Abstract:
A process and an apparatus are disclosed for recovering ethane, ethylene, and heavier hydrocarbon components from a hydrocarbon gas stream. The stream is cooled, expanded to lower pressure, and supplied to a first fractionation tower at a mid-column feed position. A distillation liquid stream is withdrawn from the first fractionation tower below the feed position of the expanded stream, heated, and directed into a second fractionation tower that produces an overhead vapor stream and a bottom liquid stream. The overhead vapor stream is cooled to condense it, with a portion of the condensed stream directed to the second fractionation tower as its top feed and the remainder directed to the first fractionation tower at a lower column feed position. The bottom liquid stream from the second fractionation tower is cooled and directed to the first fractionation tower as its top feed.

Description:
The applicants claim the benefits under Title 35, U.S. Code, Section 119( e ) of prior U.S. Provisional Application No. 61/295,119 which was filed on Jan. 14, 2010. 
    
    
     BACKGROUND OF THE INVENTION 
     This invention relates to a process for the separation of a hydrocarbon bearing gas stream containing significant quantities of components more volatile than methane (e.g., hydrogen, nitrogen, etc.) into two fractions: a first fraction containing predominantly methane and the more volatile components, and a second fraction containing the recovered desirable ethane/ethylene and heavier hydrocarbon components. 
     Ethylene, ethane, propylene, propane, and/or heavier hydrocarbons can be recovered from a variety of gases, such as natural gas, refinery gas, and synthetic gas streams obtained from other hydrocarbon materials such as coal, crude oil, naphtha, oil shale, tar sands, and lignite. Hydrocarbon bearing gas typically contains components more volatile than methane (e.g., hydrogen, nitrogen, etc.) and often unsaturated hydrocarbons (e.g., ethylene, propylene, etc.) and aromatic hydrocarbons (e.g., benzene, toluene, etc.) in addition to methane, ethane, and hydrocarbons of higher molecular weight such as propane, butane, and pentane. Sulfur-containing gases and carbon dioxide are also sometimes present. 
     The present invention is generally concerned with the recovery of ethylene, ethane, and heavier (C 2 +) hydrocarbons from such gas streams. Recent changes in ethylene demand have created increased markets for ethylene and derivative products. In addition, fluctuations in the prices of both natural gas and its natural gas liquid (NGL) constituents have increased the value of ethane and heavier components as liquid products. These market conditions have resulted in the demand for processes which can provide high ethylene and ethane recovery and more efficient recovery of all these products. Available processes for separating these materials include those based upon cooling and refrigeration of gas, oil absorption, and refrigerated oil absorption. Additionally, cryogenic processes have become popular because of the availability of economical equipment that produces power while simultaneously expanding and extracting heat from the gas being processed. Depending upon the pressure of the gas source, the richness (ethane, ethylene, and heavier hydrocarbons content) of the gas, and the desired end products, each of these processes or a combination thereof may be employed. 
     The cryogenic expansion process is now generally preferred for natural gas liquids recovery because it provides maximum simplicity with ease of startup, operating flexibility, good efficiency, safety, and good reliability. U.S. Pat. Nos. 3,292,380; 4,061,481; 4,140,504; 4,157,904; 4,171,964; 4,185,978; 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; 5,555,748; 5,566,554; 5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664; 6,182,469; 6,578,379; 6,712,880; 6,915,662; 7,191,617; 7,219,513; reissue U.S. Pat. No. 33,408; and co-pending application Ser. Nos. 11/430,412; 11/839,693; 11/971,491; 12/206,230; 12/689,616; 12/717,394; 12/750,862; 12/772,472; 12/781,259; 12/868,993; 12/869,007; and 12/869,139 describe relevant processes (although the description of the present invention in some cases is based on different processing conditions than those described in the cited U.S. patents and applications). 
     In a typical cryogenic expansion recovery process, a feed gas stream under pressure is cooled by heat exchange with other streams of the process and/or external sources of refrigeration such as a propane compression-refrigeration system. As the gas is cooled, liquids may be condensed and collected in one or more separators as high-pressure liquids containing some of the desired C 2 + components. Depending on the richness of the gas and the amount of liquids formed, the high-pressure liquids may be expanded to a lower pressure and fractionated. The vaporization occurring during expansion of the liquids results in further cooling of the stream. Under some conditions, pre-cooling the high pressure liquids prior to the expansion may be desirable in order to further lower the temperature resulting from the expansion. The expanded stream, comprising a mixture of liquid and vapor, is fractionated in a distillation (demethanizer or deethanizer) column. In the column, the expansion cooled stream(s) is (are) distilled to separate residual methane, hydrogen, nitrogen, and other volatile gases as overhead vapor from the desired C 2  components, C 3  components, and heavier hydrocarbon components as bottom liquid product, or to separate residual methane, C 2  components, hydrogen, nitrogen, and other volatile gases as overhead vapor from the desired C 3  components and heavier hydrocarbon components as bottom liquid product. 
