Patent Publication Number: US-2019170435-A1

Title: Hydrocarbon Gas Processing

Description:
BACKGROUND OF THE INVENTION 
     This invention relates to a process and an apparatus for the separation of a gas containing hydrocarbons. This application is a continuation of U.S. patent application Ser. No. 14/828,093 filed Aug. 17, 2015 and claims the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application No. 62/045,908 which was filed on Sep. 4, 2014. 
    
    
     Ethylene, ethane, propylene, propane, and/or heavier hydrocarbons can be recovered from a variety of gases, such as natural gas, refinery gas, and synthetic gas streams obtained from other hydrocarbon materials such as coal, crude oil, naphtha, oil shale, tar sands, and lignite. Natural gas usually has a major proportion of methane and ethane, i.e., methane and ethane together comprise at least 50 mole percent of the gas. The gas also contains relatively lesser amounts of heavier hydrocarbons such as propane, butanes, pentanes, and the like, as well as hydrogen, nitrogen, carbon dioxide, and other gases. 
     The present invention is generally concerned with the recovery of ethylene, ethane, propylene, propane and heavier hydrocarbons from such gas streams. A typical analysis of a gas stream to be processed in accordance with this invention would be, in approximate mole percent, 69.4% methane, 11.8% ethane and other C 2  components, 5.6% propane and other C 3  components, 0.9% iso-butane, 1.8% normal butane, and 1.1% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present. 
     The historically cyclic fluctuations in the prices of both natural gas and its natural gas liquid (NGL) constituents have at times reduced the incremental value of ethane, ethylene, propane, propylene, and heavier components as liquid products. This has resulted in a demand for processes that can provide more efficient recoveries of these products, for processes that can provide efficient recoveries with lower capital investment, and for processes that can be easily adapted or adjusted to vary the recovery of a specific component over a broad range. 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; 8,590,340; 8,881,549; 8,919,148; 9,021,831; 9,021,832; 9,052,136; 9,052,137; 9,057,558; 9,068,774; 9,074,814; 9,080,810; and 9,080,811; reissue U.S. Pat. No. 33,408; and co-pending application Ser. Nos. 11/839,693; 12/750,862; 12/772,472; 12/781,259; 12/868,993; 12/869,007; 12/869,139; 14/462,056; and 14/462,083 describe relevant processes (although the description of the present invention in some cases is based on different processing conditions than those described in the listed 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, 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, 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 split into two streams. One portion of the vapor is 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. 
     The remaining portion of the vapor is cooled to substantial condensation by heat exchange with other process streams, e.g., the cold fractionation tower overhead. Some or all of the high-pressure liquid may be combined with this vapor portion prior to cooling. The resulting cooled stream is then expanded through an appropriate expansion device, such as an expansion valve, to the pressure at which the demethanizer is operated. During expansion, a portion of the liquid will vaporize, resulting in cooling of the total stream. The flash expanded stream is then supplied as top feed to the demethanizer. Typically, the vapor portion of the flash expanded stream and the demethanizer overhead vapor combine in an upper separator section in the fractionation tower as residual methane product gas. Alternatively, the cooled and expanded stream may be supplied to a separator to provide vapor and liquid streams. The vapor is combined with the tower overhead and the liquid is supplied to the column as a top column feed. 
     In the ideal operation of such a separation process, the residue gas leaving the process will contain substantially all of the methane in the feed gas with essentially none of the heavier hydrocarbon components, and the bottoms fraction leaving the demethanizer will contain substantially all of the heavier 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 C 2 , C 3 , and C 4 + components occur because the top liquid feed contains substantial quantities of these components and heavier hydrocarbon components, resulting in corresponding equilibrium quantities of C 2  components, C 3  components, C 4  components, and heavier hydrocarbon components in the vapors leaving the top fractionation stage of the demethanizer. The loss of these desirable 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, C 3  components, C 4  components, and heavier hydrocarbon components from the vapors. 
     In recent years, the preferred processes for hydrocarbon separation use an upper absorber section to provide additional rectification of the rising vapors. The source of the reflux stream for the upper rectification section is typically a recycled stream of residue gas supplied under pressure. The recycled residue gas stream is usually cooled to substantial condensation by heat exchange with other process streams, e.g., the cold fractionation tower overhead. The resulting substantially condensed stream is then expanded through an appropriate expansion device, such as an expansion valve, to the pressure at which the demethanizer is operated. During expansion, a portion of the liquid will usually vaporize, resulting in cooling of the total stream. The flash expanded stream is then supplied as top feed to the demethanizer. Typically, the vapor portion of the expanded stream and the demethanizer overhead vapor combine in an upper separator section in the fractionation tower as residual methane product gas. Alternatively, the cooled and expanded stream may be supplied to a separator to provide vapor and liquid streams, so that thereafter the vapor is combined with the tower overhead and the liquid is supplied to the column as a top column feed. Typical process schemes of this type are disclosed in U.S. Pat. Nos. 4,889,545; 5,568,737; 5,881,569; and 9,052,137, assignee&#39;s co-pending application Ser. No. 12/717,394, 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, Texas, Mar. 11-13, 2002. Unfortunately, these processes require the use of a large amount of compression power 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. 
     The present invention also employs an upper rectification section (or a separate rectification column in some embodiments). However, the reflux for this rectification section is provided by cooling two streams derived from the feed gas to substantial condensation and then expanding both streams to the operating pressure of the fractionation tower. During expansion, a portion of each stream is vaporized, resulting in cooling of each total stream. One cooled, expanded stream is then supplied to the fractionation tower at a top column feed point and the other cooled, expanded stream is supplied to the tower at an upper mid-column feed point. The condensed liquid in the top column feed, which is predominantly liquid methane, can then be used to absorb C 2  components, C 3  components, C 4  components, and heavier hydrocarbon components from the vapors rising through the upper rectification section and thereby capture these valuable components in the bottom liquid product from the demethanizer. 
     In accordance with the present invention, it has been found that C 2  recovery in excess of 93% and C 3  and C 4 + recoveries in excess of 99% can be obtained. In addition, the present invention makes possible essentially 100% separation of methane and lighter components from the C 2  components and heavier components at higher recovery levels compared to the prior art while maintaining the same 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 NGL recovery column overhead temperatures of −50° F. [−46° C.] or colder. 
    
