Abstract:
A process and an apparatus are disclosed for a compact processing assembly to remove C 5  and heavier hydrocarbon components from a hydrocarbon gas stream. The hydrocarbon gas stream is expanded to lower pressure and supplied to the processing assembly between an absorbing means and a mass transfer means. A distillation vapor stream is collected from the upper region of the absorbing means and cooled in a first heat and mass transfer means inside the processing assembly to partially condense it, forming a residual vapor stream and a condensed stream. The condensed, stream is supplied to the absorbing means at its top feed point. A distillation liquid stream is collected from the lower region of the mass transfer means and directed into a second beat and mass transfer means inside the processing assembly to heat it and strip out its volatile components.

Description:
BACKGROUND OF THE INVENTION 
       [0001]    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. 61/876,415 which was filed on Sep. 11, 2013 and No. 61/879,308 which was filed on Sep. 18, 2013. 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. 
         [0002]    Natural gas is typically recovered from wells drilled into underground reservoirs. It usually has a major proportion of methane, i.e., methane comprises at least 50 mole percent of the gas. Depending on the particular underground reservoir, the natural gas also contains relatively lesser amounts of heavier hydrocarbons such as ethane, propane, butanes, pentanes and the like, as well as water, hydrogen, nitrogen, carbon dioxide, and other gases. A typical analysis of a natural gas stream to be processed in accordance with this invention would be, in approximate mole percent, 89.2% methane, 4.9% ethane and other C 2  components, 2.6% propane and other C 3  components, 0.4% iso-butane, 1.3% normal butane, and 0.6% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present. 
         [0003]    Most natural gas is handled in gaseous form. The Most common means for transporting natural gas from the wellhead to gas processing plants and thence to the natural gas consumers is in high-pressure gas transmission pipelines. In a number of circumstances, however, it has been found necessary and/or desirable to liquefy the natural gas either for transport or for use. In remote locations, for instance, there is often no pipeline infrastructure that would allow for convenient transportation of the natural gas to market, in such cases, the much lower specific volume of liquefied natural gas (LNG) relative to natural gas in the gaseous state can greatly reduce transportation costs by allowing delivery of the LNG using cargo ships and transport trucks. 
         [0004]    A relatively recent concept for commercializing natural gas from remote locations is to install a liquefaction plant on an offshore platform or on a ship (commonly referred to as floating LNG or FLNG) to allow moving the facility to another location when the gas reservoir is depleted. Deck space is at a premium for both of these, because each increment of deck space requires a very large quantity of supporting structure (and hull volume in the case of FLNG). As a result, great emphasis is placed on minimizing the “footprint” of each processing step in order to minimize the investment cost and thereby maximize the number of gas reservoirs in remote locations that can be economically produced. 
         [0005]    For remote locations such as those contemplated here, recovery of the various hydrocarbons heavier than methane as separate products is generally not economically viable since there is usually no means of transporting and selling the resultant hydrocarbon product streams. Instead, to the largest extent possible, these heavier hydro-carbons are liquefied along with the methane and sold as part of the LNG product. However, some degree of heavier hydrocarbon removal is often required prior to liquefying the natural gas because there are usually limitations on the heating value of the re-vaporized gas when it is subsequently distributed from the LNG receiving terminal. In addition, hydrocarbons heavier than C 5  or C 6  (particularly aromatic hydrocarbons) generally must be removed upstream of the liquefaction step to avoid plugging inside the liquefaction plant caused by freezing of these heavier hydrocarbons. For these reasons, it is typical to include a processing step to remove these hydrocarbons (“heavy ends removal”) before liquefying the natural gas. 
         [0006]    Available processes for removing these heavier hydrocarbons 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 he employed, 
         [0007]    The cryogenic expansion process is now generally preferred for removing heavy hydrocarbons from natural gas 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; 4,5149,824; 4,061,039; 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; 6,742,358; 6,915,662: 6,945,075; 7,010,937; 7,191,617; 7,204,100; 7,210,311; 7,219,513; 7,565,815; 8,590,340; reissue U.S. Pat. No.  33 , 408 ; and co-pending application Ser. Nos. 11/430,412: 11/839,693; 12/206,230; 12/487,078; 12/689,616; 12/717,394; 12/750,862; 12/772,472; 12/781,259; 12/868,993; 12/869,007; 12/869,139; 12/979,563; 13/048,315; 13/051,682; 13/052,348; 13/052,575; and 13/053,792 describe relevant processes (although the description of the present invention in some cases is based on different processing conditions than those described in the cited U.S. Patents and co-pending applications). 
         [0008]    The present invention is a novel means of removing heavier hydrocarbon components from natural gas that combines what heretofore have been individual equipment items into a common housing, thereby reducing the plot space requirements, the capital cost of the plant, and (more importantly) the capital cost of the associated platform or ship. 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 may damage the environment. 
     
