Patent Publication Number: US-7216507-B2

Title: Liquefied natural gas processing

Description:
BACKGROUND OF THE INVENTION 
     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 lean LNG stream and a less volatile natural gas liquids (NGL) or liquefied petroleum gas (LPG) stream. The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application Nos. 60/584,668 which was filed on Jul. 1, 2004, 60/646,903 which was filed on Jan. 24, 2005, Ser. No. 60/669,642 which was filed on Apr. 8, 2005, and Ser. No. 60/671,930 which was filed on Apr. 15, 2005. 
     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 because these hydrocarbons have a higher value as liquid products (for use as petrochemical feedstocks, as an example) than their value as fuel. 
     Although there are many processes which may be used to separate ethane 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; and 5,114,451 and co-pending application Ser. No. 10/675,785 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. Patent Application Publication No. US 2003/0158458 A1 describes such a process. 
     The present invention is generally concerned with the recovery of ethylene, ethane, propylene, propane, and heavier hydrocarbons from such LNG streams. It uses a novel process arrangement to allow high ethane or 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. 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. 
    
    
     
       For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings: 
         FIG. 1  is a flow diagrams of a prior art LNG processing plant; 
         FIG. 2  is a flow diagram of a prior art LNG processing plant in accordance with U.S. Patent Application Publication No. US 2003/0158458 A1; 
         FIG. 3  is a flow diagram of an LNG processing plant in accordance with the present invention; and 
         FIGS. 4 through 13  are flow diagrams illustrating alternative means of application of the present invention to an LNG processing plant. 
     
    
    
     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 
     Referring now to  FIG. 1 , for comparison purposes we begin with an example of a prior art LNG processing plant adapted to produce an NGL product containing the majority of the C 2  components and heavier hydrocarbon components present in the feed stream. The LNG to be processed (stream  41 ) from LNG tank  10  enters pump  11  at −255° F. [−159° C.]. Pump  11  elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to separator  15 . Stream  41   a  exiting the pump is heated in heat exchangers  12  and  13  by heat exchange with gas stream  52  at −120° F. [−84° C.] and demethanizer bottom liquid product (stream  51 ) at 80° F. [27° C.]. 
     The heated stream  41   c  enters separator  15  at −163° F. [−108° C.] and 230 psia [1,586 kPa(a)] where the vapor (stream  46 ) is separated from the remaining liquid (stream  47 ). Stream  47  is pumped by pump  28  to higher pressure, then expanded to the operating pressure (approximately 430 psia [2,965 kPa(a)]) of fractionation tower  21  by control valve  20  and supplied to the tower as the top column feed (stream  47   b ). 
     Fractionation column or tower  21 , commonly referred to as a demethanizer, 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 trays and/or packing provide the necessary contact between the liquids falling downward in the column and the vapors rising upward. The column also includes one or more reboilers (such as reboiler  25 ) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column. These vapors strip the methane from the liquids, so that the bottom liquid product (stream  51 ) is substantially devoid of methane and comprised of the majority of the C 2  components and heavier hydrocarbons contained in the LNG feed stream. (Because of the temperature level required in the column reboiler, a high level source of utility heat is typically required to provide the heat input to the reboiler, such as the heating medium used in this example.) The liquid product stream  51  exits the bottom of the tower at 80° F. [27° C.], based on a typical specification of a methane fraction of 0.005 on a volume basis in the bottom product. After cooling to 43° F. [6° C.] in heat exchanger  13  as described previously, the liquid product (stream  51   a ) flows to storage or further processing. 
     Vapor stream  46  from separator  15  enters compressor  27  (driven by an external power source) and is compressed to higher pressure. The resulting stream  46   a  is combined with the demethanizer overhead vapor, stream  48 , leaving demethanizer  21  at −130° F. [−90° C.] to produce a methane-rich residue gas (stream  52 ) at −120° F. [−84° C.], which is thereafter cooled to −143° F. [−97° C.] in heat exchanger  12  as described previously to totally condense the stream. Pump  32  then pumps the condensed liquid (stream  52   a ) to 1365 psia [9,411 kPa(a)] (stream  52   b ) for subsequent vaporization and/or transportation. 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 1  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE I 
               
               
                   
               
             
            
               
                 (FIG. 1) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 41 
                 9,524 
                 977 
                 322 
                 109 
                 10,979 
               
               
                 46 
                 3,253 
                 20 
                 1 
                 0 
                 3,309 
               
               
                 47 
                 6,271 
                 957 
                 321 
                 109 
                 7,670 
               
               
                 48 
                 6,260 
                 78 
                 5 
                 0 
                 6,355 
               
               
                 52 
                 9,513 
                 98 
                 6 
                 0 
                 9,664 
               
               
                 51 
                 11 
                 879 
                 316 
                 109 
                 1,315 
               
            
           
           
               
            
               
                 Recoveries* 
               
            
           
           
               
               
               
            
               
                   
                 Ethane 
                 90.00% 
               
               
                   
                 Propane 
                 98.33% 
               
               
                   
                 Butanes+ 
                 99.62% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Feed Pump 
                 123 
                 HP 
                 [202 
                 kW] 
               
               
                 Demethanizer Feed Pump 
                 132 
                 HP 
                 [217 
                 kW] 
               
               
                 LNG Product Pump 
                 773 
                 HP 
                 [1,271 
                 kW] 
               
               
                 Vapor Compressor 
                 527 
                 HP 
                 [867 
                 kW] 
               
               
                 Totals 
                 1,555 
                 HP 
                 [2,557 
                 kW] 
               
            
           
           
               
            
               
                 High Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 Demethanizer Reboiler 
                 23,271 
                 MBTU/Hr 
                 [15,032 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
       FIG. 2  shows an alternative prior art process in accordance with U.S. Patent Application Publication No. US 2003/0158458 A1 that can achieve somewhat higher recovery levels with lower utility consumption than the prior art process used in  FIG. 1 . The process of  FIG. 2 , adapted here to produce an NGL product containing the majority of the C 2  components and heavier hydrocarbon components present in the feed stream, has been applied to the same LNG composition and conditions as described previously for  FIG. 1 . 
     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.]. Pump  11  elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to fractionation tower  21 . Stream  41   a  exiting the pump is heated in heat exchangers  12  and  13  by heat exchange with column overhead vapor stream  48  at −130° F. [−90° C.], compressed vapor stream  52   a  at −122° F. [−86° C.], and demethanizer bottom liquid product (stream  51 ) at 85° F. [29° C.]. The partially heated stream  41   c  is then further heated to −120° F. [−84° C.] (stream  41   d ) in heat exchanger  14  using low level utility heat. (High level utility heat is normally more expensive than low level utility heat, so lower operating cost is usually achieved when the use of low level heat, such as the sea water used in this example, is maximized and the use of high level heat is minimized.) After expansion to the operating pressure (approximately 450 psia [3,103 kPa(a)]) of fractionation tower  21  by control valve  20 , stream  41   e  flows to a mid-column feed point at −123° F. [−86° C.]. 
     The demethanizer 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. As is often the case in natural gas processing plants, the fractionation tower may consist of two sections. The upper absorbing (rectification) section  21   a  contains the trays and/or packing to provide the necessary contact between the vapors rising upward and cold liquid falling downward to condense and absorb the ethane and heavier components; the lower stripping (demethanizing) section  21   b  contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section also includes one or more reboilers (such as reboiler  25 ) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column. These vapors strip the methane from the liquids, so that the bottom liquid product (stream  51 ) is substantially devoid of methane and comprised of the majority of the C 2  components and heavier hydrocarbons contained in the LNG feed stream. 
     Overhead stream  48  leaves the upper section of fractionation tower  21  at −130° F. [−90° C.] and flows to heat exchanger  12  where it is cooled to −135° F. [−93° C.] and partially condensed by heat exchange with the cold LNG (stream  41   a ) as described previously. The partially condensed stream  48   a  enters reflux separator  26  wherein the condensed liquid (stream  53 ) is separated from the uncondensed vapor (stream  52 ). The liquid stream  53  from reflux separator  26  is pumped by reflux pump  28  to a pressure slightly above the operating pressure of demethanizer  21  and stream  53   b  is then supplied as cold top column feed (reflux) to demethanizer  21  by control valve  30 . This cold liquid reflux absorbs and condenses the C 2  components and heavier hydrocarbon components from the vapors rising in the upper absorbing (rectification) section  21   a  of demethanizer  21 . 
     The liquid product stream  51  exits the bottom of fractionation tower  21  at 85° F. [29° C.], based on a methane fraction of 0.005 on a volume basis in the bottom product. After cooling to 0° F. [−18° C.] in heat exchanger  13  as described previously, the liquid product (stream  51   a ) flows to storage or further processing. The methane-rich residue gas (stream  52 ) leaving reflux separator  26  is compressed to 493 psia [3,400 kPa(a)] (stream  52   a ) by compressor  27  (driven by an external power source), so that the stream can be totally condensed as it is cooled to −136° F. [−93° C.] in heat exchanger  12  as described previously. Pump  32  then pumps the condensed liquid (stream  52   b ) to 1365 psia [9,411 kPa(a)] (stream  52   c ) for subsequent vaporization and/or transportation. 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 2  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE II 
               
               
                   
               
             
            
               
                 (FIG. 2) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 41 
                 9,524 
                 977 
                 322 
                 109 
                 10,979 
               
               
                 48 
                 10,540 
                 177 
                 0 
                 0 
                 10,766 
               
               
                 53 
                 1,027 
                 79 
                 0 
                 0 
                 1,108 
               
               
                 52 
                 9,513 
                 98 
                 0 
                 0 
                 9,658 
               
               
                 51 
                 11 
                 879 
                 322 
                 109 
                 1,321 
               
            
           
           
               
            
               
                 Recoveries* 
               
            
           
           
               
               
               
            
               
                   
                 Ethane 
                  90.01% 
               
               
                   
                 Propane 
                 100.00% 
               
               
                   
                 Butanes+ 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Feed Pump 
                 298 
                 HP 
                 [490 
                 kW] 
               
               
                 Reflux Pump 
                 5 
                 HP 
                 [8 
                 kW] 
               
               
                 LNG Product Pump 
                 762 
                 HP 
                 [1,253 
                 kW] 
               
