Patent Publication Number: US-11643604-B2

Title: Hydrocarbon gas processing

Description:
This invention relates to a process and apparatus for the separation of a gas containing hydrocarbons and significant quantities of components more volatile than methane (e.g., hydrogen, nitrogen, etc.). The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application No. 62/923,075 which was filed on Oct. 18, 2019. 
    
    
     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. Hydrocarbon bearing gas often contains components more volatile than methane as well as unsaturated hydrocarbons (e.g., ethylene, propylene, etc.) in addition to methane, ethane and hydrocarbons of higher molecular weight such as propane, butane, and pentane. 
     The present invention is generally concerned with improving 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, 42.4% methane, 22.8% ethane and other C 2  components, 7.6% propane and other C 3  components, 3.1% iso-butane, 2.7% normal butane, and 2.7% pentanes plus, with the balance made up of hydrogen, nitrogen, carbon monoxide, and carbon dioxide. Sulfur-containing gases are also sometimes present. 
     Recent changes in ethylene demand have created increased markets for ethylene and derivative products. In addition, fluctuations in the prices of both natural gas and its natural gas liquid (NGL) constituents have increased the incremental value of ethane, ethylene, propane, propylene, and heavier components as liquid products. These market conditions have resulted in a demand for processes that can provide high recoveries and 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; 8,590,340; 8,881,549; 8,919,148; 9,021,831; 9,021,832; 9,052,136; 9,052,137; 9,057,558; 9,068,774; 9,074,814; 9,080,810; 9,080,811; 9,476,639; 9,637,428; 9,783,470; 9,927,171; 9,933,207; 9,939,195; 10,227,273; 10,553,794; 10,551,118; 10,551,119; and 10,753,678; reissue U.S. Pat. No. 33,408; published US applications US20080078205A1; US20110067441A1; US20110067443A1; US2016/0069610A1; US2016/0377341A1; US 2018-0347898 A1; US2018/0347899A1; US 2019/0170435A1; and US 2020/0292230A1 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). 
     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) 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. 
     If the feed gas is not totally condensed (typically it is not), the vapor remaining from the partial condensation can be passed through a work expansion machine or engine, or an expansion valve, to a lower pressure at which additional liquids are condensed as a result of further cooling of the stream. The pressure after expansion is essentially the same as the pressure at which the distillation column is operated. The combined vapor-liquid phases resulting from the expansion are supplied as feed to the column. 
     In the ideal operation of such a separation process, the residue gas leaving the process will contain substantially all of the methane and more volatile components in the feed gas with essentially none of the heavier hydrocarbon components and the bottoms fraction leaving the demethanizer will contain substantially all of the heavier hydrocarbon components with essentially no methane or more volatile components. In practice, however, this ideal situation is not obtained because the conventional demethanizer is operated largely as a stripping column. The residue gas from 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. This problem is exacerbated if the gas stream(s) being processed contain relatively large quantities of components more volatile than methane (e.g., hydrogen, nitrogen, etc.), because the volatile vapors rising up the column strip C 2 + components from the liquids flowing downward. The loss of these desirable C 2 + components could be significantly reduced if the rising vapors could be brought into contact with a significant quantity of liquid (reflux) capable of absorbing the C 2  components, 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. For many of these processes, a portion of the vapor remaining from the partial condensation of the feed gas 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. Unfortunately, this feed stream is not very effective at capturing the desired C 2 + components when the feed gas contains significant quantities of components more volatile than methane because the stream cannot be substantially condensed. This results in large amounts of flash vapor in the stream, which carries away equilibrium quantities of the C 2 + components rather than recovering them in the column. 
     Many processes combine this condensed flash expanded stream with another source of top reflux to the column. The source of this reflux stream is 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. Typical process schemes of this type are disclosed in U.S. Pat. Nos. 4,889,545; 5,568,737; 5,881,569; 9,052,137; and 9,080,811, and in Mowrey, E. Ross, “Efficient, High Recovery of Liquids from Natural Gas Utilizing a High Pressure Absorber”, Proceedings of the Eighty-First Annual Convention of the Gas Processors Association, Dallas, Tex., Mar. 11-13, 2002. Unfortunately, this method is not efficient when the residue gas contains significant quantities of components more volatile than methane because the recycle stream cannot be cooled to substantial condensation. 
     Another means of providing a reflux stream for the upper rectification section is to withdraw a distillation vapor stream from a lower location on the tower (and perhaps combine it with a portion of the tower overhead vapor). This vapor (or combined vapor) stream is compressed to higher pressure, then cooled to substantial condensation, expanded to the tower operating pressure, and supplied as top feed to the tower. Typical process schemes of this type are disclosed in U.S. Pat. Nos. 5,275,005 and 9,476,639 and in published application nos. US 2008-0078205 A1 and US 2011-0067443 A1. However, this stream is generally not sufficient to provide the desired rectification by itself and combining it with the condensed flash expanded stream described earlier is not effective when there is a significant quantity of components more volatile than methane in the feed gas. 
     The present invention also employs an upper rectification section (or a separate rectification column in some embodiments). However, the reflux for this upper rectification section is provided by cooling a liquid stream derived from the feed gas and then expanding the stream to the operating pressure of the fractionation tower. During expansion, a portion of the stream is vaporized, resulting in cooling of the total stream. The cooled, expanded stream is supplied to the tower at the top column feed point, where it can then be used to absorb C 2  components, C 3  components, C 4  components, and heavier hydrocarbon components from the vapors rising through the rectification section and thereby capture these valuable components in the bottom liquid product from the demethanizer. Surprisingly, the present invention can achieve high recovery without using a second reflux stream that is predominantly methane as required in some of the prior art. This eliminates the reflux compressor used in that prior art, reducing capital cost and operating cost compared to that prior art. 
     In accordance with the present invention, it has been found that C 2  component recovery in excess of 95% and C 3  component and C 4 + component recoveries in excess of 99% can be obtained. In addition, the present invention makes possible essentially 100% separation of methane and lighter components from the C 2  components and heavier components at the same energy requirements compared to the prior art while increasing the recovery level. The present invention is particularly advantageous when processing feed gases that contain more than 10 mole % of components more volatile than methane (e.g., hydrogen, nitrogen, etc.). 
     For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings: 
    
