Patent Document

BACKGROUND OF THE INVENTION 
     This invention relates to a process for processing natural gas or other methane-rich gas streams to produce a liquefied natural gas (LNG) stream that has a high methane purity and a liquid stream containing predominantly hydrocarbons heavier than methane. 
     Natural gas is typically recovered from wells drilled into underground reservoirs. It usually has a major proportion of methane, i.e., methane comprises at least 50 mole percent of the gas. Depending on the particular underground reservoir, the natural gas also contains relatively lesser amounts of heavier hydrocarbons such as ethane, propane, butanes, pentanes and the like, as well as water, hydrogen, nitrogen, carbon dioxide, and other gases. 
     Most natural gas is handled in gaseous form. The most common means for transporting natural gas from the wellhead to gas processing plants and thence to the natural gas consumers is in high pressure gas transmission pipelines. In a number of circumstances, however, it has been found necessary and/or desirable to liquefy the natural gas either for transport or for use. In remote locations, for instance, there is often no pipeline infrastructure that would allow for convenient transportation of the natural gas to market. In such cases, the much lower specific volume of LNG relative to natural gas in the gaseous state can greatly reduce transportation costs by allowing delivery of the LNG using cargo ships and transport trucks. 
     Another circumstance that favors the liquefaction of natural gas is for its use as a motor vehicle fuel. In large metropolitan areas, there are fleets of buses, taxi cabs, and trucks that could be powered by LNG if there were an economic source of LNG available. Such LNG-fueled vehicles produce considerably less air pollution due to the clean-burning nature of natural gas when compared to similar vehicles powered by gasoline and diesel engines which combust higher molecular weight hydrocarbons. In addition, if the LNG is of high purity (i.e., with a methane purity of 95 mole percent or higher), the amount of carbon dioxide (a “greenhouse gas”) produced is considerably less due to the lower carbon:hydrogen ratio for methane compared to all other hydrocarbon fuels. 
     The present invention is generally concerned with the liquefaction of natural gas while producing as a co-product a liquid stream consisting primarily of hydrocarbons heavier than methane, such as natural gas liquids (NGL) composed of ethane, propane, butanes, and heavier hydrocarbon components, liquefied petroleum gas (LPG) composed of propane, butanes, and heavier hydrocarbon components, or condensate composed of butanes and heavier hydrocarbon components. Producing the co-product liquid stream has two important benefits: the LNG produced has a high methane purity, and the co-product liquid is a valuable product that may be used for many other purposes. A typical analysis of a natural gas stream to be processed in accordance with this invention would be, in approximate mole percent, 84.2% methane, 7.9% ethane and other C 2  components, 4.9% propane and other C 3  components, 1.0% iso-butane, 1.1% normal butane, 0.8% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present. 
     There are a number of methods known for liquefying natural gas. For instance, see Finn, Adrian J., Grant L. Johnson, and Terry R. Tomlinson, “LNG Technology for Offshore and Mid-Scale Plants”, Proceedings of the Seventy-Ninth Annual Convention of the Gas Processors Association, pp. 429-450, Atlanta, Ga., Mar. 13-15, 2000 and Kikkawa, Yoshitsugi, Masaaki Ohishi, and Noriyoshi Nozawa, “Optimize the Power System of Baseload LNG Plant”, Proceedings of the Eightieth Annual Convention of the Gas Processors Association, San Antonio, Tex., Mar. 12-14, 2001 for surveys of a number of such processes. U.S. Pat. Nos. 4,445,917; 4,525,185; 4,545,795; 4,755,200; 5,291,736; 5,363,655; 5,365,740; 5,600,969; 5,615,561; 5,651,269; 5,755,114; 5,893,274; 6,014,869; 6,062,041; 6,119,479; 6,125,653; 6,250,105 B1; 6,269,655 B1; 6,272,882 B1; 6,308,531 B1; 6,324,867 B1; 6,347,532 B1; and our co-pending U.S. patent application Ser. No. 10/161,780 filed Jun. 4, 2002 also describe relevant processes. These methods generally include steps in which the natural gas is purified (by removing water and troublesome compounds such as carbon dioxide and sulfur compounds), cooled, condensed, and expanded. Cooling and condensation of the natural gas can be accomplished in many different manners. “Cascade refrigeration” employs heat exchange of the natural gas with several refrigerants having successively lower boiling points, such as propane, ethane, and methane. As an alternative, this heat exchange can be accomplished using a single refrigerant by evaporating the refrigerant at several different pressure levels. “Multi-component refrigeration” employs heat exchange of the natural gas with one or more refrigerant fluids composed of several refrigerant components in lieu of multiple single-component refrigerants. Expansion of the natural gas can be accomplished both isenthalpically (using Joule-Thomson expansion, for instance) and isentropically (using a work-expansion turbine, for instance). 
     Regardless of the method used to liquefy the natural gas stream, it is common to require removal of a significant fraction of the hydrocarbons heavier than methane before the methane-rich stream is liquefied. The reasons for this hydrocarbon removal step are numerous, including the need to control the heating value of the LNG stream, and the value of these heavier hydrocarbon components as products in their own right. Unfortunately, little attention has been focused heretofore on the efficiency of the hydrocarbon removal step. 
     In accordance with the present invention, it has been found that careful integration of the hydrocarbon removal step into the LNG liquefaction process can produce both LNG and a separate heavier hydrocarbon liquid product using significantly less energy than prior art processes. The present invention, although applicable at lower pressures, is particularly advantageous when processing feed gases in the range of 400 to 1500 psia [2,758 to 10,342 kPa(a)] or higher. 
    