               
                 Vapor Compressor 
                 226 
                 HP 
                 [371 
                 kW] 
               
               
                 Totals 
                 1,291 
                 HP 
                 [2,122 
                 kW] 
               
            
           
           
               
            
               
                 Low Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Heater 
                 6,460 
                 MBTU/Hr 
                 [4,173 
                 kW] 
               
            
           
           
               
            
               
                 High Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 Demethanizer Reboiler 
                 17,968 
                 MBTU/Hr 
                 [11,606 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     Comparing the recovery levels displayed in Table II above for the  FIG. 2  prior art process with those in Table I for the  FIG. 1  prior art process shows that the  FIG. 2  process can achieve essentially the same ethane recovery and slightly higher propane and butanes+ recoveries. Comparing the utilities consumptions in Table II with those in Table I shows that the  FIG. 2  process requires less power and less high level utility heat than the  FIG. 1  process. The reduction in power is achieved through the use of reflux for demethanizer  21  in the  FIG. 2  process to provide more efficient recovery of the ethane and heavier components in the tower. This in turn allows for a higher tower feed temperature than the  FIG. 1  process, reducing the reboiler heating requirements in demethanizer  21  (which uses high level utility heat) through the use of low level utility heat in heat exchanger  14  to heat the tower feed. (Note that the  FIG. 1  process cools bottom product stream  51   a  to 43° F. [6° C.], versus the desired 0° F. [−18° C.] for the  FIG. 2  process. For the  FIG. 1  process, attempting to cool stream  51   a  to a lower temperature does reduce the high level utility heat requirement of reboiler  25 , but the resulting higher temperature for stream  41   c  entering separator  15  causes the power usage of vapor compressor  27  to increase disproportionately, because the operating pressure of separator  15  must be lowered if the same recovery efficiencies are to be maintained.) 
     DESCRIPTION OF THE INVENTION 
     EXAMPLE 1 
       FIG. 3  illustrates a flow diagram of a process in accordance with the present invention. The LNG composition and conditions considered in the process presented in  FIG. 3  are the same as those in  FIGS. 1 and 2 . Accordingly, the  FIG. 3  process can be compared with that of the  FIGS. 1 and 2  processes to illustrate the advantages of the present invention. 
     In the simulation of the  FIG. 3  process, the LNG to be processed (stream  41 ) from LNG tank  10  enters pump  11  at −255° F. [−159° C.]. Pump  11  elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to separator  15 . Stream  41   a  exiting the pump is split into two portions, streams  42  and  43 . The first portion, stream  42 , is expanded to the operating pressure (approximately 450 psia [3,103 kPa(a)]) of fractionation column  21  by expansion valve  17  and supplied to the tower at an upper mid-column feed point. The second portion, stream  43 , is heated prior to entering separator  15  so that all or a portion of it is vaporized. In the example shown in  FIG. 3 , stream  43  is first heated to −106° F. [−77° C.] in heat exchangers  12  and  13  by cooling compressed overhead vapor stream  48   a  at −112° F. [−80° C.], reflux stream  53  at −129° F. [−90° C.], and the liquid product from the column (stream  51 ) at 85° F. [29° C.]. The partially heated stream  43   b  is then further heated (stream  43   c ) in heat exchanger  14  using low level utility heat. Note that in all cases exchangers  12 ,  13 , and  14  are 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 heating services will depend on a number of factors including, but not limited to, inlet LNG flow rate, heat exchanger size, stream temperatures, etc.) 
     The heated stream  43   c  enters separator  15  at −62° F. [−52° C.] and 625 psia [4,309 kPa(a)] where the vapor (stream  46 ) is separated from any remaining liquid (stream  47 ). The vapor from separator  15  (stream  46 ) enters a work expansion machine  18  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  18  expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream  46   a  to a temperature of approximately −85° F. [−65° 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  19 ) that can be used to re-compress the column overhead vapor (stream  48 ), for example. The partially condensed expanded stream  46   a  is thereafter supplied as feed to fractionation column  21  at a mid-column feed point. The separator liquid (stream  47 ) is expanded to the operating pressure of fractionation column  21  by expansion valve  20 , cooling stream  47   a  to −77° F. [−61° C.] before it is supplied to fractionation tower  21  at a lower mid-column feed point. 
     The demethanizer in fractionation column  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. Similar to the fractionation tower shown in  FIG. 2 , the fractionation tower in  FIG. 3  may consist of two sections. The upper absorbing (rectification) section contains the trays and/or packing to provide the necessary contact between the vapors rising upward and cold liquid falling downward to condense and absorb the ethane and heavier components; the lower stripping (demethanizing) section contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section also includes one or more reboilers (such as reboiler  25 ) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column. The liquid product stream  51  exits the bottom of the tower at 85° F. [29° C.], based on a methane fraction of 0.005 on a volume basis in the bottom product. After cooling to 0° F. [−18° C.] in heat exchanger  13  as described previously, the liquid product (stream  51   a ) flows to storage or further processing. 
     Overhead distillation stream  48  is withdrawn from the upper section of fractionation tower  21  at −134° F. [−92° C.] and flows to compressor  19  driven by expansion machine  18 , where it is compressed to 550 psia [3,789 kPa(a)] (stream  48   a ). At this pressure, the stream is totally condensed as it is cooled to −129° F. [−90° C.] in heat exchanger  12  as described previously. The condensed liquid (stream  48   b ) is then divided into two portions, streams  52  and  53 . The first portion (stream  52 ) is the methane-rich lean LNG stream, which is then pumped by pump  32  to 1365 psia [9,411 kPa(a)] (stream  52   a ) for subsequent vaporization and/or transportation. 
     The remaining portion is reflux stream  53 , which flows to heat exchanger  12  where it is subcooled to −166° F. [−110° C.] by heat exchange with a portion of the cold LNG (stream  43 ) as described previously. The subcooled reflux stream  53   a  is expanded to the operating pressure of demethanizer  21  by expansion valve  30  and the expanded stream  53   b  is then supplied as cold top column feed (reflux) to demethanizer  21 . This cold liquid reflux absorbs and condenses the C 2  components and heavier hydrocarbon components from the vapors rising in the upper rectification section of demethanizer  21 . 
     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 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 41 
                 9,524 
                 977 
                 322 
                 109 
                 10,979 
               
               
                 42 
                 1,743 
                 179 
                 59 
                 20 
                 2,009 
               
               
                 43 
                 7,781 
                 798 
                 263 
                 89 
                 8,970 
               
               
                 46 
                 7,291 
                 554 
                 96 
                 14 
                 7,993 
               
               
                 47 
                 490 
                 244 
                 167 
                 75 
                 977 
               
               
                 48 
                 10,318 
                 105 
                 0 
                 0 
                 10,474 
               
               
                 53 
                 805 
                 8 
                 0 
                 0 
                 817 
               
               
                 52 
                 9,513 
                 97 
                 0 
                 0 
                 9,657 
               
               
                 51 
                 11 
                 880 
                 322 
                 109 
                 1,322 
               
            
           
           
               
            
               
                 Recoveries* 
               
            
           
           
               
               
               
            
               
                   
                 Ethane 
                  90.05% 
               
               
                   
                 Propane 
                  99.89% 
               
               
                   
                 Butanes+ 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Feed Pump 
                 396 
                 HP 
                 [651 
                 kW] 
               
               
                 LNG Product Pump 
                 756 
                 HP 
                 [1,243 
                 kW] 
               
               
                 Totals 
                 1,152 
                 HP 
                 [1,894 
                 kW] 
               
            
           
           
               
            
               
                 Low Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Heater 
                 18,077 
                 MBTU/Hr 
                 [11,677 
                 kW] 
               
            
           
           
               
            
               