    
     
         [0009]    For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings: 
           [0010]      FIGS. 1 and 2  are flow diagrams of prior art heavy ends removal processes for a natural gas liquefaction plant; 
           [0011]      FIG. 3  is a flow diagram of a heavy ends removal plant for a natural gas liquefaction plant in accordance with the present invention; and 
           [0012]      FIGS. 4 through 6  are flow diagrams illustrating alternative means of application of the present invention to heavy ends removal from a natural gas stream. 
       
    
    
       [0013]    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. 
         [0014]    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 
       [0015]      FIG. 1  is a process flow diagram showing the design of a processing plant to remove heavier hydrocarbon components from a natural gas stream using a prior art process. In this simulation of the process, inlet gas enters the plant at 60° F. [15° C.] and 995 psia [6,858 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. 
         [0016]    The feed stream  31  is divided into two portions, streams  32  and  33 . Stream  32  is cooled in heat exchanger  10  by heat exchange with cool residue gas stream  37 , while stream  33  is cooled in heat exchanger  11  by heat exchange with flash expanded liquids (stream  35   a ). streams  32   a  and  33   a  recombine to form stream  31   a , which enters separator  12  at 4° F. [−16° C.]find 980 psia [6,755 kPa(a)] where the vapor (stream  34 ) is separated from the condensed liquid (stream  35 ). 
         [0017]    The vapor from separator  12  (stream  34 ) enters a work expansion machine  13  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  13  expands the vapor substantially isentropically to the operating pressure (approximately 470 psia [3,238 kPa(a)]) of fractionation tower  16 , with the work expansion cooling the expanded stream  34   a  to a temperature of approximately −59° F. [−51° C]. The typical commercially available expanders are capable of recovering on the order of 80-85% of the work theoretically available in an ideal iseniropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item  14 ) that can be used to re-compress the heated residue gas stream, (stream  37   a ), for example. The partially condensed expanded stream  34   a  is thereafter supplied as feed to fractionation tower  16  at an upper mid-column feed point. The separator liquid (stream  35 ) is expanded to slightly above the operating pressure of fractionation tower  16  by expansion valve  15 , then heated from −17° F. [−27° C.] to 54° F. [12° C.] in heat exchanger  11  as described earlier before stream  35   b  is supplied to fractionation tower  16  at a lower mid-column feed point. 
         [0018]    Fractionation tower  16  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 ease, the .fractionation tower may consist of two sections, an upper rectifying section  16   a  and a lower stripping section  16   b.  The upper rectifying section  16   a  contains trays and/or packing and provides the necessary contact between the vapor rising from the lower distillation or stripping section  16   b  and a liquid stream (reflux) to remove the heavier hydrocarbon components from the vapor. The lower, stripping section  16   b  also contains trays and/or packing and provides die necessary contact between the liquids failing downward and the vapors rising upward. The aripplug section  16   b  also includes at least one reboiler (such as the reboiler  17 ) which heats and vaporizes 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  39 , of lighter hydrocarbon components so that it contains only the heavier hydrocarbon components that were in the natural gas feed stream (stream  31 ), whereupon it exits the bottom of the tower at 358° F. [181° C.]. 
         [0019]    The column overhead vapor (stream  36 ) is withdrawn from the top of tower  16  at −6° F. [−21° C.] and is cooled to −18° F. [−28° C.] and partially condensed (stream  36   a ) in heat exchanger  18  using a refrigerant. The operating pressure in reflux separator  19  is maintained slightly below the operating pressure of tower  16 . This provides the driving force which causes overhead vapor stream  36  to flow through reflux condenser  18  and thence into the reflux separator  19  wherein the condensed liquid (stream  38 ) is separated front the uncondensed vapor (stream  37 ). The liquid stream  38  from reflux separator  19  is pumped by reflux pump  20  to a pressure slightly above the operating pressure of tower  16 , and stream  38   a  is then supplied as cold top column feed (reflux) to tower  16 . This cold liquid reflux absorbs and condenses the heavier hydrocarbon components in the vapors rising up in rectifying section  16   a  of tower  16 . The residue gas (vapor stream  37 ) passes countercurrently to the incoming feed gas in heat exchanger  10  where it is heated to 55° F. [13° C.] as it provides cooling as previously described. The residue gas is then re-compressed in two stages, compressor  14  driven by expansion machine  13  and compressor  21 . driven by a supplemental power source. (In the  FIG. 1  process, compressor  21  consists of two compression stages with intercooling between the stages.) After stream  37   c  is cooled in discharge cooler  22 , the residue gas product (stream  37   d ) flows to the sales gas pipeline or to the liquefaction plant at  1603  psia [11,050 kPa(a)]. 
         [0020]    A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 1  is set forth in the following table: 
         [0000]    
       
