Patent Publication Number: US-8110023-B2

Title: Configurations and methods for offshore LNG regasification and BTU control

Description:
This application claims priority to our copending U.S. provisional patent application with the Ser. No. 60/636,960, filed Dec. 16, 2004, and which is incorporated by reference herein. 
    
    
     FIELD OF THE INVENTION 
     The field of the invention is natural gas processing, especially as it relates to LNG (liquefied natural gas) regasification and processing in a combined on-/offshore facility. 
     BACKGROUND OF THE INVENTION 
     Offshore LNG regasification has become an increasingly attractive option in LNG import. Among other advantages, offshore regasification terminals or terminals in a relatively remote location help to reduce various safety and security concerns of local communities nearby a terminal that would otherwise be onshore or in a location near human habitation and/or activity. 
     Unfortunately, offshore installations are generally significantly more expensive than onshore installations, and numerous additional technical challenges arise from offshore LNG storage, unloading and regasification. Several solutions have recently been proposed to overcome at least some of these difficulties. However, all or almost all of the presently known offshore configurations fail to provide a mechanism by which chemical composition of the LNG can be altered to a desirable composition (e.g., processing of lower quality LNG with heating values higher than the North American pipeline specifications). As pipeline transport of natural gas in North America and other countries must typically meet hydrocarbon dew point and gross heating value requirements of the associated distribution systems, presence of heavier components in LNG is generally not desirable. 
     In many presently known configurations, heavy hydrocarbons are removed from LNG in a process that includes vaporizing the LNG in a demethanizer using a reboiler, and re-condensing the demethanizer overhead to a liquid that is then pumped and vaporized. For example, McCartney describes in U.S. Pat. No. 6,564,579 such regasification process and configurations. While these configurations and methods typically operate satisfactorily under onshore conditions, offshore installation would be unacceptable under most scenarios as these configurations require relatively substantial space. 
     In currently known offshore LNG regasification terminals, LNG is typically heated to pipeline specification (e.g., about 50° F. and 1200 psig) in offshore vaporizers using seawater or submerged combustion vaporizers. Commonly, fractionation facilities are not provided due to the space limitation in an offshore environment, and the regasified LNG is then sent via an undersea pipeline to an onshore consumer gas pipeline. Thus, while offshore regasification is realized, change in chemical composition is typically not possible using such configurations. It should be noted that when the LNG is fully vaporized, BTU reduction and/or recovery of non-methane components (e.g., ethane, propane, etc.) is generally not economical as these processes would require significant refrigeration and recompression. Consequently, and at least for these reasons, only high quality LNG with acceptable heating value content and/or desirable chemical composition are imported, while lower quality LNG (e.g., LNG with relatively high BTU) is often rejected. 
     Thus, while numerous configurations and methods to separate heavier components from LNG or to reduce BTU of LNG are known in the art, all or almost all of them fail to provide economically attractive operation, especially in an offshore environment. Therefore, there is still a need to provide improved configurations and methods for LNG regasification that allows for simple and cost-effective removal of non-methane components to thereby produce LNG with a desirable BTU and/or chemical composition. 
     SUMMARY OF THE INVENTION 
     The present invention is directed to configurations and methods in which LNG is first pumped to supercritical pressure and then vaporized, preferably in an offshore vaporizer or a vaporizer that is in a location that is remote (e.g., more than 1 km) from a populated area, to a temperature that is a function of the concentration of non-methane components in the LNG (e.g., between about −20° F. to about 15° F.). The so formed supercritical vaporized natural gas is then transported to an onshore facility and split into a first and second portion, wherein the split ratio is once more a function of the concentration of non-methane components in the LNG. 
     The first portion is then processed to remove at least some non-methane components from the natural gas. Most preferably, work is produced by expanding the regasified natural gas to thereby power recompression of the lean natural gas, which is then combined with the second portion to thereby form a processed LNG. 