                 High Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 Demethanizer Reboiler 
                 8,441 
                 MBTU/Hr 
                 [5,452 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     Comparing the recovery levels displayed in Table III above for the  FIG. 3  process with those in Table I for the  FIG. 1  prior art process shows that the present invention matches the ethane recovery and achieves slightly higher propane recovery (99.89% versus 98.33%) and butanes+recovery (100.00% versus 99.62%) of the  FIG. 1  process. However, comparing the utilities consumptions in Table III with those in Table I shows that both the power required and the high level utility heat required for the present invention are much lower than for the  FIG. 1  process (26% lower and 64% lower, respectively). 
     Comparing the recovery levels displayed in Table III with those in Table II for the  FIG. 2  prior art process shows that the present invention essentially matches the liquids recovery of the  FIG. 2  process. (Only the propane recovery is slightly lower, 99.89% versus 100.00%.) However, comparing the utilities consumptions in Table III with those in Table II shows that both the power required and the high level utility heat required for the present invention are significantly lower than for the  FIG. 2  process (11% lower and 53% lower, respectively). 
     There are three primary factors that account for the improved efficiency of the present invention. First, compared to the  FIG. 1  prior art process, 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  12  to generate a liquid reflux stream (stream  53 ) that contains very little of the C 2  components and heavier hydrocarbon components that are to be recovered, resulting in efficient rectification in the upper absorbing section of fractionation tower  21  and avoiding the equilibrium limitations of the prior art  FIG. 1  process. Second, compared to the  FIGS. 1 and 2  prior art processes, splitting the LNG feed into two portions before feeding fractionation column  21  allows more efficient use of low level utility heat, thereby reducing the amount of high level utility heat consumed by reboiler  25 . The relatively colder portion of the LNG feed (stream  42   a  in  FIG. 3 ) serves as a supplemental reflux stream for fractionation tower  21 , providing partial rectification of the vapors in the expanded vapor and liquid streams (streams  46   a  and  47   a  in  FIG. 3 ) so that heating and partially vaporizing this portion (stream  43 ) of the LNG feed does not unduly increase the condensing load in heat exchanger  12 . Third, compared to the  FIG. 2  prior art process, using a portion of the cold LNG feed (stream  42   a  in  FIG. 3 ) as a supplemental reflux stream allows using less top reflux for fractionation tower  21 , as can be seen by comparing stream  53  in Table III with stream  53  in Table II. The lower top reflux flow, plus the greater degree of heating using low level utility heat in heat exchanger  14  (as seen by comparing Table III with Table II), results in less total liquid feeding fractionation column  21 , reducing the duty required in reboiler  25  and minimizing the amount of high level utility heat needed to meet the specification for the bottom liquid product from the demethanizer. 
     EXAMPLE 2 
     An alternative embodiment of the present invention is shown in  FIG. 4 . The LNG composition and conditions considered in the process presented in  FIG. 4  are the same as those in  FIG. 3 , as well as those described previously for  FIGS. 1 and 2 . Accordingly, the  FIG. 4  process of the present invention can be compared to the embodiment displayed in  FIG. 3  and to the prior art processes displayed in  FIGS. 1 and 2 . 
     In the simulation of the  FIG. 4  process, the LNG to be processed (stream  41 ) from LNG tank  10  enters pump  11  at −255° F. [−159° C.]. Pump  11  elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to separator  15 . Stream  41   a  exiting the pump is heated prior to entering separator  15  so that all or a portion of it is vaporized. In the example shown in  FIG. 4 , stream  41   a  is first heated to −99° F. [−73° C.] in heat exchangers  12  and  13  by cooling compressed overhead vapor stream  48   b  at −63° F. [−53° C.], reflux stream  53  at −135° F. [−93° C.], and the liquid product from the column (stream  51 ) at 85° F. [29° C.]. The partially heated stream  41   c  is then further heated (stream  41   d ) in heat exchanger  14  using low level utility heat. 
     The heated stream  41   d  enters separator  15  at −63° F. [−53° C.] and 658 psia [4,537 kPa(a)] where the vapor (stream  44 ) is separated from any remaining liquid (stream  47 ). The separator liquid (stream  47 ) is expanded to the operating pressure (approximately 450 psia [3,103 kPa(a)]) of fractionation column  21  by expansion valve  20 , cooling stream  47   a  to −82° F. [−63° C.] before it is supplied to fractionation tower  21  at a lower mid-column feed point. 
     The vapor (stream  44 ) from separator  15  is divided into two streams,  45  and  46 . Stream  45 , containing about 30% of the total vapor, passes through heat exchanger  16  in heat exchange relation with the cold demethanizer overhead vapor at −134° F. [−92° C.] (stream  48 ) where it is cooled to substantial condensation. The resulting substantially condensed stream  45   a  at −129° F. [−89° C.] is then flash expanded through expansion valve  17  to the operating pressure of fractionation tower  21 . During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in  FIG. 4 , the expanded stream  45   b  leaving expansion valve  17  reaches a temperature of −133° F. [−92° C.] and is supplied to fractionation tower  21  at an upper mid-column feed point. 
     The remaining 70% of the vapor from separator  15  (stream  46 ) enters a work expansion machine  18  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  18  expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream  46   a  to a temperature of approximately −90° F. [−68° C.]. The partially condensed expanded stream  46   a  is thereafter supplied as feed to fractionation column  21  at a mid-column feed point. 
     The liquid product stream  51  exits the bottom of the tower at 85° F. [29° C.], based on a methane fraction of 0.005 on a volume basis in the bottom product. After cooling to 0° F. [−18° C.] in heat exchanger  13  as described previously, the liquid product (stream  51   a ) flows to storage or further processing. 
     Overhead distillation stream  48  is withdrawn from the upper section of fractionation tower  21  at −134° F. [−92° C.] and passes countercurrently to the incoming feed gas in heat exchanger  16  where it is heated to −78° F. [−61° C.]. The heated stream  48   a  flows to compressor  19  driven by expansion machine  18 , where it is compressed to 498 psia [3,430 kPa(a)] (stream  48   b ). At this pressure, the stream is totally condensed as it is cooled to −135° F. [−93° C.] in heat exchanger  12  as described previously. The condensed liquid (stream  48   c ) is then divided into two portions, streams  52  and  53 . The first portion (stream  52 ) is the methane-rich lean LNG stream, which is then pumped by pump  32  to 1365 psia [9,411 kPa(a)] (stream  52   a ) for subsequent vaporization and/or transportation. 
     The remaining portion is reflux stream  53 , which flows to heat exchanger  12  where it is subcooled to −166° F. [−110° C.] by heat exchange with the cold LNG (stream  41   a ) as described previously. The subcooled reflux stream  53   a  is expanded to the operating pressure of demethanizer  21  by expansion valve  30  and the expanded stream  53   b  is then supplied as cold top column feed (reflux) to demethanizer  21 . This cold liquid reflux absorbs and condenses the C 2  components and heavier hydrocarbon components from the vapors rising in the upper rectification section of demethanizer  21 . 
     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 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 41 
                 9,524 
                 977 
                 322 
                 109 
                 10,979 
               
               
                 44 
                 8,789 
                 647 
                 111 
                 16 
                 9,609 
               
               
                 47 
                 735 
                 330 
                 211 
                 93 
                 1,370 
               
               
                 45 
                 2,663 
                 196 
                 34 
                 5 
                 2,911 
               
               
                 46 
                 6,126 
                 451 
                 77 
                 11 
                 6,698 
               
               
                 48 
                 10,547 
                 108 
                 0 
                 0 
                 10,706 
               
               
                 53 
                 1,034 
                 11 
                 0 
                 0 
                 1,049 
               
               
                 52 
                 9,513 
                 97 
                 0 
                 0 
                 9,657 
               
               
                 51 
                 11 
                 880 
                 322 
                 109 
                 1,322 
               
            
           
           
               
            
               
                 Recoveries* 
               
            
           
           
               
               
               
            
               
                   
                 Ethane 
                  90.06% 
               
               
                   
                 Propane 
                  99.96% 
               
               
                   
                 Butanes+ 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Feed Pump 
                 419 
                 HP 
                 [688 
                 kW] 
               
               
                 LNG Product Pump 
                 761 
                 HP 
                 [1,252 
                 kW] 
               
               
                 Totals 
                 1,180 
                 HP 
                 [1,940 
                 kW] 
               
            
           
           
               
            
               
                 Low Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Heater 
                 16,119 
                 MBTU/Hr 
                 [10,412 
                 kW] 
               
            
           
           
               
            
               
                 High Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 Demethanizer Reboiler 
                 8,738 
                 MBTU/Hr 
                 [5,644 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     Comparing Table IV above for the  FIG. 4  embodiment of the present invention with Table III for the  FIG. 3  embodiment of the present invention shows that the liquids recovery is essentially the same for the  FIG. 4  embodiment. Since the  FIG. 4  embodiment uses the tower overhead (stream  48 ) to generate the supplemental reflux (stream  45   b ) for fractionation column  21  by condensing and subcooling a portion of the separator  15  vapor (stream  45 ) in heat exchanger  16 , the gas entering compressor  19  (stream  48   a ) is considerably warmer than the corresponding stream in the  FIG. 3  embodiment (stream  48 ). Depending on the type of compression equipment used in this service, the warmer temperature may offer advantages in terms of metallurgy, etc. However, since supplemental reflux stream  45   b  supplied to fractionation column  21  is not as cold as stream  42   a  in the  FIG. 3  embodiment, more top reflux (stream  53   b ) is required and less low level utility heating can be used in heat exchanger  14 . This increases the load on reboiler  25  and increases the amount of high level utility heat required by the  FIG. 4  embodiment of the present invention compared to the  FIG. 3  embodiment. The higher top reflux flow rate also increases the power requirements of the  FIG. 4  embodiment slightly (by about 2%) compared to the  FIG. 3  embodiment. The choice of which embodiment to use for a particular application will generally be dictated by the relative costs of power and high level utility heat and the relative capital costs of pumps, heat exchangers, and compressors. 
     EXAMPLE 3 
     A simpler alternative embodiment of the present invention is shown in  FIG. 5 . The LNG composition and conditions considered in the process presented in  FIG. 5  are the same as those in  FIGS. 3 and 4 , as well as those described previously for  FIGS. 1 and 2 . Accordingly, the  FIG. 5  process of the present invention can be compared to the embodiments displayed in  FIGS. 3 and 4  and to the prior art processes displayed in  FIGS. 1 and 2 . 
     In the simulation of the  FIG. 5  process, the LNG to be processed (stream  41 ) from LNG tank  10  enters pump  11  at −255° F. [−159° C.]. Pump  11  elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to separator  15 . Stream  41   a  exiting the pump is heated prior to entering separator  15  so that all or a portion of it is vaporized. In the example shown in  FIG. 5 , stream  41   a  is first heated to −102° F. [−75° C.] in heat exchangers  12  and  13  by cooling compressed overhead vapor stream  48   a  at −110° F. [−79° C.], reflux stream  53  at −128° F. [−89° C.], and the liquid product from the column (stream  51 ) at 85° F. [29° C.]. The partially heated stream  41   c  is then further heated (stream  41   d ) in heat exchanger  14  using low level utility heat. 
     The heated stream  41   d  enters separator  15  at −74° F. [−59° C.] and 715 psia [4,930 kPa(a)] where the vapor (stream  46 ) is separated from any remaining liquid (stream  47 ). The separator vapor (stream  46 ) enters a work expansion machine  18  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  18  expands the vapor substantially isentropically to the tower operating pressure (approximately 450 psia [3,103 kPa(a)]), with the work expansion cooling the expanded stream  46   a  to a temperature of approximately −106° F. [−77° C.]. The partially condensed expanded stream  46   a  is thereafter supplied as feed to fractionation column  21  at a mid-column feed point. The separator liquid (stream  47 ) is expanded to the operating pressure of fractionation tower  21  by expansion valve  20 , cooling stream  47   a  to −99° F. [−73° C.] before it is supplied to fractionation column  21  at a lower mid-column feed point. 
     The liquid product stream  51  exits the bottom of the tower at 85° F. [29° C.], based on a methane fraction of 0.005 on a volume basis in the bottom product. After cooling to 0° F. [−18° C.] in heat exchanger  13  as described previously, the liquid product (stream  51   a ) flows to storage or further processing. 
     Overhead distillation stream  48  is withdrawn from the upper section of fractionation tower  21  at −134° F. [−92° C.] and flows to compressor  19  driven by expansion machine  18 , where it is compressed to 563 psia [3,882 kPa(a)] (stream  48   a ). At this pressure, the stream is totally condensed as it is cooled to −128° F. [−89° C.] in heat exchanger  12  as described previously. The condensed liquid (stream  48   b ) is then divided into two portions, streams  52  and  53 . The first portion (stream  52 ) is the methane-rich lean LNG stream, which is then pumped by pump  32  to 1365 psia [9,411 kPa(a)] (stream  52   a ) for subsequent vaporization and/or transportation. 
     The remaining portion is reflux stream  53 , which flows to heat exchanger  12  where it is subcooled to −184° F. [−120° C.] by heat exchange with the cold LNG (stream  41   a ) as described previously. The subcooled reflux stream  53   a  is expanded to the operating pressure of demethanizer  21  by expansion valve  30  and the expanded stream  53   b  is then supplied as cold top column feed (reflux) to demethanizer  21 . This cold liquid reflux absorbs and condenses the C 2  components and heavier hydrocarbon components from the vapors rising in the upper rectification section of demethanizer  21 . 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 5  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE V 
               