         
               
             
               
               
               
               
               
               
               
             
               
             
               
               
               
             
               
             
               
               
               
               
             
           
               
                 TABLE I 
               
               
                   
               
               
                 (FIG. 1) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
             
               
                   
               
             
          
           
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes 
                 Pentanes+ 
                 Total 
               
               
                   
               
               
                 31 
                 29,885 
                 1,626 
                 869 
                 574 
                 189 
                 33,501 
               
               
                 34 
                 29,570 
                 1,555 
                 769 
                 428 
                 77 
                 32,756 
               
               
                 35 
                 315 
                 71 
                 100 
                 146 
                 112 
                 745 
               
               
                 36 
                 29,987 
                 1,666 
                 959 
                 586 
                 7 
                 33,564 
               
               
                 38 
                 102 
                 40 
                 91 
                 164 
                 4 
                 402 
               
               
                 37 
                 29,885 
                 1,626 
                 868 
                 422 
                 3 
                 33,162 
               
               
                 39 
                 0 
                 0 
                 1 
                 152 
                 186 
                 339 
               
               
                   
               
             
          
           
               
                 Recoveries* 
               
               
                   
               
             
          
           
               
                   
                 Butanes 
                 26.52% 
               
               
                   
                 Pentanes+ 
                 98.42% 
               
               
                   
                   
               
             
          
           
               
                 Power 
               
               
                   
               
             
          
           
               
                   
                 Residue Gas Compression 
                 15,290 HP 
                 [25,137 kW] 
               
               
                   
                 Refrigerant Compression 
                   861 HP 
                  [1,415 kW] 
               
               
                   
                 Total Compression 
                 16,151 HP 
                 [26,552 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
             
          
         
       
     
