Abstract:
A process and an apparatus are disclosed for the recovery of ethane, ethylene, propane, propylene, and heavier hydrocarbon components from a hydrocarbon gas stream in a compact processing assembly. The gas stream is cooled and divided into first and second streams. The first stream is further cooled to condense substantially all of it and is thereafter expanded to lower pressure and supplied as a feed between first and second absorbing means inside the processing assembly. The second stream is expanded to lower pressure and supplied as the bottom feed to the second absorbing means. A distillation vapor stream is collected from the upper region of the first absorbing means and directed into one or more heat exchange means inside the processing assembly to heat it while cooling the gas stream and the first stream. The heated distillation vapor stream is compressed to higher pressure and divided into a volatile residue gas fraction and a compressed recycle stream. The compressed recycle stream is cooled to condense substantially all of it by the distillation vapor stream in the one or more heat exchange means inside the processing assembly, and is thereafter expanded to lower pressure and supplied as top feed to the first absorbing means. A distillation liquid stream is collected from the lower region of the second absorbing means and directed into a heat and mass transfer means inside the processing assembly to heat it and strip out its volatile components while cooling the gas stream. The quantities and temperatures of the feeds to the first and second absorbing means are effective to maintain the temperature of the upper region of the first absorbing means at a temperature whereby the major portions of the desired components are recovered in the stripped distillation liquid stream.

Description:
This invention relates to a process and apparatus for the separation of a gas containing hydrocarbons. The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application No. 61/186,361 which was filed on Jun. 11, 2009. The applicants also claim the benefits under Title 35, United States Code, Section 120 as a continuation-in-part of U.S. patent application Ser. No. 12/689,616 which was filed on Jan. 19, 2010, and as a continuation-in-part of U.S. patent application Ser. No. 12/372,604 which was filed on Feb. 17, 2009. Assignees S.M.E. Products LP and Ortloff Engineers, Ltd. were parties to a joint research agreement that was in effect before the invention of this application was made. 
    
    
     BACKGROUND OF THE INVENTION 
     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, 90.3% methane, 4.0% ethane and other C 2  components, 1.7% propane and other C 3  components, 0.3% iso-butane, 0.5% normal butane, and 0.8% 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 and for processes that can provide efficient recoveries with lower capital investment. 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; and 12/206,230 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). 
     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; and 5,881,569, co-pending application Ser. Nos. 11/430,412 and 11/971,491, 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. 
     The present invention employs a novel means of performing the various steps described above more efficiently and using fewer pieces of equipment. This is accomplished by combining what heretofore have been individual equipment items into a common housing, thereby reducing the plot space required for the processing plant and reducing the capital cost of the facility. Surprisingly, applicants have found that the more compact arrangement also significantly reduces the power consumption required to achieve a given recovery level, thereby increasing the process efficiency and reducing the operating cost of the facility. In addition, the more compact arrangement also eliminates much of the piping used to interconnect the individual equipment items in traditional plant designs, further reducing capital cost and also eliminating the associated flanged piping connections. Since piping flanges are a potential leak source for hydrocarbons (which are volatile organic compounds, VOCs, that contribute to greenhouse gases and may also be precursors to atmospheric ozone formation), eliminating these flanges reduces the potential for atmospheric emissions that can damage the environment. 
     In accordance with the present invention, it has been found that C 2  recoveries in excess of 95% can be obtained. Similarly, in those instances where recovery of C 2  components is not desired, C 3  recoveries in excess of 95% can be maintained. In addition, the present invention makes possible essentially 100% separation of methane (or C 2  components) and lighter components from the C 2  components (or C 3  components) and heavier components at lower energy requirements compared to the prior art while maintaining the same recovery level. 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. 5,568,737; 
         FIG. 2  is a flow diagram of a natural gas processing plant in accordance with the present invention; and 
         FIGS. 3 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&#39;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. 5,568,737. In this simulation of the process, inlet gas enters the plant at 110° F. [43° C.] and 915 psia [6,307 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 −26° F. [−32° C.] in heat exchanger  10  by heat exchange with cool distillation vapor stream  41   a , while stream  33  is cooled to −32° F. [−35° C.] in heat exchanger  11  by heat exchange with demethanizer reboiler liquids at 41° F. [5° C.] (stream  43 ) and side reboiler liquids at −49° F. [−45° C.] (stream  42 ). Streams  32   a  and  33   a  recombine to form stream  31   a , which enters separator  12  at −28° F. [−33° C.] and 893 psia [6,155 kPa(a)] where the vapor (stream  34 ) is separated from the condensed liquid (stream  35 ). 