     In one aspect of the inventive subject matter, a method of providing a natural gas product includes a step in which vaporized supercritical LNG is provided, most preferably from offshore to an onshore terminal. In another step, the vaporized supercritical LNG is split into a first and second stream, wherein the first stream is processed to remove at least some non-methane components from the first stream to form a lean natural gas product, and wherein the step of processing further includes a first turbo-expansion of at least a portion of the first stream. In yet another step, the lean natural gas product is compressed using at least in part energy from the first turbo-expansion, and the compressed lean natural gas product is then combined with the second stream to thereby form a sales gas with predetermined content of non-methane components. 
     Preferably, the vaporized supercritical LNG is at a predetermined temperature and split ratio between first and second streams is at a predetermined ratio, wherein both the temperature and ratio are a function of a concentration of non-methane components in the LNG. It is further preferred that in such methods the first stream is processed in an absorber that further produces an absorber bottom product, wherein the bottom product is further processed in at least one downstream column (typically operated at a lower pressure than the absorber pressure) to produce at least one of an ethane product and a propane-containing product. In at least some of such configurations, it is preferred that the downstream column is operated as a demethanizer and provides an overhead product to the absorber as a reflux stream and/or a bottom feed stream. A second turbo-expansion may be included that expands at least a portion of the first stream, wherein the first turbo-expansion provides reflux condenser duty, and wherein the second turbo-expansion provides refrigeration duty in the absorber. 
     Accordingly, it is contemplated that an offshore facility may include a source of LNG (e.g., LNG carrier, submerged or floating LNG tank) and a pump fluidly coupled to the source, wherein the pump pumps LNG to supercritical pressure. A regasification unit (e.g., open rack seawater vaporizer, submerged combustion fuel fired vaporizer, intermediate fluid vaporizer, and/or Rankine cycle vaporizer) is then coupled to the pump and operated to regasify the supercritical LNG to a predetermined temperature (about −20° F. to about 20° F.), wherein a controller is operationally linked with the regasification unit and enabled to set the temperature of the regasified LNG as a function of the concentration of non-methane components in the LNG. Most preferably, the controller comprises a central processing unit programmed to control the temperature as a function of previously provided information on chemical composition of the LNG. 
     In another aspect of the inventive subject matter, a LNG processing plant includes an onshore or offshore portion that is configured to pump LNG to supercritical pressure and to regasify the pressurized LNG. An onshore portion of such plants is configured to process one portion of the regasified LNG to remove at least some non-methane content in the LNG to thereby form a lean natural gas product, wherein the onshore portion is configured to produce a sales gas from the lean natural gas product and another portion of the regasified LNG. Typically, the onshore portion comprises an absorber that receives the one portion of the regasified LNG to thereby produce the lean regasified. Similar to configurations above, contemplated plants include a turbo-expander that expands the one portion of the regasified LNG before entry into the absorber, and still further include a compressor coupled to the expander and compresses the lean natural gas product. A downstream column will typically be configured to receive an absorber bottom product and to produce an ethane and propane-containing product, or may be configured as a demethanizer to receive an absorber bottom product and to produce a reflux stream and/or a bottom feed stream to the absorber. 
     Viewed from another perspective, contemplated plants may include a source (e.g., onshore or offshore) that provides regasified LNG at supercritical pressure, wherein the LNG has a first quantity of non-methane components. An onshore flow divider may be provided that produces a first and a second stream from the regasified LNG, and an onshore absorber is configured to produce a lean natural gas product from a turbo-expanded portion of the first stream. An onshore compressor will then compress the lean natural gas product, wherein the compressor uses energy from the turbo-expansion of the first stream. An onshore flow combining element is configured to produce a sales gas from the compressed lean natural gas product and the second stream, wherein the sales gas has a quantity of non-methane components that is less than the first quantity. 
     Various objects, features, aspects and advantages of the present invention will become more apparent from the following detailed description of preferred embodiments of the invention. 
    