    
     
       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 diagram of a natural gas liquefaction plant adapted for co-production of LPG in accordance with the present invention; 
         FIGS. 2 and 3  are diagrams of alternative fractionation systems which may be employed in the process of the present invention; 
         FIG. 4  is a pressure-enthalpy phase diagram for methane used to illustrate the advantages of the present invention over prior art processes; and 
         FIGS. 5 ,  6 ,  7 ,  8 ,  9 , and  10  are flow diagrams of alternative natural gas liquefaction plants adapted for co-production of a liquid stream in accordance with the present invention. 
     
    
    
     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 International System of Units (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. The production rates reported as pounds per hour (Lb/Hr) correspond to the stated molar flow rates in pound moles per hour. The production rates reported as kilograms per hour (kg/Hr) correspond to the stated molar flow rates in kilogram moles per hour. 
     DESCRIPTION OF THE INVENTION 
     Referring now to  FIG. 1 , we begin with an illustration of a process in accordance with the present invention where it is desired to produce an LPG co-product containing the majority of the propane and heavier components in the natural gas feed stream. In this simulation of the present invention, inlet gas enters the plant at 90° F. [32° C.] and 1285 psia [8,860 kPa(a)] as stream  31 . If the inlet gas contains a concentration of carbon dioxide and/or sulfur compounds which would prevent the product streams from meeting specifications, these compounds are removed by appropriate pretreatment of the feed gas (not illustrated). In addition, the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccant has typically been used for this purpose. 
     The feed stream  31  is cooled in heat exchanger  10  by heat exchange with refrigerant streams and flashed separator liquids at −14° F. [−26° C.] (stream  40   a ). Note that in all cases heat exchanger  10  is representative of either a multitude of individual heat exchangers or a single multi-pass heat exchanger, or any combination thereof. (The decision as to whether to use more than one heat exchanger for the indicated cooling services will depend on a number of factors including, but not limited to, inlet gas flow rate, heat exchanger size, stream temperatures, etc.) The cooled stream  31   a  enters separator  11  at 23° F. [−5° C.] and 1278 psia [8,812 kPa(a)] where the vapor (stream  32 ) is separated from the condensed liquid (stream  33 ). 
     The vapor (stream  32 ) from separator  11  is divided into two streams,  34  and  36 , with stream  34  containing about 42% of the total vapor. Some circumstances may favor combining stream  34  with some portion of the condensed liquid (stream  39 ) to form stream  35 , but in this simulation there is no flow in stream  39 . Combined stream  35  passes through heat exchanger  13  in heat exchange relation with refrigerant stream  71   e , resulting in cooling and substantial condensation of stream  35   a . The substantially condensed stream  35   a  at −90° F. [−68° C.] is then flash expanded through an appropriate expansion device, such as expansion valve  14 , to slightly above the operating pressure (approximately 450 psia [3,103 kPa(a)]) of fractionation tower  19 . During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in  FIG. 1 , the expanded stream  35   b  leaving expansion valve  14  reaches a temperature of −123° F. [−86° C.]. The expanded stream  35   b  is warmed to −78° F. [−61° C.] and further vaporized in heat exchanger  21  as it provides cooling and partial condensation of vapor distillation stream  37  rising from the fractionation stages of fractionation tower  19 . The warmed stream  35   c  is then supplied at an upper mid-point feed position in deethanizing section  19   b  of fractionation tower  19 . 
     The remaining 58% of the vapor from separator  11  (stream  36 ) enters a work expansion machine  15  in which mechanical energy is extracted from this portion of the high pressure feed. The machine  15  expands the vapor substantially isentropically from a pressure of about 1278 psia [8,812 kPa(a)] to the tower operating pressure, with the work expansion cooling the expanded stream  36   a  to a temperature of approximately −57° F. [−49° C.]. The typical commercially available expanders are capable of recovering on the order of 80-85% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item  16 ) that can be used to re-compress the tower overhead gas (stream  49 ), for example. The expanded and partially condensed stream  36   a  is supplied as feed to distillation column  19  at a lower mid-column feed point. Stream  40 , the remaining portion of the separator liquid (stream  33 ) is flash expanded to slightly above the operating pressure of deethanizer  19  by expansion valve  12 , cooling stream  40  to −14° F. [−26° C.] (stream  40   a ) before it provides cooling to the incoming feed gas as described earlier. Stream  40   b , now at 75° F. [24° C.], then enters deethanizer  19  at a second lower mid-column feed point. 
     The deethanizer in fractionation tower  19  is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. As is often the case in natural gas processing plants, the fractionation tower may consist of two sections. The upper section  19   a  is a separator wherein the top feed is divided into its respective vapor and liquid portions, and wherein the vapor rising from the lower distillation or deethanizing section  19   b  is combined with the vapor portion (if any) of the top feed to form the deethanizer overhead vapor (stream  37 ) which exits the top of the tower. The lower, deethanizing section  19   b  contains the trays and/or packing and provides the necessary contact between the liquids falling downward and the vapors rising upward. The deethanizing section also includes one or more reboilers (such as reboiler  20 ) 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  41  exits the bottom of the tower at 213° F. [101° C.], based on a typical specification of an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product. 
