Abstract:
A process and apparatus for the recovery of ethane, ethylene, propane, propylene, and heavier hydrocarbons from a liquefied natural gas (LNG) stream is disclosed. The LNG feed stream is divided into two portions. The first portion is supplied to a fractionation column at an upper mid-column feed point. The second portion is directed in heat exchange relation with a warmer distillation stream rising from the fractionation stages of the column, whereby this portion of the LNG feed stream is partially vaporized and the distillation stream is totally condensed. The condensed distillation stream is divided into a “lean” LNG product stream and a reflux stream, whereupon the reflux stream is supplied to the column at a top column feed position. The partially vaporized portion of the LNG feed stream is separated into vapor and liquid streams which are thereafter supplied to the column at lower mid-column feed positions. The quantities and temperatures of the feeds to the column are effective to maintain the column overhead temperature at a temperature whereby the major portion of the desired components is recovered in the bottom liquid product from the column.

Description:
CROSS REFERENCE TO RELATED APPLICATIONS  
       [0001]     The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application No. 60/810,244 which was filed on Jun. 2, 2006 and 60/812,686 which was filed on Jun. 8, 2006. 
     
    
     BACKGROUND OF THE INVENTION  
       [0002]     This invention relates to a process for the separation of ethane and heavier hydrocarbons or propane and heavier hydrocarbons from liquefied natural gas, hereinafter referred to as LNG, to provide a volatile methane-rich lean LNG stream and a less volatile natural gas liquids (NGL) or liquefied petroleum gas (LPG) stream.  
         [0003]     As an alternative to transportation in pipelines, natural gas at remote locations is sometimes liquefied and transported in special LNG tankers to appropriate LNG receiving and storage terminals. The LNG can then be re-vaporized and used as a gaseous fuel in the same fashion as natural gas. Although LNG usually has a major proportion of methane, i.e., methane comprises at least 50 mole percent of the LNG, it also contains relatively lesser amounts of heavier hydrocarbons such as ethane, propane, butanes, and the like, as well as nitrogen. It is often necessary to separate some or all of the heavier hydrocarbons from the methane in the LNG so that the gaseous fuel resulting from vaporizing the LNG conforms to pipeline specifications for heating value. In addition, it is often also desirable to separate the heavier hydrocarbons from the methane because these hydrocarbons have a higher value as liquid products (for use as petrochemical feedstocks, as an example) than their value as fuel.  
         [0004]     Although there are many processes which may be used to separate ethane and heavier hydrocarbons from LNG, these processes often must compromise between high recovery, low utility costs, and process simplicity (and hence low capital investment). U.S. Pat. Nos. 2,952,984; 3,837,172; 5,114,451; and 7,155,931 describe relevant LNG processes capable of ethane or propane recovery while producing the lean LNG as a vapor stream that is thereafter compressed to delivery pressure to enter a gas distribution network. However, lower utility costs may be possible if the lean LNG is instead produced as a liquid stream that can be pumped (rather than compressed) to the delivery pressure of the gas distribution network, with the lean LNG subsequently vaporized using a low level source of external heat or other means. U.S. Pat. Nos. 7,069,743 and 7,216,507 describe such processes.  
         [0005]     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, and also offers significant reduction in capital investment. A typical analysis of an LNG stream to be processed in accordance with this invention would be, in approximate mole percent, 89.8% methane, 6.5% ethane and other C 2  components, 2.2% propane and other C 3  components, and 1.0% butanes plus, with the balance made up of nitrogen. 
     
    
       [0006]     For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings:  
         [0007]      FIG. 1  is a flow diagram of an LNG processing plant in accordance with the present invention; and  
         [0008]      FIGS. 2, 3 , and  4  are flow diagrams illustrating alternative means of application of the present invention to an LNG processing plant. 
