LNG FACILITY WITH INTEGRATED NGL EXTRACTION TECHNOLOGY FOR ENHANCED NGL RECOVERY AND PRODUCT FLEXIBILITY
Process for efficiently operating a natural gas liquefaction system with integrated heavies removal/natural gas liquids recovery to produce liquefied natural gas (LNG) and/or natural gas liquids (NGL) products with varying characteristics, such as, for example higher heating value (HHV) and/or propane content. Resulting LNG and/or NGL products are capable of meeting the significantly different specifications of two or more markets.
This application claims priority benefit under 35 U.S.C. Section 119(e) of U.S. Provisional Patent Ser. No. 60/698,402 filed Jul. 12, 2005, the entire disclosure of which is incorporated herein by reference.
BACKGROUND OF THE INVENTION1. Field of the Invention
This invention relates generally to a method and apparatus for liquefying natural gas. In another aspect, the invention concerns an improved liquefied natural gas (LNG) facility capable of efficiently supplying LNG products meeting significantly different product specifications.
2. Description of the Prior Art
The cryogenic liquefaction of natural gas is routinely practiced as a means of converting natural gas into a more convenient form for transportation and/or storage. Generally, liquefaction of natural gas reduces its volume by about 600-fold, thereby resulting in a liquefied product that can be readily stored and transported at near atmospheric pressure.
Natural gas is frequently transported by pipeline from the supply source to a distant market. It is desirable to operate the pipeline under a substantially constant and high load factor, but often the deliverability or capacity of the pipeline will exceed demand while at other times the demand will exceed the deliverability of the pipeline. In order to shave off the peaks where demand exceeds supply or the valleys where supply exceeds demand, it is desirable to store the excess gas in such a manner that it can be delivered as the market dictates. Such practice allows future demand peaks to be met with material from storage. One practical means for doing this is to convert the gas to a liquefied state for storage and to then vaporize the liquid as demand requires.
The liquefaction of natural gas is of even greater importance when transporting gas from a supply source that is separated by great distances from the candidate market, and a pipeline either is not available or is impractical. This is particularly true where transport must be made by ocean-going vessels. Ship transportation of natural gas in the gaseous state is generally not practical because appreciable pressurization is required to significantly reduce the specific volume of the gas, and such pressurization requires the use of more expensive storage containers.
In view of the foregoing, it would be advantageous to store and transport natural gas in the liquid state at approximately atmospheric pressure. In order to store and transport natural gas in the liquid state, the natural gas is cooled to −240° F. to −260° F. where the liquefied natural gas (LNG) possesses a near-atmospheric vapor pressure.
Numerous systems exist in the prior art for the liquefaction of natural gas in which the gas is liquefied by sequentially passing the gas at an elevated pressure through a plurality of cooling stages whereupon the gas is cooled to successively lower temperatures until the liquefaction temperature is reached. Cooling is generally accomplished by indirect heat exchange with one or more refrigerants such as propane, propylene, ethane, ethylene, methane, nitrogen, carbon dioxide, or combinations of the preceding refrigerants (e.g., mixed refrigerant systems). A liquefaction methodology that may be particularly applicable to one or more embodiments of the present invention employs an open methane cycle for the final refrigeration cycle wherein a pressurized LNG-bearing stream is flashed and the flash vapors are subsequently employed as cooling agents, recompressed, cooled, combined with the processed natural gas feed stream, and liquefied, thereby producing the pressurized LNG-bearing stream.
In the past, LNG facilities have been designed and operated to provide LNG to a single market in a certain region of the world. As global demand for LNG increases, it would be advantageous for a single LNG facility to be able to supply LNG to multiple markets in different regions of the world. However, natural gas specifications vary greatly throughout the world. Typically, such natural gas specifications include criteria such as higher heating value (HHV), Wobbe index, methane content, ethane content, C3+ content, and inerts content. For example, different world markets demand an LNG product having an HHV anywhere between 950 and 1160 BTU/S CF. Existing LNG facilities are optimized to meet a certain set of specifications for a single market. Thus, changing the operating parameters of an LNG facility in an effort to make LNG that would meet the non-design specifications of a different market creates significant operating inefficiencies in the facility. These operating inefficiencies associated with producing LNG for non-design specifications generally makes it economically unfeasible to serve more than one market with a single LNG facility.
SUMMARY OF THE INVENTIONIn one embodiment of the present invention there is provided a process for producing liquefied natural gas (LNG). The process includes the following steps: (a) operating an LNG facility in a first mode of operation to thereby produce a first LNG product; (b) adjusting at least one non-feed operating parameter of the LNG facility so that the LNG facility operates in a second mode of operation; and (c) operating the LNG facility in the second mode of operation to thereby produce a second LNG product. The first and second modes of operation are not to be carried out during start-up or shut-down of the LNG facility. Steps (a) and (c) can, optionally, include producing first and second natural gas liquids (NGL) products respectively. The average higher heating value (HHV) of the second LNG product is at least about 10 BTU/SCF different than the average HHV of the first LNG product and/or the average propane content of the second NGL product is at least about 1 mole percent different than the average propane content of the first NGL product.
In another embodiment of the present invention there is provided a method of varying the heating value of LNG produced from an LNG facility. The method includes the following steps: (a) cooling natural gas by indirect heat exchange to thereby produce a first cooled stream; (b) using a first distillation column to separate at least a portion of the first cooled stream into a first relatively more volatile fraction and a first relatively less volatile fraction; (c) cooling at least a portion of the first relatively more volatile fraction to thereby produce LNG; and (d) adjusting at least one operating parameter of the first distillation column to thereby vary the HHV of the produced LNG by at least about 1 percent over a time period of less than about 72 hours.
BRIEF DESCRIPTION OF THE DRAWING FIGURESA preferred embodiment of the present invention is described in detail below with reference to the attached drawing figures, wherein:
The present invention can be implemented in a process/facility used to cool natural gas to its liquefaction temperature, thereby producing liquefied natural gas (LNG). The LNG process generally employs one or more refrigerants to extract heat from the natural gas and then reject the heat to the environment. In one embodiment, the LNG process employs a cascade-type refrigeration process that uses a plurality of multi-stage cooling cycles, each employing a different refrigerant composition, to sequentially cool the natural gas stream to lower and lower temperatures. In another embodiment, the LNG process is a mixed refrigerant process that employs at least one refrigerant mixture to cool the natural gas stream.
Natural gas can be delivered to the LNG process at an elevated pressure in the range of from about 500 to about 3,000 pounds per square in absolute (psia), about 500 to about 1,000 psia, or 600 to 800 psia. Depending largely upon the ambient temperature, the temperature of the natural gas delivered to the LNG process can generally be in the range of from about 0 to about 180° F., about 20 to about 150° F., or 60 to 125° F.
In one embodiment, the present invention can be implemented in an LNG process that employs cascade-type cooling followed by expansion-type cooling. In such a liquefaction process, the cascade-type cooling may be carried out at an elevated pressure (e.g., about 650 psia) by sequentially passing the natural gas stream through first, second, and third refrigeration cycles employing respective first, second, and third refrigerants. In one embodiment, the first and second refrigeration cycles are closed refrigeration cycles, while the third refrigeration cycle is an open refrigeration cycle that utilizes a portion of the processed natural gas as a source of the refrigerant. The third refrigeration cycle can include a multi-stage expansion cycle to provide additional cooling of the processed natural gas stream and reduce its pressure to near atmospheric pressure.