               
                   
               
             
            
               
                 (FIG. 5) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 41 
                 9,524 
                 977 
                 322 
                 109 
                 10,979 
               
               
                 46 
                 7,891 
                 475 
                 72 
                 10 
                 8,493 
               
               
                 47 
                 1,633 
                 502 
                 250 
                 99 
                 2,486 
               
               
                 48 
                 11,861 
                 121 
                 0 
                 0 
                 12,042 
               
               
                 53 
                 2,348 
                 24 
                 0 
                 0 
                 2,385 
               
               
                 52 
                 9,513 
                 97 
                 0 
                 0 
                 9,657 
               
               
                 51 
                 11 
                 880 
                 322 
                 109 
                 1,322 
               
            
           
           
               
            
               
                 Recoveries* 
               
            
           
           
               
               
               
            
               
                   
                 Ethane 
                  90.02% 
               
               
                   
                 Propane 
                 100.00% 
               
               
                   
                 Butanes+ 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Feed Pump 
                 457 
                 HP 
                 [752 
                 kW] 
               
               
                 LNG Product Pump 
                 756 
                 HP 
                 [1,242 
                 kW] 
               
               
                 Totals 
                 1,213 
                 HP 
                 [1,994 
                 kW] 
               
            
           
           
               
            
               
                 Low Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Heater 
                 16,394 
                 MBTU/Hr 
                 [10,590 
                 kW] 
               
            
           
           
               
            
               
                 High Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 Demethanizer Reboiler 
                 10,415 
                 MBTU/Hr 
                 [6,728 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     Comparing Table V above for the  FIG. 5  embodiment of the present invention with Table III for the  FIG. 3  embodiment and Table IV for the  FIG. 4  embodiment of the present invention shows that the liquids recovery is essentially the same for the  FIG. 5  embodiment. Since the  FIG. 5  embodiment does not use supplemental reflux for fractionation column  21  like the  FIGS. 3 and 4  embodiments do (streams  42   a  and  45   b , respectively), more top reflux (stream  53   b ) is required and less low level utility heating can be used in heat exchanger  14 . This increases the load on reboiler  25  and increases the amount of high level utility heat required by the  FIG. 5  embodiment of the present invention compared to the  FIGS. 3 and 4  embodiments. The higher top reflux flow rate also increases the power requirements of the  FIG. 5  embodiment slightly (by about 5% and 3%, respectively) compared to the  FIGS. 3 and 4  embodiments. The choice of which embodiment to use for a particular application will generally be dictated by the relative costs of power and high level utility heat and the relative capital costs of columns, pumps, heat exchangers, and compressors. 
     EXAMPLE 4 
     A slightly more complex design that maintains the same C 2  component recovery with lower power consumption can be achieved using another embodiment of the present invention as illustrated in the  FIG. 6  process. The LNG composition and conditions considered in the process presented in  FIG. 6  are the same as those in  FIGS. 3 through 5 , as well as those described previously for  FIGS. 1 and 2 . Accordingly, the  FIG. 6  process of the present invention can be compared to the embodiments displayed in  FIGS. 3 through 5  and to the prior art processes displayed in  FIGS. 1 and 2 . 
     In the simulation of the  FIG. 6  process, the LNG to be processed (stream  41 ) from LNG tank  10  enters pump  11  at −255° F. [−159° C.]. Pump  11  elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to absorber column  21 . In the example shown in  FIG. 6 , stream  41   a  exiting the pump is first heated to −120° F. [−84° C.] in heat exchanger  12  by cooling the overhead vapor (distillation stream  48 ) withdrawn from contacting and separating device absorber column  21  at −129° F. [−90° C.] and the overhead vapor (distillation stream  50 ) withdrawn from fractionation stripper column  24  at −83° F. [−63° C.]. The partially heated liquid stream  41   b  is then divided into two portions, streams  42  and  43 . The first portion, stream  42 , is expanded to the operating pressure (approximately 495 psia [3,413 kPa(a)]) of absorber column  21  by expansion valve  17  and supplied to the tower at a lower mid-column feed point. 
     The second portion, stream  43 , is heated prior to entering absorber column  21  so that all or a portion of it is vaporized. In the example shown in  FIG. 6 , stream  43  is first heated to −112° F. [−80° C.] in heat exchanger  13  by cooling the liquid product from fractionation stripper column  24  (stream  51 ) at 88° F. [31° C.]. The partially heated stream  43   a  is then further heated (stream  43   b ) in heat exchanger  14  using low level utility heat. The partially vaporized stream  43   b  is expanded to the operating pressure of absorber column  21  by expansion valve  20 , cooling stream  43   c  to −67° F. [−55° C.] before it is supplied to absorber column  21  at a lower column feed point. The liquid portion (if any) of expanded stream  43   c  commingles with liquids falling downward from the upper section of absorber column  21  and the combined liquid stream  49  exits the bottom of absorber column  21  at −79° F. [−62° C.]. The vapor portion of expanded stream  43   c  rises upward through absorber column  21  and is contacted with cold liquid falling downward to condense and absorb the C 2  components and heavier hydrocarbon components. 
     The combined liquid stream  49  from the bottom of contacting device absorber column  21  is flash expanded to slightly above the operating pressure (465 psia [3,206 kPa(a)]) of stripper column  24  by expansion valve  22 , cooling stream  49  to −83° F. [−64° C.] (stream  49   a ) before it enters fractionation stripper column  24  at a top column feed point. In the stripper column  24 , stream  49   a  is stripped of its methane by the vapors generated in reboiler  25  to meet the specification of a methane fraction of 0.005 on a volume basis. The resulting liquid product stream  51  exits the bottom of stripper column  24  at 88° F. [31° C.], is cooled to 0° F. [−18° C.] in heat exchanger  13  (stream  51   a ) as described previously, and then flows to storage or further processing. 
     The overhead vapor (stream  50 ) from stripper column  24  exits the column at −83° F. [−63° C.] and flows to heat exchanger  12  where it is cooled to −132° F. [−91° C.] as previously described, totally condensing the stream. Condensed liquid stream  50   a  then enters overhead pump  33 , which elevates the pressure of stream  50   b  to slightly above the operating pressure of absorber column  21 . After expansion to the operating pressure of absorber column  21  by control valve  35 , stream  50   c  at −130° F. [−90° C.] is then supplied to absorber column  21  at an upper mid-column feed point where it commingles with liquids falling downward from the upper section of absorber column  21  and becomes part of liquids used to capture the C 2  and heavier components in the vapors rising from the lower section of absorber column  21 . 
     Overhead distillation stream  48 , withdrawn from the upper section of absorber column  21  at −129° F. [−90° C.], flows to heat exchanger  12  and is cooled to −135° F. [−93° C.] as described previously, totally condensing the stream. The condensed liquid (stream  48   a ) is pumped to a pressure somewhat above the operating pressure of absorber column  21  by pump  31  (stream  48   b ), then divided into two portions, streams  52  and  53 . The first portion (stream  52 ) is the methane-rich lean LNG stream, which is then pumped by pump  32  to 1365 psia [9,411 kPa(a)] (stream  52   a ) for subsequent vaporization and/or transportation. 
     The remaining portion is reflux stream  53 , which is expanded to the operating pressure of absorber column  21  by control valve  30 . The expanded stream  53   a  is then supplied at −135° F. [−93° C.] as cold top column feed (reflux) to absorber column  21 . This cold liquid reflux absorbs and condenses the C 2  components and heavier hydrocarbon components from the vapors rising in the upper section of absorber column  21 . 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 6  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE VI 
               
               
                   
               
             
            
               
                 (FIG. 6) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 41 
                 9,524 
                 977 
                 322 
                 109 
                 10,979 
               
               
                 42 
                 2,769 
                 284 
                 94 
                 32 
                 3,192 
               
               
                 43 
                 6,755 
                 693 
                 228 
                 77 
                 7,787 
               
               
                 48 
                 10,546 
                 108 
                 0 
                 0 
                 10,706 
               
               
                 49 
                 1,373 
                 994 
                 329 
                 109 
                 2,808 
               
               
                 50 
                 1,362 
                 114 
                 7 
                 0 
                 1,486 
               
               
                 53 
                 1,033 
                 11 
                 0 
                 0 
                 1,049 
               
               
                 52 
                 9,513 
                 97 
                 0 
                 0 
                 9,657 
               
               
                 51 
                 11 
                 880 
                 322 
                 109 
                 1,322 
               
            
           
           
               
            
               
                 Recoveries* 
               
            
           
           
               
               
               
            
               
                   
                 Ethane 
                  90.04% 
               
               
                   
                 Propane 
                  99.88% 
               
               
                   
                 Butanes+ 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Feed Pump 
                 359 
                 HP 
                 [590 
                 kW] 
               
               
                 Absorber Overhead Pump 
                 48 
                 HP 
                 [79 
                 kW] 
               
               
                 Stripper Overhead Pump 
                 11 
                 HP 
                 [18 
                 kW] 
               
               
                 LNG Product Pump 
                 717 
                 HP 
                 [1,179 
                 kW] 
               
               
                 Totals 
                 1,135 
                 HP 
                 [1,866 
                 kW] 
               
            
           
           
               
            
               
                 Low Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Heater 
                 16,514 
                 MBTU/Hr 
                 [10,667 
                 kW] 
               
            
           
           
               
            
               