    
     
       BRIEF DESCRIPTION OF THE DRAWINGS 
         FIG.  1    is a flow diagram of a prior art gas processing plant in accordance with co-pending application Ser. No. 11/839,693; 
         FIG.  2    is a flow diagram of a prior art gas processing plant in accordance with U.S. Pat. No. 5,275,005; 
         FIG.  3    is a flow diagram of a prior art gas processing plant in accordance with U.S. Pat. No. 4,171,964; 
         FIG.  4    is a flow diagram of a gas processing plant in accordance with the present invention; and 
         FIG.  5    is a flow diagram illustrating alternative means of application of the present invention to a gas stream. 
     
    
    
     In the following explanation of the above figures, tables are provided summarizing flow rates calculated for representative process conditions. In the tables appearing herein, the values for flow rates (in moles per hour) have been rounded to the nearest whole number for convenience. The total stream rates shown in the tables include all non-hydrocarbon components and hence are generally larger than the sum of the stream flow rates for the hydrocarbon components. Temperatures indicated are approximate values rounded to the nearest degree. It should also be noted that the process design calculations performed for the purpose of comparing the processes depicted in the figures are based on the assumption of no heat leak from (or to) the surroundings to (or from) the process. The quality of commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art. 
     For convenience, process parameters are reported in both the traditional British units and in the units of the Système International d&#39;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 a gas stream using prior art according to co-pending application US 2008-0078205 A1. In this simulation of the process, inlet gas enters the plant at 142° F. [61° C.] and 798 psia [5,500 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  enters compressor  16  driven by expansion machine  15  and is boosted to higher pressure. After cooling to 142° F. [61° C.] in cooler  20 , stream  31   b  at 1146 psia [7,901 kPa(a)] is cooled in heat exchanger  10  by heat exchange with cool residue gas (stream  39   a ) at −57° F. [−49° C.], pumped liquid product (stream  42   a ) at 12° F. [−11° C.], demethanizer reboiler liquids (stream  41 ) at −11° F. [−24° C.], and propane refrigerant. Note that in all cases exchanger  10  is representative of either a multitude of individual heat exchangers or a single multi-pass heat exchanger, or any combination thereof. (The decision as to whether to use more than one heat exchanger for the indicated cooling services will depend on a number of factors including, but not limited to, inlet gas flow rate, heat exchanger size, stream temperatures, etc.) The cooled stream  31   c  enters separator  11  at −14° F. [−26° C.] and 1123 psia [7,743 kPa(a)] where the vapor (stream  32 ) is separated from the condensed liquid (stream  35 ). The separator liquid (stream  35 ) is expanded to the operating pressure (approximately 141 psia [972 kPa(a)]) of fractionation tower  19  by expansion valve  18 , cooling stream  35   a  to −64° F. [−53° C.] before it is supplied to fractionation tower  19  at a lower mid-column feed point. 
     The vapor (stream  32 ) from separator  11  is divided into two streams,  33  and  34 . Stream  33 , containing about 5% of the total vapor, passes through heat exchanger  12  in heat exchange relation with cold distillation stream  39  at −143° F. [−97° C.] where it is cooled to substantial condensation. The resulting substantially condensed stream  33   a  at −119° F. [−84° C.] is then flash expanded through an appropriate expansion device, such as expansion valve  17 , to the operating pressure of fractionation tower  19 . During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in  FIG.  1   , the expanded stream  33   b  leaving expansion valve  17  reaches a temperature of −178° F. [−117° C.] and is supplied to fractionation tower  19  at an upper mid-column feed point. 
     The remaining vapor from separator  11  (stream  34 ) enters work expansion machine  15  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  15  expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream  34   a  to a temperature of approximately −132° F. [−91° 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 compress the feed gas (stream  31 ), for example. The partially condensed expanded stream  34   a  is thereafter supplied as feed to fractionation tower  19  at a mid-column feed point. 