In the sequence of first, second, and third refrigeration cycles, the refrigerant having the highest boiling point can be utilized first, followed by a refrigerant having an intermediate boiling point, and finally by a refrigerant having the lowest boiling point. In one embodiment, the first refrigerant has a mid-boiling point within about 20, about 10, or 5° F. of the boiling point of pure propane at atmospheric pressure. The first refrigerant can contain predominately propane, propylene, or mixtures thereof. The first refrigerant can contain at least about 75 mole percent propane, at least 90 mole percent propane, or can consist essentially of propane. In one embodiment, the second refrigerant has a mid-boiling point within about 20, about 10, or 5° F. of the boiling point of pure ethylene at atmospheric pressure. The second refrigerant can contain predominately ethane, ethylene, or mixtures thereof. The second refrigerant can contain at least about 75 mole percent ethylene, at least 90 mole percent ethylene, or can consist essentially of ethylene. In one embodiment, the third refrigerant has a mid-boiling point within about 20, about 10, or 5° F. of the boiling point of pure methane at atmospheric pressure. The third refrigerant can contain at least about 50 mole percent methane, at least about 75 mole percent methane, at least 90 mole percent methane, or can consist essentially of methane. At least about 50, about 75, or 95 mole percent of the third refrigerant can originate from the processed natural gas stream.
The first refrigeration cycle can cool the natural gas in a plurality of cooling stages/steps (e.g., two to four cooling stages) by indirect heat exchange with the first refrigerant. Each indirect cooling stage of the refrigeration cycles can be carried out in a separate heat exchanger. In one embodiment, core-and-kettle heat exchangers are employed to facilitate indirect heat exchange in the first refrigeration cycle. After being cooled in the first refrigeration cycle, the temperature of the natural gas can be in the range of from about −45 to about −10° F., about −40 to about −15° F., or −20 to −30° F. A typical decrease in the natural gas temperature across the first refrigeration cycle may be in the range of from about 50 to about 210° F., about 75 to about 180° F., or 100 to 140° F.
The second refrigeration cycle can cool the natural gas in a plurality of cooling stages/steps (e.g., two to four cooling stages) by indirect heat exchange with the second refrigerant. In one embodiment, the indirect heat exchange cooling stages in the second refrigeration cycle can employ separate, core-and-kettle heat exchangers. Generally, the temperature drop across the second refrigeration cycle can be in the range of from about 50 to about 180° F., about 75 to about 150° F., or 100 to 120° F. In the final stage of the second refrigeration cycle, the processed natural gas stream can be condensed (i.e., liquefied) in major portion, preferably in its entirety, thereby producing a pressurized LNG-bearing stream. Generally, the process pressure at this location is only slightly lower than the pressure of the natural gas fed to the first stage of the first refrigeration cycle. After being cooled in the second refrigeration cycle, the temperature of the natural gas may be in the range of from about −205 to about −70° F., about −175 to about −95° F., or −140 to −125° F.
The third refrigeration cycle can include both an indirect heat exchange cooling section and an expansion-type cooling section. To facilitate indirect heat exchange, the third refrigeration cycle can employ at least one brazed-aluminum plate-fin heat exchanger. The total amount of cooling provided by indirect heat exchange in the third refrigeration cycle can be in the range of from about 5 to about 60° F., about 7 to about 50° F., or 10 to 40° F.
The expansion-type cooling section of the third refrigeration cycle can further cool the pressurized LNG-bearing stream via sequential pressure reduction to approximately atmospheric pressure. Such expansion-type cooling can be accomplished by flashing the LNG-bearing stream to thereby produce a two-phase vapor-liquid stream. When the third refrigeration cycle is an open refrigeration cycle, the expanded two-phase stream can be subjected to vapor-liquid separation and at least a portion of the separated vapor phase (i.e., the flash gas) can be employed as the third refrigerant to help cool the processed natural gas stream. The expansion of the pressurized LNG-bearing stream to near atmospheric pressure can be accomplished by using a plurality of expansion steps (i.e., two to four expansion steps) where each expansion step is carried out using an expander. Suitable expanders include, for example, either Joule-Thomson expansion valves or hydraulic expanders. In one embodiment, the third refrigeration cycle can employ three sequential expansion cooling steps, wherein each expansion step can be followed by a separation of the gas-liquid product. Each expansion-type cooling step can cool the LNG-bearing stream in the range of from about 10 to about 60° F., about 15 to about 50° F., or 25 to 35° F. The reduction in pressure across the first expansion step can be in the range of from about 80 to about 300 psia, about 130 to about 250 psia, or 175 to 195 psia. The pressure drop across the second expansion step can be in the range of from about 20 to about 110 psia, about 40 to about 90 psia, or 55 to 70 psia. The third expansion step can further reduce the pressure of the LNG-bearing stream by an amount in the range of from about 5 to about 50 psia, about 10 to about 40 psia, or 15 to 30 psia. The liquid fraction resulting from the final expansion stage is the final LNG product. Generally, the temperature of the final LNG product can be in the range of from about −200 to about −300° F., about −225 to about −275° F., or −240 to −260° F. The pressure of the final LNG product can be in the range of from about 0 to about 40 psia, about 10 to about 20 psia, or 12.5 to 17.5 psia.
The natural gas feed stream to the LNG process usually contains such quantities of C2+ components so as to result in the formation of a C2+ rich liquid in one or more of the cooling stages of the second refrigeration cycle. Generally, the sequential cooling of the natural gas in each cooling stage is controlled so as to remove as much of the C2 and higher molecular weight hydrocarbons as possible from the gas, thereby producing a vapor stream predominating in methane and a liquid stream containing significant amounts of ethane and heavier components. This liquid can be further processed via gas-liquid separators employed at strategic locations downstream of the cooling stages. In one embodiment, one objective of the gas/liquid separators is to maximize the rejection of the C5+ material to avoid freezing in downstream processing equipment. The gas/liquid separators may also be utilized to vary the amount of C2 through C4 components that remain in the natural gas product to affect certain characteristics of the finished LNG product. The exact configuration and operation of gas-liquid separators may be dependant on a number of parameters, such as the C2+ composition of the natural gas feed stream, the desired BTU content (i.e., heating value) of the LNG product, the value of the C2+ components for other applications, and other factors routinely considered by those skilled in the art of LNG plant and gas plant operation.
In one embodiment of the present invention, the LNG process can include natural gas liquids (NGL) integration within the LNG facility. One may significantly enhance the efficiency of LNG production and NGL recovery by integrating the two functions in one facility. In addition, the present invention can employ an integrated heavies removal/NGL recovery system that allows for prompt and economical variation in the BTU content (i.e., higher heating value (HHV)) of the LNG product stream so that various LNG markets can be served by one facility.
Accordingly, in one embodiment of the present invention, an LNG facility is provided that can be operated in different modes of operation to produce LNG and/or NGL products that meet different product specifications. For example, the LNG facility can be operated in a low-BTU mode to produce an LNG product having a low BTU content (e.g., 950-1060 BTU/SCF) or in a high-BTU mode to produce an LNG product having a high BTU content (e.g., 1070-1160 BTU/SCF). The LNG facility can also be operated in different modes of operation to produce different NGL products. For example, the LNG facility can be operated in a propane rejection mode to produce an NGL product having a low propane content (e.g., 0-20 mole percent) or in a propane recovery mode to produce an NGL product having a high propane content (e.g., 40-85 mole percent).
