Abstract:
A process and an apparatus are disclosed for removing carbon dioxide from a hydrocarbon gas stream. The gas stream is cooled, expanded to intermediate pressure, and supplied to a fractionation tower at a top column feed position. The tower overhead vapor stream is compressed to higher pressure and cooled to partially condense it, forming a condensed stream. The condensed stream is expanded to intermediate pressure, used to subcool a portion of the tower bottom liquid product, then supplied to the tower at a mid-column feed position. The subcooled portion of the tower bottom liquid product is expanded to lower pressure and used to cool the compressed overhead vapor stream. The quantities and temperatures of the feeds to the fractionation tower are effective to maintain the overhead temperature of the fractionation tower at a temperature whereby the major portion of the carbon dioxide is recovered in the tower bottom liquid product.

Description:
[0001]    This invention relates to a process and apparatus for the separation of a gas containing hydrocarbons and carbon dioxide. The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application No. 61/351,059 which was filed on Jun. 3, 2010. 
     
    
     BACKGROUND OF THE INVENTION 
       [0002]    Hydrocarbons are found in 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. In many cases, the gas streams from these sources are contaminated with high concentrations of carbon dioxide, making the gas streams unsuitable for use as fuel, chemical plant feedstock, or other purposes. There are a variety of processes that have been developed to remove the carbon dioxide using chemical, physical, and hybrid solvents. Other processes have been developed that use a refrigerated absorbent stream composed of heavy (C 4 -C 10  typically) hydrocarbons to remove the carbon dioxide in a distillation column, such as the process described in U.S. Pat. No. 4,318,723. All of these processes have increasingly higher capital cost and operating cost as the carbon dioxide concentration in the gas stream increases, which often makes processing of such gas streams uneconomical. 
         [0003]    One method for improving the economics of processing gas streams containing high concentrations of carbon dioxide is to provide bulk separation of the carbon dioxide from the gas stream prior to processing with solvents or absorbents, so that only a minor fraction of the carbon dioxide must then be removed from the gas stream. For example, semi-permeable membranes have often been used for bulk removal of carbon dioxide. However, a significant fraction of the lighter hydrocarbons in the gas stream are often “lost” in the carbon dioxide stream that is separated by bulk removal processes of this type. 
         [0004]    A better alternative for bulk removal of carbon dioxide is to use distillation to fractionate the gas stream into a light hydrocarbon stream and a carbon dioxide stream, so that removal of the residual carbon dioxide from the light hydrocarbon stream is all that is required to produce pipeline-quality gas for use as fuel, chemical plant feedstock, etc. The majority of the carbon dioxide that is removed is recovered as a liquid rather than a vapor, allowing the carbon dioxide to be pumped (rather than compressed) for subsequent use in tertiary oil recovery operations or for other purposes, resulting in substantial reductions in capital cost and operating cost. 
         [0005]    The present invention is generally concerned with the removal of the majority of the carbon dioxide 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, 44.3% hydrogen, 13.0% carbon monoxide, 4.0% methane, and 38.5% carbon dioxide, with the balance made up of nitrogen and argon. Sulfur containing gases are also sometimes present. 
         [0006]    In a typical distillation process for removing carbon dioxide, 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. The gas is condensed as it is cooled, and the high-pressure liquid is expanded to an intermediate pressure, resulting in further cooling of the stream due to the vaporization occurring during expansion of the liquids. The expanded stream, comprising a mixture of liquid and vapor, is fractionated in a distillation column to separate residual methane, nitrogen, and other volatile gases as overhead vapor from the carbon dioxide and the heavier hydrocarbon components as bottom liquid product. A portion of the liquid carbon dioxide can be flash expanded to lower pressure and thereafter used to provide low-level refrigeration to the process streams if desired. 
         [0007]    The present invention employs a novel means of condensing the distillation column overhead vapor to increase the carbon dioxide removal efficiency. Instead of cooling the column overhead vapor to condense reflux for the fractionation column, the overhead vapor is compressed to higher pressure and then cooled to partially condense it. The resulting condensate is mostly liquid carbon dioxide, which can be flash expanded to intermediate pressure and used to provide mid-level refrigeration to the process streams before being returned to the distillation column at a mid-column feed point. In addition, the residue gas that remains after the condensate has been removed is suitable to be sent to treating without requiring further compression. Surprisingly, applicants have found that this novel process arrangement not only allows removing more of the carbon dioxide, but also reduces the power consumption required to achieve a given level of carbon dioxide removal, thereby increasing the process efficiency and reducing the operating cost of the facility. 
         [0008]    In accordance with the present invention, it has been found that more than 75% of the carbon dioxide can be removed while leaving more than 99.8% of the methane and lighter components in the residue gas stream. 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 distillation column overhead temperatures of −50° F. [−46° C.] or colder. 
     