         [0021]      FIG. 2  is a process flow diagram showing the design of a processing plant to remove heavier hydrocarbon components from a natural gas stream using another prior art process. The process of  FIG. 2  has been applied to the same feed gas composition and conditions as described previously for  FIG. 1 . 
         [0022]    In this simulation of the process, inlet gas enters the plant at 60° F. [15° C. ] and 995 psia [6,858 kPa(a)] as stream  31  and flows directly to work expansion machine  13  in which mechanical energy is extracted from the high pressure feed. The machine  13  expands the vapor substantially isentropically to the operating pressure (approximately 355 psia [2,446 kPa(a)]) of fractionation tower  16 , with the work expansion cooling the expanded stream  31   a  to a temperature of approximately −34° F. [−37° C.] before it is supplied as feed to fractionation tower  16  at. a mid-column feed point. 
         [0023]    Liquid product stream  39  is stripped of the lighter hydrocarbon components and exits the bottom of tower  16  at 320° F. [160° C.]. The column overhead vapor (stream  36 ) is withdrawn from the top of tower  16  at −8° F. [−22° C.] and is cooled to −24° F. [−31° C. ] and partially condensed (stream  36   a ) in heat exchange  18  using a refrigerant, and the condensed liquid stream  38  is separated from the uncondensed vapor (stream  37 ) in reflux separator  19 . Liquid stream  38  is pumped by reflux pump  20  to a pressure slightly above the operating pressure of tower  16  and stream  38   a  is then supplied as cold top column feed (reflux) to tower  16 . The residue gas (vapor stream  37 ) is re-compressed in two stages, compressor  14  driven by expansion machine  13  and compressor  21  driven by a supplemental power source. (In the  FIG. 2  process, compressor  21  consists of two compression stages with intercooling between the stages.) After stream  37   b  is cooled m discharge cooler  22 , the residue gas product (stream  37   c ) flows to the sales gas pipeline or to the liquefaction plant at 1603 psia [11,050 kPa(a)]. 
         [0024]    A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 2  is set forth in the following table: 
         [0000]    
       
         
               
             
               
               
               
               
               
               
               
             
               
             
               
               
               
             
               
             
               
               
               
               
             
           
               
                 TABLE II 
               
               
                   
               
               
                 (FIG. 2) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
             
               
                   
               
             
          
           
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes 
                 Pentanes+ 
                 Total 
               
               
                   
               
               
                 31 
                 29,885 
                 1,626 
                 869 
                 574 
                 189 
                 33,501 
               
               
                 36 
                 29,987 
                 1,673 
                 986 
                 664 
                 9 
                 33,677 
               
               
                 38 
                 102 
                 47 
                 118 
                 239 
                 6 
                 512 
               
               
                 37 
                 29,885 
                 1,626 
                 868 
                 425 
                 3 
                 33,165 
               
               
                 39 
                 0 
                 0 
                 1 
                 149 
                 186 
                 336 
               
               
                   
               
             
          
           
               
                 Recoveries* 
               
               
                   
               
             
          
           
               
                   
                 Butanes 
                 25.92% 
               
               
                   
                 Pentanes+ 
                 98.42% 
               
               
                   
                   
               
             
          
           
               
                 Power 
               
               
                   
               
             
          
           
               
                   
                 Residue Gas Compression 
                 13,805 HP 
                 [22,695 kW] 
               
               
                   
                 Refrigerant Compression 
                  1,261 HP 
                  [2,073 kW] 
               
               
                   
                 Total Compression 
                 15,066 HP 
                 [24,768 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
             
          
         
       
     