     The demethanizer in tower  19  is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. As is often the case in gas processing plants, the fractionation tower may consist of three sections. The upper section  19   a  is a separator wherein the partially vaporized top feed is divided into its respective vapor and liquid portions, and wherein the vapor rising from the middle absorbing (rectifying) section  19   b  is combined with the vapor portion of the top feed to form the cold demethanizer overhead vapor (stream  39 ) which exits the top of the tower at −143° F. [−97° C.]. The lower stripping (demethanizing) section  19   c  contains the trays and/or packing and provides the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section  19   c  also includes reboilers (such as the 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  42 , of methane and lighter components. Stream  34   a  enters demethanizer  19  at an intermediate feed position located in the lower region of rectifying section  19   b  of demethanizer  19 . The liquid portion of the expanded stream commingles with liquids falling downward from rectifying section  19   b  and the combined liquid continues downward into demethanizing section  19   c  of demethanizer  19 . The vapor portion of the expanded stream commingles with vapors arising from demethanizing section  19   c  and the combined vapor rises upward through rectifying section  19   b  and is contacted with cold liquid falling downward to condense and absorb the C 2  components, C 3  components, and heavier components. 
     A portion of the distillation vapor (stream  40 ) is withdrawn from an upper region of demethanizing section  19   c  in fractionation column  19 , below the feed position of expanded stream  34   a  in the lower region of rectifying section  19   b . The distillation vapor stream  40  at −93° F. [−77° C.] is heated to 132° F. [55° C.] in heat exchanger  22  as it cools compressed stream  40   e . Heated stream  40   a  is compressed to 412 psia [2,839 kPa(a)] (stream  40   d ) in two stages by reflux compressors  23  and  25 , with cooling to 142° F. [61° C.] after each stage in coolers  24  and  26 . The cooled compressed stream  40   e  is cooled to −55° F. [−48° C.] (stream  40   f ) in heat exchanger  22  as described earlier, then further cooled to −119° F. [−84° C.] and substantially condensed (stream  40   g ) in heat exchanger  12  by heat exchange with cold demethanizer overhead stream  39  as described previously. The cold residue gas stream is warmed (stream  39   a ) as it provides cooling to compressed distillation vapor stream  40   f.    
     The substantially condensed stream  40   g  is flash expanded to the operating pressure of demethanizer  19  by expansion valve  14 . A portion of the stream is vaporized, further cooling stream  40   h  to −157° F. [−105° C.] before it is supplied as cold top column feed (reflux) to separator section  19   a  in the upper region of fractionation tower  19 . The liquids separated therein become the top feed to rectifying section  19   b  and the cold liquid reflux absorbs and condenses the C 2  components, C 3  components, and heavier components rising in the upper region of rectifying section  19   b  of demethanizer  19 . 
     Liquid product stream  42  exits the bottom of the tower at 8° F. [−13° C.], based on a typical specification of a methane to ethane ratio of 0.010:1 on a molar basis in the bottom product. It is pumped to a pressure of approximately 455 psia [3,135 kPa(a)] in demethanizer bottoms pump  21 , and the pumped liquid product is then warmed to 83° F. [28° C.] as it provides cooling of stream  31   b  in exchanger  10  before flowing to storage. 
     The residue gas (demethanizer overhead vapor stream  39 ) passes countercurrently to the incoming feed gas in heat exchanger  12  where it is heated to −57° F. [−49° C.] (stream  39   a ) and in heat exchanger  10  where it is heated to 137° F. [58° C.] (stream  39   b ). The residue gas product (stream  39   b ) then flows to the fuel gas system at 116 psia [801 kPa(a)]. 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG.  1    is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE I 
               
               
                   
               
               
                 (FIG. 1) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
            
               
                   
               
            
           
           
               
               
               
               
               
               
               
            
               
                 Stream 
                 Hydrogen 
                 Methane 
                 C 2  Comp. 
                 C 3  Comp. 
                 C 4 + Comp. 
                 Total 
               
               
                   
               
               