     
    
       [0009]     In the following explanation of the above figures, tables are provided summarizing flow rates calculated for representative process conditions. In the tables appearing herein, the values for flow rates (in moles per hour) have been rounded to the nearest whole number for convenience. The total stream rates shown in the tables include all non-hydrocarbon components and hence are generally larger than the sum of the stream flow rates for the hydrocarbon components. Temperatures indicated are approximate values rounded to the nearest degree. It should also be noted that the process design calculations performed for the purpose of comparing the processes depicted in the figures are based on the assumption of no heat leak from (or to) the surroundings to (or from) the process. The quality of commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art.  
         [0010]     For convenience, process parameters are reported in both the traditional British units and in the units of the Systeme International d&#39;Unités (SI). The molar flow rates given in the tables may be interpreted as either pound moles per hour or kilogram moles per hour. The energy consumptions reported as horsepower (HP) and/or thousand British Thermal Units per hour (MBTU/Hr) correspond to the stated molar flow rates in pound moles per hour. The energy consumptions reported as kilowatts (kW) correspond to the stated molar flow rates in kilogram moles per hour.  
       DESCRIPTION OF THE INVENTION  
     EXAMPLE 1  
       [0011]      FIG. 1  illustrates a flow diagram of a process in accordance with the present invention adapted to produce an NGL product containing the majority of the C 2  components and heavier hydrocarbon components present in the feed stream.  
         [0012]     In the simulation of the  FIG. 1  process, the LNG to be processed (stream  41 ) from LNG tank  10  enters pump  11  at −255° F. [−159° C.], which elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to separator  13 . Stream  41   a  exiting the pump is split into two portions, streams  42  and  43 . The first portion, stream  42 , is heated to −220° F. [−140° C.] (stream  42   a ) in heat exchanger  12  and then is pumped to higher pressure by pump  18 . Pumped stream  42   b  at −219° F. [−140° C.] is then supplied to fractionation column  21  at an upper mid-column feed point.  
         [0013]     The second portion of stream  41   a  (stream  43 ) is heated prior to entering separator  13  so that at least a portion of it is vaporized. In the example shown in  FIG. 1 , stream  43  is heated in heat exchanger  12  by cooling overhead vapor distillation stream  48  and reflux stream  53 . The heated stream  43   a  enters separator  13  at −171° F. [−113° C.] and 192 psia [1,324 kPa(a)] where the vapor (stream  46 ) is separated from any remaining liquid (stream  47 ). Stream  46  enters compressor  14  (driven by an external power source) and is compressed to a pressure high enough to enter fractionation tower  21 , operating at approximately 265 psia [1,825 kPa(a)]. The compressed vapor stream  46   a  is thereafter supplied as feed to fractionation column  21  at a mid-column feed point.  
         [0014]     The separator liquid (stream  47 ) is pumped to higher pressure by pump  15 , and stream  47   a  is then heated to −156° F. [−104° C.] in heat exchanger  16  by providing cooling of the liquid product from the column (stream  51 ). The partially heated stream  47   b  is then further heated to −135° F. [−93° C.] (stream  47   c ) in heat exchanger  17  using low level utility heat before it is supplied to fractionation tower  21  at a lower mid-column feed point. (High level utility heat, such as the heating medium used in tower reboiler  25 , is normally more expensive than low level utility heat, so lower operating cost is usually achieved when 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.)  
         [0015]     Note that in all cases heat exchangers  12 ,  16 , and  17  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.) Alternatively, heat exchangers  16  and/or  17  could be replaced by other heating means, such as a heater using sea water as illustrated in  FIG. 1 , a heater using a utility stream rather than a process stream (like stream  51  used in  FIG. 1 ), an indirect fired heater, or a heater using a heat transfer fluid warmed by ambient air, as warranted by the particular circumstances.  
         [0016]     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. 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 in the vapors; 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. The liquid product stream  51  exits the bottom of the tower at 40° F. [4° C.], based on a methane fraction of 0.008 on a molar basis in the bottom product. After cooling to 0° F. [−18° C.] in heat exchanger  16  as described previously, the liquid product (stream  51   a ) flows to storage or further processing.  
