Abstract:
A process and apparatus for the recovery of heavier hydrocarbons from a liquefied natural gas (LNG) stream is disclosed. The LNG feed stream is heated to vaporize at least part of it, then supplied to a fractionation column at a mid-column feed position. A vapor distillation stream is withdrawn from the fractionation column below the mid-column feed position and directed in heat exchange relation with the LNG feed stream, cooling the vapor distillation stream as it supplies at least part of the heating of the LNG feed stream. The vapor distillation stream is cooled sufficiently to condense at least a part of it, forming a condensed stream. At least a portion of the condensed stream is directed to the fractionation column as its top feed. The quantities and temperatures of the feeds to the column are effective to maintain the column overhead temperature at a temperature whereby the major portion of the desired components is recovered in the bottom liquid product from the column.

Description:
[0001]    The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application No. 60/938,489 which was filed on May 17, 2007. 
     
    
     BACKGROUND OF THE INVENTION 
       [0002]    This invention relates to a process for the separation of ethane and heavier hydrocarbons or propane and heavier hydrocarbons from liquefied natural gas, hereinafter referred to as LNG, to provide a volatile methane-rich gas stream and a less volatile natural gas liquids (NGL) or liquefied petroleum gas (LPG) stream. 
         [0003]    As an alternative to transportation in pipelines, natural gas at remote locations is sometimes liquefied and transported in special LNG tankers to appropriate LNG receiving and storage terminals. The LNG can then be re-vaporized and used as a gaseous fuel in the same fashion as natural gas. Although LNG usually has a major proportion of methane, i.e., methane comprises at least 50 mole percent of the LNG, it also contains relatively lesser amounts of heavier hydrocarbons such as ethane, propane, butanes, and the like, as well as nitrogen. It is often necessary to separate some or all of the heavier hydrocarbons from the methane in the LNG so that the gaseous fuel resulting from vaporizing the LNG conforms to pipeline specifications for heating value. In addition, it is often also desirable to separate the heavier hydrocarbons from the methane and ethane because these hydrocarbons have a higher value as liquid products (for use as petrochemical feedstocks, as an example) than their value as fuel. 
         [0004]    Although there are many processes which may be used to separate ethane and/or propane and heavier hydrocarbons from LNG, these processes often must compromise between high recovery, low utility costs, and process simplicity (and hence low capital investment). U.S. Pat. Nos. 2,952,984; 3,837,172; 5,114,451; and 7,155,931 describe relevant LNG processes capable of ethane or propane recovery while producing the lean LNG as a vapor stream that is thereafter compressed to delivery pressure to enter a gas distribution network. However, lower utility costs may be possible if the lean LNG is instead produced as a liquid stream that can be pumped (rather than compressed) to the delivery pressure of the gas distribution network, with the lean LNG subsequently vaporized using a low level source of external heat or other means. U.S. Pat. Nos. 7,069,743 and 7,216,507 and co-pending application Ser. No. 11/749,268 describe such processes. 
         [0005]    The present invention is generally concerned with the recovery of propylene, propane, and heavier hydrocarbons from such LNG streams. It uses a novel process arrangement to allow high propane recovery while keeping the processing equipment simple and the capital investment low. Further, the present invention offers a reduction in the utilities (power and heat) required to process the LNG to give lower operating cost than the prior art processes, and also offers significant reduction in capital investment. A typical analysis of an LNG stream to be processed in accordance with this invention would be, in approximate mole percent, 86.7% methane, 8.9% ethane and other C 2  components, 2.9% propane and other C 3  components, and 1.0% butanes plus, with the balance made up of nitrogen. 
     
    
     
         [0006]    For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings: 
           [0007]      FIG. 1  is a flow diagram of an LNG processing plant in accordance with the present invention where the vaporized LNG product is to be delivered at a relatively low pressure; and 
           [0008]      FIG. 2  is a flow diagram illustrating an alternative means of application of the present invention to an LNG processing plant where the vaporized LNG product must be delivered at relatively higher pressure. 
       