    
     
       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 a prior art natural gas processing plant in accordance with U.S. Pat. No. 4,278,457; 
         FIG. 2  is a flow diagram of a prior art natural gas processing plant in accordance with U.S. Pat. No. 5,568,737; 
         FIG. 3  is a flow diagram of a prior art natural gas processing plant in accordance with U.S. Pat. No. 7,191,617; 
         FIG. 4  is a flow diagram of a prior art natural gas processing plant in accordance with assignee&#39;s co-pending application Ser. No. 11/839,693; 
         FIG. 5  is a flow diagram of a natural gas processing plant in accordance with the present invention; and 
         FIGS. 6 through 9  are flow diagrams illustrating alternative means of application of the present invention to a natural 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′Unités (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 PRIOR ART 
       FIG. 1  is a process flow diagram showing the design of a processing plant to recover C 2 + components from natural gas using prior art according to U.S. Pat. No. 4,278,457. In this simulation of the process, inlet gas enters the plant at 104° F. [40° C.] and 896 psia [6,181 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. 
     The feed stream  31  is divided into two portions, streams  32  and  33 . Stream  32  is cooled to substantial condensation in heat exchanger  13  by heat exchange with cold residue gas (stream  41 ), flashed separator liquids (stream  35   a ), and propane refrigerant. The resulting substantially condensed stream  32   a  at −132° F. [−91° C.] is then flash expanded through expansion valve  17  to the operating pressure (approximately 316 psia [2,181 kPa(a)]) of fractionation tower  19 . During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in  FIG. 1 , the expanded stream  32   b  leaving expansion valve  17  reaches a temperature of −159° F. [−106° C.] before it is supplied to fractionation tower  19  as the top column feed. 
     The remaining portion of feed stream  31 , stream  33 , is cooled in heat exchanger  10  by heat exchange with cool residue gas (stream  41   a ) and propane refrigerant. Note that in all cases exchangers  10  and  13  are 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  33   a  enters separator  12  at −30° F. [−35° C.] and 870 psia [6,002 kPa(a)] where the vapor (stream  34 ) is separated from the condensed liquid (stream  35 ). The separator liquid (stream  35 ) is expanded to slightly above the operating pressure of fractionation tower  19  by expansion valve  18 . Expanded stream  35   a  is heated from −63° F. [−53° C.] to 23° F. [−5° C.] in heat exchanger  13  as described earlier before heated stream  35   b  is supplied to fractionation tower  19  at a lower mid-column feed point. 
     The vapor (stream  34 ) from separator  12  enters a work expansion machine  15  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  15  expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream  34   a  to a temperature of approximately −99° F. [−73° 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  16 ) that can be used to re-compress the residue gas (stream  41   b ), for example. The partially condensed expanded stream  34   a  is thereafter supplied as feed to fractionation tower  19  at an upper mid-column feed point. 
     The demethanizer in tower  19  is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. As is often the case in natural gas processing plants, the fractionation tower may consist of three sections. The upper section  19   a  is a separator wherein the partially vaporized top feed is separated into its respective vapor and liquid portions, and wherein the vapor rising from the rectifying section  19   b  below is combined with the vapor portion of the top feed to form the cold demethanizer overhead vapor (stream  41 ) which exits the top of the tower at −137° F. [−94° C.]. The middle absorbing (rectifying) section  19   b  contains the trays and/or packing to provide the necessary contact between the vapor portions of the expanded streams  34   a  and  35   b  rising upward and cold liquid falling downward. The lower stripping (demethanizing) section  19   c  contains additional trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. Demethanizing section  19   c  also includes one or more reboilers (such as the reboiler  21  and side reboiler  20  shown in  FIG. 1 ) 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  44 , of methane and lighter components. Stream  34   a  enters demethanizer  19  at an intermediate feed position located below rectifying section  19   b  and above demethanizing section  19   c . The liquid portion of the expanded stream  34   a  commingles with liquids falling downward from rectifying section  19   b  and the combined liquid continues downward into demethanizing section  19   c  of column  19 . The vapor portion of the expanded stream  34   a  commingles with vapors rising upward from demethanizing section  19   c  and the combined vapor rises upward through rectifying section  19   b  and is contacted with cold liquid falling downward to condense and absorb the C 2  components, C 3  components, and heavier components from the vapor. 
     The liquid product (stream  44 ) exits the bottom of tower  19  at 79° F. [26° C.], based on a typical specification of a methane concentration of 0.07% on a volume basis in the bottom product. The cold residue gas stream  41  passes countercurrently to one portion of the feed gas in heat exchanger  13  where it is heated to −31° F. [−35° C.] (stream  41   a ), and countercurrently to the other portion of the feed gas in heat exchanger  10  where it is heated to 88° F. [31° C.] (stream  41   b ). The residue gas is then re-compressed in two stages. The first stage is compressor  16  driven by expansion machine  15 . The second stage is compressor  22  driven by a supplemental power source which compresses the residue gas (stream  41   d ) to sales line pressure. After cooling to 110° F. [43° C.] in discharge cooler  23 , the residue gas product (stream  41   e ) flows to the sales gas pipeline at 918 psia [6,330 kPa(a)], sufficient to meet line requirements (usually on the order of the inlet pressure). 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 1  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE I 
               
               
                   
               
               
                 (FIG. 1) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
            
               
                   
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 31 
                 22,875 
                 3,898 
                 1,843 
                 1,240 
                 32,972 
               
               
                 32 
                 6,428 
                 1,095 
                 518 
                 348 
                 9,265 
               
               
                 33 
                 16,447 
                 2,803 
                 1,325 
                 892 
                 23,707 
               
               
                 34 
                 14,127 
                 1,566 
                 382 
                 94 
                 18,283 
               
               
                 35 
                 2,320 
                 1,237 
                 943 
                 798 
                 5,424 
               
               
                 41 
                 22,867 
                 504 
                 20 
                 2 
                 26,507 
               
               
                 44 
                 8 
                 3,394 
                 1,823 
                 1,238 
                 6,465 
               