The average higher heating value (HHV) of LNG produced during different modes of operation of the LNG facility can differ from one another by at least about 10 BTU/SCF, at least about 20 BTU/SCF, or at least 50 BTU/SCF. Further, the average HHV of the LNG products produce by different modes of operation can vary by at least about 1 percent, at least about 3 percent, or at least 5 percent in the different modes of operation. In one embodiment, the difference in the average propane content of NGL produced during different modes of operation can be at least about 1 mole percent, at least about 2 mole percent, or at least 5 mole percent. The different modes of operation discussed herein are steady-state modes of operation, not operation during start-up or shut-down of the LNG facility. In one embodiment, each of the different steady-state modes of operation is carried out over a time period of at least one week, at least two weeks, or at least four weeks (as opposed to a lesser period of time that would typically be required for start-up or shut-down).
It is known that the HHV of produced LNG in conventional LNG plants may vary slightly over long periods of time do to changes in feed composition and/or changes in ambient conditions. However, in one embodiment, the present invention allows for relatively large and rapid adjustments in the HHV value of the LNG product and/or the propane content of the NGL product. To accomplish the relatively large and rapid adjustment in the HHV of the LNG product and/or the propane content of the NGL product, the LNG facility can be transition between the different modes of operation over a time period of less than 1 week, less than 3 days, less than 1 day, or less than 12 hours. In accordance with an embodiment of the present invention, the production of LNG does not cease during transitioning between different modes of operation. Rather, the LNG facility can be rapidly transitioned from one steady-state operating mode to another steady-state operating mode without requiring shut-down of the facility.
To transition the LNG facility from a first mode of operation to a second mode of operation, one or more operating parameters of the LNG facility can be adjusted. The operating parameter adjusted to transition the LNG facility between different modes of operation can be a non-feed operating parameter of the LNG facility (i.e., the transition between modes of operation is not caused by adjusting the composition of the feed to the LNG facility). For example, when the LNG facility includes a heavies removal/NGL recovery system that employs a distillation column to separate the processed natural gas stream into different components based on relative volatilities, the operating parameter adjusted to transition the LNG facility between different modes of operation can be an operating parameter of the distillation column. Such distillation column operating parameters may include, for example, column feed composition, column feed temperature, column overhead pressure, reflux stream flow rate, reflux stream composition, reflux stream temperature, stripping gas flow rate, stripping gas composition, and stripping gas temperature.
In one embodiment, the heavies removal/NGL recovery system of the LNG facility can employ a two column configuration. Such a system can include a first distillation column (e.g., a heavies removal column) and a second distillation column (e.g., a demethanizer, deethanizer, or depropanizer). Heavy liquids can be concentrated and removed from the bottom of the heavies removal column and can thereafter be routed to the second distillation column. The second column can be operated to stabilize the bottoms product and send lighter components overhead, eventually ending up in the LNG product. In accordance with one embodiment, the distillation columns are operated in a manner that produces only enough heavy material in the overhead to provide the LNG BTU content desired, as well as to stabilize the bottoms stream by removing undesired light components. In such a two column configuration, one or more operating parameters of one or both of the distillation columns can be adjusted to transition the LNG facility between different modes of operation. The various operating parameters that can be adjusted to transition the LNG facility between different modes of operation are discussed in detail below with reference to
LNG facilities capable of being operated in accordance with the present invention can have a variety of configurations. The flow schematics and apparatuses illustrated in
To facilitate an understanding of
The inventive LNG facilities illustrated in
As illustrated in
The operation of the LNG facility illustrate in
The cooled natural gas stream from high-stage propane chiller 14, also referred to herein as the methane-rich stream, flows via conduit 102 to a separation vessel 58 wherein gas and liquid phases are separated. The liquid phase, which can be rich in C3+ components, is removed via conduit 303. The vapor phase is removed via conduit 104 and fed to intermediate-stage propane chiller 16 wherein the stream is cooled via an indirect heat exchange means 62. The resultant vapor/liquid stream is then routed to low-stage propane chiller 18 via conduit 112 wherein it is cooled by an indirect heat exchange means 64. The cooled methane-rich stream then flows through conduit 114 and enters high-stage ethylene chiller 24, which will be discussed further in a subsequent section.
The propane gas from high-stage propane chiller 14 is returned to the high-stage inlet port of propane compressor 10 via conduit 306. The residual liquid propane is passed via conduit 308 through a pressure reduction means, illustrated here as expansion valve 72, whereupon an additional portion of the liquefied propane is flashed or vaporized. The resulting cooled, two-phase stream enters intermediate-stage propane chiller 16 by means of conduit 310, thereby providing coolant for chiller 16. The vapor portion of the propane refrigerant exits intermediate-stage propane chiller 16 via conduit 312 and is fed to the intermediate-stage inlet port of propane compressor 10. The liquid portion flows from intermediate-stage propane chiller 16 through conduit 314 and is passed through a pressure-reduction means, illustrated here as expansion valve 73, whereupon a portion of the propane refrigerant stream is vaporized. The resulting vapor/liquid stream then enters low-stage propane chiller 18 via conduit 316, wherein it acts as a coolant. The vaporized propane refrigerant stream then exits low-stage propane chiller 18 via conduit 318 and is routed to the low-stage inlet port of propane compressor 10, whereupon it is compressed and recycled through the previously described propane refrigeration cycle.
As previously noted, the ethylene refrigerant stream in conduit 202 is cooled in high-stage propane chiller 14 via indirect heat exchange means 8. The cooled ethylene refrigerant stream then exits high-stage propane chiller 14 via conduit 204. The partially condensed stream enters intermediate-stage propane chiller 16, wherein it is further cooled by an indirect heat exchange means 66. The two-phase ethylene stream is then routed to low-stage propane chiller 18 by means of conduit 206 wherein the stream is totally condensed or condensed nearly in its entirety via indirect heat exchange means 68. The ethylene refrigerant stream is then fed via conduit 208 to a separation vessel 70 wherein the vapor portion, if present, is removed via conduit 210. The liquid ethylene refrigerant is then fed to the ethylene economizer 30 by means of conduit 212. The ethylene refrigerant at this location in the process is generally at a temperature of about −24° F. and a pressure of about 285 psia.
Turning now to the ethylene refrigeration cycle illustrated in
The ethylene refrigerant vapor exits high-stage ethylene chiller 24 via conduit 216 and is routed back to the ethylene economizer 30, warmed via an indirect heat exchange means 76, and subsequently fed via conduit 218 to the high-stage inlet port of ethylene compressor 20. The liquid portion of the ethylene refrigerant stream exits high-stage ethylene chiller 24 via conduit 220 and is then further cooled in an indirect heat exchange means 78 of ethylene economizer 30. The resulting cooled ethylene stream exits ethylene economizer 30 via conduit 222 and passes through a pressure reduction means, illustrated here as expansion valve 80, whereupon a portion of the ethylene is flashed.
In a manner similar to high-stage ethylene chiller 24, the two-phase refrigerant stream enters intermediate-stage ethylene chiller 26 via conduit 224, wherein it acts as a coolant for the natural gas stream flowing through an indirect heat exchange means 84. The cooled methane-rich stream exiting intermediate-stage ethylene chiller 24 via conduit A is totally condensed or condensed nearly in its entirety. The stream is then routed to the heavies removal/NGL recovery system of the process in
The vapor and liquid portions of the ethylene refrigerant stream exit intermediate-stage ethylene chiller 26 via conduits 226 and 228, respectively. The gaseous stream in conduit 226 combines with a yet to be described ethylene vapor stream in conduit 238. The combined ethylene refrigerant stream enters ethylene economizer 30 via conduit 239, is warmed by an indirect heat exchange means 86, and is fed to the low-stage inlet port of ethylene compressor 20 via conduit 230. The effluent from the low-stage of the ethylene compressor 20 is routed to an inter-stage cooler 88, cooled, and returned to the high-stage port of the ethylene compressor 20. Preferably, the two compressor stages are a single module although they may each be a separate module, and the modules may be mechanically coupled to a common driver. The compressed ethylene product flows to ethylene cooler 22 via conduit 236 wherein it is cooled via indirect heat exchange with an external fluid (e.g., air or water). The resulting condensed ethylene stream is then introduced via conduit 202 to high-stage propane chiller 14 for additional cooling as previously noted.