    
     
         [0009]    For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings: 
           [0010]      FIG. 1  is a flow diagram of a prior art synthesis gas processing plant; and 
           [0011]      FIG. 2  is a flow diagram of a synthesis gas processing plant in accordance with the present invention. 
       
    
    
       [0012]    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. 
         [0013]    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 
       [0014]      FIG. 1  is a process flow diagram showing the design of a processing plant to remove carbon dioxide from synthesis gas using a prior art process. In this simulation of the process, inlet gas enters the plant at 120° F. [49° C.] and 1080 psia [7,446 kPa(a)] as stream  31 . The feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid and liquid desiccants have both been used for this purpose. 
         [0015]    The feed stream  31  is cooled to −20° F. [−29° C.] in heat exchanger  10  by heat exchange with column reboiler liquids at 49° F. [9° C.] (stream  37 ), column side reboiler liquids at 34° F. [1° C.] (stream  42 ), and propane refrigerant. Stream  31   a  is further cooled in heat exchanger  50  by heat exchange with cool carbon dioxide vapor at −56° F. [−49° C.] (stream  43 ), cold residue gas at −60° F. [−51° C.] (stream  35 ), and pumped liquid at −60° F. [−51° C.] (stream  36   a ). The further cooled stream  31   b  enters separator  11  at −27° F. [−33° C.] and 1049 psia [7,233 kPa(a)] where the vapor (stream  32 ) is separated from the condensed liquid (stream  33 ). 
         [0016]    The vapor from separator  11  (stream  32 ) enters a work expansion machine  12  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  12  expands the vapor substantially isentropically to the operating pressure (approximately 665 psia [4,583 kPa(a)]) of fractionation tower  15 , with the work expansion cooling the expanded stream  32   a  to a temperature of approximately −48° F. [−45° C.]. The typical commercially available expanders are capable of recovering on the order of 80-88% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item  13 ) that can be used to re-compress the residue gas (stream  35   b ), for example. The partially condensed expanded stream  32   a  is thereafter supplied to fractionation tower  15  at its top column feed point. The separator liquid (stream  33 ) is expanded to the operating pressure of fractionation tower  15  by expansion valve  14 , cooling stream  33   a  to −28° F. [−33° C.] before it is supplied to fractionation tower  15  at an upper mid-column feed point. 
         [0017]    Overhead vapor stream  34  leaves fractionation tower  15  at −48° F. [−45° C.] and is cooled and partially condensed in heat exchanger  18 . The partially condensed stream  34   a  enters separator  19  at −60° F. [−51° C.] and 658 psia [4,535 kPa(a)] where the vapor (cold residue gas stream  35 ) is separated from the condensed liquid (stream  36 ). Liquid stream  36  is pumped to slightly above the operating pressure of fractionation tower  15  by pump  51  before stream  36   a  enters heat exchanger  50  and is heated to −26° F. [−32° C.] by heat exchange with the feed gas as described previously. The heated stream  36   b  is, thereafter supplied as feed to fractionation tower  15  at a lower mid-column feed point. 
         [0018]    Fractionation tower  15  is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. It 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 column bottom liquid product (stream  38 ) of methane and lighter components. The trays and/or packing provide the necessary contact between the stripping vapors rising upward and cold liquid falling downward, so that the bottom product stream  38  exits the bottom of the tower at 50° F. [10° C.], based on reducing the methane concentration in the bottom product to 0.47% on a molar basis. 
         [0019]    Column bottom product stream  38  is predominantly liquid carbon dioxide. A small portion (stream  39 ) is subcooled in heat exchanger  21  by cool residue gas stream  35   a . The subcooled liquid (stream  39   a ) at −20° F. [−29° C.] is expanded to lower pressure by expansion valve  22  and partially vaporized, further cooling stream  39   b  to −65° F. [−54° C.] before it enters heat exchanger  18 . The residual liquid in stream  39   b  functions as refrigerant in heat exchanger  18  to provide cooling of stream  34  as described previously, with the resulting carbon dioxide vapor leaving at −56° F. [−49° C.] as stream  43 . Since stream  39   b  could contain small amounts of heavier hydrocarbons, a small liquid purge (stream  44 ) may be drawn off from heat exchanger  18  to prevent an accumulation of heavier hydrocarbons in the refrigerant liquid that could elevate its boiling point and reduce the cooling efficiency of heat exchanger  18 . 
         [0020]    The cool carbon dioxide vapor from heat exchanger  18  (stream  43 ) is heated to −28° F. [−33° C.] in heat exchanger  50  by heat exchange with the feed gas as described previously. The warm carbon dioxide vapor (stream  43   a ) at 74 psia [508 kPa(a)] is then compressed to high pressure in three stages by compressors  23 ,  25 , and  27 , with cooling to 120° F. [49° C.] after each stage of compression by discharge coolers  24 ,  26 , and  28 . The remaining portion (stream  40 ) of column bottom product stream  38  is pumped to high pressure by pump  29  so that stream  40   a  can combine with the high pressure gas (stream  43   g ) leaving discharge cooler  28 , forming high pressure carbon dioxide stream  41  which then flows to reinjection at 82° F. [28° C.] and 1115 psia [7,688 kPa(a)]. 
         [0021]    The cool residue gas (stream  35   a ) leaves heat exchanger  50  at −28° F. [−33° C.] after heat exchange with the feed gas as described previously, and is further heated to −8° F. [−22° C.] in heat exchanger  21  by heat exchange with liquid carbon dioxide stream  39  as described previously. The warm residue gas stream  35   b  is then re-compressed in two stages, compressor  13  driven by expansion machine  12  and compressor  17  driven by a supplemental power source. Residue gas stream  35   d  then flows to treating at 90° F. [32° C.] and 1115 psia [7,688 kPa(a)]. 
         [0022]    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 
                 Hydrogen 
                 C. Monoxide 
                 Methane 
                 C. Dioxide 
                 Total 
               