                 High Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 Demethanizer Reboiler 
                 8,358 
                 MBTU/Hr 
                 [5,399 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     Comparing Table VI above for the  FIG. 6  embodiment of the present invention with Tables III through V for the  FIGS. 3 through 5  embodiments of the present invention shows that the liquids recovery is essentially the same for the  FIG. 6  embodiment. However, comparing the utilities consumptions in Table VI with those in Tables III through V shows that both the power required and the high level utility heat required for the  FIG. 6  embodiment of the present invention are lower than for the  FIGS. 3 through 5  embodiments. The power requirement for the  FIG. 6  embodiment is 1%, 4%, and 6% lower, respectively and the high level utility heat requirement is 1%, 4%, and 20% lower, respectively. 
     The reductions in utilities requirements for the  FIG. 6  embodiment of the present invention relative to the  FIGS. 3 through 5  embodiments can be attributed mainly to two factors. First, by splitting fractionation column  21  in the  FIGS. 3 through 5  embodiments into a separate absorber column  21  and stripper column  24 , the operating pressures of the two columns can be optimized independently for their respective services. The operating pressure of fractionation column  21  in the  FIGS. 3 through 5  embodiments cannot be raised much above the values shown without incurring the detrimental effect on distillation performance that would result from the higher operating pressure. This effect is manifested by poor mass transfer in fractionation column  21  due to the phase behavior of its vapor and liquid streams. Of particular concern are the physical properties that affect the vapor-liquid separation efficiency, namely the liquid surface tension and the differential in the densities of the two phases. With the operating pressures of the rectification operation (absorber column  21 ) and the stripping operation (stripper column  24 ) no longer coupled together as they are in the  FIGS. 3 through 5  embodiments, the stripping operation can be conducted at a reasonable operating pressure while conducting the rectification operation at a higher pressure that facilitates the condensation of its overhead stream (stream  48  in the  FIG. 6  embodiment) in heat exchanger  12 . 
     Second, in addition to the portion of the LNG feed stream used as a supplemental reflux stream in the  FIGS. 3 and 4  embodiments (stream  42   a  in  FIG. 3  and stream  45   b  in  FIG. 4 ), the  FIG. 6  embodiment of the present invention uses a second supplemental reflux stream (stream  50   c ) for absorber column  21  to help rectify the vapors in stream  43   c  entering the lower section of absorber column  21 . This allows for more optimal use of low level utility heat in heat exchanger  14  to reduce the load on reboiler  25 , reducing the high level utility heat requirement. The choice of which embodiment to use for a particular application will generally be dictated by the relative costs of power and high level utility heat and the relative capital costs of columns, pumps, heat exchangers, and compressors. 
     EXAMPLE 5 
     The present invention can also be adapted to produce an LPG product containing the majority of the C 3  components and heavier hydrocarbon components present in the feed stream as shown in  FIG. 7 . The LNG composition and conditions considered in the process presented in  FIG. 7  are the same as described previously for  FIGS. 1 through 6 . Accordingly, the  FIG. 7  process of the present invention can be compared to the prior art processes displayed in  FIGS. 1 and 2  as well as to the other embodiments of the present invention displayed in  FIGS. 3 through 6 . 
     In the simulation of the  FIG. 7  process, the LNG to be processed (stream  41 ) from LNG tank  10  enters pump  11  at −255° F. [−159° C.]. Pump  11  elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to absorber column  21 . In the example shown in  FIG. 7 , stream  41   a  exiting the pump is first heated to −99° F. [−73° C.] in heat exchangers  12  and  13  by cooling the overhead vapor (distillation stream  48 ) withdrawn from contacting and separating device absorber column  21  at −90° F. [−68° C.], the compressed overhead vapor (stream  50   a ) at 57° F. [14° C.] which was withdrawn from fractionation stripper column  24 , and the liquid product from fractionation stripper column  24  (stream  51 ) at 190° F. [88° C.]. 
     The partially heated stream  41   c  is then further heated (stream  41   d ) to −43° F. [−42° C.] in heat exchanger  14  using low level utility heat. The partially vaporized stream  41   d  is expanded to the operating pressure (approximately 465 psia [3,206 kPa(a)]) of absorber column  21  by expansion valve  20 , cooling stream  41   e  to −48° F. [−44° C.] before it is supplied to absorber column  21  at a lower column feed point. The liquid portion (if any) of expanded stream  41   e  commingles with liquids falling downward from the upper section of absorber column  21  and the combined liquid stream  49  exits the bottom of absorber column  21  at −50° F. [−46° C.]. The vapor portion of expanded stream  41   e  rises upward through absorber column  21  and is contacted with cold liquid falling downward to condense and absorb the C 3  components and heavier hydrocarbon components. 
     The combined liquid stream  49  from the bottom of contacting device absorber column  21  is flash expanded to slightly above the operating pressure (430 psia [2,965 kPa(a)]) of stripper column  24  by expansion valve  22 , cooling stream  49  to −53° F. [−47° C.] (stream  49   a ) before it enters fractionation stripper column  24  at a top column feed point. In the stripper column  24 , stream  49   a  is stripped of its methane and C 2  components by the vapors generated in reboiler  25  to meet the specification of an ethane to propane ratio of 0.020:1 on a molar basis. The resulting liquid product stream  51  exits the bottom of stripper column  24  at 190° F. [88° C.], is cooled to 0° F. [−18° C.] in heat exchanger  13  (stream  51   a ) as described previously, and then flows to storage or further processing. 
     The overhead vapor (stream  50 ) from stripper column  24  exits the column at 30° F. [−1° C.] and flows to overhead compressor  34  (driven by a supplemental power source), which elevates the pressure of stream  50   a  to slightly above the operating pressure of absorber column  21 . Stream  50   a  enters heat exchanger  12  where it is cooled to −78° F. [−61° C.] as previously described, totally condensing the stream. Condensed liquid stream  50   b  is expanded to the operating pressure of absorber column  21  by control valve  35 , and the resulting stream  50   c  at −84° F. [−64° C.] is then supplied to absorber column  21  at a mid-column feed point where it commingles with liquids falling downward from the upper section of absorber column  21  and becomes part of liquids used to capture the C 3  and heavier components in the vapors rising from the lower section of absorber column  21 . 
     Overhead distillation stream  48 , withdrawn from the upper section of absorber column  21  at −90° F. [−68° C.], flows to heat exchanger  12  and is cooled to −132° F. [−91° C.] as described previously, totally condensing the stream. The condensed liquid (stream  48   a ) is pumped to a pressure somewhat above the operating pressure of absorber column  21  by pump  31  (stream  48   b ), then divided into two portions, streams  52  and  53 . The first portion (stream  52 ) is the methane-rich lean LNG stream, which is then pumped by pump  32  to 1365 psia [9,411 kPa(a)] (stream  52   a ) for subsequent vaporization and/or transportation. 
     The remaining portion is reflux stream  53 , which is expanded to the operating pressure of absorber column  21  by control valve  30 . The expanded stream  53   a  is then supplied at −131° F. [−91° C.] as cold top column feed (reflux) to absorber column  21 . This cold liquid reflux absorbs and condenses the C 3  components and heavier hydrocarbon components from the vapors rising in the upper section of absorber column  21 . 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 7  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE VII 
               
               
                   
               
             
            
               
                 (FIG. 7) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 41 
                 9,524 
                 977 
                 322 
                 109 
                 10,979 
               
               
                 48 
                 11,475 
                 1,170 
                 4 
                 0 
                 12,705 
               
               
                 49 
                 426 
                 326 
                 396 
                 116 
                 1,266 
               
               
                 50 
                 426 
                 320 
                 77 
                 7 
                 832 
               
               
                 53 
                 1,951 
                 199 
                 1 
                 0 
                 2,160 
               
               
                 52 
                 9,524 
                 971 
                 3 
                 0 
                 10,545 
               
               
                 51 
                 0 
                 6 
                 319 
                 109 
                 434 
               
            
           
           
               
            
               
                 Recoveries* 
               
            
           
           
               
               
               
            
               
                   
                 Propane 
                  99.00% 
               
               
                   
                 Butanes+ 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Feed Pump 
                 325 
                 HP 
                 [535 
                 kW] 
               
               
                 Absorber Overhead Pump 
                 54 
                 HP 
                 [89 
                 kW] 
               
               
                 LNG Product Pump 
                 775 
                 HP 
                 [1,274 
                 kW] 
               
               
                 Stripper Ovhd Compressor 
                 67 
                 HP 
                 [110 
                 kW] 
               
               
                 Totals 
                 1,221 
                 HP 
                 [2,008 
                 kW] 
               
            
           
           
               
            
               
                 Low Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Heater 
                 15,139 
                 MBTU/Hr 
                 [9,779 
                 kW] 
               
            
           
           
               
            
               