     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 essentially the same as the pressure at which the distillation column is operated. The combined vapor-liquid phases resulting from the expansion are supplied as feed to the column. 
     In the ideal operation of such a separation process, the residue gas leaving the process will contain substantially all of the methane and more volatile components in the feed gas with essentially none of the heavier hydrocarbon components, and the bottoms fraction leaving the demethanizer will contain substantially all of the heavier hydrocarbon components with essentially no methane or more volatile components. In practice, however, this ideal situation is not obtained because the conventional demethanizer is operated largely as a stripping column. The methane product of the process, therefore, typically comprises vapors leaving the top fractionation stage of the column, together with vapors not subjected to any rectification step. Considerable losses of ethylene and ethane occur because the top liquid feed contains substantial quantities of C 2 + components and heavier hydrocarbon components, resulting in corresponding equilibrium quantities of C 2 + components in the vapors leaving the top fractionation stage of the demethanizer. This problem is exacerbated if the gas stream(s) being processed contain relatively large quantities of components more volatile than methane (e.g., hydrogen, nitrogen, etc.) because the volatile vapors rising up the column strip C 2 + components from the liquids flowing downward. The loss of these desirable C 2 + components could be significantly reduced if the rising vapors could be brought into contact with a significant quantity of liquid (reflux) capable of absorbing the C 2 + components from the vapors. 
     A number of processes have been developed to use a cold liquid that is predominantly methane as the reflux stream to contact the rising vapors in a rectification section in the distillation column. Typical process schemes of this type are disclosed in U.S. Pat. Nos. 4,889,545; 5,568,737; and 5,881,569, and in Mowrey, E. Ross, “Efficient, High Recovery of Liquids from Natural Gas Utilizing a High Pressure Absorber”, Proceedings of the Eighty-First Annual Convention of the Gas Processors Association, Dallas, Tex., Mar. 11-13, 2002. Unfortunately, these processes require the use of a compressor to provide the motive force for recycling the reflux stream to the demethanizer, adding to both the capital cost and the operating cost of facilities using these processes. In addition, the cold methane reflux creates temperatures within the distillation column that are −112° F. [−80° C.] and colder. Many gas streams of this type contain significant quantities of nitrous oxides (NO X ) at times, which can accumulate in cold sections of a processing plant as NO X  gums (commonly referred to as “blue ice”) at temperatures lower than this. “Blue ice” can become explosive upon warming, and has been identified as the cause of a number of deflagrations and/or explosions in processing plants. 
     Other processes have been developed that use a heavy (C 4 -C 10  typically) hydrocarbon absorbent stream to reflux the distillation column. Examples of processes of this type are U.S. Pat. Nos. 4,318,723; 5,546,764; 7,273,542; and 7,714,180. While such processes generally operate at temperatures warm enough to avoid concerns about “blue ice”, the absorbent stream is typically produced from the distillation column bottoms stream, with the result that any aromatic hydrocarbons present in the feed gas will concentrate in the distillation column. Aromatic hydrocarbons such as benzene can freeze solid at normal processing temperatures, causing frequent disruptions in the processing plant. 
     In accordance with the present invention, it has been found that ethane recovery in excess of 88% can be obtained without requiring any temperatures to be lower than −112° F. [−80° C.]. The present invention is particularly advantageous when processing feed gases containing more than 10 mole % of components more volatile than methane. 
    
    
     
       For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings: 
         FIG. 1  is a flow diagram of gas processing plant in accordance with the present invention; and 
         FIG. 2  is a flow diagrams illustrating alternative means of application of the present invention to a gas stream. 
     