         [0017]     Overhead vapor distillation stream  48  is withdrawn from the upper section of fractionation tower  21  at −166° F. [−110° C.] and is totally condensed as it is cooled to −170° F. [−112° C.] in heat exchanger  12  as described previously. The condensed liquid (stream  48   a ) 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  20  to 1365 psia [9,411 kPa(a)] (stream  52   a ) for subsequent vaporization and/or transportation.  
         [0018]     The remaining portion is reflux stream  53 , which flows to heat exchanger  12  where it is subcooled to −220° F. [−140° C.] by heat exchange with the portions of the cold LNG (streams  42  and  43 ) as described previously. The subcooled reflux stream  53   a  is pumped to the operating pressure of demethanizer  21  by pump  19  and stream  53   b  at −220° F. [−140° C.] 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 .  
         [0019]     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    9,859   710   245   115   10,980       42     789    57    20    9     878       43    9,070   653   225   106   10,102       46    5,213    26    1    0    5,282       47    3,857   627   224   106    4,820       48   10,369    7    0    0   10,430       53     519    0    0    0     522       52    9,850    7    0    0    9,908       51      9   703   245   115    1,072                Recoveries*                   Ethane    98.98%           Propane   100.00%           Butanes+   100.00%           Power           LNG Booster Pump     123 HP   [  203 kW]           Reflux Pump      1 HP   [   1 kW]           Supplemental Reflux Pump      4 HP   [   7 kW]           Liquid Feed Pump     38 HP      63 kW]           Vapor Feed Compressor     453 HP   [  745 kW]           LNG Product Pump     821 HP   [1,349 kW]           Totals   1,440 HP   [2,368 kW]           Low Level Utility Heat           Liquid Feed Heater   7,890 MBTU/Hr   [5,097 kW]           High Level Utility Heat           Demethanizer Reboiler   8,450 MBTU/Hr   [5,458 kW]                         *(Based on un-rounded flow rates)             
 
         [0020]     There are four primary factors that account for the improved efficiency of the present invention. First, compared to many prior art processes, the present invention does not depend on the LNG feed itself to directly serve as the reflux for fractionation column  21 . Rather, the refrigeration inherent in the cold LNG is used in heat exchanger  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 such prior art processes. Second, compared to many 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   b  in  FIG. 1 ) serves as a supplemental reflux stream for fractionation tower  21 , providing partial rectification of the vapors in the vapor and liquid feed streams (streams  46   a  and  47   c  in  FIG. 1 ) so that heating and partially vaporizing the other portion (stream  43 ) of the LNG feed does not unduly increase the condensing load in heat exchanger  12 . Third, compared to many prior art processes, using a portion of the cold LNG feed (stream  42   b  in  FIG. 1 ) as a supplemental reflux stream allows using less top reflux (stream  53   b  in  FIG. 1 ) for fractionation tower  21 . The lower top reflux flow, plus the greater degree of heating using low level utility heat in heat exchanger  17 , 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. Fourth, compared to many prior art processes, the initial separation of the LNG into vapor and liquid fractions in separator  13  is performed at relatively low pressure. The relative volatilities between the lighter components (i.e., methane) and the desirable heavier components that are to be recovered (i.e., the C 2  and heavier components) are more favorable at lower pressure, resulting in less of the desirable components being present in stream  46   a  and subsequently requiring rectification in fractionation tower  21 .  
       EXAMPLE 2  
       [0021]     An alternative embodiment of the present invention is shown in  FIG. 2 . The LNG composition and conditions considered in the process presented in  FIG. 2  are the same as those in  FIG. 1 . Accordingly, the  FIG. 2  process of the present invention can be compared to the embodiment displayed in  FIG. 1 .  