               
                   
               
            
           
           
               
            
               
                 Recoveries* 
               
               
                   
               
            
           
           
               
               
               
            
               
                   
                 Ethane 
                 87.06% 
               
               
                   
                 Propane 
                 98.92% 
               
               
                   
                 Butanes+ 
                 99.88% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
               
                   
               
            
           
           
               
               
               
               
               
            
               
                   
                 Residue Gas Compression 
                 16,044 
                 HP 
                 [26,376 kW] 
               
               
                   
                 Refrigerant Compression 
                 7,492 
                 HP 
                 [12,317 kW] 
               
               
                   
                 Total Compression 
                 23,536 
                 HP 
                 [38,693 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
       FIG. 2  represents an alternative prior art process according to U.S. Pat. No. 5,568,737. The process of  FIG. 2  has been applied to the same feed gas composition and conditions as described above for  FIG. 1 . In the simulation of this process, as in the simulation for the process of  FIG. 1 , operating conditions were selected to maximize the recovery level for a given energy consumption. 
     After feed stream  31  is divided into two portions, streams  32  and  33 , stream  32  is cooled to substantial condensation in heat exchanger  13  by heat exchange with cold residue gas (stream  41 ), flashed separator liquids (stream  35   a ), and propane refrigerant. The resulting substantially condensed stream  32   a  at −140° F. [−96° C.] is then flash expanded through expansion valve  17  to the operating pressure (approximately 340 psia [2,346 kPa(a)]) of fractionation tower  19 . During expansion a portion of the stream is vaporized, resulting in cooling of expanded stream  32   b  to −159° F. [−106° C.] before it is supplied to fractionation tower  19  at an upper mid-column feed point. 
     The remaining portion of feed stream  31 , stream  33 , is cooled in heat exchanger  10  by heat exchange with a portion (stream  47 ) of the cool demethanizer overhead vapor (stream  45   a ) and propane refrigerant, and cooled stream  33   a  enters separator  12  at −25° F. [−32° C.] and 870 psia [6,002 kPa(a)] where the vapor (stream  34 ) is separated from the condensed liquid (stream  35 ). The separator liquid (stream  35 ) is expanded to slightly above the operating pressure of fractionation tower  19  by expansion valve  18 . Expanded stream  35   a  is heated from −54° F. [−48° C.] to 23° F. [−5° C.] in heat exchanger  13  as described earlier before heated stream  35   b  is supplied to fractionation tower  19  at a lower mid-column feed point. 
     The vapor (stream  34 ) from separator  12  enters a work expansion machine  15  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  15  expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream  34   a  to a temperature of approximately −89° F. [−67° C.]. The partially condensed expanded stream  34   a  is thereafter supplied as feed to fractionation tower  19  at an intermediate mid-column feed point. 
     The recompressed and cooled distillation vapor stream  45   e  is divided into two streams. One portion, stream  41 , is the volatile residue gas product. The other portion, recycle stream  48 , flows to heat exchanger  11  where it is cooled to 0° F. [−18° C.] by heat exchange with the remaining portion (stream  46 ) of cool demethanizer overhead vapor stream  45   a . The cooled recycle stream  48   a  then flows to exchanger  13  where it is further cooled to −140° F. [−96° C.] and substantially condensed by heat exchange with cold distillation vapor stream  45  and propane refrigeration. The substantially condensed stream  48   b  is then expanded through an appropriate expansion device, such as expansion valve  14 , to the demethanizer operating pressure, resulting in cooling of the total stream to −167° F. [−111° C.]. The expanded stream  48   c  is then supplied to fractionation tower  19  as the top column feed. The vapor portion of stream  48   c  combines with the vapors rising from the top fractionation stage of the column to form distillation vapor stream  45 , which is withdrawn from an upper region of the tower. 
     The liquid product (stream  44 ) exits the bottom of tower  19  at 88° F. [31° C.], based on a methane concentration of 0.07% on a volume basis in the bottom product. The demethanizer overhead vapor stream  45  passes countercurrently to one portion of the feed gas (stream  32 ) and the partially cooled recycle stream (stream  48   a ) in heat exchanger  13  where it is heated to −27° F. [−33° C.] (stream  45   a ) and then divided into two portions, stream  46  and stream  47 . Stream  46  passes countercurrently to recycle stream  48  in heat exchanger  11  and is heated to 105° F. [41° C.] (stream  46   a ), while stream  47  passes countercurrently to the other portion of the feed gas in heat exchanger  10  where it is heated to 91° F. [33° C.] (stream  47   a ). Streams  46   a  and  47 a recombine as stream  45   b  at 92° F. [33° C.], which is then re-compressed in two stages, compressor  16  driven by expansion machine  15  and compressor  22  driven by a supplemental power source. After cooling to 110° F. [43° C.] in discharge cooler  23 , stream  45   e  is split into the residue gas product (stream  41 ) and the recycle stream  48  as described earlier. Residue gas stream  41  then flows to the sales gas pipeline 918 psia [6,330 kPa(a)]. 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 2  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE II 
               
               
                   
               
               
                 (FIG. 2) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
            
               
                   
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 31 
                 22,875 
                 3,898 
                 1,843 
                 1,240 
                 32,972 
               
               
                 32 
                 5,513 
                 939 
                 444 
                 299 
                 7,946 
               
               
                 33 
                 17,362 
                 2,959 
                 1,399 
                 941 
                 25,026 
               
               
                 34 
                 15,187 
                 1,761 
                 450 
                 114 
                 19,761 
               
               
                 35 
                 2,175 
                 1,198 
                 949 
                 827 
                 5,265 
               
               
                 48 
                 3,035 
                 69 
                 0 
                 0 
                 3,518 
               
               
                 45 
                 25,902 
                 592 
                 0 
                 0 
                 30,022 
               
               
                 41 
                 22,867 
                 523 
                 0 
                 0 
                 26,504 
               
               
                 44 
                 8 
                 3,375 
                 1,843 
                 1,240 
                 6,468 
               
               
                   
               
            
           
           
               
            
               
                 Recoveries* 
               
               
                   
               
            
           
           
               
               
               
            
               
                   