               
                   
               
               
                 31 
                 22,177 
                 6,499 
                 2,014 
                 19,288 
                 50,115 
               
               
                 32 
                 21,992 
                 6,311 
                 1,901 
                 10,700 
                 41,036 
               
               
                 33 
                 185 
                 188 
                 113 
                 8,588 
                 9,079 
               
               
                 34 
                 22,201 
                 6,535 
                 1,981 
                 9,654 
                 40,509 
               
               
                 36 
                 24 
                 36 
                 24 
                 2,536 
                 2,622 
               
               
                 38 
                 0 
                 0 
                 57 
                 12,170 
                 12,228 
               
               
                 39 
                 0 
                 0 
                 15 
                 3,235 
                 3,250 
               
               
                 43 
                 0 
                 0 
                 15 
                 3,235 
                 3,250 
               
               
                 44 
                 0 
                 0 
                 0 
                 0 
                 0 
               
               
                 40 
                 0 
                 0 
                 42 
                 8,935 
                 8,978 
               
               
                 35 
                 22,177 
                 6,499 
                 1,957 
                 7,118 
                 37,887 
               
               
                 41 
                 0 
                 0 
                 57 
                 12,170 
                 12,228 
               
               
                   
               
             
          
           
               
                 Recovery/Removal* 
                   
                   
                   
                   
               
             
          
           
               
                 Methane and Lighter 
                 99.34% 
                 (recovered in the Residue Gas) 
               
               
                 Carbon Dioxide 
                 63.10% 
                 (removed from the Residue Gas) 
               
             
          
           
               
                 Carbon Dioxide Concentrations* 
                   
                   
                   
                   
               
               
                 Residue Gas 
                 18.79% 
                   
                   
                   
               
               
                 Carbon Dioxide Product 
                 99.50% 
                   
                   
                   
               
               
                 Power 
                   
                   
                   
                   
               
             
          
           
               
                 Carbon Dioxide Compression 
                 4,955 
                 HP 
                 [8,146 
                 kW] 
                   
                   
               
               
                 Residue Gas Compression 
                 5,717 
                 HP 
                 [9,398  
                 kW] 
                   
                   
               
               
                 Refrigerant Compression 
                 14,960 
                 HP 
                 [24,594 
                 kW] 
                   