     The overhead distillation stream  37  leaves deethanizer  19  at −73° F. [−59° C.] and is cooled and partially condensed in reflux condenser  21  as described earlier. The partially condensed stream  37   a  enters reflux drum  22  at −94° F. [−70° C.] where the condensed liquid (stream  44 ) is separated from the uncondensed vapor (stream  43 ). The condensed liquid (stream  44 ) is pumped by pump  23  to a top feed point on deethanizer  19  as reflux stream  44   a.    
     When the deethanizing section forms the lower portion of a fractionation tower, reflux condenser  21  may be located inside the tower above column  19  as shown in FIG.  2 . This eliminates the need for reflux drum  22  and reflux pump  23  because the distillation stream is then both cooled and separated in the tower above the fractionation stages of the column. Alternatively, use of a dephlegmator (such as dephlegmator  21  in  FIG. 3 ) in place of reflux condenser  21  in  FIG. 1  eliminates the reflux drum and reflux pump and also provides concurrent fractionation stages to replace those in the upper section of the deethanizer column. If the dephlegmator is positioned in a plant at grade level, it is connected to a vapor/liquid separator and the liquid collected in the separator is pumped to the top of the distillation column. The decision as to whether to include the reflux condenser inside the column or to use a dephlegmator usually depends on plant side and heat exchanger surface requirements. 
     The uncondensed vapor (stream  43 ) from reflux drum  22  is warmed to 93° F. [34° C.] in heat exchanger  24 , and a portion (stream  48 ) is then withdrawn to serve as fuel gas for the plant. (The amount of fuel gas that must be withdrawn is largely determined by the fuel required for the engines and/or turbines driving the gas compressors in the plant, such as refrigerant compressors  64 ,  66 , and  68  in this example.) The remainder of the warmed vapor (stream  49 ) is compressed by compressor  16  driven by expansion machines  15 ,  61 , and  63 . After cooling to 100° F. [38° C.] in discharge cooler  25 , stream  49   b  is further cooled to −83° F. [−64° C.] in heat exchanger  24  by cross exchange with the cold vapor, stream  43 . 
     Stream  49   c  then enters heat exchanger  60  and is further cooled by refrigerant stream  71   d  to −255° F. [−160° C.] to condense and subcool it, whereupon it enters a work expansion machine  61  in which mechanical energy is extracted from the stream. The machine  61  expands liquid stream  49   d  substantially isentropically from a pressure of about 593 psia [4,085 kPa(a)] to the LNG storage pressure (15.5 psia [107 kPa(a)]), slightly above atmospheric pressure. The work expansion cools the expanded stream  49   e  to a temperature of approximately −256° F. [−160° C.], whereupon it is then directed to the LNG storage tank  62  which holds the LNG product (stream  50 ). 
     All of the cooling for streams  35  and  49   c  is provided by a closed cycle refrigeration loop. The working fluid for this cycle is a mixture of hydrocarbons and nitrogen, with the composition of the mixture adjusted as needed to provide the required refrigerant temperature while condensing at a reasonable pressure using the available cooling medium. In this case, condensing with cooling water has been assumed, so a refrigerant mixture composed of nitrogen, methane, ethane, propane, and heavier hydrocarbons is used in the simulation of the  FIG. 1  process. The composition of the stream, in approximate mole percent, is 8.7% nitrogen, 31.7% methane, 47.0% ethane, and 8.6% propane, with the balance made up of heavier hydrocarbons. 
     The refrigerant stream  71  leaves discharge cooler  69  at 100° F. [38° C.] and 607 psia [4,185 kPa(a)]. It enters heat exchanger  10  and is cooled to −34° F. [−37° C.] and partially condensed by the partially warmed expanded refrigerant stream  71   f  and by other refrigerant streams. For the  FIG. 1  simulation, it has been assumed that these other refrigerant streams are commercial-quality propane refrigerant at three different temperature and pressure levels. The partially condensed refrigerant stream  71   a  then enters heat exchanger  13  for further cooling to −90° F. [−68° C.] by partially warmed expanded refrigerant stream  71   e , further condensing the refrigerant (stream  71   b ). The refrigerant is condensed and then subcooled to −255° F. [−160° C.] in heat exchanger  60  by expanded refrigerant stream  71   d . The subcooled liquid stream  71   c  enters a work expansion machine  63  in which mechanical energy is extracted from the stream as it is expanded substantially isentropically from a pressure of about 586 psia [4,040 kPa(a)] to about 34 psia [234 kPa(a)]. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream to −264° F. [−164° C.] (stream  71   d ). The expanded stream  71   d  then reenters heat exchangers  60 ,  13 , and  10  where it provides cooling to stream  49   c , stream  35 , and the refrigerant (streams  71 ,  71   a , and  71   b ) as it is vaporized and superheated. 
     The superheated refrigerant vapor (stream  71   g ) leaves heat exchanger  10  at 90° F. [32° C.] and is compressed in three stages to 617 psia [4,254 kPa(a)]. Each of the three compression stages (refrigerant compressors  64 ,  66 , and  68 ) is driven by a supplemental power source and is followed by a cooler (discharge coolers  65 ,  67 , and  69 ) to remove the heat of compression. The compressed stream  71  from discharge cooler  69  returns to heat exchanger  10  to complete the cycle. 
     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 
               