                 Ethane 
                 86.57% 
               
               
                   
                 Propane 
                 100.00% 
               
               
                   
                 Butanes+ 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
               
                   
               
            
           
           
               
               
               
               
               
            
               
                   
                 Residue Gas Compression 
                 17,363 
                 HP 
                 [28,544 kW] 
               
               
                   
                 Refrigerant Compression 
                 6,214 
                 HP 
                 [10,216 kW] 
               
               
                   
                 Total Compression 
                 23,577 
                 HP 
                 [38,760 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     A comparison of Tables I and II shows that, compared to the  FIG. 1  process, the  FIG. 2  process has slightly lower ethane recovery (86.57% versus 87.06%), but improves propane recovery from 98.92% to 100.00% and butanes+recovery from 99.88% to 100.00%. Comparison of Tables I and II further shows that the improvement in yields for the  FIG. 2  process was achieved using essentially the same power requirements. 
       FIG. 3  represents an alternative prior art process according to U.S. Pat. No. 7,191,617. The process of  FIG. 3  has been applied to the same feed gas composition and conditions as described above for  FIGS. 1 and 2 . In the simulation of this process, as in the simulation for the processes of  FIGS. 1 and 2 , operating conditions were selected to maximize the recovery level for a given energy consumption. 
     After feed stream  31  is divided into two portions, streams  32  and  33 , stream  32  is cooled to substantial condensation in heat exchanger  13  by heat exchange with cold residue gas (stream  41 ), flashed separator liquids (stream  35   a ), and propane refrigerant. The resulting substantially condensed stream  32   a  at −120° F. [−84° C.] is then flash expanded through expansion valve  17  to the operating pressure (approximately 324 psia [2,235 kPa(a)]) of fractionation tower  19 . During expansion a portion of the stream is vaporized, resulting in cooling of expanded stream  32   b  to −153° F. [−103° C.] before it is supplied to fractionation tower  19  at an upper mid-column feed point. 
     The remaining portion of feed stream  31 , stream  33 , is cooled in heat exchanger  10  by heat exchange with cool residue gas (stream  41   a ) and propane refrigerant, and cooled stream  33   a  enters separator  12  at −34° F. [−36° C.] and 870 psia [6,002 kPa(a)] where the vapor (stream  34 ) is separated from the condensed liquid (stream  35 ). The separator liquid (stream  35 ) is expanded to slightly above the operating pressure of fractionation tower  19  by expansion valve  18 . Expanded stream  35   a  is heated from −66° F. [−54° C.] to 21° F. [−6° C.] in heat exchanger  13  as described earlier before heated stream  35   b  is supplied to fractionation tower  19  at a lower mid-column feed point. 
     The vapor (stream  34 ) from separator  12  enters a work expansion machine  15  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  15  expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream  34   a  to a temperature of approximately −100° F. [−74° C.]. The partially condensed expanded stream  34   a  is thereafter supplied as feed to fractionation tower  19  at an intermediate mid-column feed point. 
     A portion of the distillation vapor (stream  49 ) is withdrawn from an upper region of the demethanizing section in fractionation column  19 , below the feed position of expanded stream  34   a . Distillation vapor stream  49  is then cooled from −100° F. [−73° C.] to −146° F. [−99° C.] and partially condensed (stream  49   a ) in heat exchanger  24  by heat exchange with the cold demethanizer overhead stream  48  exiting the top of demethanizer  19  at −150° F. [−101° C.]. The cold demethanizer overhead stream is warmed to −118° F. [−84° C.] (stream  48   a ) as it cools and condenses a portion of stream  49 . 
     The operating pressure in reflux separator  25  is maintained slightly below the operating pressure of demethanizer  19 . This provides the driving force which causes distillation vapor stream  49  to flow through heat exchanger  24  and thence into the reflux separator  25  where the condensed liquid (stream  51 ) is separated from the uncondensed vapor (stream  50 ). Stream  50  then combines with the warmed demethanizer overhead stream  48   a  from heat exchanger  24  to form cold residue gas stream  41  at −123° F. [−86° C.]. 
     The liquid stream  51  from reflux separator  25  is pumped by pump  26  to a pressure slightly above the operating pressure of demethanizer  19 , and stream  51   a  is then supplied as cold top column feed (reflux) to demethanizer  19  at −145° F. [−98° C.]. This cold liquid reflux absorbs and condenses the C 2  components, C 3  components, and heavier components rising in the upper region of the rectifying section of demethanizer  19 . 
     The liquid product (stream  44 ) exits the bottom of tower  19  at 81° F. [27° C.], based on a methane concentration of 0.07% on a volume basis in the bottom product. The cold residue gas stream  41  passes countercurrently to one portion of the feed gas in heat exchanger  13  where it is heated to −35° F. [−37° C.] (stream  41   a ), and countercurrently to the other portion of the feed gas in heat exchanger  10  where it is heated to 88° F. [31° C.] (stream  41   b ). The residue gas is then re-compressed in two stages, compressor  16  driven by expansion machine  15  and compressor  22  driven by a supplemental power source. After cooling to 110° F. [43° C.] in discharge cooler  23 , the residue gas product (stream  41   e ) flows to the sales gas pipeline at 918 psia [6,330 kPa(a)]. 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 3  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE III 
               
               
                   
               
               
                 (FIG. 3) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
            
               
                   
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 31 
                 22,875 
                 3,898 
                 1,843 
                 1,240 
                 32,972 
               
               
                 32 
                 6,062 
                 1,033 
                 488 
                 329 
                 8,737 
               
               
                 33 
                 16,813 
                 2,865 
                 1,355 
                 911 
                 24,235 
               
               
                 34 
                 14,234 
                 1,525 
                 360 
                 86 
                 18,356 
               
               
                 35 
                 2,579 
                 1,340 
                 995 
                 825 
                 5,879 
               
               
                 49 
                 4,853 
                 364 
                 18 
                 1 
                 5,443 
               
               
                 50 
                 3,384 
                 49 
                 0 
                 0 
                 3,622 
               
               
                 51 
                 1,469 
                 315 
                 18 
                 1 
                 1,821 
               
               
                 48 
                 19,483 
                 288 
                 2 
                 0 
                 22,698 
               