                   
               
               
                 Carbon Dioxide Pump 
                 324 
                 HP 
                 [533 
                 kW] 
                   
                   
               
               
                 Totals 
                 25,956 
                 HP 
                 [42,671 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
             
          
         
       
     
       DESCRIPTION OF THE INVENTION 
       [0023]      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. 
         [0024]    In the simulation of the  FIG. 2  process, inlet gas enters the plant at 120° F. [49° C.] and 1080 psia [7,446 kPa(a)] as stream  31  and is cooled in heat exchanger  10  by heat exchange with column reboiler liquids at 47° F. [8° C.] (stream  37 ), residue gas at 30° F. [−1° C.] (stream  35   a ), cool expanded liquids at 20° F. [−7° C.] (stream  36   b ), and propane refrigerant. The cooled stream  31   a  enters separator  11  at −30° F. [−34° C.] and 1049 psia [7,233 kPa(a)] where the vapor (stream  32 ) is separated from the condensed liquid (stream  33 ). 
         [0025]    The vapor from separator  11  (stream  32 ) enters a work expansion machine  12  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  12  expands the vapor substantially isentropically to the operating pressure (approximately 640 psia [4,413 kPa(a)]) of fractionation tower  15 , with the work expansion cooling the expanded stream  32   a  to a temperature of approximately −54° F. [−48° C.]. The partially condensed expanded stream  32   a  is thereafter supplied to fractionation tower  15  at its top column feed point. The separator liquid (stream  33 ) is expanded to the operating pressure of fractionation tower  15  by expansion valve  14 , cooling stream  33   a  to −30° F. [−35° C.] before it is supplied to fractionation tower  15  at an upper mid-column feed point. 
         [0026]    Overhead vapor stream  34  leaves fractionation tower  15  at −52° F. [−47° C.] and is compressed in two stages, compressor  13  driven by expansion machine  12  and compressor  17  driven by a supplemental power source. The compressed stream  34   b  is then cooled and partially condensed in heat exchanger  18 . The partially condensed stream  34   c  enters separator  19  at −60° F. [−51° C.] and 1130 psia [7,791 kPa(a)] where the vapor (cold residue gas stream  35 ) is separated from the condensed liquid (stream  36 ). Liquid stream  36  is expanded to slightly above the operating pressure of fractionation tower  15  by expansion valve  20  before stream  36   a  enters heat exchanger  21 . The expanded stream  36   a  is heated from −59° F. [−51° C.] to 20° F. [−7° C.] and partially vaporized by heat exchange with liquid carbon dioxide stream  39  (which is described further below in paragraph [0028]). The partially vaporized stream  36   b  is further vaporized in heat exchanger  10  by heat exchange with the feed gas as described previously, and stream  36   c  at 38° F. [3° C.] is thereafter supplied as feed to fractionation tower  15  at a lower mid-column feed point. 
         [0027]    Fractionation tower  15  is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. It also includes reboilers (such as the reboiler described previously, and optionally a reboiler  16  heated by an external source of heat) 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 column bottom liquid product (stream  38 ) of methane and lighter components. The trays and/or packing provide the necessary contact between the stripping vapors rising upward and cold liquid falling downward, so that the bottom product stream  38  exits the bottom of the tower at 48° F. [9° C.], based on reducing the methane concentration in the bottom product to 0.30% on a molar basis. 
         [0028]    Column bottom product stream  38  is predominantly liquid carbon dioxide. A minor portion (stream  39 ) is subcooled in heat exchanger  21  by flash expanded liquid stream  36   a  as described previously. The subcooled liquid (stream  39   a ) at −33° F. [−36° C.] is expanded to lower pressure by expansion valve  22  and partially vaporized, further cooling stream  39   b  to −65° F. [−54° C.] before it enters heat exchanger  18 . The residual liquid in stream  39   b  functions as refrigerant in heat exchanger  18  to provide a portion of the cooling of compressed overhead vapor stream  34   b  as described previously, with the resulting carbon dioxide vapor leaving at 22° F. [−6° C.] (stream  39   c ). 
         [0029]    The warm carbon dioxide vapor (stream  39   c ) at 78 psia [536 kPa(a)] is then compressed to high pressure in three stages by compressors  23 ,  25 , and  27 , with cooling to 120° F. [49° C.] after each stage of compression by discharge coolers  24 ,  26 , and  28 . The remaining portion (stream  40 ) of column bottom product stream  38  is pumped to high pressure by pump  29  so that stream  40   a  can combine with the high pressure gas (stream  39   i ) leaving discharge cooler  28 , forming high pressure carbon dioxide stream  41  which then flows to reinjection at 84° F. [29° C.] and 1115 psia [7,688 kPa(a)]. 
         [0030]    The cold residue gas (stream  35 ) from separator  19  enters heat exchanger  18  and is heated to 30° F. [−1° C.] by heat exchange with compressed overhead vapor stream  34   b  as described previously. Cool residue gas stream  35   a  is further heated to 72° F. [22° C.] in heat exchanger  10  by heat exchange with the feed gas as described previously. The warm residue gas stream  35   b  then flows to treating at 1115 psia [7,688 kPa(a)]. 
         [0031]    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 
                 Hydrogen 
                 C. Monoxide 
                 Methane 
                 C. Dioxide 
                 Total 
               
