3,451   1,700    76    0    5,306       39    1,816   1,300    61    0    3,229       38     454     325    15    0     807       34   81,340   4,113    87    0   86,594       37      0     16   791   439    1,246            Recoveries*                   Propane   90.09%       Butanes+   99.99%       Power       Residue Gas Compression   18,911 HP   [31,089 kW]       Utility Heat       Deethanizer Reboiler   17,844 MBTU/Hr   [11,526 kW]                          
 
         [0048]    In the prior art illustrated in FIG. 2, the much lower feed gas pressure results in much less refrigeration (compared to FIG. 1) being generated by work expansion machine  13 . Consequently, much better process heat integration is required to achieve the desired C 3  component recovery efficiency than was the case for the FIG. 1 processing conditions. The process cooling available from the residue gas (stream  34   a ) in heat exchanger  10  must be supplemented by the process cooling available in the cold separator liquids (stream  33 ) and the cold liquid (stream  35 ) leaving the bottom of separator/absorber  15  in order to accomplish the necessary inlet gas cooling. In fact, the process heat integration is so complete that the only means available to increase the C 3  component recovery with the prior art process for these processing conditions would be to increase the expansion ratio across work expansion machine  13  (by reducing the operating pressure of separator/absorber  15 ) to increase the refrigeration generated by the machine, and/or to add external refrigeration. Of course, this would have the undesired consequence of increasing the external power requirements, either in compressor  22 , in the external refrigeration system, or both.  
       DESCRIPTION OF THE INVENTION  
     EXAMPLE 1  
       [0049]    [0049]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.  
         [0050]    In the simulation of the FIG. 3 process, inlet gas enters the plant at 80° F. [27° C.] and 1215 psia [8,377 kPa(a)] as stream  31 . The feed stream  31  is cooled in exchanger  10  by heat exchange with cool residue gas at −56° F. [−49° C.] (stream  34   a ) and with separator/absorber liquids at −113° F. [−80° C.] (stream  35   a ). The cooled stream  31   a  (a dense-phase fluid at these conditions) is supplied directly to work expansion machine  13  at −35° F. [−37° C.].  
         [0051]    The work expansion machine  13  extracts mechanical energy from the high pressure feed by expanding the stream substantially isentropically from a pressure of about 1210 psia [8,343 kPa(a)] to a pressure of about 575 psia [3,964 kPa(a)] (the operating pressure of separator/absorber tower  15 ), with the work expansion cooling the expanded stream  31   b  to a temperature of approximately −98° F. [−72° C.]. The expanded and partially condensed stream  31   b  enters the lower section of separator/absorber  15 . The liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section and the combined liquid stream  35  exits the bottom of separator/absorber  15  at −100° F. [−73° C.]. The vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C 3  components and heavier components.  
         [0052]    Unlike the prior art process illustrated in FIG. 1, in the present invention the operating pressure of deethanizer  17  is maintained below (not above) the operating pressure of separator/absorber  15 . Consequently, a pump is not required for the combined liquid stream  35  from the bottom of the separator/absorber  15 . Instead, the stream is flash expanded to slightly above the 450 psia [3,103 kPa(a)] operating pressure of deethanizer  17  by expansion valve  27 , cooling stream  35  to −113° F. [−80° C.] (stream  35   a ) before it provides cooling to the incoming feed gas as describe earlier. Stream  35   b , now at −73° F. [−58° C.], then enters deethanizer  17  at a mid-column feed point. In the deethanizer, stream  35   b  is stripped of its methane and C 2  components. The resulting liquid product stream  37  exits the bottom of the deethanizer at 207° F. [97° C.] and is cooled to 110° F. [43° C.] (stream  37   a ) in heat exchanger  19  before flowing to storage.  
         [0053]    The deethanizer overhead vapor (stream  36 ) exits deethanizer  17  at −56° F. [−49° C.] and is warmed to 105° F. [41° C.] (stream  36   a ) in heat exchanger  24  before entering compressor  25  (driven by a supplemental power source). Stream  36   b  leaves compressor  25  at 600 psia [4,137 kPa(a)] and is cooled to 110° F. [43° C.] (stream  36   c ) in heat exchanger  26 . Stream  36   c  is then directed in exchange relation with the deethanizer overhead vapor (stream  36 ) in heat exchanger  24  to cool it (stream  36   d ) and conserve process cooling.  
         [0054]    With the increase in pressure provided by compressor  25 , stream  36   d  can now pressure flow through heat exchanger  20  and thence to the upper feed point of separator/absorber  15 . In heat exchanger  20 , the compressed deethanizer overhead at −41° F. [−40° C.] is directed in heat exchange relation with the overhead (stream  34 ) from separator/absorber  15 , cooling the stream to −98° F. [−72° C.] (stream  36   e ) and partially condensing it. The partially condensed stream is then supplied to the separator section in separator/absorber tower  15  where the condensed liquid is separated from the uncondensed vapor. The uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream  34  leaving the upper region of separator/absorber  15 . The condensed liquid is divided into two portions. One portion, stream  39 , is routed to the lower absorbing section of separator/absorber  15  as the cold liquid that contacts the vapors rising upward through the absorbing section. The other portion, stream  38 , is flash expanded to slightly above the operating pressure of deethanizer  17  by expansion valve  28  and the resulting stream  38   a  is then supplied at −112° F. [−80° C.] to the separator section in deethanizer  17  where its condensed liquid is separated from its uncondensed vapor. The uncondensed vapor combines with the vapor rising from the lower distillation section to form the deethanizer overhead stream  36  leaving the upper region of deethanizer  17 , while the condensed liquid is routed to the lower distillation section of deethanizer  17  as reflux for the vapors rising upward through the distillation section.  
         [0055]    The distillation stream leaving the top of separator/absorber  15  at −103° F. [−75° C.] is the cold residue gas stream  34 . The residue gas stream passes countercurrently to compressed deethanizer overhead stream  36   d  in heat exchanger  20  and is warmed to −56° F. [−49° C.] (stream  34   a ) as it provides cooling and partial condensation of the compressed deethanizer overhead stream. The residue gas is further warmed to 75° F. [24° C.] (stream  34   b ) as it passes countercurrently to the incoming feed gas in heat exchanger  10 . The residue gas is then re-compressed in two stages. The first stage is compressor  14  driven by expansion machine  13 . The second stage is compressor  22  driven by a supplemental power source which compresses the residue gas (stream  34   d ) to sales line pressure. After cooling in discharge cooler  23 , the residue gas product (stream  34   e ) flows to the sales gas pipeline at 110° F. [43° C.] and 1215 psia [8,377 kPa(a)].  