    
     
       BRIEF DESCRIPTION OF THE DRAWINGS 
         FIG. 1  is one exemplary configuration of offshore LNG regasification with onshore processing using a two-column design. 
         FIG. 2  is another exemplary configuration of offshore LNG regasification with onshore processing using a three-column design. 
     
    
    
     DETAILED DESCRIPTION 
     The inventors have discovered that non-methane components (i.e., those having two or more carbon atoms (C2+)) can be separated from LNG in an economically desirable manner in which the LNG is pumped to supercritical pressure, preferably in an offshore or remote onshore location and in which the supercritical LNG is regasified to an intermediate temperature in the offshore or remote location. The so heated supercritical natural gas is then transferred to a processing unit (e.g., onshore location). Alternatively, at least one of the offshore functions may also be performed onshore. 
     Depending on the chemical composition of the LNG, a variable fraction of the heated and vaporized natural gas is then processed in an onshore location to form a lean natural gas product that is then combined with another fraction of the heated and vaporized natural gas to thereby produce a sales gas with predetermined composition and/or heating value. Thus, it should be recognized that such configurations may be employed for BTU control of import LNG that fails to meet pipeline specification. Onshore processing will most typically take advantage of the relatively high pressure of the vaporized natural gas, which is expanded in a turboexpander to generate power for recompression of the residue gas, and/or to supply at least part of the refrigeration (cooling) requirements of reflux condensers in downstream fractionation columns (demethanizer and/or deethanizer). Thus, cooling for the separation process is provided by the vaporized LNG, and it should therefore be recognized that the temperature of the vaporized supercritical natural gas will be a function of the non-methane content in the LNG. 
     In one especially preferred configuration, a portion of the flashed vapor from the first turboexpander is processed in a second turboexpander that is configured for varying levels of BTU reduction (the ratio of turbo expanded to non-turbo expanded vapor will determine the level of C2+ removal). At least part of the power generated by the second turboexpander is used to recompress the residue gas. It should be especially noted that two turboexpanders operating in series can provide significant power to recompress the residue gas to the pipeline pressure. However, where desirable, one or more additional compressors can be added where a high pipeline delivery pressure is required. It is also noted that by bypassing a portion of the onshore vapor around the first turboexpander, the size of the downstream processing unit can be reduced, lowering the capital cost of the onshore BTU reduction unit. Of course, the actual quantity of bypassed material will predominantly depend on the BTU content of the import LNG, the pipeline gas heating value requirement, and/or the desire for C2 and C3+ products. 
     In such configurations, contemplated plants are built as a two column plant in which a first column operates as a reflux demethanizer that receives two reflux streams, and in which a second column operates as a deethanizer producing an ethane overhead vapor and a bottom C3+ product (i.e., product comprising compounds having three or more carbon atoms). Such configurations will advantageously allow change in component separation and varying levels of BTU control by changing process temperatures and split ratios of the reflux streams. 
     An exemplary scheme of a two column plant configuration is depicted in  FIG. 1 . Here, the plant comprises an offshore LNG receiving terminal that receives LNG from an LNG carrier  51 . LNG is unloaded from the carrier via unloading arms to the offshore LNG storage tank  52 . The LNG storage tanks can be a gravity based structure, or a floating LNG vessel. A typical LNG composition (stream  1 ) is shown in Table 1. LNG from the storage tanks is pumped by the primary pump  53  to an intermediate pressure, typically at 100 psig. The pressurized LNG is further pumped by the secondary pump  54  to supercritical pressure, typically 1500 psig to 2200 psig forming stream  2 . It should be noted that the secondary pump discharge pressure will be typically increased with increasing content of non-methane components in the LNG and/or with increased onshore pipeline gas delivery pressure. The supercritical ING is then heated in LNG vaporizers  55  to an intermediate temperature typically at −10° F. to 10° F., forming stream  3 . The intermediate temperature is selected as a function of the LNG composition and the level of BTU reduction. Most typically, stream  3  will have a lower temperature when higher levels of C2+ extraction are required onshore. Conventional LNG vaporizers can be used for the regasification facility, including open rack seawater vaporizers, submerged combustion fuel fired vaporizers, intermediate fluid vaporizers, Rankine cycle vaporizers and/or other suitable heat sources (which may also come from an onshore location). The heated LNG is then transported via an undersea pipeline  56  to the onshore facility. 