     The vapor (stream  34 ) from separator  12  is divided into two streams,  36  and  39 . Stream  36 , containing about 27% of the total vapor, is combined with the separator liquid (stream  35 ), and the combined stream  38  passes through heat exchanger  13  in heat exchange relation with cold distillation vapor stream  41  where it is cooled to substantial condensation. The resulting substantially condensed stream  38   a  at −139° F. [−95° C.] is then flash expanded through expansion valve  14  to the operating pressure (approximately 396 psia [2,730 kPa(a)]) of fractionation tower  18 . 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  38   b  leaving expansion valve  14  reaches a temperature of −140° F. [−95° C.] and is supplied to fractionation tower  18  at a first mid-column feed point. 
     The remaining 73% 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 −95° F. [−71° 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 heated distillation vapor stream (stream  41   b ), for example. The partially condensed expanded stream  39   a  is thereafter supplied as feed to fractionation tower  18  at a second mid-column feed point. 
     The recompressed and cooled distillation vapor stream  41   e  is divided into two streams. One portion, stream  46 , is the volatile residue gas product. The other portion, recycle stream  45 , flows to heat exchanger  10  where it is cooled to −26° F. [−32° C.] by heat exchange with cool distillation vapor stream  41   a . The cooled recycle stream  45   a  then flows to exchanger  13  where it is cooled to −139° F. [−95° C.] and substantially condensed by heat exchange with cold distillation vapor stream  41 . The substantially condensed stream  45   b  is then expanded through an appropriate expansion device, such as expansion valve  22 , to the demethanizer operating pressure, resulting in cooling of the total stream to −147° F. [−99° C.]. The expanded stream  45   c  is then supplied to fractionation tower  18  as the top column feed. The vapor portion (if any) of stream  45   c  combines with the vapors rising from the top fractionation stage of the column to form distillation vapor stream  41 , which is withdrawn from an upper region of the tower. 
     The demethanizer in tower  18  is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. As is often the case in natural gas processing plants, the fractionation tower may consist of two sections. The upper section  18   a  is a separator wherein the partially vaporized top feed is divided into its respective vapor and liquid portions, and wherein the vapor rising from the lower distillation or demethanizing section  18   b  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 −144° F. [−98° C.]. The lower, demethanizing section  18   b  contains the trays and/or packing and provides the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section  18   b  also includes reboilers (such as the reboiler and the 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  44 , of methane and lighter components. 
     The liquid product stream  44  exits the bottom of the tower at 64° F. [18° C.], based on a typical specification of a methane to ethane ratio of 0.010:1 on a mass basis in the bottom product. The demethanizer overhead vapor stream  41  passes countercurrently to the incoming feed gas and recycle stream in heat exchanger  13  where it is heated to −40° F. [−40° C.] (stream  41   a ) and in heat exchanger  10  where it is heated to 104° F. [40° C.] (stream  41   b ). The distillation vapor stream is then re-compressed in two stages. The first stage is compressor  16  driven by expansion machine  15 . The second stage is compressor  20  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  21 , stream  41   e  is split into the residue gas product (stream  46 ) and the recycle stream  45  as described earlier. Residue gas stream  46  flows to the sales gas pipeline at 915 psia [6,307 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   12,398   546   233   229   13,726       32   8,431   371   159   156   9,334       33   3,967   175   74   73   4,392       34   12,195   501   179   77   13,261       35   203   45   54   152   465       36   3,317   136   49   21   3,607       38   3,520   181   103   173   4,072       39   8,878   365   130   56   9,654       41   13,765   30   0   0   13,992       45   1,377   3   0   0   1,400       46   12,388   27   0   0   12,592       44   10   519   233   229   1,134                    
Recoveries*
 
                                                 Ethane   94.99%           Propane   99.99%           Butanes+   100.00%                        
Power
 
                                         Residue Gas Compression   6,149 HP   [10,109 kW]                    
* (Based on un-rounded flow rates)
 
     DESCRIPTION OF THE INVENTION 
       FIG. 2  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. 2  are the same as those in  FIG. 1 . Accordingly, the  FIG. 2  process can be compared with that of the  FIG. 1  process to illustrate the advantages of the present invention. 