               
                 41 
                 22,867 
                 337 
                 2 
                 0 
                 26,320 
               
               
                 44 
                 8 
                 3,561 
                 1,841 
                 1,240 
                 6,652 
               
               
                   
               
            
           
           
               
            
               
                 Recoveries* 
               
               
                   
               
            
           
           
               
               
               
            
               
                   
                 Ethane 
                 91.35% 
               
               
                   
                 Propane 
                 99.88% 
               
               
                   
                 Butanes+ 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
               
                   
               
            
           
           
               
               
               
               
               
            
               
                   
                 Residue Gas Compression 
                 15,932 
                 HP 
                 [26,192 kW] 
               
               
                   
                 Refrigerant Compression 
                 7,640 
                 HP 
                 [12,560 kW] 
               
               
                   
                 Total Compression 
                 23,572 
                 HP 
                 [38,752 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     A comparison of Tables I, II, and III shows that the  FIG. 3  process improves the ethane recovery from 87.06% (for  FIGS. 1 ) and 86.57% (for  FIGS. 2 ) to 91.35%. The propane recovery for the  FIG. 3  process (99.88%) is significantly higher than that of the  FIG. 1  process (98.92%) but slightly lower than that of the  FIG. 2  process (100.00%). The butanes+recovery for the  FIG. 3  process (100.00%) is slightly higher than that of the  FIG. 1  process (99.88%) and the same as that of the  FIG. 2  process (100.00%). Comparison of Tables I, II, and III further shows that the improvement in yields for the  FIG. 3  process was achieved using essentially the same power requirements. 
       FIG. 4  represents an alternative prior art process according to co-pending application Ser. No. 11/839,693. The process of  FIG. 4  has been applied to the same feed gas composition and conditions as described above for  FIGS. 1 through 3 . In the simulation of this process, as in the simulation for the process of  FIGS. 1 through 3 , operating conditions were selected to maximize the recovery level for a given energy consumption. 
     After feed stream  31  is divided into two portions, streams  32  and  33 , stream  32  is cooled to substantial condensation in heat exchanger  13  by heat exchange with cold residue gas (stream  41 ), flashed separator liquids (stream  35   a ), and propane refrigerant. The resulting substantially condensed stream  32   a  at −151° F. [−101° C.] is then flash expanded through expansion valve  17  to the operating pressure (approximately 319 psia [2,202 kPa(a)]) of fractionation tower  19 . During expansion a portion of the stream is vaporized, resulting in cooling of expanded stream  32   b  to −165° F. [−109° C.] before it is supplied to fractionation tower  19  at an upper mid-column feed point. 
     The remaining portion of feed stream  31 , stream  33 , is cooled in heat exchanger  10  by heat exchange with cool residue gas (stream  41   a ) and propane refrigerant, and cooled stream  33   a  enters separator  12  at −40° F. [−40° C.] and 870 psia [6,002 kPa(a)] where the vapor (stream  34 ) is separated from the condensed liquid (stream  35 ). The separator liquid (stream  35 ) is expanded to slightly above the operating pressure of fractionation tower  19  by expansion valve  18 . Expanded stream  35   a  is heated from −73° F. [−59° C.] to 4° F. [−16° C.] in heat exchanger  13  as described earlier before heated stream  35   b  is supplied to fractionation tower  19  at a lower mid-column feed point. 
     The vapor (stream  34 ) from separator  12  enters a work expansion machine  15  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  15  expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream  34   a  to a temperature of approximately −107° F. [−77° C.]. The partially condensed expanded stream  34   a  is thereafter supplied as feed to fractionation tower  19  at an intermediate mid-column feed point. 
     A portion of the distillation vapor (stream  49 ) is withdrawn from an upper region of the demethanizing section in fractionation column  19  below expanded stream  34   a  at −102° F. [−75° C.] and is compressed to approximately 486 psia [3,353 kPa(a)] by vapor compressor  27 . The compressed stream  49   a  is then cooled from −52° F. [−47° C.] to −151° F. [−101° C.] and substantially condensed (stream  49   b ) in heat exchanger  13  as described earlier. 
     Since substantially condensed stream  49   b  is at a pressure greater than the operating pressure of demethanizer  19 , it is flash expanded through expansion valve  14  to the operating pressure of fractionation tower  19 . During expansion a portion of the stream is vaporized, resulting in cooling of the total stream to −159° F. [−106° C.]. The expanded stream  49   c  is then supplied as cold top column feed (reflux) to demethanizer  19 . The vapor portion of stream  49   c  combines with the distillation vapor rising from the upper fractionation stage to form residue gas stream  41  exiting the top of demethanizer  19  at −154° F. [−103° C.], while the cold liquid reflux portion absorbs and condenses the C 2  components, C 3  components, and heavier components rising in the upper region of the rectifying section of demethanizer  19 . 
     The liquid product (stream  44 ) exits the bottom of tower  19  at 79° F. [26° C.], based on a methane concentration of 0.07% on a volume basis in the bottom product. The cold residue gas stream  41  passes countercurrently to one portion of the feed gas in heat exchanger  13  where it is heated to −56° F. [−49° C.] (stream  41   a ), and countercurrently to the other portion of the feed gas in heat exchanger  10  where it is heated to 92° F. [33° C.] (stream  41   b ). The residue gas is then re-compressed in two stages, compressor  16  driven by expansion machine  15  and compressor  22  driven by a supplemental power source. After cooling to 110° F. [43° C.] in discharge cooler  23 , the residue gas product (stream  41   e ) flows to the sales gas pipeline at 918 psia [6,330 kPa(a)]. 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 4  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE IV 
               
               
                   
               
               
                 (FIG. 4) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
            
               
                   
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 31 
                 22,875 
                 3,898 
                 1,843 
                 1,240 
                 32,972 
               
               
                 32 
                 4,529 
                 772 
                 365 
                 245 
                 6,529 
               
               
                 33 
                 18,346 
                 3,126 
                 1,478 
                 995 
                 26,443 
               
               
                 34 
                 15,105 
                 1,522 
                 341 
                 79 
                 19,366 
               