       DESCRIPTION OF THE INVENTION 
     Example 1 
       [0025]      FIG. 3  illustrates a flow diagram of a process in accordance with the present invention. The feed gas composition and conditions considered in the process presented in  FIG. 3  are the same as those in  FIG. 1 . Accordingly, the  FIG. 3  process can be compared with that of the  FIG. 1  process to illustrate the advantages of the present invention. 
         [0026]    In the process illustrated in  FIG. 3 , inlet gas enters the plant at 60° F. [15° C] and 995 psia [6,858 kPa(a)] as stream  31  and is directed to a heat exchange means in feed cooling section  116   a  inside processing assembly  116 . 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  31  flowing through one pass of the heat exchange means and flash expanded separator liquids (stream  35   a ) and a residue gas stream, from condensing section  116   b  inside processing assembly  116 . Stream  31  is cooled while sealing die flash expanded separator liquids and the residue gas stream. 
         [0027]    Separator section  116   e  has an internal head or other means to divide it from stripping section  116   d , so that the two sections inside processing assembly  116  can operate at different pressures. The cooled stream  31   a  enters separator section  116   c  at 4° F. [−16° C.] and 980 psia [6,755 kPa(a)] where any condensed liquid (stream  35 ) is separated from the vapor (stream  34 ). Stream  35  exits separator section  116   e  and is expanded by expansion valve  15  to slightly above the operating pressure (470 psia [3,238 kPa(a)]) of stripping section  116   d  inside processing assembly  116 , cooling stream  35   a  to −17° F. [−27° C.]. Stream  35   a  enters the heat exchange means in feed, cooling section  116   a  to supply cooling to the feed gas as described previously, heating stream  35   b  to 54° F. [12° C.] before it enters below a mass transfer means inside stripping section  116   d  of processing assembly  116 . 
         [0028]    The vapor (stream  34 ) from separator section  116   e  enters a work expansion machine  13  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  13  expands the vapor substantially isentropically to the operating pressure of rectifying section  116   c  inside processing assembly  116 , with the work expansion cooling the expanded stream  34   a  to −59° F. [−51° C. ]. The partially condensed expanded stream  34   a  is thereafter supplied as feed between an absorbing means inside rectifying section  116   c  and the mass transfer means inside stripping section  116   d  of processing assembly  116 . 
         [0029]    A heat and mass transfer means is located below the mass frontier means inside stripping section  116   d . The 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 heating medium flowing through one pass of the heat and mass transfer means and a distillation liquid stream flowing downward from the lower region of the mass transfer means, so that the distillation liquid stream is heated. As the distillation liquid stream is heated, a portion of it is vaporized to form shipping vapors that rise upward to the mass transfer means 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 shipping 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  39  of lighter hydrocarbon components. The stripping vapors produced in the heat and mass transfer means continue upward to the mass transfer means in stripping section  116   d  to provide partial stripping of the lighter hydrocarbon components in the liquids flowing downward from the upper part of processing assembly  116 . 
         [0030]    Another heat and mass transfer means is located inside condensing section  116   b , above the absorbing means inside rectifying section  116   c  of processing assembly  116 . 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 a refrigerant stream flowing through one pass of the heat and mass transfer means and a distillation vapor stream arising from the upper region of the absorbing means flowing upward through the other pass, so that the distillation vapor stream is cooled by the refrigerant. As the distillation vapor stream is cooled, a portion of it is condensed and falls downward while the remaining distillation vapor stream continues flowing upward through die heat and. mass transfer means. The heat and mass transfer means provides continuous contact between the condensed liquid and the distillation vapor stream so that it also functions to provide mass transfer between the vapor and liquid phases, thereby absorbing heavier hydrocarbon components from the distillation vapor stream to rectify it. The condensed liquid is collected from the bottom of the heat and mass transfer means and directed to the upper region of the absorbing means inside rectifying section  16   c  to provide partial rectification of the heavier hydrocarbon components in the vapors flowing upward from the lower part of processing assembly  116 . 
         [0031]    The absorbing means inside rectifying section  116   c  and the mass transfer means inside stripping section  116   d  each consist 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  116   c  and stripping section  116   d  provide the necessary contact between the vapors rising upward and liquid falling downward. The liquid portion of expanded stream  34   a  commingles with liquids falling downward from rectifying section  116   c  and the combined liquids continue downward into stripping section  116   d , which vaporizes and strips the lighter hydrocarbon components from these liquids. The vapors arising from stripping section  116   d  combine with, the vapor portion of expanded stream  34   a  and rise upward through rectifying section  116   c,  to be contacted with the cold liquid failing downward to condense and absorb the heavier hydrocarbon components from these vapors. 
         [0032]    The distillation liquid flowing downward from the hear and mass transfer means in stripping section  116   d  inside processing assembly  116  has been stripped of the lighter hydrocarbon components so that it contains only the heavier hydrocarbon components that were in the natural gas feed stream (stream  31 ) whereupon it exits the lower region of stripping section  116   d  as stream  39  and leaves processing assembly  116  at 356° F. [180° C. ]. The distillation vapor stream arising from condensing section  116   b  is heated in feed cooling section  116   a  as it provides cooling to stream  31  as described previously, whereupon residue gas stream  37  leaves processing assembly  116  at 55° F. [13° C. ]. The residue gas stream is then re-compressed in two stages, compressor  14  driven by expansion machine  13  and compressor  21  driven by a supplemental power source. (In this example, compressor  21  consists of two compression stages with intercooling between the stages.) After cooling in discharge cooler  22 , residue gas stream  37   c  flows to the sales gas pipeline or to the liquefaction plant at 1603 psia [11,050 kPa(a)]. 
         [0033]    A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 3  is set forth in the following table: 
         [0000]    
       