    
    
     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. 
     For convenience, process parameters are reported in both the traditional British units and in the units of the Système International d&#39;Unitès (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/or thousand British Thermal Units per hour (MBTU/Hr) 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 INVENTION 
       FIG. 1  illustrates a flow diagram of a process in accordance with the present invention. In the simulation of the  FIG. 1  process, inlet gas enters the plant at 100° F. [38° C.] and 77 psia [531 kPa(a)] as stream  51  If the inlet gas contains a concentration of sulfur compounds and/or carbon dioxide which would prevent the product streams from meeting specifications, the sulfur compounds and/or carbon dioxide are removed by appropriate pretreatment of the feed gas (not illustrated). 
     The inlet gas is compressed to higher pressure in three stages before processing (compressors  10  and  15  driven by an external power source and compressor  13  driven by work expansion machine  14 ). Discharge coolers  11  and  16  are used to cool the gas between stages, and separators  12  and  17  are used to remove any water or other liquids that condense from the gas stream as it is cooled. The cooled compressed gas stream  54  leaving separator  17  is dehydrated in dehydration unit  18  to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccant has typically been used for this purpose. 
     The dehydrated gas stream  61  at 100° F. [38° C.] and 560 psia [3,859 kPa(a)] enters heat exchanger  20  and is cooled by heat exchange with cool residue gas (stream  68   a ), liquid product at 28° F. [−2° C.] (stream  71   a ), demethanizer reboiler liquids at 13° F. [−11° C.] (stream  70 ), and propane refrigerant. Note that in all cases exchanger  20  is representative of either a multitude of individual heat exchangers or a single multi-pass heat exchanger, or any combination thereof (The decision as to whether to use more than one heat exchanger for the indicated cooling services will depend on a number of factors including, but not limited to, inlet gas flow rate, heat exchanger size, stream temperatures, etc.) The cooled stream  61   a  enters separator  21  at 40° F. [4° C.] and 550 psia [3,790 kPa(a)] where the vapor (stream  62 ) is separated from the condensed liquid (stream  63 ). The separator liquid (stream  63 ) is expanded to the operating pressure (approximately 175 psia [1,207 kPa(a)]) of fractionation tower  28  by expansion valve  22 , cooling stream  63   a  to 16° F. [−9° C.] before it is supplied to fractionation tower  28  at a lower column feed point. 
     The vapor (stream  62 ) from separator  21  is further cooled in heat exchanger  23  by heat exchange with cold residue gas (stream  68 ), demethanizer side reboiler liquids at −10° F. [−23° C.] (stream  69 ), flashed liquids (stream  65   a ), and propane refrigerant. The cooled stream  62   a  enters separator  24  at −42° F. [−41° C.] and 535 psia [3,686 kPa(a)] where the vapor (stream  64 ) is separated from the condensed liquid (stream  65 ). The separator liquid (stream  65 ) is expanded to slightly above the tower operating pressure by expansion valve  25 , cooling stream  65   a  to −63° F. [−53° C.] before it is heated to −40° F. [−40° C.] in heat exchanger  23 . The heated stream  65   b  is then supplied to fractionation tower  28  at a lower mid-column feed point. 
     The vapor (stream  64 ) from separator  24  enters work expansion machine  14  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  14  expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream  64   a  to a temperature of approximately −105° F. [−76° C.]. The typical commercially available expanders are capable of recovering on the order of 80-85% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item  13 ) that can be used to compress the inlet gas (stream  52 ), for example. The partially condensed expanded stream  64   a  is thereafter supplied as feed to fractionation tower  28  at an upper mid-column feed point. 
     The demethanizer in tower  28  is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The demethanizer tower consists of two sections: an upper absorbing (rectification) section that contains the trays and/or packing to provide the necessary contact between the vapor portion of the expanded stream  64   a  rising upward and cold liquid falling downward to condense and absorb the C 2  components, C 3  components, and heavier components from the vapors rising upward; and a lower, stripping (demethanizing) section that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section also includes one or more reboilers (such as the reboiler and side reboiler described previously) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product, stream  71 , of methane and lighter components. Stream  64   a  enters demethanizer  28  at an intermediate feed position located in the lower region of the absorbing section of demethanizer  28 . The liquid portion of the expanded stream commingles with liquids falling downward from the absorbing section and the combined liquid continues downward into the stripping section of demethanizer  28 . 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 2  components, C 3  components, and heavier components. 