    
    
       [0009]    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. 
         [0010]    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 INVENTION 
     Example 1 
       [0011]      FIG. 1  illustrates a flow diagram of a process in accordance with the present invention adapted to produce an LPG product containing the majority of the C 3  components and heavier hydrocarbon components present in the feed stream. 
         [0012]    In the simulation of the  FIG. 1  process, the LNG to be processed (stream  41 ) from LNG tank  10  enters pump  11  at −255° F. [−159° C.], which elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers  13  and  14  and thence to fractionation column  21 . Stream  41   a  exiting the pump at −253° F. [−158° C.] and 440 psia [3,032 kPa(a)] is heated to −196° F. [−127° C.] (stream  41   b ) in heat exchanger  13  by cooling and partially condensing distillation vapor stream  50  which has been withdrawn from a mid-column region of fractionation tower  21 . The heated stream  41   b  is then further heated to −87° F. [−66° C.] in heat exchanger  14  using low level utility heat. (High level utility heat, such as the heating medium used in tower reboiler  25 , is normally more expensive than low level utility heat, so lower operating cost is usually achieved when use of low level heat, such as sea water, is maximized and the use of high level utility heat is minimized.) The further heated stream  41   c , now partially vaporized, is then supplied to fractionation column  21  at an upper mid-column feed point. Under some circumstances, it may be desirable to separate stream  41   c  into vapor stream  42  and liquid stream  43  via separator  15  and route each stream separately to fractionation column  21  as indicated by the dashed lines in  FIG. 1 . 
         [0013]    The deethanizer in tower  21  is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The deethanizer tower consists of two sections: an upper absorbing (rectification) section  21   a  that contains the necessary trays or packing to provide the necessary contact between the vapor portion of stream  41   c  rising upward and cold liquid falling downward to condense and absorb propane and heavier components from the vapor portion; and a lower, stripping section  21   b  that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The deethanizer stripping section  21   b  also includes one or more reboilers (such as reboiler  25 ) which heat and vaporize a portion of the liquid at the bottom of the column to provide the stripping vapors which flow up the column. These vapors strip the methane and C 2  components from the liquids, so that the bottom liquid product (stream  51 ) is substantially devoid of methane and C 2  components and is comprised of the majority of the C 3  components and heavier hydrocarbons contained in the LNG feed stream. 
         [0014]    Stream  41   c  enters fractionation column  21  at an upper mid-column feed position located in the lower region of absorbing section  21   a  of fractionation column  21 . The liquid portion of stream  41   c  comingles with the liquids falling downward from the absorbing section and the combined liquid proceeds downward into stripping section  21   b  of deethanizer  21 . The vapor portion of stream  41   c  rises upward through absorbing section  21   a  and is contacted with cold liquid falling downward to condense and absorb the C 3  components and heavier components. 
         [0015]    A liquid stream  49  from deethanizer  21  is withdrawn from the lower region of absorbing section  21   a  and is routed to heat exchanger  13  where it is heated as it provides cooling of distillation vapor stream  50  as described earlier. Typically, the flow of this liquid from the deethanizer is via a thermosiphon circulation, but a pump could be used. The liquid stream is heated from −86° F. [−65° C.] to −65° F. [−54° C.], partially vaporizing stream  49   c  before it is returned as a mid-column feed to deethanizer  21 , typically in the middle region of stripping section  21   b . Alternatively, the liquid stream  49  may be routed directly without heating to the lower mid-column feed point in the stripping section  21   b  of deethanizer  21  as shown by dashed line  49   a.    
         [0016]    A portion of the distillation vapor (stream  50 ) is withdrawn from the upper region of stripping section  21   b  at −10° F. [−23° C.]. This stream is then cooled and partially condensed (stream  50   a ) in exchanger  13  by heat exchange with LNG stream  41   a  and liquid stream  49  (if applicable) as described previously. The partially condensed stream  50   a  then flows to reflux separator  19  at −85° F. [−65° C.]. 
         [0017]    The operating pressure in reflux separator  19  (406 psia [2,797 kPa(a)]) is maintained slightly below the operating pressure of deethanizer  21  (415 psia [2,859 kPa(a)]). This provides the driving force which causes distillation vapor stream  50  to flow through heat exchanger  13  and thence into reflux separator  19  wherein the condensed liquid (stream  53 ) is separated from any uncondensed vapor (stream  52 ). Stream  52  then combines with the deethanizer overhead stream  48  to form cold residue gas stream  56  at −95° F. [−71° C.], which is then heated to 40° F. [4° C.] using low level utility heat in heat exchanger  27  before flowing to the sales gas pipeline at 381 psia [2,625 kPa(a)]. 
         [0018]    The liquid stream  53  from reflux separator  19  is pumped by pump  20  to a pressure slightly above the operating pressure of deethanizer  21 , and the pumped stream  53   a  is then divided into at least two portions. One portion, stream  54 , is supplied as top column feed (reflux) to deethanizer  21 . This cold liquid reflux absorbs and condenses the C 3  components and heavier components rising in the upper rectification region of absorbing section  21   a  of deethanizer  21 . The other portion, stream  55 , is supplied to deethanizer  21  at a mid-column feed position located in the upper region of stripping section  21   b , in substantially the same region where distillation vapor stream  50  is withdrawn, to provide partial rectification of stream  50 . 
         [0019]    The deethanizer overhead vapor (stream  48 ) exits the top of deethanizer  21  at −94° F. [−70° C.] and is combined with vapor stream  52  as described previously. The liquid product stream  51  exits the bottom of the tower at 185° F. [85° C.] based on an ethane:propane ratio of 0.02:1 on a molar basis in the bottom product, and flows to storage or further processing. 
         [0020]    A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 1  is set forth in the following table: 
         [0000]    
       