         [0056]    A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 3 is set forth in the table below:  
                                                           TABLE III                           (FIG. 3)       Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]            Stream   Methane   Ethane   Propane   Butanes+   Total               31   81,340   4,128   878   439   87,840       35   11,825   2,920   879   439   16,261       36   17,501   4,827    6    0   22,642       39    7,839   2,658    3    0   10,653       38    5,676   1,925    2    0    7,714       34   81,340   4,111    3    0   86,506       37      0     18   876   439    1,333            Recoveries*                   Propane    99.70%       Butanes+   100.00%       Power       Residue Gas Compression   28,422 HP   [46,725 kW]       Overhead Vapor Compression    3,810 HP    [6,264 kW]       Total Compression   32,232 HP   [52,989 kW]       Utility Heat       Deethanizer Reboiler   51,073 MBTU/Hr   [32,990 kW]                          
 
         [0057]    Comparison of the utility consumptions of the prior art process displayed in Table I with the utility consumptions of the present invention displayed in Table III shows that the present invention maintains the desired C 3  component recovery while reducing the utility heat requirement and substantially reducing the compression horsepower. The utility heat requirement is more than four percent lower than the prior art process, while the compression horsepower is more than fourteen percent lower than the prior art process.  
         [0058]    Comparing the present invention to the prior art process displayed in FIG. 1, note that while the operating pressure of deethanizer  17  is the same in both cases, the operating pressure of separator/absorber  15  in the present invention is significantly higher than in the FIG. 1 process, 575 psia [3,964 kPa(a)] versus  435  psia [2,999 kPa(a)]. Accordingly, the residue gas enters compressor  14  at a higher pressure in the FIG. 3 process and less compression horsepower is therefore needed to deliver the residue gas to pipeline pressure. Further, with separator/absorber  15  operating at a higher pressure than deethanizer  17 , it is no longer necessary to pump the absorber bottom liquid (stream  35 ) and the reflux stream (stream  38 ) to feed deethanizer  17 , eliminating the capital and operating cost of pumps  16  and  21  in the FIG. 1 process.  
         [0059]    In essence, work expansion machine  13  and compressors  14  and  22  represent an open cycle mechanical-compression refrigeration loop that provides the process cooling in the prior art process of FIG. 1, with a working fluid (streams  31  and  34 ) that is predominantly methane. In the present invention illustrated in FIG. 3, the refrigeration provided by this cycle has been reduced by the addition of a second open cycle refrigeration loop powered by compressor  25 . Examination of Table III shows that the working fluid for this second cycle (the deethanizer overhead, stream  36 ) has a substantially lower concentration of methane and a substantially higher concentration of ethane than the working fluid in the first cycle. In general, the efficiency of mechanical-compression refrigeration cycles improves as the molecular weight of the working fluid increases. This effect, together with the much lower flow rate of stream  36  compared to streams  31 / 34  and the lower compression ratio needed from compressor  25  compared to compressors  14 / 22 , accounts for most of the improvement in energy efficiency of the present invention relative to the prior art of FIG. 1. As a measure of this increase in efficiency, note that the total reflux generated for the two columns in the present invention shown in FIG. 3 (the sum of streams  38  and  39 ) is nearly twice that for the prior art of FIG. 1, and yet this is accomplished using 14% less power in the mechanical-refrigeration cycles.  
         [0060]    With compressor  25  supplying the motive force to cause the deethanizer overhead (stream  36  in FIG. 3) to flow through heat exchanger  20  and thence to separator/absorber  15 , the operating pressures of separator/absorber  15  and deethanizer  17  are no longer coupled together as they are in the prior art process. Instead, the operating pressures of the two columns can be optimized independently. In the case of deethanizer  17 , the pressure can be selected to insure good distillation characteristics, while for separator/absorber  15  the pressure can be selected to optimize the process cooling versus the residue gas compression requirements.  
       EXAMPLE 2  
       [0061]    [0061]FIG. 3 represents the preferred embodiment of the present invention for the temperature and pressure conditions shown because it typically provides the simplest plant arrangement for a given C 3  component recovery level. A slightly more complex design that maintains the same C 3  component recovery with lower utility consumption can be achieved using another embodiment of the present invention as illustrated in the FIG. 4 process. 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.  
         [0062]    In the simulation of the FIG. 4 process, the feed gas cooling and expansion scheme is much the same as that used in FIG. 3. The difference lies in the manner in which the vapor distillation stream  36  leaving the overhead of deethanizer  17  is used to generate reflux for deethanizer  17  and separator/absorber  15 . Referring to FIG. 4, the deethanizer overhead vapor (stream  36 ) exits deethanizer  17  at −39° F. [−39° C.] and is warmed to 105° F. [41° ] (stream  36   a ) in heat exchanger  24  before entering compressor  25  (driven by a supplemental power source). Stream  36   b  leaves compressor  25  at 600 psia [4,137 kPa(a)] and is cooled to 110° F. [43° C.](stream  36   c ) in heat exchanger  26 . Stream  36   c  is then directed in heat exchange relation with the deethanizer overhead vapor (stream  36 ) in heat exchanger  24  to cool it to −24° F. [31° C.] (stream  36   d ) an conserve process cooling.  
         [0063]    In heat exchanger  20 , the compressed deethanizer overhead (stream  36   d ) is directed in heat exchange relation with the overhead (stream  34 ) from separator/absorber  15 , cooling the stream to −50° F. [−46° C.] (stream  36   e ) and partially condensing it before it is withdrawn. The partially condensed stream  36   e  enters separator  30  where the condensed liquid is separated from the uncondensed vapor. The condensed liquid (stream  38 ) from separator  30  is flash expanded to slightly above the operating pressure of deethanizer  17  by expansion valve  28  (stream  38   a ), which partially vaporizes the stream and cools it further to −63° F. [−53° C.]. It is then supplied to the separator section in deethanizer  17  where the liquid is separated from the flash vapor. The flash vapor combines with the vapor rising from the lower distillation section to form the deethanizer overhead stream  36  leaving the upper region of deethanizer  17 , while the liquid is routed to the lower distillation section of deethanizer  17  as reflux for the vapors rising upward through the distillation section.  
         [0064]    The uncondensed vapor (stream  39 ) from separator  30  is routed back to heat exchanger  20  to also direct it in heat exchange relation with the overhead (stream  34 ) from separator/absorber  15 , cooling the stream to −98° F. [−72° C.] (stream  39   a ) and partially condensing it. The partially condensed stream is then supplied to the separator section in separator/absorber tower  15  where its condensed liquid is separated from its uncondensed vapor. The uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream  34  leaving the upper region of separator/absorber  15 , while the condensed liquid is routed to the lower absorbing section of separator/absorber  15  as the cold liquid that contacts the vapors rising upward through the absorbing section.  