The liquid portion of the ethylene refrigerant stream from intermediate-stage ethylene chiller 26 in conduit 228 enters low-stage ethylene chiller/condenser 28 and cools the methane-rich stream in conduit 120 via an indirect heat exchange means 90. The stream in conduit 120 is a combination of a heavies-depleted (i.e., light hydrocarbon rich) stream from the heavies removal/NGL recovery system of the process in conduit C and a recycled methane refrigerant stream in conduit 158. As noted previously, details of the heavies removal/NGL recovery system will be described in further detail below. The vaporized ethylene refrigerant from low-stage ethylene chiller/condenser 28 flows via conduit 238 and joins the ethylene vapors from the intermediate-stage ethylene chiller in conduit 226. The combined ethylene refrigerant vapor stream is then heated by the indirect heat exchange means 86 in the ethylene economizer 30 as described previously. The pressurized, LNG-bearing stream exiting the ethylene refrigeration cycle via conduit 122 can be at a temperature in the range of from about −200 to about −50° F., about −175 to about −100° F., or −150 to −125° F. and a pressure in the range from about 500 to about 700 psia, or 550 to 725 psia.
The pressurized, LNG-bearing stream is then routed to main methane economizer 36, wherein it is further cooled by an indirect heat exchange means 92. The stream exits through conduit 124 and enters the expansion-cooling section of the methane refrigeration cycle. The liquefied methane-rich stream is then passed through a pressure-reduction means, illustrated here as high-stage methane expander 40, whereupon a portion of the stream is vaporized. The resulting two-phase product enters high-stage methane flash drum 42 via conduit 163 and the gaseous and liquid phases are separated. The high-stage methane flash gas is transported to main methane economizer 36 via conduit 155 wherein it is heated via an indirect heat exchange means 93 and exits main methane economizer 36 via conduit 168 and enters the high-stage inlet port of methane compressor 32.
The liquid product from high-stage flash drum 42 enters secondary methane economizer 38 via conduit 166, wherein the stream is cooled via an indirect heat exchange means 39. The resulting cooled stream flows via conduit 170 to a pressure reduction means, illustrated here as intermediate-stage methane expander 44, wherein a portion of the liquefied methane stream is vaporized. The resulting two-phase stream in conduit 172 then enters intermediate-stage methane flash drum 46 wherein the liquid and vapor phases are separated and exit via conduits 176 and 178, respectively. The vapor portion enters secondary methane economizer 38, is heated by an indirect heat exchange means 41, and then reenters main methane economizer 36 via conduit 188. The stream is further heated by indirect heat exchange means 95 before being fed into the intermediate-stage inlet port of methane compressor 32 via conduit 190.
The liquid product from the bottom of intermediate-stage methane flash drum 46 then enters the final stage of the expansion cooling section as it is routed via conduit 176 through a pressure reduction means, illustrated here as low-stage methane expander 48, whereupon a portion of the liquid stream is vaporized. The cooled, mixed-phase product is routed via conduit 186 to low-stage methane flash drum 50, wherein the vapor and liquid portions are separated. The LNG product, which is at approximately atmospheric pressure, exits low-stage methane flash drum 50 via conduit 198 and is routed to storage, represented by LNG storage vessel 99.
As shown in
As previously noted, the methane refrigerant stream from high-stage propane chiller 14 in conduit 154 enters main methane economizer 36. The stream is then further cooled via indirect heat exchange means 98. The resulting methane refrigerant stream flows via conduit 158 and is combined with the heavies-depleted vapor stream in conduit C prior to entering low-stage ethylene chiller/condenser 28 via conduit 120, as previously discussed.
The operation of the heavies removal/NGL recovery system illustrated in
A predominantly methane overhead product exits first distillation column 452 via conduit 462 and passes through a pressure control means 412, which is preferably a flow control valve, and reenters the liquefaction stage via conduit C.
As shown in
As shown in
The overhead vapor product of second distillation column 454 flows via conduit 482 through a pressure control means 420, which is preferably a flow control valve, to economizing heat exchanger 402 via conduit 483. The stream is cooled and partially condensed via an indirect heat exchange means 422. This two-phase stream is then passed to a second distillation column reflux separation vessel 424 via conduit 486 wherein the liquid and vapor phases are separated. The liquid stream is refluxed back to second distillation column 454 by means of conduit 488. The vapor stream passes through conduit 490 and into economizing heat exchanger 402, wherein the vapor is cooled and partially condensed via an indirect heat exchange means 426. The stream exits economizing heat exchanger 402 via conduit 492 and is routed to cooler 428, wherein it is further cooled and condensed, preferably condensed in its entirety, via indirect heat exchange. Cooler 428 can be an external cooler, or can be a pass in one of the chillers (e.g., ethylene chiller 28) illustrated in
Generally, the characteristics of the final LNG product can be altered to meet the different specifications of two or more markets by manipulating one or more key process parameters, such as, for example, the temperature or pressure of process vessels or the temperature, pressure, flow, or composition of streams associated with the process vessels. Such associated streams include, for example, a column reflux stream, a column stripping gas stream, and a column feed stream. In order to affect changes to process variables, the configuration of related process equipment may be modified. For example, the number, arrangement, operation, and/or type of equipment utilized can be changed to achieve the desired result.
In accordance with one embodiment of the present invention, the higher heating value (HHV) of the LNG product can be adjusted by varying one or more operating parameters of the system illustrated in
There are a number of ways to affect the adjustments of items (1)-(15) listed above. For example, the amount of C2+ components contained in feed stream(s) 456 and/or 460 to first distillation column 452 can be adjusted using additional upstream separation techniques. For example, the temperature of feed streams 456,460 to first distillation column 452 can be lowered at least about 1° F. or at least 3° F. by adjusting flow rates in heat exchanger 402 or other upstream heat exchangers. For example, the flow rate of reflux stream 498 to first distillation column 452 can be increased by providing more cooling of overhead stream 149 of second distillation column 454 in heat exchanger 402 (pass 422). For example, the temperature of reflux stream 498 to first distillation column 452 can be lowered by at least 5° F. by providing more cooling in heat exchanger 402 (pass 426) or heat exchanger 428. For example, the amount of C2+ components contained in reflux stream 498 to first distillation column 452 can be increased by at least 10 mole percent by altering the operation of second distillation column 454. For example, the flow rate of stripping gas stream 466 to first distillation column 452 can be lowered via control valves (not shown). For example, the temperature of stripping gas stream 466 to first distillation column 452 can be lowered at least 5° F. by providing less heating in heat exchanger 402 (pass 414). For example, the overhead pressure of first distillation column can be increased by restricting overhead flow in line 462 via valve 412. For example, the amount of C3+ components contained in feed stream 470 to second distillation column 454 can be lowered by including additional separation means or combining a methane-rich stream between columns 452 and 454. For example, the temperature of feed stream 470 to second distillation column 454 can be lowered by providing additional cooling to the stream in conduit 470. For example, the flow rate of reflux stream 488 to second distillation column 454 can be increased by providing more cooling to overhead stream 482 of second distillation column 454 in heat exchanger 402 (pass 422). For example, the temperature of reflux stream 488 to second distillation column 454 can be lowered by providing more cooling to overhead stream 482 of second distillation column 454 in heat exchanger 402 (pass 422). For example, the flow rate of reboil stream 472 to second distillation column 454 can be lowered by decreasing the amount of heat transfer taking place in the reboiler of second distillation column 454. For example, the temperature of reboil stream 472 to second distillation column 454 can be lowered by decreasing the amount of heat transfer taking place in the reboiler of second distillation column 454. For example, the overhead pressure of second distillation column 454 can be increased by restricting overhead flow in line 482 via valve 420.