     Therefore, it should be appreciated that contemplated configurations will include an offshore facility comprising a source of LNG and a pump that is fluidly coupled to the source, wherein the pump is configured to produce LNG at supercritical pressure (typically between about 1500 psig and 220 psig, and even higher). A regasification unit is coupled to the pump and configured to regasify the supercritical LNG to a predetermined temperature, wherein a controller (e.g., CPU, or human operator) is operationally linked to the regasification unit and enabled to set the temperature of the regasified LNG as a function of a concentration of non-methane components in the LNG. 
     Most typically, the LNG source is a LNG carrier, a submerged and/or floating LNG tank. In less preferred aspects, the LNG source may also be a pipeline (preferably undersea pipeline). It should further be appreciated that the regasification unit need not be limited to a specific type, but that all known types and especially those suitable for offshore operation are deemed suitable for use herein. Therefore, contemplated regasification units include open rack seawater vaporizers, submerged combustion fuel fired vaporizers, intermediate fluid vaporizers, Rankine cycle vaporizers, etc. With respect to the temperature of the vaporized supercritical natural gas it should be noted that the particular temperature will depend on the chemical composition of the LNG, and especially on the content of non-methane components in the LNG. However, it is generally preferred that the temperature will be below normal pipeline operating conditions, and especially preferred temperatures are between about −20° F. to about 20° F. However, and especially where the LNG is relatively rich and/or where it is desired to produce a particularly lean sales gas, the temperature may also be between 60° F. and −10° F. Thus, it is generally preferred that the controller has a central processing unit that is programmed to control the temperature as a function of entered or otherwise previously provided information on chemical composition of the LNG. Alternatively, pumping to supercritical pressure and/or vaporization of the supercritical LNG may also be performed in an onshore location using components well known in the art. However, where vaporization is performed onshore, it is generally preferred that the heat for the vaporization is provided at least in part by thermal integration with a power cycle (e.g., using heat exchange fluids coupled to a steam cycle or HRSG). 
     Alternatively, the LNG source and/or the regasification unit may also be located in an area that is relatively remote from human habitation and/or activity and will provide the onshore facility with the regasified supercritical natural gas. For example, storage and/or regasification may be done in a configuration in which the storage and/or regasification are at least 1 km, more typically at least 5 km, and most typically at least 10 km away from the onshore facility. 
     Once the supercritical vaporized LNG  3  reaches the onshore facility, stream  3  is split into two portions, stream  4  and stream  5 , wherein the ratio between streams is a function of the desired level of BTU reduction (and/or concentration of non-methane components). Stream  4  bypasses the BTU reduction unit and is mixed with residue gas stream  20  forming sales gas stream  21  that is fed to the gas pipeline. Stream  5  is letdown in pressure in a first turboexpander  57  forming stream  6 , typically at about 1100 psig and a temperature of about −10° F. to −60° F. The first turboexpander  57  provides a portion of the compression power to operate the residue compressor, which is operationally coupled to the expander. Stream  6  is heated in exchanger  68  to 0° F. to −25° F. to form stream  7  by supplying refrigeration duties for the reflux condenser  68 . The two phase stream is separated in the separator  59  into a liquid stream  9  and a vapor stream  8 . Vapor stream  8  is further split into stream  11  and stream  12 . it should be noted that the split between streams  11  and  12  is adjusted as necessary to meet the varying levels of BTU reduction or C2+ recovery (infra). The liquid stream  9  is letdown in pressure in a JT valve  60  to about 450 psig forming stream  10  that enters the lower section of the first column  63 . 
     When a high C2+ removal is required, the flow of stream  12  relative to stream  11  is increased, resulting in an increase in reflux flow to the overhead exchanger  64  where stream  12  is chilled to typically −90° F. to −110° F. forming stream  14 . Stream  14  is then letdown in pressure by JT valve  62  forming stream  15  to about 450 psig to 500 psig and fed to the upper section of the first column (here: operating as a demethanizer). Stream  11  is letdown in pressure to about 450 psig to 500 psig in the second turboexpander  61  forming stream  13 , typically at −40° F. to −60° F. and fed to the mid section of column  63 . The power generated by the second turboexpander is preferably used to provide a portion of the residue gas compression requirement. The turboexpander  61  also chills the feed gas, supplying a portion of the rectification duty in the first column. 