               
                 35 
                 3,241 
                 1,604 
                 1,137 
                 916 
                 7,077 
               
               
                 49 
                 2,429 
                 179 
                 8 
                 0 
                 2,722 
               
               
                 41 
                 22,867 
                 263 
                 1 
                 0 
                 26,244 
               
               
                 44 
                 8 
                 3,635 
                 1,842 
                 1,240 
                 6,728 
               
               
                   
               
            
           
           
               
            
               
                 Recoveries* 
               
               
                   
               
            
           
           
               
               
               
            
               
                   
                 Ethane 
                 93.26% 
               
               
                   
                 Propane 
                 99.95% 
               
               
                   
                 Butanes+ 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
               
                   
               
            
           
           
               
               
               
               
               
               
            
               
                   
                 Residue Gas Compression 
                 15,859 
                 HP 
                 [26,072 
                 kW] 
               
               
                   
                 Vapor Compression 
                 351 
                 HP 
                 [577 
                 kW] 
               
               
                   
                 Refrigerant Compression 
                 7,366 
                 HP 
                 [12,110 
                 kW] 
               
               
                   
                 Total Compression 
                 23,576 
                 HP 
                 [38,759 
                 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     A comparison of Tables I, II, III, and IV shows that the  FIG. 4  process improves the ethane recovery from 87.06% (for  FIG. 1 ), 86.57% (for  FIGS. 2 ), and 91.35% (for  FIGS. 3 ) to 93.26%. The propane recovery for the  FIG. 4  process (99.95%) is significantly higher than that of the  FIG. 1  process (98.92%) and higher than that of the  FIG. 3  process (99.88%), but slightly lower than that of the  FIG. 2  process (100.00%). The butanes+recovery for the  FIG. 4  process (100.00%) is slightly higher than that of the  FIG. 1  process (99.88%) and the same as that of the  FIG. 2  process and the  FIG. 3  process (100.00% for both). Comparison of Tables I, II, III, and IV further shows that the improvement in yields for the  FIG. 4  process was achieved using essentially the same power requirements. 
     DESCRIPTION OF THE INVENTION 
       FIG. 5  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. 5  are the same as those in  FIGS. 1 through 4 . Accordingly, the  FIG. 5  process can be compared with that of the  FIGS. 1 through 4  processes to illustrate the advantages of the present invention. 
     In the simulation of the  FIG. 5  process, inlet gas enters the plant at 104° F. [40° C.] and 896 psia [6,181 kPa(a)] as stream  31 . After feed stream  31  is divided into two portions, streams  32  and  33 , stream  32  is cooled to substantial condensation in heat exchanger  13  by heat exchange with cold residue gas (stream  41 ), flashed separator liquids (stream  35   a ), and propane refrigerant. The resulting substantially condensed stream  32   a  at −148° F. [−100° C.] is then flash expanded through expansion valve  17  to the operating pressure (approximately 324 psia [2,235 kPa(a)]) of fractionation tower  19 . During expansion a portion of the stream is vaporized, resulting in cooling of expanded stream  32   b  to −164° F. [−109° C.] before it is supplied to fractionation tower  19  at an upper mid-column feed point. 
     The remaining portion of feed stream  31 , stream  33 , is cooled in heat exchanger  10  by heat exchange with cool residue gas (stream  41   a ) and propane refrigerant, and cooled stream  33   a  enters separator  12  at −36° F. [−38° C.] and 870 psia [6,002 kPa(a)] where the vapor (stream  34 ) is separated from the condensed liquid (stream  35 ). The separator liquid (stream  35 ) is expanded to slightly above the operating pressure of fractionation tower  19  by expansion valve  18 . Expanded stream  35   a  is heated from −68° F. [−56° C.] to 10° F. [−12° C.] in heat exchanger  13  as described earlier before heated stream  35   b  is supplied to fractionation tower  19  at a lower mid-column feed point. 
     The vapor (stream  34 ) from separator  12  is divided into two streams,  36  and  39 . Stream  36 , containing about 16% of the total vapor, is cooled to −148° F. [−100° C.] and substantially condensed (stream  36   a ) in heat exchanger  13  as described earlier. Substantially condensed stream  36   a  is flash expanded through expansion valve  14  to the operating pressure of fractionation tower  19 . During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in  FIG. 5 , the expanded stream  36   b  leaving expansion valve  14  reaches a temperature of −169° F. [−112° C.] before it is supplied as cold top column feed (reflux) to demethanizer  19 . This cold liquid reflux absorbs and condenses the C 2  components, C 3  components, and heavier components rising in the upper region of the rectifying section of demethanizer  19 . 
     The remaining 84% of the vapor from separator  12  (stream  39 ) enters a work expansion machine  15  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  15  expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream  39   a  to a temperature of approximately −103° F. [−75° C.]. The partially condensed expanded stream  39   a  is thereafter supplied as feed to fractionation tower  19  at an intermediate mid-column feed point. 
     The demethanizer in tower  19  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 three sections. The upper section  19   a  is a separator wherein the partially vaporized top feed (stream  36   b ) is separated into its respective vapor and liquid portions, and wherein the vapor rising from the rectifying section  19   b  below is combined with the vapor portion of the top feed to form the cold demethanizer overhead vapor (stream  41 ) which exits the top of the tower at −151° F. [−102° C.]. The middle absorbing (rectifying) section  19   b  contains the trays and/or packing to provide the necessary contact between the vapor portions of the expanded streams  32   b ,  39   a , and  35   b  rising upward and cold liquid falling downward. The lower stripping (demethanizing) section  19   c  contains additional trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. Demethanizing section  19   c  also includes one or more reboilers (such as the reboiler  21  and side reboiler  20  shown in  FIG. 5 ) 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  44 , of methane and lighter components. Stream  39   a  enters demethanizer  19  at an intermediate feed position located below rectifying section  19   b  and above demethanizing section  19   c . The liquid portion of the expanded stream  39   a  commingles with liquids falling downward from rectifying section  19   b  and the combined liquid continues downward into demethanizing section  19   c  of column  19 . The vapor portion of the expanded stream  39   a  commingles with vapors rising upward from demethanizing section  19   c  and the combined vapor rises upward through rectifying section  19   b  and is contacted with cold liquid falling downward to condense and absorb the C 2  components, C 3  components, and heavier components from the vapor. 
     In demethanizing section  19   c  of demethanizer  19 , the feed streams are stripped of their methane and lighter components. The resulting liquid product (stream  44 ) exits the bottom of tower  19  at 80° F. [27° C.] based on a methane concentration of 0.07% on a volume basis in the bottom product. The cold residue gas stream  41  leaves demethanizer  19  and passes countercurrently to one portion of the feed gas and a portion of the separator vapor in heat exchanger  13  where it is heated to −44° F. [−42° C.] (stream  41   a ) and countercurrently to the other portion of the feed gas in heat exchanger  10  where it is heated to 91° F. [33° C.] (stream  41   b ) as it provides cooling as previously described. The residue gas is then re-compressed in two stages, compressor  16  driven by expansion machine  15  and compressor  22  driven by a supplemental power source. After cooling to 110° F. [43° C.] in discharge cooler  23 , the residue gas product (stream  41   e ) flows to the sales gas pipeline at 918 psia [6,330 kPa(a)]. 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 5  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE V 
               