     A portion of the distillation liquid (stream  72 ) is withdrawn from an intermediate region of the stripping section in fractionation column  28 , below the feed position of expanded stream  64   a  in the lower region of the absorbing section but above the feed position of expanded liquid stream  63   a  in the stripping section. Withdrawing the distillation liquid at this location provides a liquid stream that is predominantly C 2 -C 5  hydrocarbons containing very little of the volatile components (e.g., methane, hydrogen, nitrogen, etc.) and little of the aromatic hydrocarbons and heavier hydrocarbon components. This distillation vapor stream  72  is pumped to higher pressure by pump  30  (stream  72   a ) and then heated from −25° F. [−32° C.] to 77° F. [25° C.] and partially vaporized in heat exchanger  31  by heat exchange with the hot depropanizer bottom stream  78 . The heated stream  72   b  then enters depropanizer  32  (operating at 265 psia [1,828 kPa(a)]) at a mid-column feed point. 
     The depropanizer in tower  32  is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The depropanizer tower consists of two sections: an upper absorbing (rectification) section that contains the trays and/or packing to provide the necessary contact between the vapor portion of the heated stream  72   b  rising upward and cold liquid falling downward to condense and absorb the C 4  components and heavier components; and a lower, stripping (depropanizing) section that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The depropanizing section also includes one or more reboilers (such as reboiler  33 ) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the bottom liquid product, stream  78 , of C 3  components and lighter components. Stream  72   b  enters depropanizer  32  at an intermediate feed position located between the absorbing section and the stripping section of depropanizer  32 . The liquid portion of the heated stream commingles with liquids falling downward from the absorbing section and the combined liquid continues downward into the stripping section of depropanizer  32 . The vapor portion of the heated stream rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C 4  components and heavier components. 
     The overhead vapor (stream  73 ) from depropanizer  32  enters reflux condenser  34  and is cooled by propane refrigerant from 59° F. [15° C.] to −33° F. [−36° C.] to condense it before entering reflux separator  35  at 260 psia [1,793 kPa(a)]. If there is any uncondensed vapor (stream  74 ), it is expanded to the operating pressure of demethanizer  28  by expansion valve  38  and returned to demethanizer  28  at a lower column feed point. In the simulation of  FIG. 1 , however, all of the overhead vapor is condensed and leaves reflux separator  35  in liquid stream  75 . Stream  75  is pumped by pump  36  to a pressure slightly above the operating pressure of depropanizer  32 , and a portion (stream  76 ) of stream  75   a  is then supplied as top column feed (reflux) to depropanizer  32  to absorb and condense the C 4  components and heavier components rising in the absorbing section of the column. The remaining portion (stream  77 ) contains the C 3  and lighter components stripped from distillation liquid stream  72 . It is expanded to the operating pressure of demethanizer  28  by expansion valve  37 , cooling stream  37   a  to −44° F. [−42° C.] before it is returned to demethanizer  28  at a lower column feed point, below the withdrawal point of distillation liquid stream  72 . 
     The bottom liquid product from depropanizer  32  (stream  78 ) has been stripped of the C 3  and lighter components, and is predominantly C 4 -C 5  hydrocarbons. It leaves the bottom of depropanizer  32  at 230° F. [110° C.] and is cooled to −20° F. [−29° C.] in heat exchanger  31  as described earlier. Stream  78   a  is further cooled to −35° F. [−37° C.] with propane refrigerant in heat exchanger  39  (stream  78   b ) and then expanded to the operating pressure of demethanizer  28  in expansion valve  40 . The expanded stream  78   c  is then supplied to demethanizer  28  as reflux, entering at the top feed location at −35° F. [−37° C.]. The C 4 -C 5  hydrocarbons in stream  78   c  act as an absorbent to capture the C 2 + components in the vapors flowing upward in the absorbing section of demethanizer  28 . 
     In the stripping section of demethanizer  28 , the feed streams are stripped of their methane and lighter components. The resulting liquid product (stream  71 ) exits the bottom of tower  28  at 24° F. [−4° C.] and is pumped to higher pressure in pump  29 . The pumped stream  71   a  is then heated to 93° F. [34° C.] in heat exchanger  20  as described previously. The cold residue gas stream  68  leaves demethanizer  28  at −32° F. [−35° C.] and passes countercurrently to the incoming feed gas in heat exchanger  23  where it is heated to 32° F. [0° C.] (stream  68   a ) and in heat exchanger  20  where it is heated to 95° F. [35° C.] (stream  68   b ) as it provides cooling as previously described. The residue gas product then flows to the fuel gas distribution header at 165 psia [1,138 kPa(a)]. 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 1  is set forth in the following table: 
     