         
               
             
               
               
               
               
               
               
             
               
             
               
               
               
             
               
             
               
               
               
               
               
             
               
             
               
               
               
               
               
             
               
             
               
               
               
               
               
             
           
               
                 TABLE I 
               
               
                   
               
             
             
               
                 (FIG. 1) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
             
          
           
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 41 
                 17,281 
                 1,773 
                 584 
                 197 
                 19,923 
               
               
                 49 
                 1,468 
                 1,154 
                 583 
                 197 
                 3,403 
               
               
                 50 
                 2,409 
                 2,456 
                 4 
                 0 
                 4,871 
               
               
                 53 
                 1,790 
                 2,371 
                 4 
                 0 
                 4,165 
               
               
                 54 
                 626 
                 830 
                 1 
                 0 
                 1,457 
               
               
                 55 
                 1,164 
                 1,541 
                 3 
                 0 
                 2,708 
               
               
                 52 
                 619 
                 85 
                 0 
                 0 
                 706 
               
               
                 48 
                 16,662 
                 1,677 
                 2 
                 0 
                 18,426 
               
               
                 56 
                 17,281 
                 1,762 
                 2 
                 0 
                 19,132 
               
               
                 51 
                 0 
                 11 
                 582 
                 197 
                 791 
               
               
                   
               
             
          
           
               
                 Recoveries* 
               
               
                   
               
             
          
           
               
                 Propane 
                 99.67% 
                   
               
               
                 Butanes+ 
                 100.00% 
               
             
          
           
               
                 Power 
               
             
          
           
               
                 Liquid Feed Pump 
                 459 
                 HP 
                 [755 
                 kW] 
               
               
                 Reflux Pump 
                  21 
                 HP 
                  [35 
                 kW] 
               
               
                 Totals 
                 480 
                 HP 
                 [790 
                 kW] 
               
             
          
           
               
                 Low Level Utility Heat 
               
             
          
           
               
                 Liquid Feed Heater 
                 71,532 
                 MBTU/Hr 
                 [46,206 
                 kW] 
               
               
                 Residue Gas Heater 
                 27,084 
                 MBTU/Hr 
                 [17,495 
                 kW] 
               
               
                 Totals 
                 98,616 
                 MBTU/Hr 
                 [63,701 
                 kW] 
               
             
          
           
               
                 High Level Utility Heat 
               
             
          
           
               
                 Deethanizer Reboiler 
                 26,816 
                 MBTU/Hr 
                 [17,322 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
             
          
         
       
     