     In the simulation of the  FIG. 2  process, inlet gas enters the plant as stream  31  and is divided into two portions, streams  32  and  33 . The first portion, stream  32 , enters a heat exchange means in the upper region of feed cooling section  118   a  inside processing assembly  118 . This heat exchange means may be comprised of a fin and tube type heat exchanger, a plate type heat exchanger, a brazed aluminum type heat exchanger, or other type of heat transfer device, including multi-pass and/or multi-service heat exchangers. The heat exchange means is configured to provide heat exchange between stream  32  flowing through one pass of the heat exchange means and a distillation vapor stream arising from separator section  118   b  inside processing assembly  118  that has been heated in a heat exchange means in the lower region of feed cooling section  118   a . Stream  32  is cooled while further heating the distillation vapor stream, with stream  32   a  leaving the heat exchange means at −25° F. [−32° C.]. 
     The second portion, stream  33 , enters a heat and mass transfer means in demethanizing section  118   e  inside processing assembly  118 . This heat and mass transfer means may also be comprised of a fin and tube type heat exchanger, a plate type heat exchanger, a brazed aluminum type heat exchanger, or other type of heat transfer device, including multi-pass and/or multi-service heat exchangers. The heat and mass transfer means is configured to provide heat exchange between stream  33  flowing through one pass of the heat and mass transfer means and a distillation liquid stream flowing downward from absorbing section  118   d  inside processing assembly  118 , so that stream  33  is cooled while heating the distillation liquid stream, cooling stream  33   a  to −47° F. [−44° C.] before it leaves the heat and mass transfer means. As the distillation liquid stream is heated, a portion of it is vaporized to form stripping vapors that rise upward as the remaining liquid continues flowing downward through the heat and mass transfer means. The heat and mass transfer means provides continuous contact between the stripping vapors and the distillation liquid stream so that it also functions to provide mass transfer between the vapor and liquid phases, stripping the liquid product stream  44  of methane and lighter components. 
     Streams  32   a  and  33   a  recombine to form stream  31   a , which enters separator section  118   f  inside processing assembly  118  at −32° F. [−36° C.] and 900 psia [6,203 kPa(a)], whereupon the vapor (stream  34 ) is separated from the condensed liquid (stream  35 ). Separator section  118   f  has an internal head or other means to divide it from demethanizing section  118   e , so that the two sections inside processing assembly  118  can operate at different pressures. 
     The vapor (stream  34 ) from separator section  118   f  is divided into two streams,  36  and  39 . Stream  36 , containing about 27% of the total vapor, is combined with the separated liquid (stream  35 , via stream  37 ), and the combined stream  38  enters a heat exchange means in the lower region of feed cooling section  118   a  inside processing assembly  118 . This heat exchange means may likewise be comprised of a fin and tube type heat exchanger, a plate type heat exchanger, a brazed aluminum type heat exchanger, or other type of heat transfer device, including multi-pass and/or multi-service heat exchangers. The heat exchange means is configured to provide heat exchange between stream  38  flowing through one pass of the heat exchange means and the distillation vapor stream arising from separator section  118   b , so that stream  38  is cooled to substantial condensation while heating the distillation vapor stream. 
     The resulting substantially condensed stream  38   a  at −138° F. [−95° C.] is then flash expanded through expansion valve  14  to the operating pressure (approximately 400 psia [2,758 kPa(a)]) of rectifying section  118   c  (an absorbing means) and absorbing section  118   d  (another absorbing means) inside processing assembly  118 . During expansion a portion of the stream may be vaporized, resulting in cooling of the total stream. In the process illustrated in  FIG. 2 , the expanded stream  38   b  leaving expansion valve  14  reaches a temperature of −139° F. [−95° C.] and is supplied to processing assembly  118  between rectifying section  118   c  and absorbing section  118   d . The liquids in stream  38   b  combine with the liquids falling from rectifying section  118   c  and are directed to absorbing section  118   d , while any vapors combine with the vapors rising from absorbing section  118   d  and are directed to rectifying section  118   c.    