                 High Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 Deethanizer Reboiler 
                 6,857 
                 MBTU/Hr 
                 [4,429 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     Comparing the utilities consumptions in Table VII above for the  FIG. 7  process with those in Tables III through VI shows that the power requirement for this embodiment of the present invention is slightly higher than that of the  FIGS. 3 through 6  embodiments. However, the high level utility heat required for the  FIG. 7  embodiment of the present invention is significantly lower than that for the  FIGS. 3 through 6  embodiments because more low level utility heat can be used in heat exchanger  14  when recovery of the C 2  components is not desired. 
     EXAMPLE 6 
     The increase in the power requirement of the  FIG. 7  embodiment relative to the  FIGS. 3 through 6  embodiments of the present invention is mainly due to compressor  34  in  FIG. 7  which provides the motive force needed to direct the overhead vapor (stream  50 ) from stripper column  24  through heat exchanger  12  and thence into absorber column  21 .  FIG. 8  illustrates an alternative embodiment of the present invention that eliminates this compressor and reduces the power requirement. The LNG composition and conditions considered in the process presented in  FIG. 8  are the same as those in  FIG. 7 , as well as those described previously for  FIGS. 1 through 6 . Accordingly, the  FIG. 8  process of the present invention can be compared to the embodiment of the present invention displayed in  FIG. 7 , to the prior art processes displayed in  FIGS. 1 and 2 , and to the other embodiments of the present invention displayed in  FIGS. 3 through 6 . 
     In the simulation of the  FIG. 8  process, the LNG to be processed (stream  41 ) from LNG tank  10  enters pump  11  at −255° F. [−159° C.]. Pump  11  elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to absorber column  21 . Stream  41   a  exiting the pump is heated first to −101° F. [−74° C.] in heat exchangers  12  and  13  as it provides cooling to the overhead vapor (distillation stream  48 ) withdrawn from contacting and separating device absorber column  21  at −90° F. [−68° C.], the overhead vapor (distillation stream  50 ) withdrawn from fractionation stripper column  24  at 20° F. [−7° C.], and the liquid product (stream  51 ) from fractionation stripper column  21  at 190° F. [88° C.]. 
     The partially heated stream  41   c  is then further heated (stream  41   d ) in heat exchanger  14  to −54° F. [−48° C.] using low level utility heat. After expansion to the operating pressure (approximately 465 psia [3,206 kPa(a)]) of absorber column  21  by expansion valve  20 , stream  41   e  flows to a lower column feed point on the column at −58° F. [−50° C.]. The liquid portion (if any) of expanded stream  41   e  commingles with liquids falling downward from the upper section of absorber column  21  and the combined liquid stream  49  exits the bottom of contacting device absorber column  21  at −61° F. [−52° C.]. The vapor portion of expanded stream  41   e  rises upward through absorber column  21  and is contacted with cold liquid falling downward to condense and absorb the C 3  components and heavier hydrocarbon components. 
     The combined liquid stream  49  from the bottom of the absorber column  21  is flash expanded to slightly above the operating pressure (430 psia [2,965 kPa(a)]) of stripper column  24  by expansion valve  22 , cooling stream  49  to −64° F. [−53° C.] (stream  49   a ) before it enters fractionation stripper column  24  at a top column feed point. In stripper column  24 , stream  49   a  is stripped of its methane and C 2  components by the vapors generated in reboiler  25  to meet the specification of an ethane to propane ratio of 0.020:1 on a molar basis. The resulting liquid product stream  51  exits the bottom of stripper column  24  at 190° F. [88° C.] and is cooled to 0° F. [−18° C.] in heat exchanger  13  (stream  51   a ) as described previously before flowing to storage or further processing. 
     The overhead vapor (stream  50 ) from stripper column  24  exits the column at 20° F. [−7° C.] and flows to heat exchanger  12  where it is cooled to −98° F. [−72° C.] as previously described, totally condensing the stream. Condensed liquid stream  50   a  then enters overhead pump  33 , which elevates the pressure of stream  50   b  to slightly above the operating pressure of absorber column  21 , whereupon it reenters heat exchanger  12  to be partially vaporized as it is heated to −70° F. [−57° C.] (stream  50   c ) by supplying part of the total cooling duty in this exchanger. After expansion to the operating pressure of absorber column  21  by control valve  35 , stream  50   d  at −75° F. [−60° C.] is then supplied to absorber column  21  at a mid-column feed point where it commingles with liquids falling downward from the upper section of absorber column  21  and becomes part of liquids used to capture the C 3  and heavier components in the vapors rising from the lower section of absorber column  21 . 
     Overhead distillation stream  48  is withdrawn from contacting device absorber column  21  at −90° F. [−68° C.] and flows to heat exchanger  12  where it is cooled to −132° F. [−91° C.] and totally condensed by heat exchange with the cold LNG (stream  41   a ) as described previously. The condensed liquid (stream  48   a ) is pumped to a pressure somewhat above the operating pressure of absorber column  21  by pump  31  (stream  48   b ), then divided into two portions, streams  52  and  53 . The first portion (stream  52 ) is the methane-rich lean LNG stream, which is then pumped by pump  32  to 1365 psia [9,411 kPa(a)] (stream  52   a ) for subsequent vaporization and/or transportation. 
     The remaining portion is reflux stream  53 , which is expanded to the operating pressure of absorber column  21  by control valve  30 . The expanded stream  53   a  is then supplied at −131° F. [−91° C.] as cold top column feed (reflux) to absorber column  21 . This cold liquid reflux absorbs and condenses the C 3  components and heavier hydrocarbon components from the vapors rising in the upper section of absorber column  21 . 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 8  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE VIII 
               
               
                   
               
             
            
               
                 (FIG. 8) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 41 
                 9,524 
                 977 
                 322 
                 109 
                 10,979 
               
               
                 48 
                 10,934 
                 1,115 
                 4 
                 0 
                 12,107 
               
               
                 49 
                 582 
                 458 
                 396 
                 116 
                 1,552 
               
               
                 50 
                 582 
                 452 
                 77 
                 7 
                 1,118 
               
               
                 53 
                 1,410 
                 144 
                 1 
                 0 
                 1,562 
               
               
                 52 
                 9,524 
                 971 
                 3 
                 0 
                 10,545 
               
               
                 51 
                 0 
                 6 
                 319 
                 109 
                 434 
               
            
           
           
               
            
               
                 Recoveries* 
               
            
           
           
               
               
               
            
               
                   
                 Propane 
                  99.03% 
               
               
                   
                 Butanes+ 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Feed Pump 
                 325 
                 HP 
                 [534 
                 kW] 
               
               
                 Absorber Overhead Pump 
                 67 
                 HP 
                 [110 
                 kW] 
               
               
                 Stripper Overhead Pump 
                 11 
                 HP 
                 [18 
                 kW] 
               
               
                 LNG Product Pump 
                 761 
                 HP 
                 [1,251 
                 kW] 
               
               
                 Totals 
                 1,164 
                 HP 
                 [1,913 
                 kW] 
               
            
           
           
               
            
               
                 Low Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Heater 
                 13,949 
                 MBTU/Hr 
                 [9,010 
                 kW] 
               
            
           
           
               
            
               
                 High Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 Deethanizer Reboiler 
                 8,192 
                 MBTU/Hr 
                 [5,292 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     Comparing Table VIII above for the  FIG. 8  embodiment of the present invention with Table VII for the  FIG. 7  embodiment of the present invention shows that the liquids recovery is essentially the same for the  FIG. 8  embodiment. Since the  FIG. 8  embodiment uses a pump (overhead pump  33  in  FIG. 8 ) rather than a compressor (overhead compressor  34  in  FIG. 7 ) to route the overhead vapor from fractionation stripper column  24  to contacting device absorber column  21 , less power is required by the  FIG. 8  embodiment. However, the high level utility heat required for the  FIG. 8  embodiment is higher (by about 19%). The choice of which embodiment to use for a particular application will generally be dictated by the relative costs of power and high level utility heat and the relative costs of pumps versus compressors. 
     EXAMPLE 7 
     A slightly more complex design that maintains the same C 3  component recovery with reduced high level utility heat consumption can be achieved using another embodiment of the present invention as illustrated in the  FIG. 9  process. The LNG composition and conditions considered in the process presented in  FIG. 9  are the same as those in  FIGS. 7 and 8 , as well as those described previously for  FIGS. 1 through 6 . Accordingly, the  FIG. 9  process of the present invention can be compared to the embodiments of the present invention displayed in  FIGS. 7 and 8 , to the prior art processes displayed in  FIGS. 1 and 2 , and to the other embodiments of the present invention displayed in  FIGS. 3 through 6 . 
     In the simulation of the  FIG. 9  process, the LNG to be processed (stream  41 ) from LNG tank  10  enters pump  11  at −255° F. [−159° C.]. Pump  11  elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to separator  15 . Stream  41   a  exiting the pump is heated prior to entering separator  15  so that all or a portion of it is vaporized. In the example shown in  FIG. 9 , stream  41   a  is first heated to −88° F. [−66° C.] in heat exchangers  12  and  13  by cooling compressed overhead vapor stream  48   a  at −70° F. [−57° C.], compressed overhead vapor stream  50   a  at 67° F. [19° C.], and the liquid product from fractionation stripper column  24  (stream  51 ) at 161° F. [72° C.]. The partially heated stream  41   c  is then further heated (stream  41   d ) in heat exchanger  14  using low level utility heat. 
     The heated stream  41   d  enters separator  15  at −16° F. [−27° C.] and 596 psia [4,109 kPa(a)] where the vapor (stream  46 ) is separated from any remaining liquid (stream  47 ). The separator vapor (stream  46 ) enters a work expansion machine  18  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  18  expands the vapor substantially isentropically to the tower operating pressure (approximately 415 psia [2,861 kPa(a)]), with the work expansion cooling the expanded stream  46   a  to a temperature of approximately −42° F. [−41° C.]. The partially condensed expanded stream  46   a  is thereafter supplied as feed to absorber column  21  at a mid-column feed point. If there is any separator liquid (stream  47 ), it is expanded to the operating pressure of absorber column  21  by expansion valve  20  before it is supplied to absorber column  21  at a lower column feed point. In the example shown in  FIG. 9 , stream  41   d  is vaporized completely in heat exchanger  14 , so separator  15  and expansion valve  20  are not needed, and expanded stream  46   a  is supplied to absorber column  21  at a lower column feed point instead. The liquid portion (if any) of expanded stream  46   a  (and expanded stream  47   a  if present) commingles with liquids falling downward from the upper section of absorber column  21  and the combined liquid stream  49  exits the bottom of absorber column  21  at −45° F. [−43° C.]. The vapor portion of expanded stream  46   a  (and expanded stream  47   a  if present) rises upward through absorber column  21  and is contacted with cold liquid falling downward to condense and absorb the C 3  components and heavier hydrocarbon components. 
     The combined liquid stream  49  from the bottom of contacting and separating device absorber column  21  is flash expanded to slightly above the operating pressure (320 psia [2,206 kPa(a)]) of fractionation stripper column  24  by expansion valve  22 , cooling stream  49  to −54° F. [−48° C.] (stream  49   a ) before it enters fractionation stripper column  24  at a top column feed point. In stripper column  24 , stream  49   a  is stripped of its methane and C 2  components by the vapors generated in reboiler  25  to meet the specification of an ethane to propane ratio of 0.020:1 on a molar basis. The resulting liquid product stream  51  exits the bottom of stripper column  24  at 161° F. [72° C.] and is cooled to 0° F. [−18° C.] in heat exchanger  13  (stream  51   a ) as described previously before flowing to storage or further processing. 
     The overhead vapor (stream  50 ) from stripper column  24  exits the column at  20 ° F. [−6° C.] flows to overhead compressor  34  (driven by a portion of the power generated by expansion machine  18 ), which elevates the pressure of stream  50   a  to slightly above the operating pressure of absorber column  21 . Stream  50   a  enters heat exchanger  12  where it is cooled to −87° F. [−66° C.] as previously described, totally condensing the stream. Condensed liquid stream  50   b  is expanded to the operating pressure of absorber column  21  by control valve  35 , and the resulting stream  50   c  at −91° F. [−68° C.] is then supplied to absorber column  21  at a mid-column feed point where it commingles with liquids falling downward from the upper section of absorber column  21  and becomes part of liquids used to capture the C 3  and heavier components in the vapors rising from the lower section of absorber column  21 . 
     Overhead distillation stream  48  is withdrawn from the upper section of absorber column  21  at −94° F. [−70° C.] and flows to compressor  19  (driven by the remaining portion of the power generated by expansion machine  18 ), where it is compressed to 508 psia [3,501 kPa(a)] (stream  48   a ). At this pressure, the stream is totally condensed as it is cooled to −126° F. [−88° C.] in heat exchanger  12  as described previously. The condensed liquid (stream  48   b ) is then divided into two portions, streams  52  and  53 . The first portion (stream  52 ) is the methane-rich lean LNG stream, which is then pumped by pump  32  to 1365 psia [9,411 kPa(a)] (stream  52   a ) for subsequent vaporization and/or transportation. 
     The remaining portion is reflux stream  53 , which is expanded to the operating pressure of absorber column  21  by expansion valve  30 . The expanded stream  53   a  is then supplied at −136° F. [−93° C.] as cold top column feed (reflux) to absorber column  21 . This cold liquid reflux absorbs and condenses the C 3  components and heavier hydrocarbon components from the vapors rising in the upper section of absorber column  21 . 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 9  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE IX 
               