     Demethanizer column  63  typically operates between about 450 psig to about 500 psig and produces an overhead stream  16  and a bottom stream  22 . It should be noted that the temperatures of these two streams will vary depending on the desired levels of C2+ recovery. For example, during high C2+ recovery, the overhead temperature is preferably maintained at about −110° F. to about −145° F., as needed for recovery of ethane and heavier components. The demethanizer column bottom temperature is maintained by reboiler  71 . During lower C2+ recovery, the overhead temperature may be increased to about −80° F. to about −100° F., as needed in rejecting some of the C2 components overhead. The refrigerant content in the first column overhead stream  16  is recovered in heat exchanger  64  by providing cooling to the reflux stream  12 . The so heated stream  17  is then compressed by the compressor that is operationally coupled to the second turboexpander forming stream  18 , typically at −10° F. to −30° F., which is further compressed by the residue gas compressor driven by the first turboexpander to form stream  19  at about 900 psig to 1200 psig. Where desirable, additional recompression with compressor  65  can be used to boost the residue gas pressure to the sales gas pipeline pressure forming stream  20  that is then mixed with bypass stream  4 . 
     The first column bottom stream  22  is letdown in pressure by JT valve  66  to about 200 to 400 psig forming stream  23  prior to entering the upper section of the second distillation column  67 , the deethanizer. The deethanizer is of a conventional column design that produces a C2 rich overhead vapor stream  24  and a C3+ bottom product stream  25 . The overhead vapor  24  is condensed in reflux condenser  68 , with cooling supplied by the feed gas stream  6 . The chilled overhead stream  26  is separated in the reflux drum  69  into an ethane product stream  27  and a liquid stream  28  that is further pumped by pump  70  forming stream  29  to be refluxed to the deethanizer column. Heating requirement in the deethanizer column is supplied with reboiler  72  using an external heat source. The overall material balance for the BTU reduction unit is shown in Table 1. 
     Therefore, the inventors contemplate a method of providing a natural gas product that includes the steps of (1) providing vaporized supercritical LNG, preferably from offshore to an onshore terminal; (2) splitting the vaporized supercritical LNG into a first and second stream; (3) processing the first stream to remove at least some non-methane components from the first stream to form a lean natural gas product, wherein the step of processing includes a first turbo-expansion of at least a portion of the first stream; (4) compressing the lean natural gas product using at least in part energy from the first turbo-expansion; and (5) combining the compressed lean natural gas product with the second stream to thereby form a sales gas with predetermined content of non-methane components. 
     As already discussed above, preferred steps of providing the vaporized supercritical LNG includes vaporizing the supercritical LNG to a predetermined temperature, wherein the temperature is a function of a concentration of non-methane components in the LNG. Similarly, the step of splitting the vaporized supercritical LNG in first and second streams is a function of a concentration of non-methane components in the LNG. Most preferably, the step of processing further includes a second turbo-expansion of at least a portion of the first stream, wherein the first turbo-expansion provides reflux condenser duty, and wherein the second turbo-expansion provides refrigeration duty in the absorber. 
     Therefore, particularly preferred plants will include a portion (preferably offshore) configured to pump LNG to supercritical pressure and to regasify the pressurized LNG, and an onshore portion configured to process one portion of the regasified LNG to remove at least a portion of non-methane content in the LNG to thereby form a lean natural gas product. In such plants, the onshore portion is typically further configured to produce a sales gas from a mixture of the lean natural gas product and another portion of the regasified LNG. Viewed from a different perspective, a plant is also contemplated having an offshore source that provides regasified LNG at supercritical pressure, wherein the LNG has a first quantity of non-methane components. An onshore flow divider is configured to produce a first and a second stream from the regasified LNG, and an onshore absorber is configured to produce a lean natural gas product from a turbo-expanded portion of the first stream. Such plants will further include an onshore compressor that compresses the lean natural gas product, wherein the compressor is configured to use energy from the turbo-expansion of the first stream, and an onshore flow combining element that is configured to produce a sales gas from the compressed lean natural gas product and the second stream, wherein the sales gas has a quantity of non-methane components that is less than the first quantity. As discussed above, it is generally preferred that a control unit (e.g., human operator, or device comprising a CPU and programmed to operate without manual or user intervention) that is configured to control the temperature of the regasified LNG and/or the ratio of first and second streams at the flow divider, wherein the temperature and/or the ratio are set as a function of a concentration of non-methane in the regasified LNG. 