               
                   
               
             
          
           
               
                 31 
                 40,977 
                 3,861 
                 2,408 
                 1,404 
                 48,656 
               
               
                 32 
                 40,193 
                 3,667 
                 2,171 
                 1,087 
                 47,123 
               
               
                 33 
                 784 
                 194 
                 237 
                 317 
                 1,533 
               
               
                 34 
                 16,680 
                 1,522 
                 901 
                 451 
                 19,556 
               
               
                 36 
                 23,513 
                 2,145 
                 1,270 
                 636 
                 27,567 
               
               
                 37 
                 44,843 
                 7,065 
                 120 
                 0 
                 52,035 
               
               
                 40 
                 784 
                 194 
                 237 
                 317 
                 1,533 
               
               
                 41 
                 0 
                 48 
                 2,385 
                 1,404 
                 3,837 
               
               
                 43 
                 40,977 
                 3,813 
                 23 
                 0 
                 44,819 
               
               
                 44 
                 3,866 
                 3,252 
                 97 
                 0 
                 7,216 
               
               
                 48 
                 2,527 
                 235 
                 1 
                 0 
                 2,765 
               
               
                 50 
                 38,450 
                 3,578 
                 22 
                 0 
                 42,054 
               
               
                   
               
             
          
         
       
     
     
       
         
               
               
               
               
             
           
               
                   
               
             
             
               
                 Recoveries in LPG* 
                   
                   
                   
               
               
                 Propane 
                 99.05% 
               
               
                 Butanes+ 
                 100.00% 
               
               
                 Production Rate 
                 197,031 
                 Lb/Hr 
                 [197,031 kg/Hr] 
               
               
                 LNG Product 
               
               
                 Production Rate 
                 725,522 
                 Lb/Hr 
                 [725,522 kg/Hr] 
               
               
                 Purity* 
                 91.43% 
               
               
                 Lower Heating Value 
                 970.4 
                 BTU/SCF 
                 [36.16 MJ/m 3 ] 
               
               
                 Power 
               
               
                 Refrigerant Compression 
                 90,714 
                 HP 
                 [149,132 kW] 
               
               
                 Propane Compression 
                 36,493 
                 HP 
                 [59,994 kW] 
               
               
                 Total Compression 
                 127,207 
                 HP 
                 [209,126 kW] 
               
               
                 Utility Heat 
               
               
                 Demethanizer Reboiler 
                 58,003 
                 MBTU/Hr 
                 [37,470 kW] 
               
               
                   
               
               
                 *(Based on un-rounded flow rates)  
               
             
          
         
       