       
         
               
             
               
               
               
               
               
               
             
               
               
               
               
               
             
               
               
               
               
             
           
               
                 TABLE I 
               
               
                   
               
               
                 (FIG. 1) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
             
               
                   
               
             
          
           
               
                   
                   
                   
                 Stream 
                 Stream 
                 Stream 
               
               
                 Component 
                 Stream 61 
                 Stream 62 
                 63 
                 64 
                 65 
               
               
                   
               
               
                 Hydrogen 
                 833 
                 823 
                 10 
                 814 
                 9 
               
               
                 Methane 
                 2,375 
                 2,225 
                 150 
                 1,980 
                 245 
               
               
                 Ethylene 
                 115 
                 95 
                 20 
                 60 
                 35 
               
               
                 Ethane 
                 944 
                 710 
                 234 
                 349 
                 361 
               
               
                 Propylene 
                 212 
                 112 
                 100 
                 23 
                 89 
               
               
                 Propane 
                 597 
                 293 
                 304 
                 51 
                 242 
               
               
                 Butylene/Butadiene 
                 135 
                 36 
                 99 
                 2 
                 34 
               
               
                 i-Butane 
                 78 
                 23 
                 55 
                 2 
                 21 
               
               
                 n-Butane 
                 166 
                 39 
                 127 
                 2 
                 37 
               
               
                 Pentanes+ 
                 46 
                 5 
                 41 
                 0 
                 5 
               
               
                 Totals 
                 5,577 
                 4,431 
                 1,146 
                 3,348 
                 1,083 
               
               
                   
               
               
                   
                   
                   
                 Stream 
                 Stream 
                 Stream 
               
               
                 Component 
                 Stream 72 
                 Stream 73 
                 75 
                 76 
                 77 
               
               
                   
               
               
                 Hydrogen 
                 0 
                 0 
                 0 
                 0 
                 0 
               
               
                 Methane 
                 186 
                 298 
                 298 
                 112 
                 186 
               
               
                 Ethylene 
                 89 
                 142 
                 142 
                 53 
                 89 
               
               
                 Ethane 
                 836 
                 1,336 
                 1,336 
                 500 
                 836 
               
               
                 Propylene 
                 129 
                 194 
                 194 
                 73 
                 121 
               
               
                 Propane 
                 353 
                 482 
                 482 
                 180 
                 302 
               
               
                 Butylene/Butadiene 
                 239 
                 24 
                 24 
                 9 
                 15 
               
               
                 i-Butane 
                 111 
                 18 
                 18 
                 7 
                 11 
               
               
                 n-Butane 
                 396 
                 16 
                 16 
                 6 
                 10 
               
               
                 Pentanes+ 
                 220 
                 0 
                 0 
                 0 
                 0 
               
               
                 Totals 
                 2,569 
                 2,515 
                 2,515 
                 941 
                 1,574 
               
               
                   
               
             
          
           
               
                   
                 Component 
                 Stream 78 
                 Stream 68 
                 Stream 71 
               
               
                   
                   
               
               
                   
                 Hydrogen 
                 0 
                 833 
                 0 
               
               
                   
                 Methane 
                 0 
                 2,352 
                 23 
               
               
                   
                 Ethylene 
                 0 
                 45 
                 70 
               
               
                   
                 Ethane 
                 0 
                 109 
                 835 
               
               
                   
                 Propylene 
                 8 
                 4 
                 208 
               
               
                   
                 Propane 
                 51 
                 21 
                 576 
               
               
                   
                 Butylene/Butadiene 
                 224 
                 22 
                 113 
               
               
                   
                 i-Butane 
                 100 
                 12 
                 66 
               
               
                   
                 n-Butane 
                 386 
                 29 
                 137 
               
               
                   
                 Pentanes+ 
                 220 
                 4 
                 42 
               
               
                   
                 Totals 
                 995 
                 3,501 
                 2,076 
               
               
                   
                   
               
             
          
           
               
                   
                 Recoveries* 
                   
                   
               
               
                   
                 Ethylene 
                 60.81% 
               
               
                   
                 Ethane 
                 88.41% 
               
               
                   
                 Propylene 
                 98.22% 
               
               
                   
                 Propane 
                 96.57% 
               
               
                   
                 Butanes+ 
                 84.03% 
               
               
                   
                 Power 
               
               
                   
                 Inlet Gas Compression 
                 6,072 HP 
                 [9,982 kW] 
               
               
                   
                 Refrigerant Compression 
                 5,015 HP 
                 [8,245 kW] 
               
               
                   