         [0022]     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 separator  13 . Stream  41   a  exiting the pump is split into two portions, streams  42  and  43 . The first portion, stream  42 , is heated to −220° F. [−140° C.] (stream  42   a ) in heat exchanger  12  and then is pumped to higher pressure by pump  18 . Pumped stream  42   b  at −219° F. [−140° C.] is then supplied to fractionation column  21  at an upper mid-column feed point.  
         [0023]     The second portion of stream  41  a (stream  43 ) is heated prior to entering separator  13  so that at least a portion of it is vaporized. In the example shown in  FIG. 2 , stream  43  is heated in heat exchanger  12  so that heated stream  43   a  enters separator  13  at −169° F. [−112° C.] and 196 psia [1,351 kPa(a)] where the vapor (stream  46 ) is separated from any remaining liquid (stream  47 ). Stream  46  is compressed by compressor  14  to a pressure high enough to enter fractionation tower  21 , operating at approximately 265 psia [1,825 kPa(a)]. The compressed vapor stream  46   a  is then divided into two portions, streams  49  and  50 . Stream  49 , comprising about 30% of the total compressed vapor, is thereafter supplied as feed to fractionation column  21  at a mid-column feed point.  
         [0024]     The separator liquid (stream  47 ) is pumped to higher pressure by pump  15 , and stream  47   a  is then heated to −153° F. [−103° C.] in heat exchanger  16  by providing cooling of the liquid product from the column (stream  51 ). The partially heated stream  47   b  is then further heated to −135° F. [−93° C.] (stream  47   c ) in heat exchanger  17  using low level utility heat before it is supplied to fractionation tower  21  at a lower mid-column feed point. The liquid product stream  51  exits the bottom of the tower at 40° F. [4° C.], and flows to storage or further processing after cooling to 0° F. [−18° C.] (stream  51   a ) in heat exchanger  16  as described previously.  
         [0025]     Overhead vapor distillation stream  48  is withdrawn from the upper section of fractionation tower  21  at −166° F. [−110° C.] and mixes with the remaining portion of the compressed vapor (stream  50 ). The combined stream  54  at −155° F. [−104° C.] is totally condensed as it is cooled to −170° F. [−112° C.] in heat exchanger  12  as described previously. The condensed liquid (stream  54   a ) 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  20  to 1365 psia [9,411 kPa(a)] (stream  52   a ) for subsequent vaporization and/or transportation.  
         [0026]     The remaining portion is reflux stream  53 , which flows to heat exchanger  12  where it is subcooled to −220° F. [−140° C.] by heat exchange with the cold LNG (streams  42  and  43 ) as described previously. The subcooled reflux stream  53   a  is pumped to the operating pressure of demethanizer  21  by pump  19  and stream  53   b  at −220° F. [−140° C.] 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 .  
         [0027]     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    9,859   710   245   115   10,980       42     789    57    20    9     878       43    9,070   653   225   106   10,102       46    5,622    31    1    0    5,698       47    3,448   622   224   106    4,404       49    1,687    10    0    0    1,710       50    3,935    21    1    0    3,988       48    6,434    2    0    0    6,458       54   10,369    23    1    0   10,446       53     518    1    0    0     522       52    9,851    22    1    0    9,924       51      8   688   244   115    1,056                Recoveries*                   Ethane   96.82%           Propane   99.76%           Butanes+   99.97%           Power           LNG Booster Pump     126 HP   [  207 kW]           Reflux Pump      1 HP   [   1 kW]           Supplemental Reflux Pump      4 HP   [   7 kW]           Liquid Feed Pump     34 HP   [  56 kW]           Vapor Compressor     462 HP   [  759 kW]           LNG Product Pump     822 HP   [1,351 kW]           Totals   1,449 HP   [2,381 kW]           Low Level Utility Heat           Liquid Feed Heater   6,519 MBTU/Hr   [4,211 kW]           High Level Utility Heat           Demethanizer Reboiler   9,737 MBTU/Hr   [6,290 kW]                         *(Based on un-rounded flow rates)             
 
         [0028]     Comparing Table II above for the  FIG. 2  embodiment of the present invention with Table I for the  FIG. 1  embodiment of the present invention shows that the liquids recovery is slightly lower for the  FIG. 2  embodiment since a significant portion of the LNG feed (stream  50 ) is not subjected to any rectification. As a result, the size of fractionation tower  21  can be significantly smaller for the  FIG. 2  embodiment, since the vapor load in the tower (represented by overhead vapor stream  48 ) is so much lower. The resulting reduction in the capital cost of the plant may justify the slightly lower liquid recovery provided by this embodiment of the present invention.  