It should be understood that the HHV of the LNG product from the LNG facility of
Table 2, below, provides a summary of broad and narrow ranges for various properties of selected streams from
The main components of the propane and ethylene refrigeration cycles of the liquefaction stage represented by
The operation of the LNG facility illustrated in
The methane-rich stream then enters intermediate-stage ethylene chiller 26 in
In the methane refrigeration cycle of
Turning now to
The operation of the heavies removal/NGL recovery system of the inventive LNG facility presented in
The liquid stream exits feed surge vessel 506 via conduit 522, whereupon it combines with the liquid product from the bottom port of first distillation column 552 in conduit 526. The composite stream travels via conduit 528 to economizing heat exchanger 502, wherein it is heated via an indirect heat exchange means 514. The resulting stream feeds second distillation column 554 via conduit 530. The liquid product from the bottom port of second distillation column 554 is the final NGL product. In
A stream is drawn from a side port of second distillation column 554 via conduit 540. The stream enters heater 512, wherein it is heated (reboiled) via indirect heat exchange with an external fluid (e.g., steam or heat transfer fluid). The resulting vapor is returned to second distillation column 554 via conduit 542, wherein it is employed as a stripping gas. The vapor stream from the overhead port of second distillation column 554 travels by way of conduit 532 to economizing heat exchanger 502, wherein it is partially condensed via indirect heat exchange means 516. The resulting, partially liquefied stream is routed via conduit 534 to the second distillation column overhead surge vessel 508, wherein the vapor and liquid are separated.
The vapor stream exits overhead surge vessel 508 via conduit P in
According to one embodiment of the present invention, the heating value of the LNG product can be adjusted by varying one or more operating parameters of the system illustrated in
As detailed previously with respect to
Similarly to
A further embodiment of the inventive LNG facility capable of efficiently supplying LNG product to meet significantly different specifications of two or more markets is illustrated in
The main components of the liquefaction step of the inventive LNG facility shown in
The partially vaporized, methane-rich stream exits low-stage propane chiller 18 via conduit 114, whereupon a portion of the stream is routed via conduit D to the heavies removal/NGL recovery system of the LNG facility illustrated in
According to
In the methane refrigeration cycle of
Turning now to
The operation of the inventive system illustrated in
Two side streams via conduits 628 and 630 are drawn from first distillation column 652. The stream in conduit 628 enters economizing heat exchanger 602, wherein it is heated (reboiled) and at least partially vaporized via an indirect heat exchange means 618. The side stream in conduit 630 acts as a coolant for a yet-to-be-discussed overhead vapor product from second distillation column 654 in a condenser 620. The resulting, at least partially, and preferably totally, vaporized streams, combine in conduit 636 prior to reentering first distillation column 652. These primarily vaporized streams then act as a stripping gas in first distillation column 652.
The liquid product from the bottom port of first distillation column 652 feeds second distillation column 654 via conduit 638. A side stream is drawn from second distillation column 654 via conduit 666 and passes through heater 612, wherein the stream is reboiled (heated) via indirect heat exchange with an external fluid (e.g., steam or other heat transfer fluid). A portion of the stream vaporizes and is routed from heater 612 via conduit 668 to second distillation column 654, wherein it is employed as stripping gas. The remaining liquid flows from heat exchanger 612 through conduit 672 and combines with the liquid product from the bottom port of second distillation column 654 in conduit 670. The composite stream is the final NGL product, which can be, in one embodiment, predominantly made up of propane and heavier components. The NGL stream is routed via conduit 676 to further processing and/or storage.
The vapor product from the overhead port of second distillation column 654 exits via conduit 640 and is thereafter condensed via condenser 620 by indirect heat exchange with the side stream from first distillation column 652 in conduit 630 as described previously. The resulting cooled, at least partially condensed stream flows via conduit 642 to second distillation column separation vessel 604, wherein the vapor and liquid phases are separated. The liquid portion flows via conduit 662 to the suction of a reflux pump 606. The stream then discharges into conduit 664 and is employed as a first distillation column 652 reflux stream.
The vapor stream exits second distillation column separation vessel 604 via conduit 634. One portion of the vapor stream can be routed by way of conduit 644 for use in other applications or as fuel. Another fraction of the vapor product can be routed via conduit G to the high-stage inlet port of methane compressor 32 in
According to
As mentioned previously, prior to entering high-stage propane chiller 14, a portion of the methane refrigerant stream in conduit 152 is routed via conduit E to the heavies removal/NGL recovery system in
Referring now to
Referring now to
The operation of the system illustrated in
Still another embodiment of the heavies removal/NGL recovery system of the LNG facility is illustrated in
The operation of the system illustrated in
According to one embodiment of the present invention, the HHV of the LNG product can be adjusted by varying one or more operating parameters of the system illustrated in
Similarly to
Still another embodiment of the inventive LNG facility is illustrated in
Referring now to
According to
In the methane refrigeration cycle of
Turning now to
The operation of the system illustrated in
As previously noted in
As also previously noted in
According to
Referring now to the remaining portion of second distillation column bottom liquid product in conduit 730 in
A side stream is drawn from second distillation column 754 via conduit 766 and sent to a heater 712, wherein the stream is heated (reboiled) via indirect heat exchange with an external fluid (e.g., steam or heat transfer fluid). The vaporized portion of the stream is returned to second distillation column 754 via conduit 768, wherein it is employed as a stripping gas. The resulting liquid portion exits second distillation column reboiler 712 via conduit 727, whereupon it combines with the liquid product from the bottom port of second distillation column 754 in conduit 770. The resulting composite stream in conduit 776 is the final NGL product. According to one embodiment, the NGL product can be rich in propane and heavier components. According to another embodiment of the present invention, second distillation column 754 may be operated in such a way as to maximize C5+ component recovery in the final NGL product. By maximizing the C5+ component recovery in the NGL product, an LNG product with a relatively higher HHV can be produced.