         
               
             
               
               
               
               
               
               
               
             
               
             
               
               
               
             
               
             
               
               
               
               
             
           
               
                 TABLE III 
               
               
                   
               
               
                 (FIG. 3) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
             
               
                   
               
             
          
           
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes 
                 Pentanes+ 
                 Total 
               
               
                   
               
               
                 31 
                 29,885 
                 1,626 
                 869 
                 574 
                 189 
                 33,501 
               
               
                 34 
                 29,571 
                 1,555 
                 769 
                 428 
                 77 
                 32,758 
               
               
                 35 
                 314 
                 71 
                 100 
                 146 
                 112 
                 743 
               
               
                 37 
                 29,885 
                 1,626 
                 868 
                 423 
                 3 
                 33,164 
               
               
                 39 
                 0 
                 0 
                 1 
                 151 
                 186 
                 337 
               
               
                   
               
             
          
           
               
                 Recoveries* 
               
               
                   
               
             
          
           
               
                   
                 Butanes 
                 26.26% 
               
               
                   
                 Pentanes+ 
                 98.42% 
               
               
                   
                   
               
             
          
           
               
                 Power 
               
               
                   
               
             
          
           
               
                   
                 Residue Gas Compression 
                 15,289 HP 
                 [25,135 kW] 
               
               
                   
                 Refrigerant Compression 
                   846 HP 
                  [1,391 kW] 
               
               
                   
                 Total Compression 
                 16,135 HP 
                 [26,526 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
             
          
         
       
     