     
     The efficiency of LNG production processes is typically compared using the “specific power consumption” required, which is the ratio of the total refrigeration compression power to the total liquid production rate. Published information on the specific power consumption for prior art processes for producing LNG indicates a range of 0.168 HP-Hr/Lb [0.276 kW-Hr/kg] to 0.182 HP-Hr/Lb [0.300 kW-Hr/kg], which is believed to be based on an on-stream factor of 340 days per year for the LNG production plant. On this same basis, the specific power consumption for the  FIG. 1  embodiment of the present invention is 0.148 HP-Hr/Lb [0.243 kW-Hr/kg], which gives an efficiency improvement of 14-23% over the prior art processes. 
     There are two primary factors that account for the improved efficiency of the present invention. The first factor can be understood by examining the thermodynamics of the liquefaction process when applied to a high pressure gas stream such as that considered in this example. Since the primary constituent of this stream is methane, the thermodynamic properties of methane can be used for the purposes of comparing the liquefaction cycle employed in the prior art processes versus the cycle used in the present invention.  FIG. 4  contains a pressure-enthalpy phase diagram for methane. In most of the prior art liquefaction cycles, all cooling of the gas stream is accomplished while the stream is at high pressure (path A-B), whereupon the stream is then expanded (path B-C) to the pressure of the LNG storage vessel (slightly above atmospheric pressure). This expansion step may employ a work expansion machine, which is typically capable of recovering on the order of 75-80% of the work theoretically available in an ideal isentropic expansion. In the interest of simplicity, fully isentropic expansion is displayed in  FIG. 4  for path B-C. Even so, the enthalpy reduction provided by this work expansion is quite small, because the lines of constant entropy are nearly vertical in the liquid region of the phase diagram. 
     Contrast this now with the liquefaction cycle of the present invention. After partial cooling at high pressure (path A-A′), the gas stream is work expanded (path A′-A″) to an intermediate pressure. (Again, fully isentropic expansion is displayed in the interest of simplicity.) The remainder of the cooling is accomplished at the intermediate pressure (path A″-B′), and the stream is then expanded (path B′-C) to the pressure of the LNG storage vessel. Since the lines of constant entropy slope less steeply in the vapor region of the phase diagram, a significantly larger enthalpy reduction is provided by the first work expansion step (path A′-A″) of the present invention. Thus, the total amount of cooling required for the present invention (the sum of paths A-A′ and A″-B′) is less than the cooling required for the prior art processes (path A-B), reducing the refrigeration (and hence the refrigeration compression) required to liquefy the gas stream. 
     The second factor accounting for the improved efficiency of the present invention is the superior performance of hydrocarbon distillation systems at lower operating pressures. The hydrocarbon removal step in most of the prior art processes is performed at high pressure, typically using a scrub column that employs a cold hydrocarbon liquid as the absorbent stream to remove the heavier hydrocarbons from the incoming gas stream. Operating the scrub column at high pressure is not very efficient, as it results in the co-absorption of a significant fraction of the methane and ethane from the gas stream, which must subsequently be stripped from the absorbent liquid and cooled to become part of the LNG product. In the present invention, the hydrocarbon removal step is conducted at the intermediate pressure where the vapor-liquid equilibrium is much more favorable, resulting in very efficient recovery of the desired heavier hydrocarbons in the co-product liquid stream. 
     Other Embodiments 
     One skilled in the art will recognize that the present invention can be adapted for use with all types of LNG liquefaction plants to allow co-production of an NGL stream, an LPG stream, or a condensate stream, as best suits the needs at a given plant location. Further, it will be recognized that a variety of process configurations may be employed for recovering the liquid co-product stream. The present invention can be adapted to recover an NGL stream containing a significant fraction of the C 2  components present in the feed gas, or to recover a condensate stream containing only the C 4  and heavier components present in the feed gas, rather than producing an LPG co-product as described earlier. 