       EXAMPLE 3  
       [0029]     Another alternative embodiment of the present invention is shown in  FIG. 3 . 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 of the present invention can be compared to the embodiments displayed in  FIGS. 1 and 2 .  
         [0030]     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  13 . Stream  41   a  exiting the pump is split into two portions, streams  42  and  43 . The first portion, stream  42 , is heated to −220° F. [−140° C.] (stream  42   a ) in heat exchanger  12  and then is pumped to higher pressure by pump  18 . Pumped stream  42   b  at −219° F. [−140° C.] is then supplied to fractionation column  21  at an upper mid-column feed point.  
         [0031]     The second portion of stream  41   a  (stream  43 ) is heated prior to entering separator  13  so that at least a portion of it is vaporized. In the example shown in  FIG. 3 , stream  43  is heated in heat exchanger  12  so that heated stream  43   a  enters separator  13  at −168° F. [−111° C.] and 198 psia [1,365 kPa(a)] where the vapor (stream  46 ) is separated from any remaining liquid (stream  47 ). Stream  47  is pumped to higher pressure by pump  15 , and stream  47   a  is then heated to −152° F. [−102° C.] in heat exchanger  16  by providing cooling of the liquid product from the column (stream  51 ). The partially heated stream  47   b  is then further heated to −135° F. [−93° C.] (stream  47   c ) in heat exchanger  17  using low level utility heat before it is supplied to fractionation tower  21  at a lower mid-column feed point. The liquid product stream  51  exits the bottom of the tower at 40° F. [5° C.], and flows to storage or further processing after cooling to 0° F. [−18° C.] (stream  51   a ) in heat exchanger  16  as described previously.  
         [0032]     Overhead vapor distillation stream  48  is withdrawn from the upper section of fractionation tower  21  at −166° F. [−110° C.]. The vapor from separator  13  (stream  46 ) enters compressor  14  and is compressed to higher pressure, allowing stream  46   a  to mix with stream  48  to form stream  54 . The combined stream  54  at −150° F. [−101° C.] is totally condensed as it is cooled to −169° F. [−112° C.] in heat exchanger  12  as described previously. The condensed liquid (stream  54   a ) 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  20  to 1365 psia [9,411 kPa(a)] (stream  52   a ) for subsequent vaporization and/or transportation.  
         [0033]     The remaining portion is reflux stream  53 , which flows to heat exchanger  12  where it is subcooled to −220° F. [−140° C.] by heat exchange with the cold LNG (streams  42  and  43 ) as described previously. The subcooled reflux stream  53   a  is pumped to the operating pressure of demethanizer  21  by pump  19  and stream  53   b  at −220° F. [−140° C.] 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 .  