     In another preferred configuration, the BTU reduction unit includes three columns, with the first column (here: absorber) operates at a higher pressure than the second column, and wherein the bottom liquid from the absorber is let down in pressure (e.g., via Joule-Thompson valve) and fed to the second column. It should be appreciated that by operating the first column at a higher pressure, the residue gas compression horsepower can be significantly reduced, especially when relatively high pipeline gas pressure is required. It should also be appreciated that the reduction in pressure of the first bottom product supplies a portion of the refrigeration for rectification function to the second column (typically via JT effect) which operates as a demethanizer. The overhead vapor from the second column is compressed in a recycle compressor and returned to the first column. The third column then operates as a deethanizer at yet lower pressure than the first and second column producing an ethane overhead vapor and bottom C3+ products. 
     It should be especially noted that the overhead vapor from the second column is split into two portions. The first portion is chilled in a reflux exchanger with overhead vapor from the absorber to thereby form a cold reflux to the top section of the first column (absorber). The second portion of the overhead vapor forms a stripping gas that is fed to the bottom of the first column. Using such split flow configurations, it is pointed out that ratio of the first portion to the second vapor portion from the second distillation column can be used to control in large part the desired level of C2+ recovery. 
     One exemplary schematic of such configurations is depicted in  FIG. 2 . Here, the plant comprises an offshore LNG receiving terminal that receives LNG from an LNG carrier  51 . LNG is unloaded from the carrier via unloading arms to the offshore LNG storage tank  52 . The LNG storage tanks can be a gravity based structure, or a floating LNG vessel. As before, a typical LNG composition (stream  1 ) is shown in Table 2. LNG from the storage tanks is pumped by the primary pump  53  to an intermediate pressure, typically at 100 psig. The pressurized LNG is further pumped by the secondary pump  54  to supercritical pressure, typically 1500 psig to 2200 psig forming stream  2 . It should be noted that the secondary pump discharge pressure is typically raised with increasing richness of the import LNG and/or onshore pipeline gas delivery pressure. 
     The supercritical LNG is then heated in LNG vaporizers  55  to an intermediate temperature typically at −10° F. to 10° F., forming stream  3 . The intermediate temperature is dependent on the LNG composition and the level of BTU reduction, and generally, lower temperature is required when higher level of C2+ extraction is required onshore. Conventional LNG vaporizers can be used for the regasification facility, including open rack seawater vaporizers, submerged combustion fuel fired vaporizers, intermediate fluid vaporizers, Rankine cycle vaporizers, or other suitable heat sources. The heated LNG is then transported via an undersea pipeline  56  to the onshore facility. 
     Once the supercritical LNG reaches the onshore facility, stream  3  is split into two portions, stream  4  and stream  5 , with the split ratio determined by the level of BTU reduction requirement. Stream  4  bypasses the BTU reduction unit and is mixed with the residue gas stream  20  forming stream  21  that is fed to the gas pipeline. Stream  5  is letdown in pressure in the first turboexpander  57  forming stream  6 , typically at 1100 psig and −20° F. to −60° F. The first turboexpander  57  provides a portion of the compression power to operate the residue compressor. Stream  6  is heated to 0° F. to −25° F. forming stream  7  by supplying the refrigeration duties for the reflux condensers  68  and  74 . The two phase stream is separated in the separator  59  into a liquid stream  9 , and a vapor stream  8  that is further split into stream  11  and stream  12 . The split is adjusted as necessary to meet the varying levels of BTU reduction or C2+ recovery (infra). The liquid stream  9  is letdown in pressure in a JT valve  60  to about 600 psig forming stream  10  that enters the lower section of the first column  63 . 
     When high C2+ removal is required, the ratio of stream  12  to stream  11  is increased, resulting in an increase in reflux flow to the overhead exchanger  64 . Stream  12  is chilled to typically −90° F. to −110° F. in exchanger  64  forming stream  14 , and is letdown in pressure by JT valve  62  forming stream  15 , to about 400 psig to 650 psig and fed to the upper section of the first column (here: absorber). Stream  11  is letdown in pressure to about 400 psig to 650 psig in the second turboexpander  61  forming stream  13 , typically at −40° F. to −60° F. and fed to the mid section of column  63 . The power generated by the second turboexpander is preferably used to provide a portion of the residue gas compression requirement. The turboexpansion also provides for chilling the feed gas, thus supplying a portion of the rectification duty in the first column. 