         [0065]    A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 4 is set forth in the table below:  
                                                           TABLE IV                           (FIG. 4)       Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]            Stream   Methane   Ethane   Propane   Butanes+   Total               31   81,340   4,128   878   439   87,840       35   11,237   2,788   878   439   15,530       36   13,548   5,755    10    0   19,583       39   11,237   2,770    2    0   14,198       38    2,311   2,985    8    0    5,385       34   81,340   4,111    3    0   86,507       37      0     18   876   439    1,333            Recoveries*                   Propane    99.70%       Butanes+   100.00%       Power       Residue Gas Compression   28,405 HP   [46,697 kW]       Overhead Vapor Compression    3,246 HP    [5,336 kW]           31,651 HP   [52,034 kW]       Utility Heat       Deethanizer Reboiler   51,255 MBTU/Hr   [33,108 kW]                          
 
         [0066]    Comparison of the utility consumptions of the prior art process displayed in Table I with the utility consumptions of the present invention displayed in Table IV shows that this embodiment of the present invention also maintains the desired C 3  component recovery while reducing the utility heat requirement and substantially reducing the compression horsepower. The utility heat requirement is more than four percent lower than the prior art process, while the compression horsepower is more than fifteen percent lower than the prior art process.  
         [0067]    Comparison of the utility consumptions displayed in Tables III and IV for the FIG. 3 and FIG. 4 processes shows that the FIG. 4 embodiment of the present invention requires slightly less compression horsepower (about 2 percent) than the FIG. 3 embodiment, but uses slightly more utility heat for the deethanizer reboiler (less than 1 percent), with the total utility requirements being about 1 percent lower for the FIG. 4 embodiment. The improvement in efficiency can be understood by comparing the reflux stream for deethanizer  17  (stream  38 ) in the FIG. 4 embodiment of the present invention with the corresponding stream in the FIG. 3 embodiment. Whereas stream  38  in FIG. 3 is predominantly methane, stream  38  in FIG. 4 is predominantly ethane because it is withdrawn after only partial cooling in heat exchanger  20  so that proportionally less of the more volatile methane has been condensed. Not only is ethane a more effective reflux liquid than methane for rectifying the C 3  and heavier components from the vapors rising upward in deethanizer  17  (as reflected by the much lower flow rate of stream  38  in the FIG. 4 embodiment), the deethanizer overhead (stream  36 ) has a lower concentration of methane (because less methane enters deethanizer  17  in the reflux) so that the mechanical-compression refrigeration efficiency of compressor  25  is improved. Although this embodiment of the present invention is more efficient than the FIG. 3 embodiment, the choice of whether to include the additional equipment that the FIG. 4 process requires will generally depend on factors which include plant size and available equipment, as well as the relative costs of compression horsepower and utility heat.  
       EXAMPLE 3  
       [0068]    A third embodiment of the present invention is shown in FIG. 5, wherein a different method of implementing the second mechanical-compression refrigeration cycle is applied to the present invention. The feed gas composition and conditions considered in the process illustrated in FIG. 5 are the same as those in FIGS. 1, 3, and  4 . Accordingly, FIG. 5 can be compared with the FIG. 1 process to illustrate the advantages of the present invention, and can likewise be compared to the embodiments displayed in FIGS. 3 and 4.  
         [0069]    In the simulation of the FIG. 5 process, inlet gas enters the plant at 80° F. [27° C.] and 1215 psia [8,377 kPa(a)] as stream  31 . The feed stream  31  is cooled in exchanger  10  by heat exchange with cool residue gas at −70° F. [−57° C.] (stream  34   a ), with cool vapor at −49° F. [−45° C.] (stream  41   a ), and with separator/absorber liquids at 112° F. [−80° C.] (stream  35   a ). The cooled stre  31   a  (a dense-phase fluid at these conditions) is supplied directly to work expansion machine  13  at −32° F. [−36° C.].  
         [0070]    The work expansion machine  13  extracts mechanical energy from the high pressure feed by expanding the stream substantially isentropically from a pressure of about 1210 psia [8,343 kPa(a)] to a pressure of about 515 psia [3,551 kPa(a)] (the operating pressure of separator/absorber tower  15 ), with the work expansion cooling the expanded stream  31   b  to a temperature of approximately −104° F. [−75° C.]. The expanded and partially condensed stream  31   b  enters the lower section of separator/absorber  15 . The liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section and the combined liquid stream  35  exits the bottom of separator/absorber  15  at −104° F. [−76° C.]. The vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C 3  components and heavier components.  
         [0071]    The combined liquid stream  35  is flash expanded to slightly above the 450 psia [3,103 kPa(a)] operating pressure of deethanizer  17  by expansion valve  27 , cooling stream  35  to −112° F. [−80° C.] (stream  35   a ) before it provides cooling to the incoming feed gas as described earlier. Stream  35   b , now at −92° F. [−69° C.], then enters deethanizer  17  at a mid-column feed point. In the deethanizer, stream  35   b  is stripped of its methane and C 2  components. The resulting liquid product stream  37  exits the bottom of the deethanizer at 207° F. [97° C.] and is cooled to 110° F. [43° C.]  37   a ) in heat exchanger  19  before flowing to storage.  
         [0072]    The deethanizer overhead vapor (stream  36 ) exits deethanizer  17  at −44° F. [−42° C.] flows through heat exchanger  20 . In heat exchanger  20 , the deethanizer overhead is directed in heat exchange relation with the overhead (stream  34 ) from separator/absorber  15  and the uncondensed vapor (stream  41 ) from separator  30 , cooling the stream to −102° F. [−74° C.] (stream  36   a ) and partially condensing it. The partially condensed stream is then supplied to separator  30  where the condensed liquid (stream  40 ) is separated from the uncondensed vapor (stream  41 ).  
         [0073]    In this embodiment of the present invention, the liquid condensed from the deethanizer overhead (stream  40 ) is at a lower pressure than the two columns (separator/absorber  15  and deethanizer  17 ), so it is pumped by pump  21  so that it can be used as reflux. After pumping, stream  40   a  is then divided into two portions. One portion, stream  39 , is supplied by control valve  29  to the separator section in separator/absorber tower  15  at −100° F. [−74° C.] (stream  39   a ) where its liquid is separated from any vapor that forms. (As the stream is at elevated pressure relative to the pressure at which it condensed, it is unlikely that any vapor will form.) Any vapor that may form combines with the vapor rising from the lower absorbing section to form the cold distillation stream  34  leaving the upper region of separator/absorber  15 , while the condensed liquid is routed to the lower absorbing section of separator/absorber  15  as the cold liquid that contacts the vapors rising upward through the absorbing section. The other portion of the pumped liquid (stream  38 ) is flash expanded to slightly above the operating pressure of deethanizer  17  by expansion valve  28  (stream  38   a ). It is then supplied at −101° F. [−74° C.] to the separator section in deethanizer  17  where its liquid is separated from any flash vapor that forms. Any flash vapor combines with the vapor rising from the lower distillation section to form the deethanizer overhead stream  36  leaving the upper region of deethanizer  17 , while the condensed liquid is routed to the lower distillation section of deethanizer  17  as reflux for the vapors rising upward through the distillation section.  