         [0034]    A comparison of Tables I and III shows that, compared to the prior art, the compact processing assembly of the present invention maintains essentially the same butanes recovery and pentane+recovery using slightly less power. However, the present invention offers at least two advantages over the prior art of  FIG. 1 . First, the compact arrangement of processing assembly  116  of the present invention replaces eight separate equipment items in the  FIG. 1  prior art (heat exchangers  10  and  11 , separator  12 , fractionation tower  16 , reboiler  17 , reflux condenser  18 , reflux separator  19 , and reflux pump  20 ) with a single equipment item (processing assembly  116  in  FIG. 3  of the present invention). This reduces the “footprint” of the processing plant and eliminates the interconnecting piping, reducing the capital cost of the processing plant itself and (more importantly) the capital cost of the platform or ship on which it is installed. Second, elimination of the interconnecting piping means that a processing plant using 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 may damage the environment. 
       Example 2 
       [0035]    The present invention also offers advantages when circumstances favor expansion of the feed gas without cooling it first, as illustrated in  FIG. 4 . The feed gas composition and conditions considered in the process presented in  FIG. 4  are the same as those in  FIG. 2 , Accordingly, the  FIG. 4  process can be compared with that of the  FIG. 2  process to further illustrate the advantages of the present invention. 
         [0036]    In the process illustrated in  FIG. 4 , inlet gas enters the plant at 60° F. [15° C.] and 995 psia [6,858 kPa(a)] as stream  31  and is directed to work expansion machine  13  as stream  34 . Mechanical energy is extracted from the high pressure feed as machine  13  expands the vapor substantially isentropically to the operating pressure (355 psia [2,446 kPa(a)]) of rectifying section  116   c  inside processing assembly  116 , with the work expansion cooling the expanded stream  34   a  to −34° F. [−37° C.]. The partially condensed expanded stream  34   a  is thereafter supplied as feed between an absorbing means inside rectifying section  116   c  and a mass transfer means inside stripping section  116   d  of processing assembly  116 . 
         [0037]    A heat and mass transfer means is located, below the mass transfer means inside stripping section  116   d.  The heat and mass transfer means is configured to provide heat exchange between a heating medium flowing through one pass of the heat and mass transfer means and a distillation liquid stream flowing downward horn die lower region of the mass transfer means, so that the distillation liquid stream is heated. As the distillation liquid stream is heated, a portion of it is vaporized to form stripping vapors that rise upward to the mass transfer means 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  39  of lighter hydrocarbon components. The stripping vapors produced in the heat and mass transfer means continue upward to the mass transfer means in stripping section  116   d  to provide partial stripping of the lighter hydrocarbon components in the liquids flowing downward from the upper part of processing assembly  116 . 
         [0038]    Another heat and mass transfer means is located inside condensing section  116   b , above the absorbing means inside rectifying section  116   c  of processing assembly  116 . The heat and mass transfer means is configured to provide heat exchange between a refrigerant stream flowing through one pass of the hear and mass transfer means and a distillation vapor stream arising irons the upper region of the absorbing means flowing upward through the other pass, so that the distillation vapor stream is cooled by the refrigerant. As the distillation vapor stream is cooled, a portion of it is condensed and falls downward while the remaining distillation vapor stream continues flowing upward through the heat and mass transfer means. The heat and mass transfer means provides continuous contact between the condensed liquid and the distillation vapor stream so that it also functions to provide mass transfer between the vapor and liquid phases, thereby absorbing heavier hydrocarbon components from the distillation vapor stream to rectify it. The condensed liquid is collected from, the bottom of the heat and mass transfer means and directed to the upper region of the absorbing means inside rectifying section  116   c  to provide partial rectification of the heavier hydrocarbon components in the vapors flowing upward from the lower part of processing assembly  116 . 
         [0039]    The distillation liquid flowing downward from the heat and mass transfer means in stripping section  116   d  inside processing assembly  116  has been shipped of the lighter hydrocarbon components so that it contains only the heavier hydrocarbon components that were in the natural gas feed stream (stream  31 ), whereupon it exits the lower region of stripping section  116   d  as stream  39  and leaves processing assembly  116  at 318° F. [159° C]. The distillation vapor stream arising from condensing section  116   b  is residue gas stream  37 , which leaves processing assembly  116  at −25° F. [−32° C]. The residue gas stream is re-compressed in two stages, compressor  14  driven by expansion machine  13  and compressor  21  driven by a supplemental power source, (in this example, compressor  21  consists of two compression stages with intercooling between the stages.) After cooling in discharge cooler  22 , residue gas stream  37   d  flows to the sales gas pipeline or to the liquefaction plant at 1603 psia [11,050 kPa(a)]. 
         [0040]    A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 4  is set forth in the following table: 
         [0000]    
       
         
               
             
               
               
               
               
               
               
               
             
               
             
               
               
               
             
               
             
               
               
               
               
             
           
               
                 TABLE IV 
               
               
                   
               
               
                 (FIG. 4) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
             
               
                   
               
             
          
           
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes 
                 Pentanes+ 
                 Total 
               
               
                   
               
               
                 31/34 
                 29,885 
                 1,626 
                 869 
                 574 
                 189 
                 33,501 
               
               
                 37 
                 29,885 
                 1,626 
                 868 
                 421 
                 3 
                 33,162 
               
               
                 39 
                 0 
                 0 
                 1 
                 153 
                 186 
                 339 
               
               
                   
               
             
          
           
               
                 Recoveries* 
               
               
                   
               
             
          
           
               
                   
                 Butanes 
                 26.70% 
               
               
                   