The vapor product from the overhead port of second distillation column 754 exits via conduit 778, whereupon the stream is cooled and at least partially condensed by condenser 720. The resulting stream exits condenser 720 via conduit 780 and enters second distillation column separation vessel 704, wherein the vapor and liquid phases are separated. The vapor portion, comprised primarily of ethane, is routed via conduit G to
According to one embodiment of the present invention, the heating values of the LNG product can be adjusted by varying one or more operating parameters of the system illustrated in
Similarly to
The main components of the system in
The remaining methane-rich stream enters high-stage ethylene chiller 24, wherein it is cooled via indirect heat exchange means 82. The resulting stream is routed via conduit B to the heavies removal/NGL recovery system in
In the methane refrigeration cycle illustrated in
Turning now to
The operation of the system illustrated in
In accordance to one embodiment of the present invention, the heating values of the LNG product can be adjusted by varying one or more operating parameters of the system illustrated in
Similarly to
Yet another embodiment of the inventive facility capable of supplying an LNG product with significantly different specifications meeting the needs of two or more different markets is presented in
The main components of the system in
The remaining methane-rich stream in
In the methane refrigeration cycle of
Turning now to
The operation of the system illustrated in
A portion of the methane recycle stream in
The liquid product from the bottom port of first distillation column 952 exits via conduit 942. A portion of the stream is then routed via conduit 944 to intermediate-stage separator 956, wherein the vapor and liquid phases are separated. The vapor phase exits via conduit 946 and is routed to intermediate stage separator economizing heat exchanger 906, wherein the stream is warmed via an indirect heat exchange means 932. The resulting stream exits intermediate stage separator economizing heat exchanger 906 and is routed via conduit G to the high-stage inlet port of methane compressor 32 in
According to
Turning now to the remainder of the bottom liquid product from first distillation column 952 in conduit 942, the stream enters first distillation column economizing heat exchanger 902, wherein it is cooled via indirect heat exchange means 918. The resulting cooled liquid is travels via conduit 976 to a condenser 920, wherein the stream in conduit 976 acts as a coolant for a yet to be discussed stream in conduit 978. After exiting condenser 920, the resulting, heated stream in conduit 968 divides into two streams in conduits 964 and 974. The portion of the stream in conduit 964 combines with the stream exiting main economizing heat exchanger 904 in conduit 960 prior to entering intermediate-stage separation vessel 956, as discussed previously. The portion of the heated stream in conduit 974 combines with the liquid phase exiting intermediate separation vessel 956 via conduit 948. The resulting composite stream enters second distillation column 954 via conduit 980.
The vapor product from the overhead port of second distillation column 954 exits via conduit 978 and enters condenser 920, wherein the stream is condensed via indirect heat exchange with the liquid stream from the bottom port of first distillation column 952 in conduit 976, as discussed previously. The at least partially condensed stream travels via conduit 982 to second distillation column separation vessel 908, wherein the vapor and liquid phases are separated. The predominantly ethane-rich vapor phase exits second distillation column separation vessel 908 and is routed for further processing and/or storage via conduit 984. The liquid phase leaves second distillation column separation vessel 908 via conduit 986 and enters the suction of a reflux pump 910. Reflux pump 910 discharges the stream as reflux to second distillation column 954 via conduit 988.
A side stream is drawn from second distillation column 954 via conduit 990. The stream is routed to a heater 912, wherein it is heated (reboiled) via indirect heat exchange with an external fluid (e.g., steam or heat transfer fluid). The vaporized portion of the stream is returned to second distillation column 954 via conduit 992, wherein it is employed as a stripping gas. The resulting liquid portion exits second distillation column reboiler 912 via conduit 994, whereupon it combines with the liquid product from the bottom port of second distillation column 954 in conduit 996. The resulting composite stream is the final NGL product. The final NGL product is comprised of ethane and heavier components and is routed to storage and/or further processing via conduit 998.
In accordance to one embodiment of the present invention, the heating values of the LNG product can be adjusted by varying one or more operating parameters of the system illustrated in
Similarly to
Still another embodiment of the inventive LNG facility is illustrated in
The main components of the system in
The remaining methane-rich stream enters high-stage ethylene chiller 24, wherein it is cooled via indirect heat exchange means 82. The resulting stream is routed via conduit B to the heavies removal/NGL recovery system in
In the methane refrigeration cycle illustrated in
Turning now to
The operation of the system illustrated in
In accordance to one embodiment of the present invention, the heating values of the LNG product can be adjusted by varying one or more operating parameters of the system illustrated in
Similarly to
In one embodiment of the present invention, the LNG production systems illustrated in
The preferred forms of the invention described above are to be used as illustration only, and should not be used in a limiting sense to interpret the scope of the present invention. Obvious modifications to the exemplary embodiments, set forth above, could be readily made by those skilled in the art without departing from the spirit of the present invention.
The inventors hereby state their intent to rely on the Doctrine of Equivalents to determine and assess the reasonably fair scope of the present invention as pertains to any apparatus not materially departing from but outside the literal scope of the invention as set forth in the following claims.
Numerical Ranges
The present description uses numerical ranges to quantify certain parameters relating to the invention. It should be understood that when numerical ranges are provided, such ranges are to be construed as providing literal support for claim limitations that only recite the lower value of the range as well as claims limitation that only recite the upper value of the range. For example, a disclosed numerical range of 10 to 100 provides literal support for a claim reciting “greater than 10” (with no upper bounds) and a claim reciting “less than 100” (with no lower bounds).
The present description uses specific numerical values to quantify certain parameters relating to the invention, where the specific numerical values are not expressly part of a numerical range. It should be understood that each specific numerical value provided herein is to be construed as providing literal support for a broad, intermediate, and narrow range. The broad range associated with each specific numerical value is the numerical value plus and minus 60 percent of the numerical value, rounded to two significant digits. The intermediate range associated with each specific numerical value is the numerical value plus and minus 30 percent of the numerical value, rounded to two significant digits. The narrow range associated with each specific numerical value is the numerical value plus and minus 15 percent of the numerical value, rounded to two significant digits. For example, if the specification describes a specific temperature of 62° F., such a description provides literal support for a broad numerical range of 25° F. to 99° F. (62° F.+/−37° F.), an intermediate numerical range of 43° F. to 81° F. (62° F.+/−19° F.), and a narrow numerical range of 53° F. to 71° F. (62° F.+/−9° F.). These broad, intermediate, and narrow numerical ranges should be applied not only to the specific values, but should also be applied to differences between these specific values. Thus, if the specification describes a first pressure of 110 psia and a second pressure of 48 psia (a difference of 62 psi), the broad, intermediate, and narrow ranges for the pressure difference between these two streams would be 25 to 99 psi, 43 to 81 psi, and 53 to 71 psi, respectively.
Definitions
As used herein, the term “natural gas” means a stream containing at least 65 mole percent methane, with the balance being ethane, higher hydrocarbons, nitrogen, carbon dioxide, and/or a minor amount of other contaminants such as mercury, hydrogen sulfide, and mercaptan.
As used herein, the term “mixed refrigerant” means a refrigerant containing a plurality of different components, where no single component makes up more than 75 mole percent of the refrigerant.
As used herein, the term “pure component refrigerant” means a refrigerant that is not a mixed refrigerant.
As used herein, the term “cascade refrigeration process” means a refrigeration process that employs a plurality of refrigeration cycles, each employing a different pure component refrigerant to successively cool natural gas.
As used herein, the term “open-cycle cascaded refrigeration process” refers to a cascaded refrigeration process comprising at least one closed refrigeration cycle and one open refrigeration cycle, where the boiling point of the refrigerant employed in the open cycle is less than the boiling point of the refrigerant employed in the closed cycle, and a portion of the cooling duty to condense the open-cycle refrigerant is provided by one or more of the closed cycles. In one embodiment of the present invention, a predominately methane stream is employed as the refrigerant in the open refrigeration cycle. This predominantly methane stream originates from the processed natural gas feed stream and can include the compressed open methane cycle gas streams.
As used herein, the term “expansion-type cooling” refers to cooling which occurs when the pressure of a gas, liquid, or two-phase system is decreased by passage through a pressure reduction means. In one embodiment, the expansion means is a Joule-Thompson expansion valve. In another embodiment of the present invention, the expansion means is a hydraulic or gas expander.
As used herein, the term “mid-boiling point” refers to the temperature at which half of the weight of a mixture of physical components has been vaporized (i.e., boiled off) at a specific pressure.
As used herein, the term “indirect heat exchange” refers to a process wherein the refrigerant cools the substance to be cooled without actual physical contact between the refrigerating agent and the substance to be cooled. Core-in-kettle heat exchangers and brazed aluminum plate-fin heat exchangers are specific examples of equipment that facilitate indirect heat exchange.