         [0074]    The distillation stream leaving the top of separator/absorber  15  at −107° F. [−77° C.] is the cold absorber overhead stream  34 . The absorber overhead stream passes countercurrently to deethanizer overhead stream  36  in heat exchanger  20  and is warmed to −70° F. [−57° C.] (stream  34   a ) as it provides cooling and partial condensation of the deethanizer overhead stream. The absorber overhead stream is further warmed to 75° F. [24° C.] (stream  34   b ) as it passes countercurrently to the incoming feed gas in heat exchanger  10 . The uncondensed vapor (stream  41 ) leaves separator  30  at −102° F. [−74° C.] and also passes countercurrently to deethanizer overhead stream  36  in heat exchanger  20  and is warmed to −49° F. [−45° C.] (stream  41   a ) as it too provides cooling and partial condensation of the deethanizer overhead stream. The vapor stream is further warmed to 65° F. [18° C.] (stream  41   b ) as it passes countercurrently to the incoming feed gas in heat exchanger  10 .  
         [0075]    The warm absorber overhead stream  34   b  and the warm vapor stream  41   b  are then re-compressed in two stages. The first stage for the absorber overhead stream is compressor  14  driven by expansion machine  13 , while the first stage for the vapor stream is compressor  25  driven by a supplemental power source. The two partially compressed streams (streams  34   c  and  41   c , respectively) combine to form the residue gas, stream  42 . The combined residue gas stream then enters compressor  22  driven by a supplemental power source, which provides the second stage of compression to raise the residue gas (stream  42   a ) to sales line pressure. After cooling in discharge cooler  23 , the residue gas product (stream  42   b ) flows to the sales gas pipeline at 110° F. [43° C.] and 1215 psia [8,377 kPa(a)].  
         [0076]    A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 5 is set forth in the table below:  
                                                           TABLE V                           (FIG. 5)       Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]            Stream   Methane   Ethane   Propane   Butanes+   Total               31   81,340   4,128   878   439   87,840       35    9,637   2,685   879   439   13,813       36   12,609   4,631    6    0   17,508       39    3,352   2,213    3    0    5,669       38    2,972   1,963    3    0    5,027       41    6,285     454    0    0    6,811       34   75,055   3,657    3    0   79,696       42   81,340   4,111    3    0   86,507       37      0     18   876   439    1,332            Recoveries*                   Propane    99.70%       Butanes+   100.00%       Power       Residue Gas Compression   32,712 HP   [53,778 kW]       Vapor Compression    1,413 HP    [2,323 kW]           34,125 HP   [56,101 kW]       Utility Heat       Deethanizer Reboiler   56,696 MBTU/Hr   [36,623 kW]                          
 
         [0077]    Comparison of the utility consumptions of the prior art process displayed in Table I with the utility consumptions of the present invention displayed in Table V shows that this embodiment of the present invention also maintains the desired C 3  component recovery while substantially reducing the compression horsepower. Although the utility heat requirement is about six percent higher than the prior art process, the compression horsepower is more than nine percent lower than the prior art process, so the total utility requirements is about four percent lower than the prior art.  
         [0078]    Comparison of the utility consumptions displayed in Tables III, IV, and V for the FIG. 3, FIG. 4, and FIG. 5 embodiments of the present invention shows that the FIG. 5 embodiment requires slightly more compression horsepower and utility heating than either the FIG. 3 or the FIG. 4 embodiment. However, if multiple stage compression or multi-wheel centrifugal compression is used to compress the residue gas stream  42 , it may be possible to compress the vapor stream  41   b  using an intermediate stage or wheel, eliminating the need for a separate compressor like compressor  25 . Thus, factors such as plant size and available equipment will determine whether the FIG. 5 embodiment would be preferable for a specific circumstance.  
       EXAMPLE 4  
       [0079]    A slightly more complex design than the FIG. 5 embodiment that maintains the same C 3  component recovery with lower utility consumption can be achieved using another embodiment of the present invention as illustrated in the FIG. 6 process. The feed gas composition and conditions considered in the process presented in FIG. 6 are the same as those in FIGS. 1 and 5. Accordingly, FIG. 6 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. 5.  
         [0080]    In the simulation of the FIG. 6 process, the feed gas cooling and expansion scheme is much the same as that used in FIG. 5. The difference lies in the manner in which the vapor distillation stream  36  leaving the overhead of deethanizer  17  is used to generate reflux for deethanizer  17  and separator/absorber  15 . Referring to FIG. 6, the deethanizer overhead vapor (stream  36 ) exits deethanizer  17  at −39° F. [−40° C.] and flows through heat exchanger  20 . In heat exchanger  20 , the deethanizer overhead is directed in heat exchange relation with the overhead (stream  34 ) from separator/absorber  15  and the uncondensed vapor (stream  41 ) from separator  30 , cooling the stream to −60° F. [−51° C.] (stream  36   a ) and partially condensing it before it is withdrawn. The partially condensed stream  36   a  enters separator  29  where the condensed liquid (stream  38 ) is separated from the uncondensed vapor (stream  40 ).  
         [0081]    In this embodiment of the present invention, the liquid condensed from the deethanizer overhead (stream  38 ) is at a lower pressure than deethanizer  17 , so it is pumped by pump  28  so that it can be used as reflux. After pumping, stream  38   a  is supplied at −60° F. [−51° C.] to the separator section in deethanizer  17  where the liquid is routed to the lower distillation section of deethanizer  17  as reflux for the vapors rising upward through the distillation section.  
         [0082]    The uncondensed vapor (stream  40 ) from separator  29  is routed back to heat exchanger  20  to also direct it in heat exchange relation with the overhead (stream  34 ) from separator/absorber  15  and the uncondensed vapor (stream  41 ) from separator  30 , cooling the stream to −102° F. [−74° C.] (stream  40   a ) and partially condensing it. The partially condensed stream is then supplied to separator  30  where the condensed liquid (stream  39 ) is separated from the uncondensed vapor (stream  41 ). Since the operating pressure of separator  30  is lower than the operating pressure of separator/absorber  15 , pump  21  is used to direct the condensed liquid (stream  39   a ) at −100° F. [−73° C.] to the separator section in separator/absorber tower  15 , where the condensed liquid is routed to the lower absorbing section of separator/absorber  15  as the cold liquid that contacts the vapors rising upward through the absorbing section.  