                 Pentanes+ 
                 98.42% 
               
               
                   
                   
               
             
          
           
               
                 Power 
               
               
                   
               
             
          
           
               
                   
                 Residue Gas Compression 
                 13,788 HP 
                 [22,667 kW] 
               
               
                   
                 Refrigerant Compression 
                  1,269 HP 
                  [2,086 kW] 
               
               
                   
                 Total Compression 
                 15,057 HP 
                 [24,753 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
             
          
         
       
     
         [0041]    A comparison of Tables II and IV shows that, compared to the prior art, the compact processing assembly of the present invention maintains essentially the same butanes recovery and pentane+recovery using slightly less power. However, the present invention offers at least two advantages over the prior art of  FIG. 2 . First, the compact arrangement of processing assembly  116  of the present invention replaces five separate equipment items in the  FIG. 2  prior art (fractionation tower  16 , reboiler  17 , reflux condenser  18 , reflux separator  19 , and reflux pump  20 ) with a single equipment item (processing assembly  116  in  FIG. 4  of the present invention). This reduces the “footprint” of the processing plant and eliminates the interconnecting piping, reducing the capital cost of the processing plant itself and (more importantly) the capital cost of the platform or shin on which it is installed. Second, elimination of the interconnecting piping means that a processing plant using die 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 sonic of which may be precursors to atmospheric ozone formation, which means the present invention reduces the potential for atmospheric releases that may damage the environment. 
       Other Embodiments 
       [0042]    Some circumstances may favor eliminating feed cooling section  116   a  from processing assembly  116  in order to reduce the height of processing assembly  116 . As shown in  FIGS. 4 and 6 , in such cases heat exchanger  10  can be used to provide the cooling of the feed gas (stream  31 ). In other circumstances, it may be advantageous not to cool the feed gas at all so that stream  34  is warmer and its work expansion in expansion machine  13  generates more power for compressor  14 . In such circumstances, heat exchanger  10  in  FIGS. 4 and 6  may not be require. 
         [0043]    In some circumstances, it may be advantageous to use an external separator vessel to separate cooled feed stream  31   a,  rather than including separator section  116   e  in processing assembly  116 . As shown in  FIGS. 4 and 5 , separator  12  can he used to separate cooled feed stream  31   a  into vapor stream  34  and liquid stream  35 . 
         [0044]    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  116   e  in  FIGS. 3 and 6  or separator  12  in FIGS-  4  and  5  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 stream  35  (as shown by the dashed lines). In such circumstances, separator section  116   e  in processing assembly  116  ( FIGS. 3 and 6 ) or separator  12  ( FIGS. 4 and 5 ) may not be required. 
         [0045]    Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine  13 , or replacement with an alternate expansion device (such as on 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 separator liquid (stream  35 ). 
         [0046]    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 separator 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  116   e  (or a collecting means in such cases when the cooled feed stream  31   a  contains no liquid) as shown by the dashed lines in  FIGS. 3 and 6 , or a heat and mass transfer means may be included in separator  12  as shown by the dashed lines in  FIGS. 4 and 5 . 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  flowing upward, so that the refrigerant further cools the vapor and condenses additional liquid, which fails downward to become part of the liquid removed in stream  35 . Alternatively, conventional gas chiller(s) could be used to cool stream  31   a  with refrigerant before stream  31   a  enters separator section  116   e  ( FIGS. 3 and 6 ) or separator  12  ( FIGS. 4 and 5 ). 
         [0047]    Depending on the type of heat transfer device selected for the heat exchange means in feed cooling section  116   a  and the type of heat and mass transfer device selected for the heat and mass transfer means in condensing section  116   b.  it may be possible to combine these in a single multi-pass and/or multi-service heat and mass transfer device. In such cases, the multi-pass and/or multi-service heat and mass transfer device will include appropriate means for distributing, segregating, end collecting streams  31 / 31   a , streams  35   a    35   b.  and the distillation vapor stream in order to accomplish the desired cooling and heating. 
         [0048]    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.