         [0021]    There are three primary factors that account for the improved efficiency of the present invention. First, compared to many prior art processes, the present invention does not depend on the LNG feed itself to directly serve as the reflux for fractionation column  21 . Rather, the refrigeration inherent in the cold LNG is used in heat exchanger  13  to generate a liquid reflux stream (stream  54 ) that contains very little of the C 3  components and heavier hydrocarbon components that are to be recovered, resulting in efficient rectification in absorbing section  21   a  of fractionation tower  21  and avoiding the equilibrium limitations of such prior art processes. Second, the partial rectification of distillation vapor stream  50  by reflux stream  55  results in a top reflux stream  54  that is predominantly liquid methane and C 2  components and contains very little C 3  components and heavier hydrocarbon components. As a result, nearly 100% of the C 3  components and substantially all of the heavier hydrocarbon components are recovered in liquid product  51  leaving the bottom of deethanizer  21 . Third, the rectification of the column vapors provided by absorbing section  21   a  allows the majority of the LNG feed to be vaporized before entering deethanizer  21  as stream  41   c  (with much of the vaporization duty provided by low level utility heat in heat exchanger  14 ). With less total liquid feeding fractionation column  21 , the high level utility heat consumed by reboiler  25  to meet the specification for the bottom liquid product from the deethanizer is minimized. 
       Example 2 
       [0022]      FIG. 1  represents the preferred embodiment of the present invention when the required delivery pressure of the vaporized LNG residue gas is relatively low. An alternative method of processing the LNG stream to deliver the residue gas at relatively high pressure is shown in another embodiment of the present invention as illustrated in  FIG. 2 . The LNG feed composition and conditions considered in the process presented in  FIG. 2  are the same as those for  FIG. 1 . Accordingly, the  FIG. 2  process of the present invention can be compared to the embodiment of  FIG. 1 . 
         [0023]    In the simulation of the  FIG. 2  process, the LNG to be processed (stream  41 ) from LNG tank  10  enters pump  11  at −255° F. [−159° C.] to elevate the pressure of the LNG to 1215 psia [8,377 kPa(a)]. The high pressure LNG (stream  41   a ) then flows through heat exchanger  12  where it is heated from −249° F. [−156° C.] to −90° F. [−68° C.] (stream  41   b ) by heat exchange with vapor stream  56   a  from booster compressor  17 . Heated stream  41   b  then flows through heat exchanger  13  where it is heated to −63° F. [−53° C.] (stream  41   c ) by cooling and partially condensing distillation vapor stream  50  which has been withdrawn from a mid-column region of fractionation tower  21 . Stream  41   c  is then further heated to −16° F. [−27° C.] in heat exchanger  14  using low level utility heat. 
         [0024]    The further heated stream  41   d  is then supplied to expansion machine  16  in which mechanical energy is extracted from the high pressure feed. The machine  16  expands the vapor substantially isentropically from a pressure of about 1190 psia [8,205 kPa(a)] to a pressure of about 415 psia [2,859 kPa(a)] (the operating pressure of fractionation column  21 ). The work expansion cools the expanded stream  42   a  to a temperature of approximately −94° F. [−70° 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  17 ) that can be used to re-compress the cold vapor stream (stream  56 ), for example. The expanded and partially condensed stream  42   a  is thereafter supplied to fractionation column  21  at an upper mid-column feed point. 
         [0025]    For the composition and conditions illustrated in  FIG. 2 , stream  41   d  is heated sufficiently to be in a completely vapor state. Under some circumstances, it may be desirable to partially vaporize stream  41   d  and then separate it into vapor stream  42  and liquid stream  43  via separator  15  as indicated by the dashed lines in  FIG. 2 . In such an instance, vapor stream  42  would enter expansion machine  16 , while liquid stream  43  would enter expansion valve  18  and the expanded liquid stream  43   a  would be supplied to fractionation column  21  at a lower mid-column feed point. 