         [0034]     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    9,859   710   245   115   10,980       42     789    57    20    9     878       43    9,070   653   225   106   10,102       46    5,742    34    1    0    5,819       47    3,328   619   224   106    4,283       48    4,627    1    0    0    4,639       54   10,369    35    1    0   10,458       53     518    2    0    0     523       52    9,851    33    1    0    9,935       51      8   677   244   115    1,045                Recoveries*                   Ethane   95.37%           Propane   99.63%           Butanes+   99.96%           Power           LNG Booster Pump     127 HP   [  209 kW]           Reflux Pump      1 HP   [   1 kW]           Supplemental Reflux Pump      4 HP   [   7 kW]           Liquid Feed Pump      32 HP   [  53 kW]           Vapor Compressor     457 HP   [  751 kW]           LNG Product Pump     826 HP   [1,358 kW]           Totals    1,447 HP   [2,379 kW]           Low Level Utility Heat           Liquid Feed Heater    6,109 MBTU/Hr   [3,946 kW]           High Level Utility Heat           Demethanizer Reboiler   10,350 MBTU/Hr   [6,686 kW]                         *(Based on un-rounded flow rates)             
 
         [0035]     Comparing Table III above for the  FIG. 3  embodiment of the present invention with Tables I and II for the  FIGS. 1 and 2 , respectively, embodiments of the present invention shows that the liquids recovery is somewhat lower for the  FIG. 3  embodiment since still more of the LNG feed (all of the compressed separator vapor, stream  46   a ) is not subjected to any rectification. Accordingly, the size of fractionation tower  21  can be still smaller for the  FIG. 3  embodiment, since the vapor load in the tower (represented by overhead vapor stream  48 ) is even lower. Thus, the capital cost of the  FIG. 3  embodiment of the present invention will likely be lower than either the  FIG. 1  or the  FIG. 2  embodiment. The choice of which embodiment to use for a particular application will generally be dictated by the relative value of the heavier hydrocarbon components, the relative costs of power and high level utility heat, and the relative capital costs of fractionation towers, pumps, heat exchangers, and compressors.  
       EXAMPLE 4  
       [0036]     Another 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  FIGS. 1 through 3 . Accordingly, the  FIG. 4  process of the present invention can be compared to the embodiments displayed in  FIGS. 1 through 3 .  
         [0037]     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 exchange and thence to separator  13  and fractionation column  21 . Stream  41   a  exiting the pump is split into two portions, streams  42  and  43 . The first portion, stream  42 , is heated to −165° F. [−109° C.] (stream  42   a ) in heat exchanger  12  and then is supplied to fractionation column  21  at an upper mid-column feed point. Depending on the discharge pressure of pump  11 , a valve  30  may be needed to reduce the pressure of stream  42   b  to that of fractionation column  21 .  
         [0038]     The second portion of stream  41   a  (stream  43 ) is heated prior to entering separator  13  so that at least a portion of it is vaporized. In the example shown in  FIG. 4 , stream  43  is heated in heat exchanger  12  so that heated stream  43   a  enters separator  13  at −168° F. [−111° C.] and 195 psia [1,342 kPa(a)] where the vapor (stream  46 ) is separated from the remaining liquid (stream  47 ). Stream  47  is pumped to higher pressure by pump  15 , and stream  47   a  is then heated to −155° F. [−104° C.] in heat exchanger  16  by providing cooling of the liquid product from the column (stream  51 ). The partially heated stream  47   b  is then further heated so that a portion of it is vaporized. In the example of  FIG. 4 , steam  47   b  is further heated in heat exchanger  17  using low level utility heat so that the further heated stream  47   c  enters separator  26  at 9° F. [−13° C.] and 750 psia [5,169 kPa where vapor stream  55  is separated from any remaining liquid stream  56 . The separator liquid stream (stream  56 ) is expanded to the operating pressure (approximately 195 psia [1,342 kPa(a)]) of fractionation column  21  by expansion valve  23 , cooling stream  56   a  to −36° F. [−38° C.] before it is supplied to fractionation column  21  at a lower mid-column feed point.  
         [0039]     The vapor from separator  26  (stream  55 ) enters a work expansion machine  27  in which mechanical energy is extracted from this portion of the higher pressure feed. The machine  27  expands the vapor substantially isentropically to the tower operating pressure with the work expansion cooling the expanded stream  55   a  to a temperature of −74° F. [−59° C.]. This partially condensed expanded stream  55   a  is thereafter supplied as feed to fractionation column  21  at a mid-column feed point.  