                 31 
                 1,998 
                 4,661 
                 2,505 
                 839 
                 939 
                 10,980 
               
               
                 32 
                 1,834 
                 2,994 
                 664 
                 85 
                 31 
                 5,640 
               
               
                 35 
                 164 
                 1,667 
                 1,841 
                 754 
                 908 
                 5,340 
               
               
                 33 
                 92 
                 150 
                 33 
                 4 
                 2 
                 282 
               
               
                 34 
                 1,742 
                 2,844 
                 631 
                 81 
                 29 
                 5,358 
               
               
                 40 
                 125 
                 1,373 
                 467 
                 6 
                 0 
                 1,976 
               
               
                 39 
                 1,998 
                 4,640 
                 459 
                 0 
                 0 
                 7,136 
               
               
                 42 
                 0 
                 21 
                 2,046 
                 839 
                 939 
                 3,844 
               
               
                   
               
            
           
           
               
               
               
            
               
                   
                 Recoveries* 
               
               
                   
                   
               
               
                   
                 C 2  Components 
                 81.68% 
               
               
                   
                 C 3  Components 
                 99.95% 
               
               
                   
                 C 4 + Components 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
               
               
               
            
               
                   
                 Power 
               
               
                   
                   
               
               
                   
                 Reflux Compression 
                 1,687 HP 
                 [2,773 kW] 
               
               
                   
                 Refrigerant Compression 
                 2,921 HP 
                 [4,802 kW] 
               
               
                   
                 Total Compression 
                 4,608 HP 
                 [7,575 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
       FIG.  2    represents an alternative prior art process in accordance with U.S. Pat. No. 5,275,005. The process of  FIG.  2    has been applied to the same feed gas composition and conditions as described above for  FIG.  1   . Accordingly, the  FIG.  2    process can be compared with that of the  FIG.  1    process. 
     The feed stream  31  enters compressor  16  driven by expansion machine  15  and is boosted to higher pressure. After cooling to 142° F. [61° C.] in cooler  20 , stream  31   b  at 1166 psia [8,042 kPa(a)] is cooled in heat exchanger  10  by heat exchange with cool residue gas (stream  39   a ) at −53° F. [−47° C.], pumped liquid product (stream  42   a ) at 11° F. [−11° C.], demethanizer reboiler liquids (stream  41 ) at −12° F. [−24° C.], and propane refrigerant. The cooled stream  31   c  enters separator  11  at −14° F. [−25° C.] and 1144 psia [7,884 kPa(a)] where the vapor (stream  32 ) is separated from the condensed liquid (stream  35 ). The separator liquid (stream  35 ) is expanded to the operating pressure (approximately 141 psia [972 kPa(a)]) of fractionation tower  19  by expansion valve  18 , cooling stream  35   a  to −64° F. [−53° C.] before it is supplied to fractionation tower  19  at a lower mid-column feed point. 
     The vapor (stream  32 ) from separator  11  enters 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  32   a  to a temperature of approximately −133° F. [−92° C.]. The partially condensed expanded stream  32   a  is thereafter supplied as feed to fractionation tower  19  at a mid-column feed point. 
     A portion of the distillation vapor (stream  40 ) is withdrawn from an intermediate region of fractionation column  19 , below the feed position of expanded stream  32   a . The distillation vapor stream  40  at −86° F. [−66° C.] is heated to 132° F. [55° C.] in heat exchanger  22  as it cools compressed stream  40   e . Heated stream  40   a  is compressed to 411 psia [2,836 kPa(a)] (stream  40   d ) in two stages by reflux compressors  23  and  25 , with cooling to 142° F. [61° C.] after each stage in coolers  24  and  26 . The cooled compressed stream  40   e  is cooled to −46° F. [−44° C.] (stream  40   f ) in heat exchanger  22  as described earlier, then further cooled to −141° F. [−96° C.] and substantially condensed (stream  40   g ) in heat exchanger  12  by heat exchange with cold demethanizer overhead stream  39  as described previously and with distillation liquid stream  43  at −127° F. [−88° C.] which is withdrawn from a region of demethanizer  19  immediately below the feed point of expanded stream  32   a . The cold residue gas stream and the distillation liquid stream are warmed as they provide cooling to compressed distillation vapor stream  40   f , and the warmed distillation liquid stream  43   a  returns to demethanizer  19  at −82° F. [−63° C.]. 
     The substantially condensed stream  40   g  is flash expanded to the operating pressure of demethanizer  19  by expansion valve  14 . A portion of the stream is vaporized, further cooling stream  40   h  to −176° F. [−116° C.] before it is supplied as cold top column feed (reflux) to fractionation tower  19 . The cold liquid reflux absorbs and condenses the C 2  components, C 3  components, and heavier components rising in the upper region of demethanizer  19 . 
     Liquid product stream  42  exits the bottom of the tower at 7° F. [−14° C.], based on a typical specification of a methane to ethane ratio of 0.010:1 on a molar basis in the bottom product. It is pumped to a pressure of approximately 455 psia [3,135 kPa(a)] in demethanizer bottoms pump  21 , and the pumped liquid product is then warmed to 83° F. [28° C.] as it provides cooling of stream  31   b  in exchanger  10  before flowing to storage. 
     The residue gas (demethanizer overhead vapor stream  39 ) passes countercurrently to cooled compressed distillation vapor stream  40   f  in heat exchanger  12  where it is heated to −53° F. [−47° C.] (stream  39   a ), and to the incoming feed gas in heat exchanger  10  where it is heated to 137° F. [58° C.] (stream  39   b ). The residue gas product (stream  39   b ) then flows to the fuel gas system at 116 psia [801 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 
                 Hydrogen 
                 Methane 
                 C 2  Comp. 
                 C 3  Comp. 
                 C 4 + Comp. 
                 Total 
               