         [0083]    The distillation stream leaving the top of separator/absorber  15  at −107° F. [−77° C.] is the cold absorber overhead stream  34 . The absorber overhead stream passes countercurrently to deethanizer overhead stream  36  and vapor stream  40  in heat exchanger  20  and is warmed to −74° F. [−59° C.] (stream  34   a ) as it provides cooling and partial condensation of the deethanizer overhead stream and the vapor stream. The absorber overhead stream is further warmed to 75° F. [24° C.] (stream  34   b ) as it passes countercurrently to the incoming feed gas in heat exchanger  10 . The uncondensed vapor (stream  41 ) leaves separator  30  at −102° F. [−74° C.] and also passes countercurrently to deethanizer overhead stream  36  and vapor stream  40  in heat exchanger  20  and is warmed to −44° F. [−42° C.] (stream  41   a ) as it too provides cooling and partial condensation of the streams. The vapor stream is further warmed to 65° F. [18° C.] (stream  41   b ) as it passes countercurrently to the incoming feed gas in heat exchanger  10 .  
         [0084]    The warm absorber overhead stream  34   b  and the warm vapor stream  41   b  are then re-compressed in two stages. The first stage for the absorber overhead stream is compressor  14  driven by expansion machine  13 , while the first stage for the vapor stream is compressor  25  driven by a supplemental power source. The two partially compressed streams (streams  34   c  and  41   c , respectively) combine to form the residue gas, stream  42 . The combined residue gas stream then enters compressor  22  driven by a supplemental power source, which provides the second stage of compression to raise the residue gas (stream  42   a ) to sales line pressure. After cooling in discharge cooler  23 , the residue gas product (stream  42   b ) flows to the sales gas pipeline at 110° F. [43° C.] and 1215 psia [8,377 kPa(a)].  
         [0085]    A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 6 is set forth in the table below:  
                                                           TABLE VI                           (FIG. 6)       Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]            Stream   Methane   Ethane   Propane   Butanes+   Total               31   81,340   4,128   878   439   87,840       35   11,196   2,812   879   439   15,518       36   12,511   5,182    13    0   17,956       40   11,195   2,795    3    0   14,185       38    1,316   2,388    10    0    3,771       41    7,955     585    0    0    8,632       39    3,240   2,210    3    0    5,553       34   73,384   3,526    2    0   77,875       42   81,340   4,111    3    0   86,507       37      0     18   876   439    1,333            Recoveries*                   Propane    99.71%       Butanes+   100.00%       Power       Residue Gas Compression   31,592 HP   [51,937 kW]       Vapor Compression    1,940 HP    [3,189 kW]           33,532 HP   [55,126 kw]       Utility Heat       Deethanizer Reboiler   54,144 MBTU/Hr   [34,974 kW]                          
 
         [0086]    Comparison of the utility consumptions of the prior art process displayed in Table I with the utility consumptions of the present invention displayed in Table VI shows that this embodiment of the present invention also maintains the desired C 3  component recovery while substantially reducing the compression horsepower. Although the utility heat requirement is about one percent higher than the prior art process, the compression horsepower is more than ten percent lower than the prior art process, so the total utility requirements is about six percent lower than the prior art.  
         [0087]    Comparison of the utility consumptions displayed in Tables V and VI for the FIG. 5 and FIG. 6 processes shows that the FIG. 6 embodiment of the present invention requires slightly less compression horsepower (about 2 percent) than the FIG. 5 embodiment, and uses slightly less utility heat for the deethanizer reboiler (about 4 percent), with the total utility requirements being about 3 percent lower for the FIG. 6 embodiment. The improvement in efficiency can be understood by comparing the reflux stream for deethanizer  17  (stream  38 ) in the FIG. 6 embodiment of the present invention with the corresponding stream in the FIG. 5 embodiment. Whereas stream  38  in FIG. 5 is predominantly methane, stream  38  in FIG. 6 is predominantly ethane because it is withdrawn after only partial cooling in heat exchanger  20  so that proportionally less of the more volatile methane has been condensed. Not only is ethane a more effective reflux liquid than methane for rectifying the C 3  and heavier components from the vapors rising upward in deethanizer  17  (as reflected by the much lower flow rate of stream  38  in the FIG. 6 embodiment), the deethanizer overhead (stream  36 ) has a lower concentration of methane (because less methane enters deethanizer  17  in the reflux) so that the mechanical-compression refrigeration efficiency of compressor  25  is improved. Although this embodiment of the present invention is more efficient than the FIG. 5 embodiment, the choice of whether to include the additional equipment that the FIG. 6 process requires will generally depend on factors which include plant size and available equipment, as well as the relative costs of compression horsepower and utility heat.  
       EXAMPLE 5  
       [0088]    [0088]FIG. 7 illustrates a flow diagram of a process in accordance with the present invention when applied to the feed gas composition and conditions considered in the process presented in FIG. 2. Accordingly, the FIG. 7 process can be compared with that of the FIG. 2 process to illustrate the advantages of the present invention.  
         [0089]    In the simulation of the FIG. 7 process, inlet gas enters the plant at 80° F. [27° C.] and 580 psia [3,999 kPa(a)] as stream  31 . The feed stream  31  is cooled in exchanger  10  by heat exchange with cool residue gas at −93° F. [−70° C.] (stream  34   a ), with separator liquids at −110° F. [−79° C.] (stream  33   a ), and with separator/absorber liquids at −121° F. [−85° C.] (stream operating pressure the feed stream is below the cricondenbar, so the cooled stream  31   a  enters separator  11  at −80° F. [−62° C.] and 570 psia [3,930 kPa(a)] where the vapor (stream  32 ) is separated from the condensed liquid (stream  33 ).  
         [0090]    The vapor (stream  32 ) from separator  11  enters work expansion machine  13  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  13  expands the vapor substantially isentropically from a pressure of about 570 psia [3,930 kPa(a)] to a pressure of about 410 psia [2,827 kPa(a)] (the operating pressure of separator/absorber  15 ), with the work expansion cooling the expanded stream  32   a  to a temperature of approximately −104° F. [−76° C.]. The expanded and partially condensed stream  32   a  enters the lower section of separator/absorber  15 . The liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section and the combined liquid stream  35  exits the bottom of separator/absorber  15  at −106° F. [−76° C.]. The vapor portion of the expanded stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C 3  components and heavier components.  