         [0040]     The liquid product stream  51  exits the bottom of the tower at 17° F. [−9° C.]. After cooling to 0° F. [−18° C.] in heat exchanger  16  as described previously, the liquid product stream  51   a  flows to storage or further processing.  
         [0041]     Overhead vapor distillation stream  48  is withdrawn from the upper section of fractionation tower  21  at −178° F. [−117° C.]. The vapor from separator  13  (stream  46 ) mixes with stream  48  to form stream  54 . The combined stream  54  at −174° F. [−114° C.] flows to compressor  28  driven by expansion machine  27 , where it is compressed to 266 psia [1,835 kPa(a)] (stream  54   a ). Stream  54   a  is totally condensed as it is cooled to −168° F. [−111° C.] in heat exchanger  12  as described previously. The condensed liquid (stream  54   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  20  to 1365 psia [9,411 kPa(a)] (stream  52   a ) for subsequent vaporization and/or transportation.  
         [0042]     The remaining portion is reflux stream  53 , which flows to heat exchanger  12  where it is subcooled to −225° F. [−143° C.] by heat exchange with the cold LNG (streams  42  and  43 ) as described previously. The subcooled reflux stream  53   a  is expanded to the operating pressure of demethanizer  21  in valve  31  and the expanded stream  53   b  at −225° F. [−143° C.] 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 .  
         [0043]     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    9,859   710   245   115   10,980       42    2,465   177    61    29    2,745       43    7,394   533   184    86    8,235       46    4,812    29    1    0    4,877       47    2,582   504   183    86    3,358       55    2,503   445   133    44    3,128       56      79    59    50    42     230       48    6,132    9    0    0    6,163       54   10,944    38    1    0   11,040       53    1,093    4    0    0    1,104       52    9,851    34    1    0    9,936       51      8   676   244   115    1,044                Recoveries*                   Ethane   95.21%           Propane   99.71%           Butanes+   99.96%           Power           LNG Booster Pump     159 HP   [  261 kW]           Liquid Feed Pump     143 HP   [  235 kW]           LNG Product Pump     826 HP   [1,358 kW]           Totals    1,128 HP   [1,854 kW]           Low Level Utility Heat           Liquid Feed Heater   14,410 MBTU/Hr   [9,308 kW]           High Level Utility Heat           Demethanizer Reboiler    2,945 MBTU/Hr   [1,902 kW]                         *(Based on un-rounded flow rates)             
 
         [0044]     Comparing Table IV above for the  FIG. 4  embodiment of the present invention with Table III for the  FIG. 3  embodiment shows that the liquids recovery is essentially the same for this  FIG. 4  embodiment, but now the Vapor Compressor has been eliminated in favor of additional liquid pumping. Because pumping is more efficient than compression, this results in a net decrease in total power consumption of approximately 22% compared to the  FIGS. 1 through 3  embodiments. The  FIG. 4  embodiment is also able to use more low level utility heat and thereby reduce the use of high level utility heat compared to the  FIGS. 1 through 3  embodiments. The high level utility heat requirement of the  FIG. 4  embodiment is only 28% to 35% of that required by the  FIGS. 1 through 3  embodiments.  
         [0045]     The size of fractionation tower  21  is somewhat larger than the  FIG. 3  embodiment, since the vapor load in the tower (represented by overhead vapor stream  48 ) is somewhat higher. However, the capital cost of this  FIG. 4  embodiment of the present invention will likely be lower than the  FIG. 3  embodiment because of the elimination of the vapor compression service. The choice of which embodiment to use for a particular application will generally be dictated by the relative value of the heavier hydrocarbon components, the relative costs of power and high level utility heat, and the relative capital costs of fractionation towers, pumps, heat exchangers, and compressors.  
