Patent Application
20120000246
This invention relates to a process for the separation of ethane and heavier hydrocarbons or propane and heavier hydrocarbons from liquefied natural gas (hereinafter referred to as LNG) combined with the separation of a gas containing hydrocarbons to provide a volatile methane-rich gas stream and a less volatile natural gas liquids (NGL) or liquefied petroleum gas (LPG) stream.
As an alternative to transportation in pipelines, natural gas at remote locations is sometimes liquefied and transported in special LNG tankers to appropriate LNG receiving and storage terminals. The LNG can then be re-vaporized and used as a gaseous fuel in the same fashion as natural gas. Although LNG usually has a major proportion of methane, i.e., methane comprises at least 50 mole percent of the LNG, it also contains relatively lesser amounts of heavier hydrocarbons such as ethane, propane, butanes, and the like, as well as nitrogen. It is often necessary to separate some or all of the heavier hydrocarbons from the methane in the LNG so that the gaseous fuel resulting from vaporizing the LNG conforms to pipeline specifications for heating value. In addition, it is often also desirable to separate the heavier hydrocarbons from the methane and ethane because these hydrocarbons have a higher value as liquid products (for use as petrochemical feedstocks, as an example) than their value as fuel.
Although there are many processes which may be used to separate ethane and/or propane and heavier hydrocarbons from LNG, these processes often must compromise between high recovery, low utility costs, and process simplicity (and hence low capital investment). U.S. Pat. Nos. 2,952,984; 3,837,172; 5,114,451; and 7,155,931 describe relevant LNG processes capable of ethane or propane recovery while producing the lean LNG as a vapor stream that is thereafter compressed to delivery pressure to enter a gas distribution network. However, lower utility costs may be possible if the lean LNG is instead produced as a liquid stream that can be pumped (rather than compressed) to the delivery pressure of the gas distribution network, with the lean LNG subsequently vaporized using a low level source of external heat or other means. U.S. Pat. Nos. 6,604,380; 6,907,752; 6,941,771; 7,069,743; and 7,216,507 and co-pending application Ser. Nos. 11/749,268 and 12/060,362 describe such processes.
Economics and logistics often dictate that LNG receiving terminals be located close to the natural gas transmission lines that will transport the re-vaporized LNG to consumers. In many cases, these areas also have plants for processing natural gas produced in the region to recover the heavier hydrocarbons contained in the natural gas. Available processes for separating these heavier hydrocarbons include those based upon cooling and refrigeration of gas, oil absorption, and refrigerated oil absorption. Additionally, cryogenic processes have become popular because of the availability of economical equipment that produces power while simultaneously expanding and extracting heat from the gas being processed. Depending upon the pressure of the gas source, the richness (ethane, ethylene, and heavier hydrocarbons content) of the gas, and the desired end products, each of these processes or a combination thereof may be employed.
The cryogenic expansion process is now generally preferred for natural gas liquids recovery because it provides maximum simplicity with ease of startup, operating flexibility, good efficiency, safety, and good reliability. U.S. Pat. Nos. 3,292,380; 4,061,481; 4,140,504; 4,157,904; 4,171,964; 4,185,978; 4,251,249; 4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063; 4,690,702; 4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,566,554; 5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664; 6,182,469; 6,578,379; 6,712,880; 6,915,662; 7,191,617; 7,219,513; reissue U.S. Pat. No. 33,408; and co-pending application Ser. Nos. 11/430,412; 11/839,693; 11/971,491; and 12/206,230 describe relevant processes (although the description of the present invention is based on different processing conditions than those described in the cited U.S. patents).
The present invention is generally concerned with the integrated recovery of propylene, propane, and heavier hydrocarbons from such LNG and gas streams. It uses a novel process arrangement to integrate the heating of the LNG stream and the cooling of the gas stream to eliminate the need for a separate vaporizer and the need for external refrigeration, allowing high C3 component recovery while keeping the processing equipment simple and the capital investment low. Further, the present invention offers a reduction in the utilities (power and heat) required to process the LNG and gas streams, resulting in lower operating costs than other processes, and also offering significant reduction in capital investment.
Heretofore, assignee's co-pending application Ser. No. 12/060,362 could be used to recover C3 components and heavier hydrocarbon components in plants processing LNG, while assignee's U.S. Pat. No. 5,799,507 has been used to recover C3 components and heavier hydrocarbon components in plants processing natural gas. Surprisingly, applicants have found that by integrating certain features of the assignee's co-pending application Ser. No. 12/060,362 with certain features of the assignee's U.S. Pat. No. 5,799,507, extremely high C3 component recovery levels can be accomplished using less energy than that required by individual plants to process the LNG and natural gas separately.
A typical analysis of an LNG stream to be processed in accordance with this invention would be, in approximate mole percent, 92.2% methane, 6.0% ethane and other C2 components, 1.1% propane and other C3 components, and traces of butanes plus, with the balance made up of nitrogen. A typical analysis of a gas stream to be processed in accordance with this invention would be, in approximate mole percent, 80.1% methane, 9.5% ethane and other C2 components, 5.6% propane and other C3 components, 1.3% iso-butane, 1.1% normal butane, 0.8% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present.
For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings:
In the following explanation of the above figures, tables are provided summarizing flow rates calculated for representative process conditions. In the tables appearing herein, the values for flow rates (in moles per hour) have been rounded to the nearest whole number for convenience. The total stream rates shown in the tables include all non-hydrocarbon components and hence are generally larger than the sum of the stream flow rates for the hydrocarbon components. Temperatures indicated are approximate values rounded to the nearest degree. It should also be noted that the process design calculations performed for the purpose of comparing the processes depicted in the figures are based on the assumption of no heat leak from (or to) the surroundings to (or from) the process. The quality of commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art.