                 Total Compression 
                 11,087 HP  
                 [18,227 kW]  
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
             
          
         
       
     
     Other Embodiments 
     In accordance with this invention, it is generally advantageous to design the absorbing (rectification) section of the demethanizer to contain multiple theoretical separation stages. However, the benefits of the present invention can be achieved with as few as two theoretical stages. For instance, all or a part of the reflux liquid (stream  78   c ) and all or a part of the expanded stream  64   a  can be combined (such as in the piping to the demethanizer) 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. Such commingling of the two streams, shall be considered for the purposes of this invention as constituting an absorbing section. 
       FIG. 2  displays another embodiment of the present invention that may be preferred in some circumstances. In the  FIG. 2  embodiment, a portion (stream  66 ) of vapor stream  64  from separator  24  is expanded to an intermediate pressure by expansion valve  26  and then combined with cooled depropanizer bottoms stream  78   b  to form a combined stream  79 . The combined stream  79  is cooled in heat exchanger  27  (stream  79   a ) by the cold demethanizer overhead stream  68 , then expanded to the operating pressure of demethanizer  28  by expansion valve  40 . The expanded stream  79   b  is then supplied as reflux to the top feed position of demethanizer  28 . The remaining portion (stream  67 ) of vapor stream  64 ) is expanded to the tower operating pressure by work expansion machine  14 , and the expanded stream  67   a  is supplied to the upper mid-column feed position on demethanizer  28 . 
     Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine  14 , or replacement with an alternate expansion device (such as an expansion valve), is feasible. Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where appropriate. For example, conditions may warrant work expansion of the reflux stream (stream  78   b  or stream  79   a ). 
     When the inlet gas is leaner, separator  21  in  FIGS. 1 and 2  may not be justified. In such cases, the feed gas cooling accomplished in heat exchangers  20  and  23  in  FIGS. 1 and 2  may be accomplished without an intervening separator. The decision of whether or not to cool and separate the feed gas in multiple steps will depend on the richness of the feed gas, plant size, available equipment, etc. Depending on the quantity of heavier hydrocarbons in the feed gas and the feed gas pressure, the cooled feed stream  61   a  leaving heat exchanger  20  and/or the cooled stream  62   a  leaving heat exchanger  23  in  FIGS. 1 and 2  may not contain any liquid (because it is above its dewpoint, or because it is above its cricondenbar), so that separator  21  and/or separator  24  shown in  FIGS. 1 and 2  are not required. 
     The expanded liquid (stream  65   a  in  FIGS. 1 and 2 ) need not be heated before it is supplied to the lower mid-column feed point on the distillation column. Instead, all or a portion of it may be supplied directly to the column. Any remaining portion of the expanded liquid may then be heated before it is fed to the distillation column. 
     In accordance with the present invention, the use of external refrigeration to supplement the cooling available to the inlet gas from other process streams may be employed, particularly in the case of a rich inlet gas. The use and distribution of separator liquids and demethanizer side draw liquids for process heat exchange, and the particular arrangement of heat exchangers for inlet gas cooling must be evaluated for each particular application, as well as the choice of process streams for specific heat exchange services. 
     In accordance with the present invention, the splitting of the vapor feed for the  FIG. 2  embodiment may be accomplished in several ways. In the process of  FIG. 2 , the splitting of vapor occurs following cooling and separation of any liquids which may have been formed. The high pressure gas may be split, however, prior to any cooling of the inlet gas or after the cooling of the gas and prior to any separation stages. In some embodiments, vapor splitting may be effected in a separator. 
     It will also be recognized that the relative amount of feed found in each branch of the split vapor feed of the  FIG. 2  embodiment will depend on several factors, including gas pressure, feed gas composition, the amount of heat which can economically be extracted from the feed, and the quantity of horsepower available. More feed to the top of the column may increase recovery while decreasing power recovered from the expander thereby increasing the compression horsepower requirements. Increasing feed lower in the column reduces the horsepower consumption but may also reduce product recovery. The relative locations of the mid-column feeds may vary depending on inlet composition or other factors such as desired recovery levels and amount of liquid formed during inlet gas cooling. Moreover, two or more of the feed streams, or portions thereof, may be combined depending on the relative temperatures and quantities of individual streams, and the combined stream then fed to a mid-column feed position. 
     The present invention provides improved recovery of C 2  components, C 3  components, and heavier hydrocarbon components per amount of utility consumption required to operate the process. An improvement in utility consumption required for operating the demethanizer 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. 
     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.