     The remaining 73% of the vapor from separator section  118   f  (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 operating pressure of absorbing section  118   d , with the work expansion cooling the expanded stream  39   a  to a temperature of approximately −99° F. [−73° C.]. The partially condensed expanded stream  39   a  is thereafter supplied as feed to the lower region of absorbing section  118   d  inside processing assembly  118 . 
     The recompressed and cooled distillation vapor stream  41   c  is divided into two streams. One portion, stream  46 , is the volatile residue gas product. The other portion, recycle stream  45 , enters a heat exchange means in the feed cooling section  118   a  inside processing assembly  118 . This heat exchange means may also be comprised of a fin and tube type heat exchanger, a plate type heat exchanger, a brazed aluminum type heat exchanger, or other type of heat transfer device, including multi-pass and/or multi-service heat exchangers. The heat exchange means is configured to provide heat exchange between stream  45  flowing through one pass of the heat exchange means and the distillation vapor stream arising from separator section  118   b , so that stream  45  is cooled to substantial condensation while heating the distillation vapor stream. 
     The substantially condensed recycle stream  45   a  leaves the heat exchange means in feed cooling section  118   a  at −138° F. [−95° C.] and is flash expanded through expansion valve  22  to the operating pressure of rectifying section  118   c  inside processing assembly  118 . During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in  FIG. 2 , the expanded stream  45   b  leaving expansion valve  22  reaches a temperature of −146° F. [−99° C.] and is supplied to separator section  118   b  inside processing assembly  118 . The liquids separated therein are directed to rectifying section  118   c , while the remaining vapors combine with the vapors rising from rectifying section  118   c  to form the distillation vapor stream that is heated in cooling section  118   a.    
     Rectifying section  118   c  and absorbing section  118   d  each contain an absorbing means consisting of a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The trays and/or packing in rectifying section  118   c  and absorbing section  118   d  provide the necessary contact between the vapors rising upward and cold liquid falling downward. The liquid portion of the expanded stream  39   a  commingles with liquids falling downward from absorbing section  118   d  and the combined liquid continues downward into demethanizing section  118   e . The stripping vapors arising from demethanizing section  118   e  combine with the vapor portion of the expanded stream  39   a  and rise upward through absorbing section  118   d , to be contacted with the cold liquid falling downward to condense and absorb most of the C 2  components, C 3  components, and heavier components from these vapors. The vapors arising from absorbing section  118   d  combine with any vapor portion of the expanded stream  38   b  and rise upward through rectifying section  118   c , to be contacted with the cold liquid portion of expanded stream  45   b  falling downward to condense and absorb most of the C 2  components, C 3  components, and heavier components remaining in these vapors. The liquid portion of the expanded stream  38   b  commingles with liquids falling downward from rectifying section  118   c  and the combined liquid continues downward into absorbing section  118   d.    
     The distillation liquid flowing downward from the heat and mass transfer means in demethanizing section  118   e  inside processing assembly  118  has been stripped of methane and lighter components. The resulting liquid product (stream  44 ) exits the lower region of demethanizing section  118   e  and leaves processing assembly  118  at 65° F. [18° C.]. The distillation vapor stream arising from separator section  118   b  is warmed in feed cooling section  118   a  as it provides cooling to streams  32 ,  38 , and  45  as described previously, and the resulting distillation vapor stream  41  leaves processing assembly  118  at 105° F. [40° C.]. The distillation vapor stream is then re-compressed in two stages, compressor  16  driven by expansion machine  15  and compressor  20  driven by a supplemental power source. After stream  41   b  is cooled to 110° F. [43° C.] in discharge cooler  21  to form stream  41   c , recycle stream  45  is withdrawn as described earlier, forming residue gas stream  46  which thereafter flows to the sales gas pipeline at 915 psia [6,307 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   12,398   546   233   229   13,726       32   8,679   382   163   160   9,608       33   3,719   164   70   69   4,118       34   12,164   495   174   72   13,213       35   234   51   59   157   513       36   3,248   132   46   19   3,528       37   234   51   59   157   513       38   3,482   183   105   176   4,041       39   8,916   363   128   53   9,685       40   0   0   0   0   0       41   13,863   30   0   0   14,095       45   1,475   3   0   0   1,500       46   12,388   27   0   0   12,595       44   10   519   233   229   1,131                    
Recoveries*
 
                                                 Ethane   95.03%           Propane   99.99%           Butanes+   100.00%                        
Power
 
                                         Residue Gas Compression   5,787 HP   [9,514 kW]                    
* (Based on un-rounded flow rates)
 
     A comparison of Tables I and II shows that the present invention maintains essentially the same recoveries as the prior art. However, further comparison of Tables I and II shows that the product yields were achieved using significantly less power than the prior art. In terms of the recovery efficiency (defined by the quantity of ethane recovered per unit of power), the present invention represents more than a 6% improvement over the prior art of the  FIG. 1  process. 