               
                   
               
               
                 31 
                 1,998 
                 4,661 
                 2,505 
                 839 
                 939 
                 10,980 
               
               
                 32 
                 1,827 
                 2,963 
                 655 
                 85 
                 31 
                 5,592 
               
               
                 35 
                 171 
                 1,698 
                 1,850 
                 754 
                 908 
                 5,388 
               
               
                 40 
                 118 
                 1,288 
                 559 
                 7 
                 0 
                 1,976 
               
               
                 43 
                 2 
                 226 
                 1,099 
                 94 
                 32 
                 1,454 
               
               
                 39 
                 1,998 
                 4,640 
                 381 
                 0 
                 0 
                 7,056 
               
               
                 42 
                 0 
                 21 
                 2,124 
                 839 
                 939 
                 3,924 
               
               
                   
               
            
           
           
               
               
               
            
               
                   
                 Recoveries* 
               
               
                   
                   
               
               
                   
                 C 2  Components 
                 84.81% 
               
               
                   
                 C 3  Components 
                 99.97% 
               
               
                   
                 C 4 + Components 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
               
               
               
            
               
                   
                 Power 
               
               
                   
                   
               
               
                   
                 Reflux Compression 
                 1,682 HP 
                 [2,765 kW] 
               
               
                   
                 Refrigerant Compression 
                 2,930 HP 
                 [4,817 kW] 
               
               
                   
                 Total Compression 
                 4,612 HP 
                 [7,582 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     A comparison of Tables I and II shows that, compared to the  FIG.  1    process, the  FIG.  2    process improves C 2  component recovery from 81.68% to 84.81% and C 3  component recovery from 99.95% to 99.97%, and the C 4 + component recovery remains the same at 100.00%. Comparison of Tables I and II further shows that these increased product yields were achieved without using additional power. 
       FIG.  3    represents another alternative prior art process in accordance with U.S. Pat. No. 4,171,964. The process of  FIG.  3    has been applied to the same feed gas composition and conditions as described above for  FIG.  1    and  FIG.  2   . Accordingly, the  FIG.  3    process can be compared with that of the  FIG.  1    and  FIG.  2    processes. 
     The feed stream  31  enters compressor  16  driven by expansion machine  15  and is boosted to higher pressure. After cooling to 142° F. [61° C.] in cooler  20 , stream  31   b  at 1060 psia [7,311 kPa(a)] is cooled in heat exchanger  10  by heat exchange with cold residue gas (stream  39 ) at −155° F. [−104° C.], pumped liquid product (stream  42   a ) at 9° F. [−13° C.], demethanizer reboiler liquids (stream  41 ) at −13° F. [−25° C.], and propane refrigerant. The cooled stream  31   c  enters separator  11  at −51° F. [−46° C.] and 1037 psia [7,153 kPa(a)] where the vapor (stream  32 ) is separated from the condensed liquid (stream  35 ). 
     The vapor (stream  32 ) from separator  11  enters 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 (approximately 141 psia [972 kPa(a)]) of fractionation tower  19 , with the work expansion cooling the expanded stream  32   a  to a temperature of approximately −162° F. [−108° C.]. The partially condensed expanded stream  32   a  is thereafter supplied as feed to fractionation tower  19  at a mid-column feed point. 
     The separator liquid (stream  35 ) is cooled to −141° F. [−96° C.] in heat exchanger  13 , and cooled liquid stream  35   a  is then divided into two streams, stream  36  and stream  37 . Stream  37  is expanded to slightly above the operating pressure of fractionation tower  19  by expansion valve  18 , cooling stream  37   a  to −152° F. [−102° C.] before it is heated as it supplies the cooling in heat exchanger  13 . The warmed stream  37   b  at −57° F. [−49° C.] is then supplied to fractionation tower  19  at a lower mid-column feed point. 
     The remaining portion of cooled liquid stream  35   a , stream  36 , is flash expanded to the operating pressure of demethanizer  19  by expansion valve  17 . A portion of the stream is vaporized, further cooling stream  36   a  to −162° F. [−108° C.] before it is supplied as cold top column feed (reflux) to fractionation tower  19 . The cold liquid reflux absorbs and condenses the C 2  components, C 3  components, and heavier components rising in the upper region of demethanizer  19 . 
     Liquid product stream  42  exits the bottom of the tower at 4° F. [−15° C.], based on a typical specification of a methane to ethane ratio of 0.010:1 on a molar basis in the bottom product. It is pumped to a pressure of approximately 455 psia [3,135 kPa(a)] in demethanizer bottoms pump  21 , and the pumped liquid product is then warmed to 83° F. [28° C.] as it provides cooling of stream  31   b  in exchanger  10  before flowing to storage. 
     The residue gas (demethanizer overhead vapor stream  39 ) passes countercurrently to the incoming feed gas in heat exchanger  10  where it is heated to 136° F. [58° C.] (stream  39   a ). The residue gas product (stream  39   a ) then flows to the fuel gas system at 116 psia [801 kPa(a)]. 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG.  3    is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE III 
               