As used herein, the terms “economizer” or “economizing heat exchanger” refer to a configuration utilizing a plurality of heat exchangers employing indirect heat exchange means to efficiently transfer heat between process streams. Generally, economizers minimize outside energy inputs by heat integrating process streams with each other.
As used herein, the term “higher heating value” or “HHV” refers to a measure of the heat released when an LNG product is combusted, accounting for the energy required to vaporize the water that results from the combustion reaction.
As used herein, the term “BTU content” is synonymous with the term “higher heating value.”
As used herein, the term “distillation column” or “separator” refer to a device for separating a stream based on relative volatility.
As used herein, the term “steady state operation” shall mean periods of relatively steady and continuous operation between start-up and shut-down.
As used herein, the term “non-feed operating parameter” shall mean any operating parameter of an item of equipment or a facility other than the composition of the main feed(s) to that item of equipment or facility.
As used herein, the terms “natural gas liquids” or “NGL” refer to mixtures of hydrocarbons whose components are, for example, typically heavier than ethane. Some examples of hydrocarbon components of NGL streams include propane, butane, and pentane isomers, benzene, toluene, and other aromatic molecules. Ethane may also be included in an NGL mixture.
As used herein, the terms “upstream” and “downstream” refer to the relative positions of various components of a natural gas liquefaction facility along the main flow path of natural gas through the plant.
As used herein, the terms “predominantly,” “primarily,” “principally,” and “in major portion,” when used to describe the presence of a particular component of a fluid stream, means that the fluid stream comprises at least 50 mole percent of the stated component. For example, a “predominantly” methane stream, a “primarily” methane stream, a stream “principally” comprised of methane, or a stream comprised “in major portion” of methane each denote a stream comprising at least 50 mole percent methane.
As used herein, the term “and/or,” when used in a list of two or more items, means that any one of the listed items can be employed by itself, or any combination of two or more of the listed items can be employed. For example, if a composition is described as containing components A, B, and/or C, the composition can contain A alone; B alone; C alone; A and B in combination; A and C in combination; B and C in combination; or A, B, and C in combination.
As used herein, the terms “comprising,” “comprises,” and “comprise” are open-ended transition terms used to transition from a subject recited before the term to one or elements recited after the term, where the element or elements listed after the transition term are not necessarily the only elements that make up of the subject.
As used herein, the terms “including,” “includes,” and “include” have the same open-ended meaning as “comprising,” “comprises,” and “comprise.”
As used herein, the terms “having,” “has,” and “have” have the same open-ended meaning as “comprising,” “comprises,” and “comprise.”
As used herein, the terms “containing,” “contains,” and “contain” have the same open-ended meaning as “comprising,” “comprises,” and “comprise.”
As used herein, the terms “a,” “an,” “the,” and “said” means one or more.
The preferred forms of the invention described above are to be used as illustration only, and should not be used in a limiting sense to interpret the scope of the present invention. Obvious modifications to the exemplary embodiments, set forth above, could be readily made by those skilled in the art without departing from the spirit of the present invention.
Claims
1. A process for producing liquefied natural gas (LNG), said process comprising:
- (a) operating an LNG facility in a first mode of operation to thereby produce a first LNG product;
- (b) adjusting at least one non-feed operating parameter of the LNG facility so that the LNG facility operates in a second mode of operation; and
- (c) operating the LNG facility in the second mode of operation to thereby produce a second LNG product, wherein said first and second modes of operation are not carried out during start-up or shut-down of the LNG facility, wherein said operating of steps (a) and (c) optionally includes producing first and second natural gas liquids (NGL) products respectively, wherein the average higher heating value (HHV) of the second LNG product is at least about 10 BTU/SCF different than the average HHV of the first LNG product and/or the average propane content of the second NGL product is at least about 1 mole percent different than the average propane content of the first NGL product.
2. The process of claim 1, wherein said adjusting of step (b) includes transitioning said LNG facility from said first mode of operation to said second mode of operation without ceasing the production of LNG.
3. The process of claim 1, wherein said adjusting of step (b) includes adjusting at least one operating parameter of a distillation column of the LNG facility.
4. The process of claim 3, wherein said operating parameter of said distillation column includes at least one operating parameter selected from the group consisting of column feed composition, column feed temperature, column overhead pressure, reflux stream flow rate, reflux stream composition, reflux stream temperature, stripping gas flow rate, stripping gas composition, and stripping gas temperature.
5. The process of claim 3, wherein said operating of steps (a) and (c) includes using said distillation column to separate a stream into a relatively more volatile fraction and a relatively less volatile fraction, wherein said first and second LNG products comprise at least a portion of said relatively more volatile fraction and/or said first and second NGL products comprise at least a portion of said relatively less volatile fraction.
6. The process of claim 5, wherein said first and second LNG products comprise at least a portion of said relatively more volatile fraction and said first and second NGL products comprise at least a portion of said relatively less volatile fraction.
7. The process of claim 1, wherein said operating of steps (a) and (c) includes cooling a natural gas feed stream, separating the cooled natural gas feed stream into a first relatively more volatile fraction and a first relatively less volatile fraction using a first distillation column, and further cooling at least a portion of the first relatively more volatile fraction to thereby produce at least a portion of the first and second LNG products.
8. The process of claim 7, wherein said operating of steps (a) and (c) includes separating at least a portion of the first relatively less volatile fraction into a second relatively more volatile fraction and a second relatively less volatile fraction using a second distillation column.
9. The process of claim 8, wherein said operating of steps (a) and (c) include further cooling at least a portion of the second relatively more volatile fraction to thereby produce at least a portion of said first and second LNG products.
10. The process of claim 8, wherein said first and second NGL products comprise at least a portion of the second relatively less volatile fraction.
11. The process of claim 7, wherein at least a portion of said cooling of the natural gas feed stream is carried out using a first refrigeration cycle employing a first refrigerant having a mid-boiling point within about 20° F. of the boiling point of pure propane at atmospheric pressure.
12. The process of claim 11, wherein at least a portion of said further cooling of the first relatively more volatile fraction is carried out using a second refrigeration cycle employing a second refrigerant having a mid-boiling point within about 20° F. of the boiling point of pure methane at atmospheric pressure.
13. The process of claim 12, wherein at least a portion of said further cooling of the first relatively more volatile fraction is carried out using a third refrigeration cycle employing a third refrigerant having a mid-boiling point within about 20° F. of pure ethylene at atmospheric pressure.
14. The process of claim 13, wherein said first, second, and third refrigerants are pure component refrigerants.
15. The process of claim 1, wherein said operating of steps (a) and (c) includes using a first distillation column to separate a first stream into a first relatively more volatile fraction and a first relatively less volatile fraction and using a second distillation column to separate at least a portion of said first relatively less volatile fraction into a second relatively more volatile fraction and a second relatively less volatile fraction.
16. The process of claim 15, wherein said first and second LNG products comprise at least a portion of the first and second relatively more volatile fractions.
17. The process of claim 15, wherein said first and second NGL products comprise at least a portion of the second relatively less volatile fraction.
18. The process of claim 15, wherein said at least one non-feed operating parameter is an operating parameter of the first and/or second distillation column.
19. The process of claim 17, wherein said first distillation column is refluxed with at least a portion of the second relatively more volatile fraction.
20. The process of claim 19, wherein said adjusting of step (b) includes adjusting the temperature and/or flow rate of the reflux into the first distillation column.
21. The process of claim 1, wherein the average HHV of the second LNG product is at least about 10 BTU/SCF different than the average HHV of the first LNG product.
22. The process of claim 1, wherein the average propane content of the second NGL product is at least about 1 mole percent different than the average propane content of the first NGL product.