               
                   
               
               
                 (FIG. 5) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
            
               
                   
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 31 
                 22,875 
                 3,898 
                 1,843 
                 1,240 
                 32,972 
               
               
                 32 
                 4,804 
                 819 
                 387 
                 260 
                 6,924 
               
               
                 33 
                 18,071 
                 3,079 
                 1,456 
                 980 
                 26,048 
               
               
                 34 
                 15,151 
                 1,588 
                 368 
                 87 
                 19,496 
               
               
                 35 
                 2,920 
                 1,491 
                 1,088 
                 893 
                 6,552 
               
               
                 36 
                 2,379 
                 249 
                 58 
                 14 
                 3,061 
               
               
                 39 
                 12,772 
                 1,339 
                 310 
                 73 
                 16,435 
               
               
                 41 
                 22,867 
                 265 
                 6 
                 0 
                 26,251 
               
               
                 44 
                 8 
                 3,633 
                 1,837 
                 1,240 
                 6,721 
               
               
                   
               
            
           
           
               
            
               
                 Recoveries* 
               
               
                   
               
            
           
           
               
               
               
            
               
                   
                 Ethane 
                 93.22% 
               
               
                   
                 Propane 
                 99.70% 
               
               
                   
                 Butanes+ 
                 99.99% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
               
                   
               
            
           
           
               
               
               
               
               
            
               
                   
                 Residue Gas Compression 
                 15,881 
                 HP 
                 [26,108 kW] 
               
               
                   
                 Refrigerant Compression 
                 7,678 
                 HP 
                 [12,623 kW] 
               
               
                   