               
                   
               
               
                 (FIG. 3) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
            
               
                   
               
            
           
           
               
               
               
               
               
               
               
            
               
                 Stream 
                 Hydrogen 
                 Methane 
                 C 2  Comp. 
                 C 3  Comp. 
                 C 4 + Comp. 
                 Total 
               
               
                   
               
               
                 31 
                 1,998 
                 4,661 
                 2,505 
                 839 
                 939 
                 10,980 
               
               
                 32 
                 1,820 
                 2,464 
                 347 
                 32 
                 9 
                 4,702 
               
               
                 35 
                 178 
                 2,197 
                 2,158 
                 807 
                 930 
                 6,278 
               
               
                 36 
                 77 
                 945 
                 928 
                 347 
                 400 
                 2,700 
               
               
                 37 
                 101 
                 1,252 
                 1,230 
                 460 
                 530 
                 3,578 
               
               
                 39 
                 1,998 
                 4,638 
                 159 
                 4 
                 0 
                 6,835 
               
               
                 42 
                 0 
                 23 
                 2,346 
                 835 
                 939 
                 4,145 
               
               
                   
               
            
           
           
               
               
               
            
               
                   
                 Recoveries* 
               
               
                   
                   
               
               
                   
                 C 2  Components 
                 93.68% 
               
               
                   
                 C 3  Components 
                 99.56% 
               
               
                   
                 C 4 + Components 
                 99.96% 
               
               
                   
                   
               
            
           
           
               
               
               
               
            
               
                   
                 Power 
               
               
                   
                   
               
               
                   
                 Refrigerant Compression 
                 4,608 HP 
                 [7,575 kW] 
               
               
                   