23. The process of claim 1, wherein the average HHV of the second LNG product is at least about 20 BTU/SCF different than the average HHV of the first LNG product.
24. The process of claim 1, wherein the average propane content of the second NGL product is at least about 2 mole percent different than the average propane content of the first NGL product.
25. The process of claim 1, wherein said first LNG product is produced over a first production time period of at least one week, wherein said second LNG product is produced over a second production time period of at least one week, wherein said first and second production time periods are separated by a transition time period of less than one week.
26. The process of claim 1, wherein said transition time period is less than one day.
27. A method of varying the heating value of LNG produced from an LNG facility, said method comprising:
- (a) cooling natural gas by indirect heat exchange to thereby produce a first cooled stream;
- (b) using a first distillation column to separate at least a portion of the first cooled stream into a first relatively more volatile fraction and a first relatively less volatile fraction;
- (c) cooling at least a portion of said first relatively more volatile fraction to thereby produce LNG; and
- (d) adjusting at least one operating parameter of the first distillation column to thereby vary the higher heating value (HHV) of the produced LNG by at least about 1 percent over a time period of less than about 72 hours.
28. The method according to claim 27, wherein said at least one operating parameter of the first distillation column is selected from the group consisting of temperature of the first cooled stream, composition of the first cooled stream, and overhead pressure of the first distillation column.
29. The method according to claim 27, wherein said at least one operating parameter of the first distillation is a non-feed operating parameter.
30. The method according to claim 27, wherein step (d) includes adjusting the overhead pressure in the first distillation column.
31. The method according to claim 27, wherein step (d) includes adjusting the temperature of the first cooled stream prior to introduction into the first distillation column.
32. The method according to claim 31, wherein step (d) includes reducing the temperature of the first cooled stream to thereby lower the HHV of the produced LNG.
33. The method according to claim 32, wherein said first cooled stream has a temperature in the range of from about −125 to about −50° F. when introduced into the first distillation column, wherein step (d) includes reducing the temperature of the first cooled stream by at least about 1° F.
34. The method according to claim 32, wherein said first cooled stream has a temperature in the range of from about −115 to about −65° F. when introduced into the first distillation column, wherein step (d) includes reducing the temperature of the first cooled stream by at least about 3° F.
35. The method according to claim 27, wherein step (b) includes introducing a predominately vapor stripping gas stream into a lower section of the first distillation column.
36. The method according to claim 35, wherein said at least one operating parameter of the first distillation column is selected from the group consisting of flow rate of the stripping gas stream, temperature of the stripping gas stream, composition of the stripping gas stream, temperature of the first cooled stream, composition of the first cooled stream, and overhead pressure of the first distillation column.
37. The method according to claim 35, wherein step (d) includes varying the flow rate of the stripping gas stream to the first distillation column.
38. The method according to claim 35, wherein step (d) includes lowering the flow rate of the stripping gas stream to the first distillation column to thereby lower the HHV of the produced LNG.
39. The method according to claim 35, wherein said first cooled stream has a temperature in the range of from about −125 to about −50° F. when introduced into the first distillation column, wherein said stripping gas stream has a temperature in the range from about −50 to about 100° F. when introduced into the first distillation column, wherein step (d) includes lowering the temperature of the stripping gas stream by at least 5° F. to thereby lower the HHV of the produced LNG.
40. The method according to claim 27, wherein step (b) includes introducing a predominately liquid reflux stream into an upper section of the first distillation column.
41. The method according to claim 40, wherein said at least one operating parameter of the first distillation column is selected from the group consisting of flow rate of the reflux stream, temperature of the reflux stream, composition of the reflux stream, temperature of the first cooled stream, composition of the first cooled stream, and overhead pressure in the first distillation column.
42. The method according to claim 40, wherein step (d) includes varying the flow rate of the reflux stream to the first distillation column.
43. The method according to claim 40, wherein step (d) includes increasing the flow rate of the reflux stream to the first distillation column to thereby lower the HHV of the produced LNG.
44. The method according to claim 40, wherein step (d) includes varying the C2+ content of the reflux stream.
45. The method according to claim 40, wherein step (d) includes increasing the C2+ content of the reflux stream from an initial C2+ content to an adjusted C2+ content to thereby lower the HHV of the LNG product.
46. The method according to claim 45, wherein said adjusted C2+ content is at least about 10 percent more than the initial C2+ content on a molar basis.
47. The method according to claim 45, wherein said initial C2+ content is less than about 75 mole percent and said adjusted C2+ content is at least about 25 mole percent.
48. The method according to claim 40, wherein step (d) includes varying the temperature of the reflux stream.
49. The method according to claim 40, wherein said first cooled stream has a temperature in the range of from about −125 to about −50° F. when introduced into the first distillation column, wherein said reflux stream has a temperature in the range from about −180 to about −80° F. when introduced into the first distillation column, wherein step (d) includes lowering the temperature of the reflux stream by a least about 5° F. to thereby lower the HHV of the produced LNG.
50. The method according to claim 40, wherein step (d) includes varying the flow rate of the LNG stream to the first distillation column.
51. The method according to claim 40, wherein step (d) includes increasing the flow rate of the LNG stream to the first distillation column to thereby lower the HHV of the produced LNG.
52. The method according to claim 27, further comprising:
- (e) using a second distillation column to separate at least a portion of said first relatively less volatile fraction into a second relatively more volatile fraction and a second relatively less volatile fraction.
53. The method according to claim 52, further comprising:
- (f) introducing at least a portion of said second relatively less volatile fraction into an upper section of said first distillation column as a reflux stream.
54. The method according to claim 52, further comprising:
- (g) adjusting at least one operating parameter of the second distillation column to thereby vary the HHV of the produced LNG.
55. The method according to claim 54, wherein said at least one operating parameter of the second distillation column is selected from the group consisting of temperature of said first relatively less volatile fraction introduced into the second distillation column, composition of said first relatively less volatile fraction introduced into the second distillation column, and overhead pressure in the second distillation column.
56. The method according to claim 54, further comprising:
- (h) introducing at least a portion of said second relatively less volatile fraction into an upper section of said first distillation column as a reflux stream, wherein step (g) includes varying the C2+ content of the reflux stream.
57. The method according to claim 27, wherein step (a) includes using a first refrigeration cycle employing a first refrigerant comprising predominantly propane to cool at least a portion of the natural gas.
58. The method according to claim 57, wherein step (c) includes using a second refrigeration cycle employing a second refrigerant comprising predominantly ethane, ethylene, and/or methane to cool at least a portion of the first separated stream.
59. The method according to claim 58, wherein said second refrigerant comprises predominately methane.
60. The method according to claim 59, further comprising:
- (i) using a third refrigeration cycle employing a third refrigerant comprising predominately ethane and/or ethylene to cool at least a portion of the first cooled stream and/or at least a portion of the first separated stream.
61. The method according to claim 27, wherein step (d) includes varying the HHV of the produced LNG by at least about 3 percent over a time period of less than about 24 hours.
62. The method according to claim 27, wherein step (d) includes varying the HHV of the produced LNG by at least 5 percent over a time period of less than 12 hours.
Type: Application
Filed: Jun 23, 2006
Publication Date: Jan 18, 2007
Inventors: Wesley Qualls (Katy, TX), Weldon Ransbarger (Houston, TX), Shawa Huang (Spring, TX), Jame Yao (Sugar Land, TX), Doug Elliot (Houston, TX), Jong Chen (Suger Land, TX), Rong-Jwyn Lee (Suger Land, TX)
Application Number: 11/426,026
International Classification: F25J 1/00 (20060101); F25J 3/00 (20060101);