                 Total Compression 
                 23,559 
                 HP 
                 [38,731 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     A comparison of Tables I through V shows that, compared to the prior art, the present invention improves the ethane recovery from 87.06% (for  FIG. 1 ), 86.57% (for  FIGS. 2 ), and 91.35% (for  FIGS. 3 ) to 93.22%, with essentially the same recovery as the  FIG. 4  process (93.26%). The propane recovery for the present invention (99.70%) is significantly higher than that of the  FIG. 1  process (98.92%) and somewhat lower than that of the  FIG. 2  process (100.00%), the  FIG. 3  process (99.88%), and the  FIG. 4  process (99.95%). The butanes+recovery for the present invention (99.99%) is slightly higher than that of the  FIG. 1  process (99.88%) and essentially the same as that of the  FIG. 2  process, the  FIG. 3  process, and the  FIG. 4  process (100.00% for all three). Comparison of Tables I through V further shows that the improvement in yields for the present invention shown in  FIG. 5  was achieved using essentially the same power requirements as the  FIG. 1 through 4  processes. 
     The improvement in the recovery efficiency of the present invention over that of the prior art processes can be understood by examining the improvement in the rectification that the present invention provides for rectifying section  19   b . Compared to the prior art of the  FIG. 1  process, the present invention produces a better top reflux stream containing more methane and less C 2 + components. Comparing reflux stream  32  in Table I for the  FIG. 1  prior art process with first reflux stream  36  in Table V for the present invention, it can be seen that the present invention provides a top reflux stream with a significantly lower concentration of C 2 + components (10.5% for the present invention versus 21.2% for the  FIG. 1  prior art process). Further, the present invention uses a second reflux stream (stream  32 ) at an intermediate feed point of rectifying section  19   b  to provide bulk recovery of the C 2  components, C 3  components, and heavier hydrocarbon components contained in expanded feed  39   a  and the vapors rising from demethanizing section  19   c . This means that less rectification is required in the upper region of rectifying section  19   b , minimizing the required flow rate for top reflux stream  36  so that more of the separator vapor (stream  34 ) flows to expansion machine  15  in stream  39 , which produces more power to drive compressor  16  and reduces the power required by compressor  22 . Note that the total reflux provided to rectification section  19   b  in streams  36  and  32  for the present invention is nearly 8% higher than that provided in the single reflux stream  32  of the  FIG. 1  process. 
     Compared to the prior art of the  FIG. 2  process, the present invention supplies its feed streams to column  19  at significantly lower temperatures, reducing the quantity of vapor entering rectification section  19   b  and reducing the quantity of reflux required. Recycle stream  48  used in the  FIG. 2  process to produce top reflux stream  48   c  adds to the cooling load imposed on cold demethanizer overhead stream  45 , reducing the cooling available in heat exchanger  10  such that the portion of the feed gas in stream  33   a  entering separator  12  for the  FIG. 2  process is much warmer than that of the present invention. As can be seen from Table II for the  FIG. 2  process, despite top reflux stream  48  having almost 15% more flow than the present invention and a much lower concentration of C 2 + components (2.0% for the  FIG. 2  prior art process versus 10.5% for the present invention), it cannot rectify the vapors rising in rectification section  19   b  as well as top reflux stream  36  of the present invention. 
     Compared to the prior art of the  FIG. 3  process, the present invention produces a much greater quantity of top reflux containing more methane and less C 2 + components. Comparing reflux stream  51  in Table III for the  FIG. 3  prior art process with top reflux stream  36  in Table V for the present invention, it can be seen that the present invention provides a top reflux stream with 68% more flow and with a significantly lower concentration of C 2 + components (10.5% for the present invention versus 18.3% for the  FIG. 3  prior art process). Note that although the total reflux provided to rectification section  19   b  in streams  51  and  32  of the  FIG. 3  process is nearly 6% higher than that of the present invention, the much higher concentration of C 2  components in top reflux stream  51  of the  FIG. 3  process (nearly twice that in top reflux stream  36  of the present invention) prevents it from providing efficient rectification. 
     Compared to the prior art of the  FIG. 4  process, the present invention achieves essentially the same rectification with its two reflux streams. Comparing reflux stream  49  in Table IV for the  FIG. 4  prior art process with top reflux stream  36  in Table V for the present invention, it can be seen that the present invention provides a top reflux stream with 12% more flow but with a higher concentration of C 2 + components (10.5% for the present invention versus 6.9% for the  FIG. 4  prior art process). Further, the total reflux provided to rectification section  19   b  for the present invention is nearly 8% higher than that in streams  49  and  32  of the  FIG. 4  process. This higher reflux flow allows the present invention to match the rectification of the  FIG. 4  prior art process despite the higher concentration of C 2 + components in its top reflux stream. 
     However, the present invention is able to match the recovery of the  FIG. 4  process without vapor compressor  27  required by the prior art process. Vapor compressors in this service are expensive to install and to operate, adding to both the capital cost and the maintenance cost of the plant, reducing revenue and reducing the return on investment. Rotating equipment like this also adds to the complexity of operating the plant, making it more difficult to optimize the process for maximum recovery with minimum energy consumption. Thus, the present invention is less expensive to build, less expensive to maintain, and easier to operate than the prior art of the  FIG. 4  process. 
     Other Embodiments 
       FIGS. 6 through 9  display other embodiments of the present invention.  FIGS. 5 through 7  depict fractionation towers constructed in a single vessel.  FIGS. 8 and 9  depict fractionation towers constructed in two vessels, absorber (rectifying) column  28  (a contacting and separating device) and partial rectification stripper (distillation) column  19 . In such cases, the overhead vapor stream  52  from partial rectification stripper column  19  flows to the lower section of absorber column  28  to be contacted and further rectified by substantially condensed stream  36   b . Pump  29  is used to route the liquids (stream  53 ) from the bottom of absorber column  28  to the top of stripper column  19  so that the two towers effectively function as one distillation system. The decision whether to construct the fractionation tower as a single vessel (such as demethanizer  19  in  FIGS. 5 through 7 ) or multiple vessels will depend on a number of factors such as plant size, the distance to fabrication facilities, etc. 
     In accordance with this invention, it is generally advantageous to design the absorbing (rectifying) 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 expanded substantially condensed stream  36   b  leaving expansion valve  14  and all or a part of expanded substantially condensed stream  32   b  ( FIG. 5 ) or  38   b  ( FIGS. 6 through 9 ) from expansion valve  17  can be combined (such as in the piping joining the expansion valves 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, combined with contacting at least a portion of expanded stream  39   a , shall be considered for the purposes of this invention as constituting an absorbing section. 
     In accordance with the present invention, the splitting of the feed gas may be accomplished in several ways. In the processes of  FIGS. 5, 6, and 8 , the splitting of the feed gas occurs before any cooling of the feed gas. In such cases, cooling and substantial condensation of one portion of the feed gas in multiple heat exchangers may be favored in some circumstances, such as heat exchangers  11  and  13  shown in  FIGS. 6 and 8 . The feed gas may also be split, however, following cooling (but prior to separation of any liquids which may have been formed) as shown in  FIGS. 7 and 9 . 
     The high pressure liquid (stream  35  in  FIGS. 5 through 9 ) need not be expanded, heated, and fed to a mid-column feed point on the distillation column. Instead, all or a portion of it may be combined with the portion of the cooled feed gas (stream  32   a  in  FIGS. 6 and 8  or stream  32  in  FIGS. 7 and 9 ) flowing to heat exchanger  13 . (This is shown by the dashed stream  37  in  FIGS. 6 through 9 .) Any remaining portion of the liquid (stream  40  in  FIGS. 6 through 9 ) may be expanded through an appropriate expansion device, such as expansion valve  18  or an expansion machine, and fed to a mid-column feed point on the distillation column (stream  40 a). Stream  40  may also be used for inlet gas cooling or other heat exchange service before or after the expansion step prior to flowing to the demethanizer. 
     As described earlier, a portion of the feed gas (stream  32 ) and a portion of the separator vapor (stream  36 ) are substantially condensed and the resulting condensate used to absorb valuable C 2  components, C 3  components, and heavier components from the vapors rising through rectifying section  19   b  of demethanizer  19  ( FIGS. 5 through 7 ), or through absorber column  28  and the upper section of partial rectification stripper column  19  ( FIGS. 8 and 9 ). However, the present invention is not limited to this embodiment. It may be advantageous, for instance, to treat only a portion of these vapors in this manner, or to use only a portion of the condensate as an absorbent, in cases where other design considerations indicate portions of the vapors or the condensate should bypass rectifying section  19   b  of demethanizer  19  ( FIGS. 5 through 7 ), or absorber column  28  and/or partial rectification stripper column  19  ( FIGS. 8 and 9 ). 
     Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine  15 , 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 substantially condensed portion of the separator vapor (stream  36   a  in  FIGS. 5 through 9 ), the substantially condensed portion of the feed stream (stream  32   a  in  FIG. 5 ), or the substantially condensed combined stream (stream  38   a  in  FIGS. 6 through 9 ). 
     In accordance with the present invention, the use of external refrigeration to supplement the cooling available to the inlet gas and/or separator vapor 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 and separator vapor cooling must be evaluated for each particular application, as well as the choice of process streams for specific heat exchange services. 
     It will also be recognized that the relative amount of feed found in each branch of the split vapor feeds 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 recompression 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 allows reduced capital expenditures and/or provides improved recovery of C 2  components, C 3  components, and heavier hydrocarbon components or of 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 or 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. 
     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.