     The improvement in recovery efficiency provided by the present invention over that of the prior art of the  FIG. 1  process is primarily due to two factors. First, the compact arrangement of the heat exchange means in feed cooling section  118   a  and the heat and mass transfer means in demethanizing section  118   e  in processing assembly  118  eliminates the pressure drop imposed by the interconnecting piping found in conventional processing plants. The result is that the portion of the feed gas flowing to expansion machine  15  is at higher pressure for the present invention compared to the prior art, allowing expansion machine  15  in the present invention to produce as much power with a higher outlet pressure as expansion machine  15  in the prior art can produce at a lower outlet pressure. Thus, rectifying section  118   c  and absorbing section  118   d  in processing assembly  118  of the present invention can operate at higher pressure than fractionation column  18  of the prior art while maintaining the same recovery level. This higher operating pressure, plus the reduction in pressure drop for the distillation vapor stream due to eliminating the interconnecting piping, results in a significantly higher pressure for the distillation vapor stream entering compressor  20 , thereby reducing the power required by the present invention to restore the residue gas to pipeline pressure. 
     Second, using the heat and mass transfer means in demethanizing section  118   e  to simultaneously heat the distillation liquid leaving absorbing section  118   d  while allowing the resulting vapors to contact the liquid and strip its volatile components is more efficient than using a conventional distillation column with external reboilers. The volatile components are stripped out of the liquid continuously, reducing the concentration of the volatile components in the stripping vapors more quickly and thereby improving the stripping efficiency for the present invention. 
     The present invention offers two other advantages over the prior art in addition to the increase in processing efficiency. First, the compact arrangement of processing assembly  118  of the present invention replaces five separate equipment items in the prior art (heat exchangers  10 ,  11 , and  13 ; separator  12 ; and fractionation tower  18  in  FIG. 1 ) with a single equipment item (processing assembly  118  in  FIG. 2 ). This reduces the plot space requirements and eliminates the interconnecting piping, reducing the capital cost of a process plant utilizing the present invention over that of the prior art. Second, elimination of the interconnecting piping means that a processing plant utilizing the present invention has far fewer flanged connections compared to the prior art, reducing the number of potential leak sources in the plant. Hydrocarbons are volatile organic compounds (VOCs), some of which are classified as greenhouse gases and some of which may be precursors to atmospheric ozone formation, which means the present invention reduces the potential for atmospheric releases that can damage the environment. 
     Other Embodiments 
     Some circumstances may favor supplying liquid stream  35  directly to the lower region of absorbing section  118   d  via stream  40  as shown in  FIGS. 2 ,  4 ,  6 , and  8 . In such cases, an appropriate expansion device (such as expansion valve  17 ) is used to expand the liquid to the operating pressure of absorbing section  118   d  and the resulting expanded liquid stream  40   a  is supplied as feed to the lower region of absorbing section  118   d  (as shown by the dashed lines). Some circumstances may favor combining a portion of liquid stream  35  (stream  37 ) with the vapor in stream  36  ( FIGS. 2 and 6 ) or with cooled second portion  33   a  ( FIGS. 4 and 8 ) to form combined stream  38  and routing the remaining portion of liquid stream  35  to the lower region of absorbing section  118   d  via streams  40 / 40   a . Some circumstances may favor combining the expanded liquid stream  40   a  with expanded stream  39   a  ( FIGS. 2 and 6 ) or expanded stream  34   a  ( FIGS. 4 and 8 ) and thereafter supplying the combined stream to the lower region of absorbing section  118   d  as a single feed. 