     The first column is also fed by the recycle stream  37  and stream  38  from the second column. By adjusting the ratio between these two streams, C2 and C3 recoveries can be adjusted as needed. The first column operating between 400 psig to 650 psig produces an overhead stream  16  and a bottom stream  22 . The temperatures of these two streams vary depending on the levels of C2+ recovery. For example, during high C2+ recovery, the overhead temperature must be maintained at −110° F. to −145° F., as needed for recovery of the ethane and heavier components. During lower C2+ recovery, the overhead temperature is increased to about −80° F. to −100° F., as needed in rejecting some of the C2 components overhead. The refrigerant content in the first column overhead stream  16  is recovered in heat exchanger  64  by providing cooling to the first and second reflux streams  37  and  12  to thereby form streams  39  and  14 , respectively. The heated stream  17  is compressed by a compressor that is at least in part driven by the second turboexpander  61  forming stream  18 , typically at −10° F. to −30° F. and is further compressed by the residue gas compressor driven by the first turboexpander  57  forming stream  19  at about 900 psig to 1200 psig. As an option, additional recompression with compressor  65  can be used to boost the residue gas pressure to the sales gas pipeline pressure forming stream  20  that can be mixed with the bypass stream  4 . 
     The first column bottom stream  22  is letdown in pressure by JT valve  66  to about 200 to 400 psig forming stream  29  prior to entering the upper section of the second distillation column  73 . Distillation column  73  operates at about 200 to 400 psig serving as a demethanizer fractionating stream  29  into C2+ bottom  31  and a C1 rich overhead stream  30 . The overhead vapor is condensed using refrigeration from the inlet feed stream  6  in reflux exchanger  74 , forming stream  32  at about 0° F. to −40° F. Stream  32  is separated in reflux drum  75  into a liquid stream  34  and a vapor stream  33 . The liquid stream  34  is pumped by reflux pump  76  forming stream  35  and returned to the top of the second column  73  as reflux. 
     The vapor stream  33  is compressed by compressor  77  forming stream  36  which is split into stream  37  and  38 , and routed to exchanger  64  providing reflux and/or to the bottom of the first column for ethane re-absorption. Heating requirement in the second column is supplied with reboiler  71  using an external heat source. The temperature of the NGL bottom product ranges from 100° F. to 200° F. depending on the level of BTU reduction. The second column bottom is sent to the third column  67  (after expansion in JT valve  78  via stream- 23 ), which is operated as a deethanizer for further fractionation. 
     The deethanizer is typically of a conventional column design that produces a C2 rich overhead vapor stream  24  and a C3+ bottom product stream  25 . The overhead vapor is condensed in reflux condenser  68 , with cooling supplied by the feed gas stream  6 . The chilled overhead stream  26  is separated in the reflux drum  69  into an ethane product stream  27  and a liquid stream  28  that is further pumped by pump  70  forming stream  29  to be refluxed to the deethanizer column. Heating requirement in the deethanizer column is supplied with reboiler  72  using an external heat source, and heating requirements of column  73  is supplied with reboiler  71  using an external heat source. The overall material balance for the BTU reduction unit is shown in Table 2. 
     Thus, specific embodiments and applications for LNG regasification and BTU control have been disclosed. It should be apparent, however, to those skilled in the art that many more modifications besides those already described are possible without departing from the inventive concepts herein. For example, the offshore portion of contemplated configurations and methods may also be positioned and/or operated in part or in toto onshore. The inventive subject matter, therefore, is not to be restricted except in the spirit of the appended claims. Moreover, in interpreting both the specification and the claims, all terms should be interpreted in the broadest possible manner consistent with the context. In particular, the terms “comprises” and “comprising” should be interpreted as referring to elements, components, or steps in a non-exclusive manner, indicating that the referenced elements, components, or steps may be present, or utilized, or combined with other elements, components, or steps that are not expressly referenced. Furthermore, where a definition or use of a term in a reference, which is incorporated by reference herein is inconsistent or contrary to the definition of that term provided herein, the definition of that term provided herein applies and the definition of that term in the reference does not apply. 
     