                 Total Compression 
                 4,608 HP 
                 [7,575 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     A comparison of Tables II and III shows that, compared to the  FIG.  2    process, the  FIG.  3    process improves C 2  component recovery from 84.81% to 93.68%, but C 3  component recovery drops from 99.97% to 99.56% and C 4 + component recovery drops from 100.00% to 99.96%. Comparison of Tables II and III further shows that these product yields were achieved using the same power. 
     DESCRIPTION OF THE INVENTION 
       FIG.  4    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.  4    are the same as those in  FIGS.  1 ,  2 , and  3   . Accordingly, the  FIG.  4    process can be compared with that of the  FIGS.  1 ,  2 , and  3    processes to illustrate the advantages of the present invention. 
     The feed stream  31  enters compressor  16  driven by expansion machine  15  and is boosted to higher pressure. After cooling to 142° F. [61° C.] in cooler  20 , stream  31   b  at 1071 psia [7,384 kPa(a)] is cooled in heat exchanger  10  by heat exchange with cool residue gas (stream  39   a ) at −134° F. [−92° C.], pumped liquid product (stream  42   a ) at 8° F. [−13° C.], demethanizer reboiler liquids (stream  41 ) at −14° F. [−25° C.], and propane refrigerant. The cooled stream  31   c  enters separator  11  at −47° F. [−44° C.] and 1045 psia [7,205 kPa(a)] where the vapor (stream  32 ) is separated from the condensed liquid (stream  35 ). 
     The vapor (stream  32 ) from separator  11  enters 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 (approximately 141 psia [972 kPa(a)]) of fractionation tower  19 , with the work expansion cooling the expanded stream  32   a  to a temperature of approximately −158° F. [−106° C.]. The partially condensed expanded stream  32   a  is thereafter supplied as feed to fractionation tower  19  at a mid-column feed point below the rectifying section and above the demethanizing section in fractionation tower  19 . 
     The separator liquid (stream  35 ) is cooled to −127° F. [−88° C.] in heat exchanger  13 , and cooled liquid stream  35   a  is then divided into two streams, stream  36  and stream  37 . Stream  37  is expanded to slightly above the operating pressure of fractionation tower  19  by expansion valve  18 , cooling stream  37   a  to −142° F. [−97° C.] before it is heated as it supplies the cooling in heat exchanger  13 . The warmed stream  37   b  at −52° F. [−47° C.] is then supplied to fractionation tower  19  at a lower mid-column feed point. 
     The remaining portion of cooled liquid stream  35   a , stream  36 , passes through heat exchanger  12  in heat exchange relation with cold distillation stream  39  at −162° F. [−108° C.] where it is further cooled to −158° F. [−106° C.]. The further cooled stream  36   a  then is flash expanded to the operating pressure of demethanizer  19  by expansion valve  17 . A portion of the stream is vaporized, further cooling stream  36   b  to −163° F. [−108° C.] before it is supplied as cold top column feed (reflux) to the separator section in fractionation tower  19 . The liquids separated therein become the top feed to the rectifying section in fractionation tower  19  and the cold liquid reflux absorbs and condenses the C 2  components, C 3  components, and heavier components rising in the rectifying section. 
     Liquid product stream  42  exits the bottom of the tower at 4° F. [−16° C.], based on a typical specification of a methane to ethane ratio of 0.010:1 on a molar basis in the bottom product. It is pumped to a pressure of approximately 455 psia [3,135 kPa(a)] in demethanizer bottoms pump  21 , and the pumped liquid product is then warmed to 83° F. [28° C.] as it provides cooling of stream  31   b  in exchanger  10  before flowing to storage. 
     The residue gas (demethanizer overhead vapor stream  39 ) passes countercurrently to cooled liquid stream  36  in heat exchanger  12  where it is heated to −134° F. [−92° C.] (stream  39   a ), and to the incoming feed gas in heat exchanger  10  where it is heated to 136° F. [58° C.] (stream  39   b ). The residue gas product (stream  39   b ) then flows to the fuel gas system at 116 psia [801 kPa(a)]. 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG.  4    is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE IV 
               
               
                   
               
               
                 (FIG. 4) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
               
                   
               
             
            
               
                   
               
            
           
           
               
               
               
               
               
               
               
            
               
                 Stream 
                 Hydrogen 
                 Methane 
                 C 2  Comp. 
                 C 3  Comp. 
                 C 4 + Comp. 
                 Total 
               
               
                   
               
               
                 31 
                 1,998 
                 4,661 
                 2,505 
                 839 
                 939 
                 10,980 
               
               
                 32 
                 1,822 
                 2,531 
                 377 
                 36 
                 11 
                 4,805 
               
               
                 35 
                 176 
                 2,130 
                 2,128 
                 803 
                 928 
                 6,175 
               
               
                 36 
                 83 
                 1,001 
                 1,000 
                 377 
                 436 
                 2,902 
               
               
                 37 
                 93 
                 1,129 
                 1,128 
                 426 
                 492 
                 3,273 
               
               
                 39 
                 1,998 
                 4,637 
                 119 
                 3 
                 0 
                 6,793 
               
               
                 42 
                 0 
                 24 
                 2,386 
                 836 
                 939 
                 4,187 
               
               
                   
               
            
           
           
               
               
               
            
               
                   
                 Recoveries* 
               
               
                   
                   
               
               
                   
                 C 2  Components 
                 95.26% 
               
               
                   
                 C 3  Components 
                 99.69% 
               
               
                   
                 C 4 + Components 
                 99.98% 
               
               
                   
                   
               
            
           
           
               
               
               
               
            
               
                   
                 Power 
               
               
                   
                   
               
               
                   
                 Refrigerant Compression 
                 4,602 HP 
                 [7,566 kW] 
               
               
                   
                 Total Compression 
                 4,602 HP 
                 [7,566 kW] 
               
               
                   
                   
               
               
                   