               
                   
               
             
            
               
                 (FIG. 9) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 41 
                 9,524 
                 977 
                 322 
                 109 
                 10,979 
               
               
                 46 
                 9,524 
                 977 
                 322 
                 109 
                 10,979 
               
               
                 48 
                 12,056 
                 1,229 
                 4 
                 0 
                 13,348 
               
               
                 49 
                 304 
                 254 
                 384 
                 115 
                 1,057 
               
               
                 50 
                 304 
                 248 
                 65 
                 6 
                 623 
               
               
                 53 
                 2,532 
                 258 
                 1 
                 0 
                 2,803 
               
               
                 52 
                 9,524 
                 971 
                 3 
                 0 
                 10,545 
               
               
                 51 
                 0 
                 6 
                 319 
                 109 
                 434 
               
            
           
           
               
            
               
                 Recoveries* 
               
            
           
           
               
               
               
            
               
                   
                 Propane 
                  98.99% 
               
               
                   
                 Butanes+ 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Feed Pump 
                 377 
                 HP 
                 [620 
                 kW] 
               
               
                 LNG Product Pump 
                 806 
                 HP 
                 [1,325 
                 kW] 
               
               
                 Totals 
                 1,183 
                 HP 
                 [1,945 
                 kW] 
               
            
           
           
               
            
               
                 Low Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Heater 
                 17,940 
                 MBTU/Hr 
                 [11,588 
                 kW] 
               
            
           
           
               
            
               
                 High Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 Deethanizer Reboiler 
                 5,432 
                 MBTU/Hr 
                 [3,509 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     Comparing Table IX above for the  FIG. 9  embodiment of the present invention with Tables VII and VIII for the  FIGS. 7 and 8  embodiments of the present invention shows that the liquids recovery is essentially the same for the  FIG. 9  embodiment. The power requirement for the  FIG. 9  embodiment is lower than that required by the  FIG. 7  embodiment by about 3% and higher than that required by the  FIG. 8  embodiment by about 2%. However, the high level utility heat required by the  FIG. 9  embodiment of the present invention is significantly lower than either the  FIG. 7  embodiment (by about  21 %) or the  FIG. 8  embodiment (by about 34%). The choice of which embodiment to use for a particular application will generally be dictated by the relative costs of power versus high level utility heat and the relative capital costs of pumps and heat exchangers versus compressors and expansion machines. 
     EXAMPLE 8 
     A slightly simpler embodiment of the present invention that maintains the same C 3  component recovery as the  FIG. 9  embodiment can be achieved using another embodiment of the present invention as illustrated in the  FIG. 10  process. The LNG composition and conditions considered in the process presented in  FIG. 10  are the same as those in  FIGS. 7 through 9 , as well as those described previously for  FIGS. 1 through 6 . Accordingly, the  FIG. 10  process of the present invention can be compared to the embodiments of the present invention displayed in  FIGS. 7 through 9 , to the prior art processes displayed in  FIGS. 1 and 2 , and to the other embodiments of the present invention displayed in  FIGS. 3 through 6 . 
     In the simulation of the  FIG. 10  process, the LNG to be processed (stream  41 ) from LNG tank  10  enters pump  11  at −255° F. [−159° C.]. Pump  11  elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to separator  15 . Stream  41   a  exiting the pump is heated prior to entering separator  15  so that all or a portion of it is vaporized. In the example shown in  FIG. 10 , stream  41   a  is first heated to −83° F. [−64° C.] in heat exchangers  12  and  13  by cooling compressed overhead vapor stream  48   a  at −61° F. [−52° C.], overhead vapor stream  50  at 40° F. [4° C.], and the liquid product from fractionation stripper column  24  (stream  51 ) at 190° F. [88° C.]. The partially heated stream  41   c  is then further heated (stream  41   d ) in heat exchanger  14  using low level utility heat. 
     The heated stream  41   d  enters separator  15  at −16° F. [−26° C.] and 621 psia [4,282 kPa(a)] where the vapor (stream  46 ) is separated from any remaining liquid (stream  47 ). The separator vapor (stream  46 ) enters a work expansion machine  18  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  18  expands the vapor substantially isentropically to the tower operating pressure (approximately 380 psia [2,620 kPa(a)]), with the work expansion cooling the expanded stream  46   a  to a temperature of approximately −50° F. [−46° C.]. The partially condensed expanded stream  46   a  is thereafter supplied as feed to absorber column  21  at a mid-column feed point. If there is any separator liquid (stream  47 ), it is expanded to the operating pressure of absorber column  21  by expansion valve  20  before it is supplied to absorber column  21  at a lower column feed point. In the example shown in  FIG. 10 , stream  41   d  is vaporized completely in heat exchanger  14 , so separator  15  and expansion valve  20  are not needed, and expanded stream  46   a  is supplied to absorber column  21  at a lower column feed point instead. The liquid portion (if any) of expanded stream  46   a  (and expanded stream  47   a  if present) commingles with liquids falling downward from the upper section of absorber column  21  and the combined liquid stream  49  exits the bottom of absorber column  21  at −53° F. [−47° C.]. The vapor portion of expanded stream  46   a  (and expanded stream  47   a  if present) rises upward through absorber column  21  and is contacted with cold liquid falling downward to condense and absorb the C 3  components and heavier hydrocarbon components. 
     The combined liquid stream  49  from the bottom of contacting and separating device absorber column  21  enters pump  23  and is pumped to slightly above the operating pressure (430 psia [2,965 kPa(a)]) of stripper column  24 . The resulting stream  49   a  at −52° F. [−47° C.] then enters fractionation stripper column  24  at a top column feed point. In stripper column  24 , stream  49   a  is stripped of its methane and C 2  components by the vapors generated in reboiler  25  to meet the specification of an ethane to propane ratio of 0.020:1 on a molar basis. The resulting liquid product stream  51  exits the bottom of stripper column  24  at 190° F. [88° C.] and is cooled to 0° F. [−18° C.] in heat exchanger  13  (stream  51   a ) as described previously before flowing to storage or further processing. 
     The overhead vapor (stream  50 ) from stripper column  24  exits the column at 40° F. [4° C.] and enters heat exchanger  12  where it is cooled to −89° F. [−67° C.] as previously described, totally condensing the stream. Condensed liquid stream  50   a  is expanded to the operating pressure of absorber column  21  by expansion valve  35 , and the resulting stream  50   b  at −94° F. [−70° C.] is then supplied to absorber column  21  at a mid-column feed point where it commingles with liquids falling downward from the upper section of absorber column  21  and becomes part of liquids used to capture the C 3  and heavier components in the vapors rising from the lower section of absorber column  21 . 
     Overhead distillation stream  48  is withdrawn from the upper section of absorber column  21  at −97° F. [−72° C.] and flows to compressor  19  driven by expansion machine  18 , where it is compressed to 507 psia [3,496 kPa(a)] (stream  48   a ). At this pressure, the stream is totally condensed as it is cooled to −126° F. [−88° C.] in heat exchanger  12  as described previously. The condensed liquid (stream  48   b ) is then divided into two portions, streams  52  and  53 . The first portion (stream  52 ) is the methane-rich lean LNG stream, which is then pumped by pump  32  to 1365 psia [9,411 kPa(a)] (stream  52   a ) for subsequent vaporization and/or transportation. 
     The remaining portion is reflux stream  53 , which is expanded to the operating pressure of absorber column  21  by expansion valve  30 . The expanded stream  53   a  is then supplied at −141° F. [−96° C.] as cold top column feed (reflux) to absorber column  21 . This cold liquid reflux absorbs and condenses the C 3  components and heavier hydrocarbon components from the vapors rising in the upper section of absorber column  21 . 
     A summary of stream flow rates and energy consumption for the process illustrated in  FIG. 10  is set forth in the following table: 
     
       
         
           
               
             
               
                 TABLE X 
               
               
                   
               
             
            
               
                 (FIG. 10) 
               
               
                 Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] 
               
            
           
           
               
               
               
               
               
               
            
               
                 Stream 
                 Methane 
                 Ethane 
                 Propane 
                 Butanes+ 
                 Total 
               
               
                   
               
               
                 41 
                 9,524 
                 977 
                 322 
                 109 
                 10,979 
               
               
                 46 
                 9,524 
                 977 
                 322 
                 109 
                 10,979 
               
               
                 48 
                 11,631 
                 1,186 
                 4 
                 0 
                 12,879 
               