       FIG. 1  represents the preferred embodiment of the present invention for the processing conditions indicated.  FIGS. 5 through 10  depict alternative embodiments of the present invention that may be considered for a particular application. Depending on the quantity of heavier hydrocarbons in the feed gas and the feed gas pressure, the cooled feed stream  31   a  leaving heat exchanger  10  may not contain any liquid (because it is above its dewpoint, or because it is above its cricondenbar), so that separator  11  shown in  FIGS. 1 and 6  through  10  is not required, and the cooled feed stream can flow directly to an appropriate expansion device, such as work expansion machine  15 . In instances where the inlet gas is richer than that heretofore described, an embodiment of the present invention such as that shown in  FIG. 5  may be employed. Condensed liquid stream  33  flows through heat exchanger  18  and is subcooled, then divided into two portions. The first portion (stream  40 ) flows through expansion valve  12  where it undergoes expansion for flash vaporization as the pressure is reduced to about the pressure of distillation column  19 . The cold stream  40   a  from expansion valve  12  then flows through heat exchanger  18  where it is partially warmed as it is used to subcool stream  33  as described earlier. Partially warmed stream  40   b  is then further warmed in heat exchanger  10  and flows to a lower mid-point feed location on fractionation column  19 . The second liquid portion (stream  39 ), still at high pressure, is (1) combined with portion  34  of the vapor stream from separator  11 , or (2) combined with substantially condensed stream  35   a , or (3) expanded in expansion valve  17  and thereafter either supplied to fractionation column  19  at an upper mid-point feed location or combined with expanded stream  35   b . Alternatively, portions of stream  39  may follow any or all of the flow paths heretofore described and depicted in FIG.  5 . 
     The disposition of the gas stream remaining after recovery of the liquid co-product stream (stream  43  in  FIGS. 1 and 6  through  10 ) before it is supplied to heat exchanger  60  for condensing and subcooling may be accomplished in many ways. In the process of  FIG. 1 , the stream is heated, compressed to higher pressure using energy derived from one or more work expansion machines, partially cooled in a discharge cooler, then further cooled by cross exchange with the original stream. As shown in  FIG. 6 , some applications may favor compressing the stream to higher pressure, using supplemental compressor  59  driven by an external power source for example. As shown by the dashed equipment (heat exchanger  24  and discharge cooler  25 ) in  FIG. 1 , some circumstances may favor reducing the capital cost of the facility by reducing or eliminating the pre-cooling of the compressed stream before it enters heat exchanger  60  (at the expense of increasing the cooling load on heat exchanger  60  and increasing the power consumption of refrigerant compressors  64 ,  66 , and  68 ). In such cases, stream  49   a  leaving the compressor may flow directly to heat exchanger  24  as shown in  FIG. 7 , or flow directly to heat exchanger  60  as shown in FIG.  8 . If work expansion machines are not used for expansion of any portions of the high pressure feed gas, a compressor driven by an external power source, such as compressor  59  shown in  FIG. 9 , may be used in lieu of compressor  16 . Other circumstances may not justify any compression of the stream at all, so that the stream flows directly to heat exchanger  60  as shown in FIG.  10  and by the dashed equipment (heat exchanger  24 , compressor  16 , and discharge cooler  25 ) in FIG.  1 . If heat exchanger  24  is not included to heat the stream before the plant fuel gas (stream  48 ) is withdrawn, a supplemental heater  58  may be needed to warm the fuel gas before it is consumed, using a utility stream or another process stream to supply the necessary heat, as shown in  FIGS. 8 through 10 . Choices such as these must generally be evaluated for each application, as factors such as gas composition, plant size, desired co-product stream recovery level, and available equipment must all be considered. 