       
         
           
               
               
               
               
               
             
               
                 TABLE 1 
               
               
                   
               
               
                 Stream 
                 LNG 
                 ETHANE 
                 LPG 
                 RESIDUE 
               
               
                 Number 
                 1 
                 27 
                 25 
                 GAS 21 
               
               
                   
               
             
            
               
                   
               
            
           
           
               
               
               
               
               
            
               
                 N2 
                 0.0034 
                 0.0000 
                 0.0000 
                 0.0037 
               
               
                 C1 
                 0.8976 
                 0.0216 
                 0.0000 
                 0.9833 
               
               
                 C2 
                 0.0501 
                 0.9584 
                 0.0100 
                 0.0116 
               
               
                 C3 
                 0.0316 
                 0.0200 
                 0.6277 
                 0.0012 
               
               
                 iC4 
                 0.0069 
                 0.0000 
                 0.1442 
                 0.0001 
               
               
                 NC4 
                 0.0103 
                 0.0000 
                 0.2160 
                 0.0001 
               
               
                 C5 
                 0.0001 
                 0.0000 
                 0.0021 
                 0.0000 
               
               
                 MMscfd 
                 1,200 
                 49 
                 57 
                 1,094 
               
               
                 BPD 
                 513,848 
                 30,827 
                 39,374 
                 443,647 
               
               
                 HHV, 
                 1123 
                 1756 
                 2765 
                 1009 
               
               
                 Btu/Scf 
               
               
                   
               
            
           
         
       
     
     
       
         
           
               
               
               
               
               
             
               
                 TABLE 2 
               
               
                   
               
               
                 Stream 
                 LNG 
                 ETHANE 
                 LPG 
                 RESIDUE 
               
               
                 Number 
                 1 
                 27 
                 25 
                 GAS 21 
               
               
                   
               
             
            
               
                   
               
            
           
           
               
               
               
               
               
            
               
                 N2 
                 0.0034 
                 0.0000 
                 0.0000 
                 0.0036 
               
               
                 C1 
                 0.8976 
                 0.0000 
                 0.0000 
                 0.9625 
               
               
                 C2 
                 0.0501 
                 0.9800 
                 0.0200 
                 0.0308 
               
               
                 C3 
                 0.0316 
                 0.0200 
                 0.6144 
                 0.0027 
               
               
                 iC4 
                 0.0069 
                 0.0000 
                 0.1449 
                 0.0002 
               
               
                 NC4 
                 0.0103 
                 0.0000 
                 0.2185 
                 0.0001 
               
               
                 C5 
                 0.0001 
                 0.0000 
                 0.0021 
                 0.0000 
               
               
                 MMscfd 
                 1,200 
                 24 
                 56 
                 1,119 
               
               
                 BPD 
                 513,848 
                 15,571 
                 38,761 
                 459,087 
               
               
                 HHV, 
                 1123 
                 1773 
                 2760 
                 1027 
               
               
                 Btu/Scf