               
                   
               
               
                 31 
                 22,177 
                 6,499 
                 2,014 
                 19,288 
                 50,115 
               
               
                 32 
                 21,984 
                 6,297 
                 1,892 
                 10,038 
                 40,343 
               
               
                 33 
                 193 
                 202 
                 122 
                 9,250 
                 9,772 
               
               
                 34 
                 22,256 
                 6,607 
                 2,041 
                 9,052 
                 40,092 
               
               
                 36 
                 79 
                 109 
                 71 
                 4,257 
                 4,517 
               
               
                 38 
                 0 
                 1 
                 44 
                 14,493 
                 14,540 
               
               
                 39 
                 0 
                 0 
                 14 
                 4,493 
                 4,507 
               
               
                 40 
                 0 
                 1 
                 30 
                 10,000 
                 10,033 
               
               
                 35 
                 22,177 
                 6,498 
                 1,970 
                 4,795 
                 35,575 
               
               
                 41 
                 0 
                 1 
                 44 
                 14,493 
                 14,540 
               
               
                   
               
             
          
           
               
                 Recovery/Removal* 
                   
                   
                   
                   
               
             
          
           
               
                 Methane and Lighter 
                 99.85% 
                 (recovered in the Residue Gas) 
               
               
                 Carbon Dioxide 
                 75.15% 
                 (removed from the Residue Gas) 
               
             
          
           
               
                 Carbon Dioxide Concentrations* 
                   
                   
                   
                   
               
               
                 Residue Gas 
                 13.47% 
                   
                   
                   
               
               
                 Carbon Dioxide Product 
                 99.69% 
                   
                   
                   
               
               
                 Power 
                   
                   
                   
                   
               
             
          
           
               
                 Carbon Dioxide Compression 
                 6,742  
                 HP 
                 [11,084  
                 kW] 
                   
                   
               
               
                 Overhead Vapor Compression 
                 5,095  
                 HP 
                 [8,376  
                 kW] 
                   
                   
               
               
                 Refrigerant Compression 
                 16,184  
                 HP 
                 [26,606 
                 kW] 
                   
                   
               
               
                 Carbon Dioxide Pump 
                 378  
                 HP 
                 [621 
                 kW] 
                   
                   
               
               
                 Totals 
                 28,399  
                 HP 
                 [46,687  
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
             
          
         
       
     