         [0026]    Expanded stream  42   a  enters fractionation column  21  at an upper mid-column feed position located in the lower region of the absorbing section of fractionation column  21 . The liquid portion of stream  42   a  comingles with the liquids falling downward from the absorbing section and the combined liquid proceeds downward into the stripping section of deethanizer  21 . The vapor portion of expanded stream  42   a  rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C 3  components and heavier components. 
         [0027]    A liquid stream  49  from deethanizer  21  is withdrawn from the lower region of the absorbing section and is routed to heat exchanger  13  where it is heated as it provides cooling of distillation vapor stream  50  as described earlier. The liquid stream is heated from −90° F. [−68° C.] to −61° F. [−52° C.], partially vaporizing stream  49   c  before it is returned as a mid-column feed to deethanizer  21 , typically in the middle region of the stripping section. Alternatively, the liquid stream  49  may be routed directly without heating to the lower mid-column feed point in the stripping section of deethanizer  21  as shown by dashed line  49   a.    
         [0028]    A portion of the distillation vapor (stream  50 ) is withdrawn from the upper region of the stripping section at −15° F. [−26° C.]. This stream is then cooled and partially condensed (stream  50   a ) in exchanger  13  by heat exchange with LNG stream  41   b  and liquid stream  49  (if applicable). The partially condensed stream  50   a  at −85° F. [−65° C.] then combines with overhead vapor stream  48  from deethanizer  21  and the combined stream  57  flows to reflux separator  19  at −95° F. [−71° C.]. (It should be noted that the combining of streams  50   a  and  48  can occur in the piping upstream of reflux separator  19  as shown in  FIG. 2 , or alternatively, streams  50   a  and  48  can flow individually to reflux separator  19  with the commingling of the streams occurring therein. 
         [0029]    The operating pressure of reflux separator  19  (406 psia [2,797 kPa(a)]) is maintained slightly below the operating pressure of deethanizer  21 . This provides the driving force which causes distillation vapor stream  50  to flow through heat exchanger  13 , combine with column overhead vapor stream  48  if appropriate, and thence flow into reflux separator  19  wherein the condensed liquid (stream  53 ) is separated from any uncondensed vapor (stream  56 ). 
         [0030]    The liquid stream  53  from reflux separator  19  is pumped by pump  20  to a pressure slightly above the operating pressure of deethanizer  21 , and the pumped stream  53   a  is then divided into at least two portions. One portion, stream  54 , is supplied as top column feed (reflux) to deethanizer  21 . This cold liquid reflux absorbs and condenses the C 3  components and heavier components rising in the upper rectification region of the absorbing section of deethanizer  21 . The other portion, stream  55 , is supplied to deethanizer  21  at a mid-column feed position located in the upper region of the stripping section in substantially the same region where distillation vapor stream  50  is withdrawn, to provide partial rectification of stream  50 . The deethanizer overhead vapor (stream  48 ) exits the top of deethanizer  21  at −98° F. [−72° C.] and is combined with partially condensed stream  50   a  as described previously. The liquid product stream  51  exits the bottom of the tower at 185° F. [85° C.] and flows to storage or further processing. 
         [0031]    The cold vapor stream  56  from separator  19  flows to compressor  17  driven by expansion machine  16  to increase the pressure of stream  56   a  sufficiently so that it can be totally condensed in heat exchanger  12 . Stream  56   a  exits the compressor at −24° F. [−31° C.] and 718 psia [4,953 kPa(a)] and is cooled to −109° F. [−79° C.] (stream  56   b ) by heat exchange with the high pressure LNG feed stream  41   a  as discussed previously. Condensed stream  56   b  is pumped by pump  26  to a pressure slightly above the sales gas delivery pressure. Pumped stream  56   c  is then heated from −95° F. [−70° C.] to 40° F. [4° C.] in heat exchanger  27  before flowing to the sales gas pipeline at 1215 psia [8,377 kPa(a)] as residue gas stream  56   d.    
         [0032]    A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 2  is set forth in the following table: 
         [0000]    
       