         [0091]    In the present invention, separator/absorber  15  operates at a higher pressure than deethanizer  17 , so the combined liquid stream  35  from the bottom of the separator/absorber  15  is flash expanded to slightly above the 290 psia [1,999 kPa(a)] operating pressure of deethanizer  17  by expansion valve  27 , cooling stream  35  to −121° F. [−85° C.] (stream  35   a ) before it provides cool the incoming feed gas as described earlier. The combined liquid stream is heated to −85° F. [−65° C.], partially vaporizing stream  35   b  before it is supplied as a mid-column feed to deethanizer  17 . The separator liquid (stream  33 ) is flash expanded to slightly above the operating pressure of deethanizer  17  by expansion valve  12 , cooling stream  33  to −110° F. [−79° C.] (stream  33   a ) before it provides cooling to the incoming feed gas as described earlier. Stream  33   b , now at 65° F. [18° C.], then enters deethanizer  17  at a lower mid-column feed point. In the deethanizer, streams  35   b  and  33   b  are stripped of their methane and C 2  components. The resulting liquid product stream  37  exits the bottom of the deethanizer at 164° F. [73° C.] and is cooled to 110° F. [43° C.] (stream  37   a ) in heat exchanger  19  before flowing to storage.  
         [0092]    The deethanizer overhead vapor (stream  36 ) exits deethanizer  17  at −47° F. [−44° C.] and is warmed to 105° F. [41° C.] (stream  36   a ) in heat exchanger  24  before entering compressor  25  (driven by a supplemental power source). Stream  36   b  leaves compressor  25  at 435 psia [2,999 kPa(a)] and is cooled to 110° F. [43° C.] (stream  36   c ) in heat exchanger  26 . Stream  36   c  is then directed in heat exchange relation with the deethanizer overhead vapor (stream  36 ) in heat exchanger  24  to cool it (stream  36   d ) and conserve process cooling.  
         [0093]    With the increase in pressure provided by compressor  25 , stream  36   d  can now pressure flow through heat exchanger  20  and thence to the upper feed point of separator/absorber  15 . In heat exchanger  20 , the compressed deethanizer overhead at −31° F. [−35° C.] is directed in heat exchange relation with the overhead (stream  34 ) from separator/absorber  15 , cooling the stream to −106° F. [−77° C.] (stream  36   e ) and partially condensing it. The partially condensed stream is then supplied to the separator section in separator/absorber tower  15  where the condensed liquid is separated from the uncondensed vapor. The uncondensed vapor combines with the vapor rising from the lower absorbing section to form the cold distillation stream  34  leaving the upper region of separator/absorber  15 . The condensed liquid is divided into two portions. One portion, stream  39 , is routed to the lower absorbing section of separator/absorber  15  as the cold liquid that contacts the vapors rising upward through the absorbing section. The other portion, stream  38 , is flash expanded to slightly above the operating pressure of deethanizer  17  by expansion valve  28  (stream  38   a ). It is then supplied at −124° F. [−87° C.] to the separator section in deethanizer  17  where its condensed liquid is separated from its uncondensed vapor. The uncondensed vapor combines with the vapor rising from the lower distillation section to form the deethanizer overhead stream  36  leaving the upper region of deethanizer  17 , while the condensed liquid is routed to the lower distillation section of deethanizer  17  as reflux for the vapors rising upward through the distillation section.  
         [0094]    The distillation stream leaving the top of separator/absorber  15  at −111° F. [−79° C.] is the cold residue gas stream  34 . The residue gas stream passes countercurrently to compressed deethanizer overhead stream  36  in heat exchanger  20  and is warmed to −93° F. [−70° C.] (stream  34   a ) as it provides cooling and partial condensation of the compressed deethanizer overhead stream. The residue gas is further warmed to 74° F. [23° C.] (stream  34   b ) as it passes countercurrently to the incoming feed gas in heat exchanger  10 . The residue gas is then re-compressed in two stages. The first stage is compressor  14  driven by expansion machine  13 . The second stage is compressor  22  driven by a supplemental power source which compresses the residue gas (stream  34   d ) to sales line pressure. After cooling in discharge cooler  23 , the residue gas product (stream  34   e ) flows to the sales gas pipeline at 110° F. [43° C.] and 613 psia [4,226 kPa(a)].  
         [0095]    A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 7 is set forth in the table below:  
                                                           TABLE VII                           (FIG. 7)       Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]            Stream   Methane   Ethane   Propane   Butanes+   Total               31   81,340   4,128   878   439   87,840       32   79,899   3,596   521   106   85,149       33    1,441     533   357   333    2,691       35    1,906     891   490   106    3,435       36    3,860   1,739    69    0    5,753       39    2,051   1,325    55    0    3,487       38     513     331    14    0     872       34   81,340   4,113    87    0   86,594       37      0     16   791   439    1,246            Recoveries*                   Propane    90.05%       Butanes+   100.00%       Power       Residue Gas Compression   15,990 HP   [26,287 kW]       Overhead Vapor Compression    1,414 HP    [2,325 kW]           17,404 HP   [28,612 kW]       Utility Heat       Deethanizer Reboiler   11,515 MBTU/Hr    [7,438 kW]                          
 
         [0096]    Comparison of the utility consumptions of the prior art process displayed in Table II with the utility consumptions of the present invention displayed in Table VII shows that the present invention maintains the desired C 3  component recovery while substantially reducing the utility heat requirement and reducing the compression horsepower. The utility heat requirement is more than thirty-five percent lower than the prior art process, while the compression horsepower is eight percent lower than the prior art process.  
         [0097]    Comparing the present invention to the prior art process displayed in FIG. 2, note that the operating pressure of deethanizer  17  is significantly lower in the present invention than in the FIG. 2 process, 290 psia [1,999 kPa(a)] versus 395 psia [2,723 kPa(a)], and the operating pressure of separator/absorber  15  is significantly higher in the present invention than in the FIG. 2 process, 410 psia [2,827 kPa(a)] versus 380 psia [2,620 kPa(a)]. Accordingly, the residue gas enters compressor  14  at a higher pressure in the FIG. 7 process and less compression horsepower is therefore needed to deliver the residue gas to pipeline pressure. Further, with separator/absorber  15  operating at a higher pressure than deethanizer  17 , it is no longer necessary to pump the absorber bottom liquid (stream  35 ) and the reflux stream (stream  38 ) to feed deethanizer  17 , eliminating the capital and operating cost of pumps  16  and  21  in the FIG. 2 process.  