               
                 49 
                 309 
                 275 
                 395 
                 117 
                 1,096 
               
               
                 50 
                 309 
                 269 
                 76 
                 8 
                 662 
               
               
                 53 
                 2,107 
                 215 
                 1 
                 0 
                 2,334 
               
               
                 52 
                 9,524 
                 971 
                 3 
                 0 
                 10,545 
               
               
                 51 
                 0 
                 6 
                 319 
                 109 
                 434 
               
            
           
           
               
            
               
                 Recoveries* 
               
            
           
           
               
               
               
            
               
                   
                 Propane 
                  99.02% 
               
               
                   
                 Butanes+ 
                 100.00% 
               
               
                   
                   
               
            
           
           
               
            
               
                 Power 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Feed Pump 
                 394 
                 HP 
                 [648 
                 kW] 
               
               
                 Absorber Bottoms Pump 
                 9 
                 HP 
                 [14 
                 kW] 
               
               
                 LNG Product Pump 
                 806 
                 HP 
                 [1,325 
                 kW] 
               
               
                 Totals 
                 1,209 
                 HP 
                 [1,987 
                 kW] 
               
            
           
           
               
            
               
                 Low Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 LNG Heater 
                 16,912 
                 MBTU/Hr 
                 [10,924 
                 kW] 
               
            
           
           
               
            
               
                 High Level Utility Heat 
               
            
           
           
               
               
               
               
               
            
               
                 Deethanizer Reboiler 
                 6,390 
                 MBTU/Hr 
                 [4,127 
                 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates) 
               
            
           
         
       
     
     Comparing Table X above for the  FIG. 10  embodiment of the present invention with Tables VII through IX for the  FIGS. 7 through 9  embodiments of the present invention shows that the liquids recovery is essentially the same for the  FIG. 10  embodiment. The power requirement for the  FIG. 10  embodiment is lower than that required by the  FIG. 7  embodiment by about 1% and higher than that required by the  FIGS. 8 and 9  embodiments by about 4% and 2%, respectively. The high level utility heat required by the  FIG. 10  embodiment of the present invention is significantly lower than both the  FIGS. 7 and 8  embodiments (by about 7% and 22%, respectively), but higher than the  FIG. 9  embodiment by about 18%. The choice of which embodiment to use for a particular application will generally be dictated by the relative costs of power versus high level utility heat and the relative capital costs of pumps, heat exchangers, compressors, and expansion machines. 
     Other Embodiments 
     Some circumstances may favor subcooling reflux stream  53  with another process stream, rather than using the cold LNG stream that enters heat exchanger  12 . In such circumstances, alternative embodiments of the present invention such as that shown in  FIGS. 11 through 13  could be employed. In the  FIGS. 11 and 12  embodiments, a portion (stream  42 ) of partially heated LNG stream  41   b  leaving heat exchanger  12  is expanded to slightly above the operating pressure of fractionation tower  21  ( FIG. 11 ) or absorber column  21  ( FIG. 12 ) by expansion valve  17  and the expanded stream  42   a  is directed into heat exchanger  29  to be heated as it provides subcooling of reflux stream  53 . The subcooled reflux stream  53   a  is then expanded to the operating pressure of fractionation tower  21  ( FIG. 11 ) or contacting and separating device absorber column  21  ( FIG. 12 ) by expansion valve  30  and the expanded stream  53   b  supplied as cold top column feed (reflux) to fractionation tower  21  ( FIG. 11 ) or absorber column  21  ( FIG. 12 ). The heated stream  42   b  leaving heat exchanger  29  is supplied to the tower at a mid-column feed point where it serves as a supplemental reflux stream. Alternatively, as shown by the dashed lines in  FIGS. 11 and 12 , stream  42  may be withdrawn from LNG stream  41   a  before it enters heat exchanger  12 . In the  FIG. 13  embodiment, the supplemental reflux stream produced by condensing overhead vapor stream  50  from fractionation stripper column  24  is used to subcool reflux stream  53  in heat exchanger  29  by expanding stream  50   b  to slightly above the operating pressure of absorber column  21  with control valve  17  and directing the expanded stream  50   c  into heat exchanger  29 . The heated stream  50   d  is then supplied to the tower at a mid-column feed point. 
     The decision regarding whether or not to subcool reflux stream  53  before it is expanded to the column operating pressure will depend on many factors, including the LNG composition, the desired recovery level, etc. As shown by the dashed lines in  FIGS. 3 through 10 , stream  53  can be routed to heat exchanger  12  if subcooling is desired, or routed directly to expansion valve  30  if no subcooling is desired. Likewise, heating of supplemental reflux stream  42  before it is expanded to the column operating pressure must be evaluated for each application. As shown by the dashed lines in  FIGS. 3 ,  6 , and  13 , stream  42  can be withdrawn prior to heating of LNG stream  41   a  and routed directly to expansion valve  17  if no heating is desired, or withdrawn from the partially heated LNG stream  41   b  and routed to expansion valve  17  if heating is desired. On the other hand, heating and partial vaporization of supplemental reflux stream  50   b  as shown in  FIG. 8  may not be advantageous, since this reduces the amount of liquid entering absorber column  21  that is used to capture the C 2  components and/or C 3  components and the heavier hydrocarbon components in the vapors rising upward from the lower section of absorber column  21 . Instead, as shown by the dashed line in  FIG. 8 , stream  50   b  can be routed directly to expansion valve  35  and thence into absorber column  21 . 
     When the LNG to be processed is leaner or when complete vaporization of the LNG in heat exchangers  12 ,  13 , and  14  is contemplated, separator  15  in  FIGS. 3 through 5  and  9  through  11  may not be justified. Depending on the quantity of heavier hydrocarbons in the inlet LNG and the pressure of the LNG stream leaving feed pump  11 , the heated LNG stream leaving heat exchanger  14  in may not contain any liquid (because it is above its dewpoint, or because it is above its cricondenbar). In such cases, separator  15  and expansion valve  20  may be eliminated as shown by the dashed lines. 
     In the examples shown, total condensation of stream  48   a  in  FIGS. 3 ,  5 , and  9  through  11 , stream  48   b  in  FIG. 4 , stream  48  in  FIGS. 6 through 8 ,  12 , and  13 , stream  50  in  FIGS. 6 ,  8 ,  10 ,  12 , and  13 , and stream  50   a  in  FIGS. 7 and 9  is shown. Some circumstances may favor subcooling either or both of these streams, while other circumstances may favor only partial condensation. Should partial condensation of either or both streams be used, processing of the uncondensed vapor may be necessary, using a compressor or other means to elevate the pressure of the vapor so that it can join the pumped condensed liquid. Alternatively, the uncondensed vapor could be routed to the plant fuel system or other such use. 
     LNG conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine  18  in  FIGS. 3 through 5  and  9  through  11 , or replacement with an alternate expansion device (such as an expansion valve), is feasible. Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where appropriate. 
     It also should be noted that expansion valves  17 ,  20 ,  22 ,  30 , and/or  35  could be replaced with expansion engines (turboexpanders) whereby work could be extracted from the pressure reduction of stream  42  in  FIGS. 3 ,  6 , and  11  through  13 , stream  45   a  in  FIG. 4 , stream  47  in  FIGS. 3 through 5  and  9  through  11 , stream  43   b  in  FIGS. 6 ,  12 , and  13 , stream  41   d  in  FIGS. 7 and 8 , stream  49  in  FIGS. 6 through 9 ,  12 , and  13 , stream  53   a  in  FIGS. 3 through 5  and  11  through  13 , stream  53  in  FIGS. 6 through 10 , stream  50   b  in  FIGS. 6 ,  7 ,  9 ,  12 , and  13 , stream  50   c  in  FIG. 8 , and/or stream  50   a  in  FIG. 10 . In such cases, the LNG (stream  41 ) and/or other liquid streams may need to be pumped to a higher pressure so that work extraction is feasible. This work could be used to provide power for pumping the LNG feed stream, for pumping the lean LNG product stream, for compression of overhead vapor streams, or to generate electricity. The choice between use of valves or expansion engines will depend on the particular circumstances of each LNG processing project. 
     In  FIGS. 3 through 13 , individual heat exchangers have been shown for most services. However, it is possible to combine two or more heat exchange services into a common heat exchanger, such as combining heat exchangers  12 ,  13 , and  14  in  FIGS. 3 through 13  into a common heat exchanger. 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. 
     It will be recognized that the relative amount of feed found in each branch of the split LNG feed to fractionation column  21  or absorber column  21  will depend on several factors, including LNG composition, the amount of heat which can economically be extracted from the feed, and the quantity of horsepower available. More feed to the top of the column may increase recovery while increasing the duty in reboiler  25  and thereby increasing the high level utility heat requirements. Increasing feed lower in the column reduces the high level utility heat consumption but may also reduce product recovery. The relative locations of the mid-column feeds may vary depending on LNG composition or other factors such as the desired recovery level and the amount of vapor formed during heating of the feed streams. 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. 
     In the examples given for the  FIGS. 3 through 6  embodiments, recovery of C 2  components and heavier hydrocarbon components is illustrated, while recovery of C 3  components and heavier hydrocarbon components is illustrated in the examples given for the  FIGS. 7 through 10  embodiments. However, it is believed that the  FIGS. 3 through 6  embodiments are also advantageous when recovery of only C 3  components and heavier hydrocarbon components is desired, and that the  FIGS. 7 through 10  embodiments are also advantageous when recovery of C 2  components and heavier hydrocarbon components is desired. Likewise, it is believed that the  FIGS. 11 through 13  embodiments are advantageous both for recovery of C 2  components and heavier hydrocarbon components and for recovery of C 3  components and heavier hydrocarbon components. 
     The present invention provides improved recovery of C 2  components and heavier hydrocarbon components or of C 3  components and heavier hydrocarbon components per amount of utility consumption required to operate the process. An improvement in utility consumption required for operating the process may appear in the form of reduced power requirements for compression or pumping, reduced energy requirements for tower reboilers, or a combination thereof. Alternatively, the advantages of the present invention may be realized by accomplishing higher recovery levels for a given amount of utility consumption, or through some combination of higher recovery and improvement in utility consumption. 
     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.