         
               
             
               
               
               
               
               
               
             
               
             
               
               
               
             
               
             
               
               
               
               
               
             
               
             
               
               
               
               
               
             
               
             
               
               
               
               
               
             
           
               
                 TABLE II 
               
               
                   
               
             
             
               
                 (FIG. 2) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
             
          
           
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 41 
                 17,281 
                 1,773 
                 584 
                 197 
                 19,923 
               
               
                 49 
                 1,800 
                 1,386 
                 584 
                 197 
                 3,969 
               
               
                 50 
                 2,585 
                 2,278 
                 5 
                 0 
                 4,871 
               
               
                 53 
                 1,927 
                 2,027 
                 6 
                 0 
                 3,962 
               
               
                 54 
                 674 
                 709 
                 2 
                 0 
                 1,387 
               
               
                 55 
                 1,253 
                 1,318 
                 4 
                 0 
                 2,575 
               
               
                 48 
                 16,623 
                 1,510 
                 2 
                 0 
                 18,222 
               
               
                 56 
                 17,281 
                 1,761 
                 1 
                 0 
                 19,131 
               
               
                 51 
                 0 
                 12 
                 583 
                 197 
                 792 
               
               
                   
               
             
          
           
               
                 Recoveries* 
               
               
                   
               
             
          
           
               
                 Propane 
                 99.84% 
                   
               
               
                 Butanes+ 
                 100.00% 
               
             
          
           
               
                 Power 
               
             
          
           
               
                 Liquid Feed Pump 
                 1,409 
                 HP 
                 [2,316 
                 kW] 
               
               
                 Reflux Pump 
                   20 
                 HP 
                   [33 
                 kW] 
               
               
                 LNG Product Pump 
                 1,024 
                 HP 
                 [1,684 
                 kW] 
               
               
                 Totals 
                 2,453 
                 HP 
                 [4,033 
                 kW] 
               
             
          
           
               
                 Low Level Utility Heat 
               
             
          
           
               
                 Liquid Feed Heater 
                 27,261 
                 MBTU/Hr 
                 [17,609 
                 kW] 
               
               
                 Residue Gas Heater 
                 54,840 
                 MBTU/Hr 
                 [35,424 
                 kW] 
               
               
                 Totals 
                 82,101 
                 MBTU/Hr 
                 [53,033 
                 kW] 
               
             
          
           
               
                 High Level Utility Heat 
               
             
          
           
               
                 Demethanizer Reboiler 
                 26,808 
                 MBTU/Hr 
                 [17,316 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
             
          
         
       
     