         [0098]    As described earlier, the deethanizer overhead (stream  36 ) in the FIG. 7 process provides a more efficient working fluid for a mechanical-compression refrigeration cycle than the inlet gas (stream  31 ) and residue gas (stream  34 ) which are predominantly methane, so that the refrigeration provided to the process by the cycle including compressor  25  not only reduces the refrigeration required from the cycle using compressors  14  and  22 , but reduces the total refrigeration energy consumption as well. Also, note that the liquid streams used to provide part of the feed gas cooling, the cold separator liquids (stream  33 ) and the cold liquid (stream  35 ) leaving the bottom of separator/absorber  15 , are cooled by flash expansion (streams  33   a  and  35   a , respectively) before entering heat exchanger  10 . As a result, these streams are considerably colder than the corresponding streams in the FIG. 2 process, allowing better heat integration and more efficient process cooling than the prior art process can provide. This is a consequence of operating deethanizer  17  at a lower pressure than separator/absorber  15  in the present invention, which is not possible with the prior art process.  
         [0099]    The substantial reduction in the utility heat required for deethanizer reboiler  18  for the present invention is a consequence of the lower operating pressure that is possible for deethanizer  17  in the FIG. 7 process. With the pressure lower in deethanizer  17 , the bubble point temperatures of all the liquid streams in the column are lower, including the bottom liquid product (164° F. [73° C.] for stream  37  in the FIG. 7 process, versus 195° F. [91° C.] for stream  37  in the FIG. 2 process). Thus, much less sensible heating is required for the column liquids in deethanizer  17 , reducing the heating load in reboiler  18  accordingly. Total energy consumption for the FIG. 7 embodiment of the present invention is only 85 percent of that required for the prior art of FIG. 2.  
       EXAMPLE 6  
       [0100]    In the embodiment of the present invention shown in FIG. 7, the process was operated to achieve the same C 3  component recovery level as the prior art process shown in FIG. 2, with the resulting reduction in the utility consumption due to the better efficiency of the present invention. Alternatively, it is also possible to adjust the operating conditions of the present invention to increase the C 3  component recovery level while keeping the utility consumption the same as the prior art process, or to provide some combination of better recovery and lower utility consumption. For example, FIG. 8 shows the present invention when applied to match the compression power used by the prior art FIG. 2 process. The feed gas composition and conditions considered in the process presented in FIG. 8 are the same as those in FIG. 2. Accordingly, the FIG. 8 process can be compared with that of the FIG. 2 process to illustrate the advantages of the present invention.  
         [0101]    In the simulation of the FIG. 8 process, the feed gas cooling and expansion scheme, the deethanizer overhead compression and cooling scheme, and the tower reflux schemes are essentially the same as those used in FIG. 7. The only difference for the FIG. 8 embodiment of the present invention is that the operating pressures of separator/absorber  15  and deethanizer  17  have been adjusted to increase the recovery level for the C 3  components, with the corresponding drops in the process operating temperatures that result from the increase in process cooling (due primarily to the increase in expansion ratio across work expansion machine  13 ). Note that relative to the FIG. 7 embodiment, in the FIG. 8 embodiment the operating pressure of separator/absorber  15  has been lowered from 410 psia [2,827 kPa(a)] to 395 psia [2,723 kPa(a)], and the operating pressure of deethanizer  17  has been lowered from 290 psia [1,999 kPa(a)] to 285 psia [1,965 kPa(a)].  
         [0102]    A summary of stream flow rates and energy consumptions for the process illustrated in FIG. 8 is set forth in the table below:  
                                                           TABLE VIII                           (FIG. 8)       Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]            Stream   Methane   Ethane   Propane   Butanes+   Total               31   81,340   4,128   878   439   87,840       32   79,899   3,596   521   106   85,149       33    1,441     533   357   333    2,691       35    2,186   1,055   531   106    3,928       36    4,209   1,944    53    0    6,299       39    2,328   1,493    42    0    3,926       38     582     373    11    0     982       34   81,340   4,112    33    0   86,538       37      0     17   846   439    1,302            Recoveries*                   Propane    96.30%       Butanes+   100.00%       Power       Residue Gas Compression   17,428 HP   [28,651 kW]       Overhead Vapor Compression    1,483 HP    [2,438 kW]           18,911 HP   [31,089 kW]       Utility Heat       Deethanizer Reboiler   12,909 MBTU/Hr    [8,339 kW]                          
 
         [0103]    Comparison of the utility consumptions of the prior art process displayed in Table II with the utility consumptions of the present invention displayed in Table VIII shows that the present invention uses the same amount of external power for compression as the prior art process while increasing the C 3  component recovery and substantially reducing the utility heat requirement. The C 3  component recovery increases from 90.09% in the prior art FIG. 2 process to 96.30% in the present invention, an increase of over six percentage points. The utility heat requirement for the present invention is more than twenty-eight percent lower than the prior art process. The choice of whether to apply the present invention to increase the C 3  component recovery level, to reduce the utility consumptions, or to provide some combination of increased recovery and reduced utility consumption will normally be governed by the specific circumstances of each application, as the optimum will depend on such factors as plant size, available equipment, and the relative values of the recovered liquid product components and the utilities consumed.  
       Other Embodiments  
       [0104]    In accordance with this invention, it is generally advantageous to design the separator/absorber to provide a contacting device composed of multiple theoretical separation stages. However, the benefits of the present invention can be achieved with as few as one theoretical stage, and it is believed that even the equivalent of a fractional theoretical stage may allow achieving these benefits. For instance, all or a part of the partially condensed stream leaving heat exchanger  20  and all or a part of the partially condensed stream from work expansion machine  13  can be combined (such as in the piping joining the expansion machine to the separator/absorber) and if thoroughly intermingled, the vapors and liquids will mix together and separate in accordance with the relative volatilities of the various components of the total combined streams. In such an embodiment, the vapor-liquid mixture from heat exchanger  20  can be used without separation, or the liquid portion thereof may be separated. Such commingling of the two streams shall be considered for the purposes of this invention as constituting a contacting device. In another variation of the foregoing, the partially condensed stream from heat exchanger  20  can be separated (using separator  30  as shown in FIG. 9, for instance), and then all or a part of the separated liquid supplied to the separator/absorber or mixed with the vapors fed thereto (with any remaining portion of the separated liquid supplied to the deethanizer).  
         [0105]    As described earlier in the preferred embodiment, the overhead vapors are partially condensed and used to absorb valuable C 3  components and heavier components from the vapors leaving the work expansion machine. However, the present invention is not limited to this embodiment. It may be advantageous, for instance, to treat only a portion of the outlet vapor from the work expansion machine in this manner, or to use only a portion of the overhead condensate as an absorbent, in cases where other design considerations indicate portions of the expansion machine outlet or overhead condensate should bypass the separator/absorber. Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine  13 , or replacement with an alternate expansion device (such as an expansion valve), is feasible, or that total (rather than partial) condensation of the overhead stream in heat exchanger  20  is possible or is preferred. It should also be noted that the separator/absorber can be constructed either as a separate vessel or as a section of the deethanizer column.  