     In accordance with the present invention, the cooling of the inlet gas stream and the feed stream to the LNG production section may be accomplished in many ways. In the processes of  FIGS. 1 and 5  through  10 , inlet gas stream  31  is cooled and condensed by external refrigerant streams and flashed separator liquids. However, the cold process streams could also be used to supply some of the cooling to the high pressure refrigerant (stream  71   a ). Further, any stream at a temperature colder than the stream(s) being cooled may be utilized. For instance, a side draw of vapor from fractionation tower  19  could be withdrawn and used for cooling. The use and distribution of tower liquids and/or vapors for process heat exchange, and the particular arrangement of heat exchangers for inlet gas and feed gas cooling, must be evaluated for each particular application, as well as the choice of process streams for specific heat exchange services. The selection of a source of cooling will depend on a number of factors including, but not limited to, feed gas composition and conditions, plant size, heat exchanger size, potential cooling source temperature, etc. One skilled in the art will also recognize that any combination of the above cooling sources or methods of cooling may be employed in combination to achieve the desired feed stream temperature(s). 
     Further, the supplemental external refrigeration that is supplied to the inlet gas stream and the feed stream to the LNG production section may also be accomplished in many different ways. In  FIGS. 1 and 6  through  10 , boiling single-component refrigerant has been assumed for the high level external refrigeration and vaporizing multi-component refrigerant has been assumed for the low level external refrigeration, with the single-component refrigerant used to pre-cool the multi-component refrigerant stream. Alternatively, both the high level cooling and the low level cooling could be accomplished using single-component refrigerants with successively lower boiling points (i.e., “cascade refrigeration”), or one single-component refrigerant at successively lower evaporation pressures. As another alternative, both the high level cooling and the low level cooling could be accomplished using multi-component refrigerant streams with their respective compositions adjusted to provide the necessary cooling temperatures. The selection of the method for providing external refrigeration will depend on a number of factors including, but not limited to, feed gas composition and conditions, plant size, compressor driver size, heat exchanger size, ambient heat sink temperature, etc. One skilled in the art will also recognize that any combination of the methods for providing external refrigeration described above may be employed in combination to achieve the desired feed stream temperature(s). 
     Subcooling of the condensed liquid stream leaving heat exchanger  60  (stream  49   d  in  FIG. 1 , stream  49   e  in FIG.,  6 , stream  49   c  in  FIG. 7 , stream  49   b  in  FIGS. 8 and 9 , and stream  49   a  in  FIG. 10 ) reduces or eliminates the quantity of flash vapor that may be generated during expansion of the stream to the operating pressure of LNG storage tank  62 . This generally reduces the specific power consumption for producing the LNG by eliminating the need for flash gas compression. However, some circumstances may favor reducing the capital cost of the facility by reducing the size of heat exchanger  60  and using flash gas compression or other means to dispose of any flash gas that may be generated. 
     Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where appropriate. For example, conditions may warrant work expansion of the substantially condensed feed stream (stream  35   a  in  FIGS. 1 and 5  through  10 ). Further, isenthalpic flash expansion may be used in lieu of work expansion for the subcooled liquid stream leaving heat exchanger  60  (stream  49   d  in  FIG. 1 , stream  49   e  in  FIG. 6 , stream  49   c  in  FIG. 7 , stream  49   b  in  FIGS. 8 and 9 , and stream  49   a  in FIG.  10 ), but will necessitate either more subcooling in heat exchanger  60  to avoid forming flash vapor in the expansion, or else adding flash vapor compression or other means for disposing of the flash vapor that results. Similarly, isenthalpic flash expansion may be used in lieu of work expansion for the subcooled high pressure refrigerant stream leaving heat exchanger  60  (stream  71   c  in  FIGS. 1 and 6  through  10 ), with the resultant increase in the power consumption for compression of the refrigerant. 
     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.

Technology Category: 2