     If the feed gas is richer, the quantity of liquid separated in stream  35  may be great enough to favor placing an additional mass transfer zone in demethanizing section  118   e  between expanded stream  39   a  and expanded liquid stream  40   a  as shown in  FIGS. 3 and 7 , or between expanded stream  34   a  and expanded liquid stream  40   a  as shown in  FIGS. 5 and 9 . In such cases, the heat and mass transfer means in demethanizing section  118   e  may be configured in upper and lower parts so that expanded liquid stream  40   a  can be introduced between the two parts. As shown by the dashed lines, some circumstances may favor combining a portion of liquid stream  35  (stream  37 ) with the vapor in stream  36  ( FIGS. 3 and 7 ) or with cooled second portion  33   a  ( FIGS. 5 and 9 ) to form combined stream  38 , while the remaining portion of liquid stream  35  (stream  40 ) is expanded to lower pressure and supplied between the upper and lower parts of the heat and mass transfer means in demethanizing section  118   e  as stream  40   a.    
     Some circumstances may favor not combining the cooled first and second portions (streams  32   a  and  33   a ) as shown in  FIGS. 4 ,  5 ,  8 , and  9 . In such cases, only the cooled first portion  32   a  is directed to separator section  118   f  inside processing assembly  118  ( FIGS. 4 and 5 ) or separator  12  ( FIGS. 8 and 9 ) where the vapor (stream  34 ) is separated from the condensed liquid (stream  35 ). Vapor stream  34  enters work expansion machine  15  and is expanded substantially isentropically to the operating pressure of absorbing section  118   d , whereupon expanded stream  34   a  is supplied as feed to the lower region of absorbing section  118   d  inside processing assembly  118 . The cooled second portion  33   a  is combined with the separated liquid (stream  35 , via stream  37 ), and the combined stream  38  is directed to the heat exchange means in the lower region of feed cooling section  118   a  inside processing assembly  118  and cooled to substantial condensation. The substantially condensed stream  38   a  is flash expanded through expansion valve  14  to the operating pressure of rectifying section  118   c  and absorbing section  118   d , whereupon expanded stream  38   b  is supplied to processing assembly  118  between rectifying section  118   c  and absorbing section  118   d . Some circumstances may favor combining only a portion (stream  37 ) of liquid stream  35  with the cooled second portion  33   a , with the remaining portion (stream  40 ) supplied to the lower region of absorbing section  118   d  via expansion valve  17 . Other circumstances may favor sending all of liquid stream  35  to the lower region of absorbing section  118   d  via expansion valve  17 . 
     In some circumstances, it may be advantageous to use an external separator vessel to separate cooled feed stream  31   a  or cooled first portion  32   a , rather than including separator section  118   f  in processing assembly  118 . As shown in  FIGS. 6 and 7 , separator  12  can be used to separate cooled feed stream  31   a  into vapor stream  34  and liquid stream  35 . Likewise, as shown in  FIGS. 8 and 9 , separator  12  can be used to separate cooled first portion  32   a  into vapor stream  34  and liquid stream  35 . 
     Depending on the quantity of heavier hydrocarbons in the feed gas and the feed gas pressure, the cooled feed stream  31   a  entering separator section  118   f  in  FIGS. 2 and 3  or separator  12  in  FIGS. 6 and 7  (or the cooled first portion  32   a  entering separator section  118   f  in  FIGS. 4 and 5  or separator  12  in  FIGS. 8 and 9 ) may not contain any liquid (because it is above its dewpoint, or because it is above its cricondenbar). In such cases, there is no liquid in streams  35  and  37  (as shown by the dashed lines), so only the vapor from separator section  118   f  in stream  36  ( FIGS. 2 and 3 ), the vapor from separator  12  in stream  36  ( FIGS. 6 and 7 ), or the cooled second portion  33   a  ( FIGS. 4 ,  5 ,  8 , and  9 ) flows to stream  38  to become the expanded substantially condensed stream  38   b  supplied to processing assembly  118  between rectifying section  118   c  and absorbing section  118   d . In such circumstances, separator section  118   f  in processing assembly  118  ( FIGS. 2 through 5 ) or separator  12  ( FIGS. 6 through 9 ) may not be required. 
     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 feed stream (stream  38   a ) or the substantially condensed recycle stream (stream  45   a ). 