         [0032]    A comparison of Tables I and II shows that, compared to the prior art, the present invention provides better methane recovery (99.85%, versus 99.44% for the prior art), much better carbon dioxide removal (75.15%, versus 63.10% for the prior art), much lower carbon dioxide concentration in the residue gas (13.47%, versus 18.79% for the prior art), and better carbon dioxide purity (99.69%, versus 99.50% for the prior art). In addition, further comparison of Tables I and II shows that this superior process performance was achieved using less power per unit of carbon dioxide removed than the prior art. In terms of the specific power consumption, the present invention represents an 8% improvement over the prior art of the  FIG. 1  process, reducing the specific power consumption from 2.13 HP-H/Lb·mole [3.51 kW-H/kg mole] of carbon dioxide removed for the prior art to 1.96 HP-H/Lb·mole [3.22 kW-H/kg mole] for the present invention. 
         [0033]    The improvement in energy efficiency provided by the present invention over that of the prior art of the  FIG. 1  process is primarily due to two factors. First, compressing overhead vapor stream  34  from fractionation tower  15  to higher pressure before supplying it to heat exchanger  18  makes it much easier to condense carbon dioxide from the stream. As can be seen by comparing stream  36  in Tables I and II, the carbon dioxide condensed in stream  36  increases from 2,536 Lb. Moles/Hr [2,536 kg moles/Hr] for the prior art to 4,257 Lb. Moles/Hr [4,257 kg moles/Hr] for the present invention. The result is that the residue gas that remains (stream  35 ) contains much less carbon dioxide, 4,795 Lb. Moles/Hr [4,795 kg moles/Hr] for the present invention versus 7,118 Lb. Moles/Hr [7,118 kg moles/Hr] for the prior art. 
         [0034]    Second, the greater quantity of liquid condensed in stream  36  for the present invention provides a process stream that can be used more effectively for mid-level refrigeration within the process. The resulting flashed stream  36   a  has 72% more flow than pumped stream  36   a  in the prior art process, allowing it to subcool a larger quantity of liquid carbon dioxide in stream  39  (39% more than the prior art) to a lower temperature (−33° F. [−36° C.], versus −20° F. [−29° C.] for the prior art), so that the resulting flashed carbon dioxide stream  39   b  for the present invention contains a much larger quantity of liquid that can be used as refrigerant to condense carbon dioxide from overhead vapor stream  34  in heat exchanger  18 . 
         [0035]    The net result of these two factors is to capture significantly more of the carbon dioxide in column bottom product stream  38  (19% more compared to the  FIG. 1  prior art process) at greater efficiency using less specific power. This also means that much less of the carbon dioxide remains in residue gas stream  35 , greatly reducing (or perhaps eliminating entirely) the downstream treating needed to condition the residue gas for subsequent processing or use, further reducing the total treating cost for a given application. 
       Other Embodiments 
       [0036]    As described earlier for the embodiment of the present invention shown in  FIG. 2 , feed stream  31  is partially condensed as it is cooled in heat exchanger  10 , and the resulting vapor stream  32  and liquid stream  33  are then expanded to the operating pressure of fractionation column  15 . However, the present invention is not limited to this embodiment. Depending on the quantity of heavier hydrocarbons in the feed gas and the feed gas pressure, the cooled feed stream  31   a  may not contain any liquid (because it is above its dewpoint, or because it is above its cricondenbar). In such cases, separator  11  is not required. Some circumstances may favor total condensation of the feed gas, followed by expanding the liquid or dense phase fluid to the operating pressure of fractionation column  15 . Such cases may likewise not require separator  11 . 
         [0037]    Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine  12 , 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 liquid streams  33 ,  36 , and/or  39   a.    
         [0038]    In accordance with the present invention, the use of external refrigeration to supplement the cooling available to the inlet gas and/or compressed overhead vapor stream  34   b  from other process streams may be employed, particularly in the case of a rich inlet gas. The use and distribution of separator liquids and/or demethanizer side draw liquids for process heat exchange, and the particular arrangement of heat exchangers for inlet gas cooling must be evaluated for each particular application, as well as the choice of process streams for specific heat exchange services. For instance, some circumstances may favor supplying partially vaporized stream  36   b  directly to fractionation tower  15  (via stream  44  in  FIG. 2 ) rather than further vaporizing it in heat exchanger  10  and then supplying the resulting stream  36   c  to fractionation tower  15 . 
         [0039]    Depending on the temperature and richness of the feed gas and the amount of methane allowed in liquid product stream  38 , there may not be sufficient heating available from feed stream  31  to cause the liquid leaving fractionation column  15  to meet the product specifications. In such cases, the fractionation column  15  may include one or more reboilers (such as reboiler  16 ) heated by an external source of heat. 
         [0040]    In some circumstances, the portion (stream  39 ) of column bottom product stream  38  that is used to provide refrigeration may not need to be restored to high pressure after it has been heated (stream  39   c ). In such cases, the compression and cooling shown (compressors  23 ,  25 , and  27  and discharge coolers  24 ,  26 , and  28 ) may not be needed, and only stream  40   a  flows to stream  41 . 
         [0041]    The present invention provides improved separation of carbon dioxide from hydrocarbon gas streams 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 pumping, reduced power requirements for external refrigeration, reduced energy requirements for tower reboiling, or a combination thereof. 
         [0042]    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.