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     A comparison of Tables I, II, III, and IV shows that, compared to the prior art processes, the present invention significantly improves C 2  component recovery (from 81.68%, 84.81%, and 93.68%, respectively, to 95.26%) and maintains much the same C 3  component and C 4 + component recoveries. Comparison of Tables I, II, III, and IV further shows that these yields were achieved using slightly less power than the prior art processes. In terms of the recovery efficiency (defined by the quantity of ethane recovered per unit of power), the present invention represents a 17%, 13%, and 2% improvement, respectively, over the prior art of the  FIG.  1   ,  FIG.  2   , and  FIG.  3    processes. 
     The superior C 2  component recovery performance of the present invention compared to that of the prior art processes is most easily understood by examining the feed streams supplied to the rectifying section in demethanizer  19 . As explained in the BACKGROUND OF THE INVENTION, the goal is to provide reflux stream(s) capable of capturing the desired C 2 + components rising from below (most of which originate in the expanded stream supplied to the mid-column feed point of demethanizer  19 , stream  34   a  in  FIG.  1   , stream  32   a  in  FIGS.  2 ,  3 , and  4   ). This usually means reflux stream(s) that are predominantly methane, with low concentrations of C 2  components and heavier components. 
     The  FIG.  1    process supplies two reflux streams to the rectifying section in demethanizer  19 , stream  40   h  at the top and stream  33   b  at the mid-point. However, the flow rate of stream  33   b  is very low and it contains a significant concentration of components more volatile than methane, adding to the rectifying load for reflux stream  40   h . This diminishes the effectiveness of stream  40   h , resulting in low recovery of the C 2  components. 
     Both the  FIG.  2    process and the  FIG.  3    process supply only a single reflux stream to the rectifying section in demethanizer  19 , stream  40   h  for  FIG.  2    and stream  36   a  for  FIG.  3   . Counterintuitively, reflux stream  36   a  in  FIG.  3    has much higher concentrations of C 2  components and C 3 + components than reflux stream  40   h  in  FIG.  2   , and yet it captures much more of the C 2  components. In the  FIG.  3    case, the heavy components in reflux stream  36   a  help to capture the C 2  components by absorption, with the tradeoff that recovery of the C 3  components and C 4 + components drops somewhat due to equilibrium losses at the top of the column. 
     In the present invention in  FIG.  4   , reflux stream  36   b  is used as the top feed to the rectifying section similar to reflux stream  36   a  in the prior art  FIG.  3    process. Comparison of Tables III and IV shows that the composition of reflux stream  36   a  in  FIG.  3    is essentially the same as that of reflux stream  36   b  in  FIG.  4   , although the flow rate of the reflux stream in the  FIG.  4    embodiment of the present invention is approximately 7% higher. However, the key difference in the present invention is the much greater degree of subcooling applied to reflux stream  36 / 36   a  upstream of expansion valve  17  in  FIG.  4    (−158° F. [−106° C.]) versus that of reflux stream  36  in  FIG.  3    (−141° F. [−96° C.]). While the temperature resulting after expansion (stream  36   a  in  FIG.  3   , stream  36   b  in  FIG.  4   ) is essentially the same, the amount of flash vapor that results is significantly lower for the present invention in  FIG.  4   . 
     There are two benefits to reducing the amount of flash vapor when the reflux stream is expanded to the operating pressure of demethanizer  19 . First, since the flash vapor leaves demethanizer  19  without being subjected to any rectification, this reduces the amount of C 2 + components lost in the flash gas. Second, the amount of liquid reflux remaining to flow to the rectifying section in demethanizer  19  is greater, allowing for better absorption of the C 2 + components and the correspondingly higher recoveries shown in Table IV versus Table III. 
     Other Embodiments 
     In accordance with this invention, it is generally advantageous to design the absorbing (rectifying) section of the demethanizer to contain multiple theoretical separation stages. However, the benefits of the present invention can be achieved with as few as one theoretical stage. For instance, all or a part of the expanded cooled liquid stream  36   b  from expansion valve  17  and at least a portion of expanded stream  32   a  can be combined (such as in the piping joining the expansion valve and the expansion machine to the demethanizer) and if thoroughly intermingled, the vapors and liquids will mix together and separate in accordance with the relative volatilities of the various components of the total combined streams. Such commingling of the two streams shall be considered for the purposes of this invention as constituting an absorbing section. 
       FIG.  4    depicts the fractionation tower constructed in a single vessel.  FIG.  5    displays another embodiment of the present invention. with the fractionation tower constructed in two vessels, absorber (rectifier) column  27  (a contacting and separating device) and stripper (distillation) column  19 . In such cases, the overhead vapor stream  44  from stripper column  19  flows to the lower section of absorber column  27  to be contacted by expanded cooled liquid stream  36   b . Pump  28  is used to route the liquids (stream  43 ) from the bottom of absorber column  27  to the top of stripper column  19  so that the two towers effectively function as one distillation system. The decision whether to construct the fractionation tower as a single vessel (such as demethanizer  19  in  FIG.  4   ) or multiple vessels will depend on a number of factors such as plant size, the distance to fabrication facilities, etc. 
     The present invention provides improved recovery of C 2  components per amount of utility consumption required to operate the process. An improvement may also be effected in lower utility consumption required for operating the process, which 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.