         [0033]    A comparison of Tables I and II shows that both the  FIG. 1  and  FIG. 2  embodiments achieve comparable recovery of C 3  and heavier components. Although the  FIG. 2  embodiment requires considerably more pumping power than the  FIG. 1  embodiment, this is a result of the much higher sales gas delivery pressure for the process conditions shown in  FIG. 2 . Nonetheless, the power required for the  FIG. 2  embodiment of the present invention is less than that of prior art processes operating under the same conditions. 
       OTHER EMBODIMENTS 
       [0034]    In accordance with this invention, it is generally advantageous to design the absorbing (rectification) section of the deethanizer to contain multiple theoretical separation stages. However, the benefits of the present invention can be achieved with as few as one theoretical stage, and it is believed that even the equivalent of a fractional theoretical stage may allow achieving these benefits. For instance, all or a part of the condensed liquid (stream  53 ) leaving reflux separator  19  and all or a part of stream  42   a  can be combined (such as in the piping to the deethanizer) 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. 
         [0035]    As described earlier, the distillation vapor stream  50  is partially condensed and the resulting condensate used to absorb valuable C 3  components and heavier components from the vapors in stream  42   a . However, the present invention is not limited to this embodiment. It may be advantageous, for instance, to treat only a portion of these vapors in this manner, or to use only a portion of the condensate as an absorbent, in cases where other design considerations indicate portions of the vapors or the condensate should bypass the absorbing section of the deethanizer. LNG conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine  16  in  FIG. 2 , or replacement with an alternate expansion device (such as an expansion valve), is feasible, or that total (rather than partial) condensation of distillation vapor stream  50  in heat exchanger  13  is possible or is preferred. 
         [0036]    In the practice of the present invention, there will necessarily be a slight pressure difference between deethanizer  21  and reflux separator  19  which must be taken into account. If the distillation vapor stream  50  passes through heat exchanger  13  and into reflux separator  19  without any boost in pressure, reflux separator  19  shall necessarily assume an operating pressure slightly below the operating pressure of deethanizer  21 . In this case, the liquid stream withdrawn from reflux separator  19  can be pumped to its feed position(s) on deethanizer  21 . An alternative is to provide a booster blower for distillation vapor stream  50  to raise the operating pressure in heat exchanger  13  and reflux separator  19  sufficiently so that the liquid stream  53  can be supplied to deethanizer  21  without pumping. 
         [0037]    Some circumstances may favor pumping the LNG stream to a higher pressure than that shown in  FIG. 1  even when the delivery pressure of the residue gas is low. In such instances, an expansion device such as expansion valve  28  or an expansion engine may be used to reduce the pressure of stream  41   c  to that of fractionation column  21 . If separator  15  is used, then an expansion device such as expansion valve  18  would also be required to reduce the pressure of separator liquid stream  43  to that of column  21 . If an expansion engine is used in lieu of expansion valve  28  and/or  18 , the work expansion could be used to drive a generator, which could in turn be used to reduce the amount of external pumping power required by the process. Similarly, the expansion engine  16  in  FIG. 2  could also be used to drive a generator, in which case compressor  17  could be driven by an electric motor. 
         [0038]    In some circumstance it may be desirable to bypass some or all of liquid stream  49  around heat exchanger  13 . If a partial bypass is desirable, the bypass stream  49   a  would then be mixed with the outlet stream  49   b  from exchanger  13  and the combined stream  49   c  returned to the stripping section of fractionation column  21 . The use and distribution of the liquid stream  49  for process heat exchange, the particular arrangement of heat exchangers for LNG stream heating and distillation vapor stream cooling, and the choice of process streams for specific heat exchange services must be evaluated for each particular application. 
         [0039]    It will also be recognized that the relative amount of feed found in each branch of the condensed liquid contained in stream  53   a  that is split between the two column feeds in  FIGS. 1 and 2  will depend on several factors, including LNG pressure, LNG stream composition, and the desired recovery levels. The optimum split cannot generally be predicted without evaluating the particular circumstances for a specific application of the present invention. It may be desirable in some cases to route all the reflux stream  53   a  to the top of the absorbing section in deethanizer  21  with no flow in dashed line  55  in  FIGS. 1 and 2 . In such cases, the quantity of liquid stream  49  withdrawn from fractionation column  21  could be reduced or eliminated. 
         [0040]    The mid-column feed positions depicted in  FIGS. 1 and 2  are the preferred feed locations for the process operating conditions described. However, the relative locations of the mid-column feeds may vary depending on the LNG composition or other factors such as desired recovery levels, etc. Moreover, two or more of the feed streams, or portions thereof, may be combined depending on the relative temperatures and quantities of individual streams, and the combined stream then fed to a mid-column feed position.  FIGS. 1 and 2  are the preferred embodiments for the compositions and pressure conditions shown. Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where appropriate. For example, conditions may warrant work expansion of the liquid stream (stream  43 ). 
         [0041]    In  FIGS. 1 and 2 , multiple heat exchanger services have been shown combined in a common heat exchanger  13 . It may be desirable in some instances to use individual heat exchangers for each service. In some cases, circumstances may favor splitting a heat exchange service into multiple exchangers. (The decision as to whether to combine heat exchange services or to use more than one heat exchanger for the indicated service will depend on a number of factors including, but not limited to, LNG flow rate, heat exchanger size, stream temperatures, etc.) Alternatively, heat exchanger  13  could be replaced by other heating means, such as a heater using sea water, a heater using a utility stream rather than a process stream (like stream  50  used in  FIGS. 1 and 2 ), an indirect fired heater, or a heater using a heat transfer fluid warmed by ambient air, as warranted by the particular circumstances. 
         [0042]    The present invention provides improved recovery of C 3  components per amount of utility consumption required to operate the process. It also provides for reduced capital expenditure in that all fractionation can be done in a single column. An improvement in utility consumption required for operating the deethanizer process may appear in the form of reduced power requirements for compression or re-compression, reduced power requirements for pumping, reduced energy requirements for tower reboilers, or a combination thereof. Alternatively, if desired, increased C 3  component recovery can be obtained for a fixed utility consumption. 
         [0043]    In the examples given for the  FIG. 1  and  FIG. 2  embodiments, recovery of C 3  components and heavier hydrocarbon components is illustrated. However, it is believed that the embodiments may also be advantageous when recovery of C 2  components and heavier hydrocarbon components is desired. 
         [0044]    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.