         [0106]    The use and distribution of the separator liquids, the separator/absorber liquids, and the reflux liquids for process heat exchange, the particular arrangement of heat exchangers for feed gas cooling, and the choice of process streams for specific heat exchange services must be evaluated for each particular application. For instance, as shown in FIG. 10, all or a part of the separator liquids (stream  33 ) may be routed directly to deethanizer  17  via an expansion device (such as expansion valve  12   a  shown in FIG. 10), with part or none of the liquid used for process cooling in heat exchanger  10 . Similarly, all or a part of the separator/absorber liquids (stream  35 ) may be routed directly to deethanizer  17  via an expansion device (such as expansion valve  27   a  shown in FIG. 10), with part or none of the liquid used for process cooling in heat exchanger  10 . Additionally, the condensed liquid that serves as reflux for deethanizer  17  (stream  38  in FIG. 10) can be used for process cooling before being supplied to the column. As shown in FIG. 10, all or a part of this liquid may be let down to slightly above the operating pressure of deethanizer  17  (using a device such as expansion valve  28 ) and used for process cooling (such as in heat exchanger  20  as shown) before being routed to deethanizer  17 , with part or none of the liquid routed directly to deethanizer  17  (via expansion valve  28   a , for example).  
         [0107]    Moreover, the use of external refrigeration to supplement the cooling available to the feed gas from other process streams may be employed as illustrated in FIG. 11, particularly in the case of an inlet gas richer than that used in Example 1. External refrigeration may also be employed to generate some or all of the reflux for the deethanizer as illustrated in FIG. 11. In such cases, all of the condensed liquid contained in the partially condensed stream leaving heat exchanger  20  (stream  36   e  in FIG. 11) might be directed only to the separator/absorber rather than a portion feeding the deethanizer. Note also in FIG. 11 (and as was shown previously in FIG. 10) that other process streams and/or external refrigeration may be used to supplement the cooling provided to the deethanizer overhead by the separator/absorber overhead (stream  34 ) in heat exchanger  20 , such as the flash expanded liquid (stream  35   a  in FIG. 11) from the bottom of the separator/absorber.  
         [0108]    Still other alternative means for generating the reflux stream for the deethanizer may be advantageous, depending of the particular application of the present invention. For example, as shown in FIG. 12, the heated flash expanded liquid (stream  35   b ) from the bottom of the separator/absorber could be used to cool a distillation stream (stream  40 ) from the deethanizer in heat exchanger  50  to partially condense the distillation stream (stream  40   a ), whereupon the condensed liquid (stream  38 ) is separated from the uncondensed vapor (stream  36 ) in separator  51 . Reflux pump  52  could then direct the condensed liquid (stream  38   a ) to deethanizer  17  to serve as its reflux, with the further heated stream  35   c  from heat exchanger  50  feeding deethanizer  17  at a mid-column feed point. Depending on the particular circumstances, the heated flash expanded liquid (stream  35   b ) from the bottom of the separator/absorber may contain an adequate quantity of liquid to serve as the reflux for the deethanizer, as shown in FIG. 13 and by the dashed lines in FIGS. 15 and 17. Further, as shown in FIGS. 13 through 18, it may be advantageous to direct some or all of the flash expanded separator liquid (stream  33   a ) to separator/absorber  15  rather than to deethanizer  17 , either to a separate fractionation zone in separator/absorber  15  or to the same fractionation zone as the outlet (stream  32   a ) from work expansion machine  13 .  
         [0109]    It will also be recognized that the relative amount of feed found in each branch of the condensed liquid contained in stream  36   e  that is split between the two towers in FIGS. 3, 7, and  8  will depend on several factors, including gas pressure, feed gas composition and the quantity of horsepower available. Similarly, the relative amount of condensation in separator  30  in FIG. 4, the relative amount of feed contained in stream  40   a  that is split between the two towers in FIG. 5, and the relative amount of condensation in separators  29  and  30  in FIG. 6 will also depend on factors such as these. The optimum split or distribution generally cannot be predicted without evaluating the particular circumstances for a specific application of the present invention. The mid-column feed positions depicted in FIGS. 3 through 8 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, etc. 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 8 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 some or all of the liquid streams (such as streams  33 ,  35 , and/or  38  in FIG. 7).  
         [0110]    It will also be recognized that the manner in which the deethanizer overhead stream (stream  36  in FIGS. 3, 4,  7  through  14 ,  17 , and  18 ) or the vapor stream (stream  41  in FIGS. 5 and 6) is compressed can be accomplished in a variety of ways. FIGS. 3 through 14,  17 , and  18  depict using a supplemental power source for compressor  25  to compress this stream, while compressing the residue gas (stream  34   b  in FIGS. 3, 4,  7 , and  8 ) or absorber overhead (stream  34   b  in FIGS. 5 and 6) using compressor  14  driven by expansion machine  13 . In other circumstances, for instance, it may be desirable to drive compressor  25  with expansion machine  13  and use a supplemental power source for compressor  14 . It may also be desirable to combine the services of compressor  25  and compressor  14  into a compound machine driven by expansion machine  13 . Further, as shown by the dashed equipment in FIGS. 3, 4,  7  through  14 ,  17 , and  18 , some circumstances may favor reducing the capital cost of the facility by eliminating heat exchanger  24  and/or heat exchanger  26  (at the expense of increasing the cooling load on heat exchanger  20  and either reducing the product recoveries or increasing the power consumption of compressor  22 ). Choices such as these must generally be evaluated for each application, as factors such as gas composition, plant size, desired recovery level, and available equipment must all be considered.  
         [0111]    The present invention provides improved recovery of C 3  components per amount of utility consumption required to operate the process. An improvement in utility consumption required for operating the deethanizer process may appear in the form of reduced power requirements for compression or re-compression, reduced power requirements for external refrigeration, reduced energy requirements for tower reboilers, or a combination thereof. Alternatively, if desired, increased C 3  component recovery can be obtained for a fixed utility consumption.  
         [0112]    It should also be noted that complete rejection of the C 2  components to the residue gas is not required by the present invention. If the project economics favor recovery of the C 2  components in the liquid product (stream  37 ), the process operating conditions can be altered to recover in the liquid product a significant portion of the C 2  components present in the feed gas. Preliminary calculations indicate that perhaps 40% of the C 2  components can be recovered in this fashion.  
         [0113]    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.