       Other Embodiments  
       [0046]     Some circumstances may favor subcooling reflux stream  53  with another process stream, rather than using the cold LNG streams that enter heat exchanger  12 . Other circumstances may favor no subcooling at all. The decision regarding whether or not to subcool reflux stream  53  before it is fed to the column will depend on many factors, including the LNG composition, the desired recovery level, etc. As shown by the dashed lines in  FIGS. 1 through 4 , stream  53  can be routed to heat exchanger  12  if subcooling is desired, but it need not be if no subcooling is desired. Likewise, heating of supplemental reflux stream  42  before it is fed to the column must be evaluated for each application. As shown by the dashed lines in  FIGS. 1 through 4 , stream  42  need not be routed to heat exchanger  12  if no heating is desired.  
         [0047]     When the LNG to be processed is leaner or when complete vaporization of the LNG in heat exchanger  17  is contemplated, separator  26  in  FIG. 4  may not be justified. Depending on the quantity of heavier hydrocarbons in the inlet LNG and the pressure of the LNG stream leaving liquid feed pump  15 , the heated LNG stream leaving heat exchanger  17  may not contain any liquid (because it is above its dewpoint, or because it is above its cricondenbar). In such cases, separator  26  may be eliminated as shown by the dashed lines.  
         [0048]     In the examples shown, total condensation of stream  48   a  in  FIG. 1 , stream  54   a  in  FIGS. 2 and 3 , and stream  54   b  in  FIG. 4  is shown. Some circumstances may favor subcooling these streams, while other circumstances may favor only partial condensation. Should partial condensation of these 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.  
         [0049]     Depending on the composition of the LNG to be processed, it may be possible to operate separator  13  at a sufficiently high pressure that compressor  14  ( FIGS. 1 through 3 ) and pump  15  ( FIGS. 1 through 4 ) are not needed to supply the vapor (stream  46 ) and liquid (stream  47 ) to fractionation tower  21 . Should the relative volatilities in separator  13  be favorable enough to allow achieving the desired recovery level with the separator pressure higher than that of the tower, compressor  14  ( FIGS. 1 through 3 ) and pump  15  ( FIGS. 1 through 4 ) may be eliminated as shown by the dashed lines.  
         [0050]     In  FIGS. 1 through 4 , 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  and  16  in  FIGS. 1 through 4  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.  
         [0051]     In  FIGS. 1 through 3 , individual pumps have been shown for the reflux pumping requirements (pumps  18  and  19 ). However, it is possible to achieve the pumping indicated by pump  19  with pump  20  alone and to achieve the pumping indicated by pump  18  with pump  11  alone at some increase in overall pumping power. If pump  19  is deleted in favor of additional pumping by pump  20 , stream  53  is taken from the discharge stream from pump  20  as shown by the dashed line. In that case, pump  19  is eliminated as shown by it being dashed in  FIGS. 1 through 3 . If pump  18  is deleted in favor of additional pumping by pump  11 , the discharge pressure from pump  11  will be higher than that shown in each of the  FIGS. 1 through 3  embodiments and an appropriate pressure reduction valve (such as dashed valve  22 ) may be required so as to maintain the operating pressure in separator  13  at the desired level. In that case, pump  18  is eliminated as shown by it being dashed in  FIGS. 1 through 3 .  
         [0052]     In  FIG. 4 , it may also be possible to further reduce pumping requirements by addition of one or more pumping services. For example, it may be possible to reduce the discharge pressure of pump  11  by adding a pump in line  42   a  that would pump that stream individually to fractionation column  21  and reduce the pressure drop taken in valve  22  in stream  43  upstream of heat exchanger  12 . The decision as to whether to combine pumping services or use more than one pump for an indicated service will depend on a number of factors including, but not limited to, LNG flow rate, stream temperatures, etc.  
         [0053]     It will be recognized that the relative amount of feed found in each branch of the split LNG feed to fractionation 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.  
         [0054]     In the examples given for the  FIGS. 1 through 4  embodiments, recovery of C 2  components and heavier hydrocarbon components is illustrated. However, it is believed that the  FIGS. 1 through 4  embodiments are also advantageous when recovery of only C 3  components and heavier hydrocarbon components is desired.  
         [0055]     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.