     In accordance with the present invention, the use of external refrigeration to supplement the cooling available to the inlet gas from the distillation vapor and liquid streams may be employed, particularly in the case of a rich inlet gas. In such cases, a heat and mass transfer means may be included in separator section  118   f  (or a gas collecting means in such cases when the cooled feed stream  31   a  or the cooled first portion  32   a  contains no liquid) as shown by the dashed lines in  FIGS. 2 through 5 , or a heat and mass transfer means may be included in separator  12  as shown by the dashed lines in  FIGS. 6  though  9 . This heat and mass transfer means may be comprised of a fin and tube type heat exchanger, a plate type heat exchanger, a brazed aluminum type heat exchanger, or other type of heat transfer device, including multi-pass and/or multi-service heat exchangers. The heat and mass transfer means is configured to provide heat exchange between a refrigerant stream (e.g., propane) flowing through one pass of the heat and mass transfer means and the vapor portion of stream  31   a  ( FIGS. 2 ,  3 ,  6 , and  7 ) or stream  32   a  ( FIGS. 4 ,  5 ,  8 , and  9 ) flowing upward, so that the refrigerant further cools the vapor and condenses additional liquid, which falls downward to become part of the liquid removed in stream  35 . Alternatively, conventional gas chiller(s) could be used to cool stream  32   a , stream  33   a , and/or stream  31   a  with refrigerant before stream  31   a  enters separator section  118   f  ( FIGS. 2 and 3 ) or separator  12  ( FIGS. 6 and 7 ) or stream  32   a  enters separator section  118   f  ( FIGS. 4 and 5 ) or separator  12  ( FIGS. 8 and 9 ). 
     Depending on the temperature and richness of the feed gas and the amount of C 2  components to be recovered in liquid product stream  44 , there may not be sufficient heating available from stream  33  to cause the liquid leaving demethanizing section  118   e  to meet the product specifications. In such cases, the heat and mass transfer means in demethanizing section  118   e  may include provisions for providing supplemental heating with heating medium as shown by the dashed lines in  FIGS. 2 through 9 . Alternatively, another heat and mass transfer means can be included in the lower region of demethanizing section  118   e  for providing supplemental heating, or stream  33  can be heated with heating medium before it is supplied to the heat and mass transfer means in demethanizing section  118   e.    
     Depending on the type of heat transfer devices selected for the heat exchange means in the upper and lower regions of feed cooling section  118   a , it may be possible to combine these heat exchange means in a single multi-pass and/or multi-service heat transfer device. In such cases, the multi-pass and/or multi-service heat transfer device will include appropriate means for distributing, segregating, and collecting stream  32 , stream  38 , stream  45 , and the distillation vapor stream in order to accomplish the desired cooling and heating. 
     Some circumstances may favor providing additional mass transfer in the upper region of demethanizing section  118   e . In such cases, a mass transfer means can be located below where expanded stream  39   a  ( FIGS. 2 ,  3 ,  6 , and  7 ) or expanded stream  34   a  ( FIGS. 4 ,  5 ,  8 , and  9 ) enters the lower region of absorbing section  118   d  and above where cooled second portion  33   a  leaves the heat and mass transfer means in demethanizing section  118   e.    
     A less preferred option for the  FIGS. 2 ,  3 ,  6 , and  7  embodiments of the present invention is providing a separator vessel for cooled first portion  32   a , a separator vessel for cooled second portion  33   a , combining the vapor streams separated therein to form vapor stream  34 , and combining the liquid streams separated therein to form liquid stream  35 . Another less preferred option for the present invention is cooling stream  37  in a separate heat exchange means inside feed cooling section  118   a  (rather than combining stream  37  with stream  36  or stream  33   a  to form combined stream  38 ), expanding the cooled stream in a separate expansion device, and supplying the expanded stream to an intermediate region in absorbing section  118   d.    
     It will be recognized that the relative amount of feed found in each branch of the split vapor feed 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 above absorbing section  118   d  may increase recovery while decreasing power recovered from the expander and thereby increasing the recompression horsepower requirements. Increasing feed below absorbing section  118   d  reduces the horsepower consumption but may also reduce product recovery. 
     The present invention 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 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 supplemental heating, 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.