For convenience, process parameters are reported in both the traditional British units and in the units of the Système International d'Unités (SI). The molar flow rates given in the tables may be interpreted as either pound moles per hour or kilogram moles per hour. The energy consumptions reported as horsepower (HP) and/or thousand British Thermal Units per hour (MBTU/Hr) correspond to the stated molar flow rates in pound moles per hour. The energy consumptions reported as kilowatts (kW) correspond to the stated molar flow rates in kilogram moles per hour.
The inlet gas stream 31 is cooled in heat exchanger 12 by heat exchange with a portion (stream 72a) of partially warmed LNG at −173° F. [−114° C.] and cool residue vapor stream 38. The cooled stream 31a enters separator 13 at −76° F. [−60° C.] and 584 psia [4,027 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35). Liquid stream 35 is flash expanded through an appropriate expansion device, such as expansion valve 17, to the operating pressure (approximately 450 psia [3,101 kPa(a)]) of fractionation tower 20. The expanded stream 35a leaving expansion valve 17 reaches a temperature of −88° F. [−67° C.] and is supplied to fractionation tower 20 at a first mid-column feed point.
The vapor from separator 13 (stream 34) enters a work expansion machine 10 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 10 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −96° F. [−71° C.]. The typical commercially available expanders are capable of recovering on the order of 80-88% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item 11) that can be used to re-compress the heated residue vapor (stream 38a), for example. The expanded stream 34a is supplied to fractionation tower 20 at a second mid-column feed point.
The deethanizer in tower 20 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The column also includes one or more reboilers (such as reboiler 19) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 41, of methane, C2 components, and lighter components. Liquid product stream 41 exits the bottom of the tower at 210° F. [99° C.], based on a typical specification of an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product.
Overhead distillation stream 43 is withdrawn from the upper section of fractionation tower 20 at −87° F. [−66° C.] and is divided into two portions, streams 44 and 47. The first portion, stream 44, flows to reflux condenser 23 where it is cooled to −237° F. [−149° C.] and totally condensed by heat exchange with a portion (stream 72) of the cold LNG (stream 71a). Condensed stream 44a enters reflux separator 24 wherein the condensed liquid (stream 46) is separated from any uncondensed vapor (stream 45). The liquid stream 46 from reflux separator 24 is pumped by reflux pump 25 to a pressure slightly above the operating pressure of deethanizer 20 and stream 46a is then supplied as cold top column feed (reflux) to deethanizer 20. This cold liquid reflux absorbs and condenses the C3 components and heavier hydrocarbon components from the vapors rising in the upper section of deethanizer 20.
The second portion (stream 47) of overhead vapor stream 43 combines with any uncondensed vapor (stream 45) from reflux separator 24 to form cool residue vapor stream 38 at −88° F. [−67° C.]. Residue vapor stream 38 passes countercurrently to inlet gas in heat exchanger 12 where it is heated to −5° F. [−21° C.] (stream 38a). The residue vapor stream is then re-compressed in two stages. The first stage is compressor 11 driven by expansion machine 10. The second stage is compressor 21 driven by a supplemental power source which compresses stream 38b to sales line pressure (stream 38c). After cooling to 126° F. [52° C.] in discharge cooler 22, stream 38d combines with warm LNG stream 71b to form the residue gas product (stream 42). Residue gas stream 42 flows to the sales gas pipeline at 1262 psia [8,701 kPa(a)], sufficient to meet line requirements.
The LNG (stream 71) from LNG tank 50 enters pump 51 at −251° F. [−157° C.]. Pump 51 elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to the sales gas pipeline. Stream 71a exits the pump 51 at −242° F. [−152° C.] and 1364 psia [9,404 kPa(a)] and is divided into two portions, streams 72 and 73. The first portion, stream 72, is heated as described previously to −173° F. [−114° C.] in reflux condenser 23 as it provides cooling to the portion (stream 44) of overhead vapor stream 43 from fractionation tower 20, and to 46° F. [8° C.] in heat exchanger 12 as it provides cooling to the inlet gas. The second portion, stream 73, is heated to 40° F. [4° C.] in heat exchanger 53 using low level utility heat. The heated streams 72b and 73a recombine to form warm LNG stream 71b, which thereafter combines with residue vapor stream 38d to form residue gas stream 42 as described previously.
A summary of stream flow rates and energy consumption for the process illustrated in
The recoveries reported in Table I are computed relative to the total quantities of propane and butanes+ contained in the gas stream being processed in the plant and in the LNG stream. Although the recoveries are quite high relative to the heavier hydrocarbons contained in the gas being processed (99.42% and 100.00%, respectively, for propane and butanes+), none of the heavier hydrocarbons contained in the LNG stream are captured in the
In the simulation of the
Expanded stream 71c enters fractionation column 62 in the lower region of the absorbing section of fractionation column 62. The liquid portion of stream 71c commingles with the liquids falling downward from the absorbing section and the combined liquid proceeds downward into the stripping section of deethanizer 62 (which includes reboiler 61). The vapor portion of expanded stream 71c rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier components.
A distillation liquid stream 72 is withdrawn from the lower region of the absorbing section in deethanizer 62 and is routed to heat exchanger 52. The distillation liquid stream is heated from −121° F. [−85° C.] to −50° F. [−45° C.], partially vaporizing stream 72a before it is returned as a lower mid-column feed to deethanizer 62, in the middle region of the stripping section.
A portion of the distillation vapor (stream 73) is withdrawn from the upper region of the stripping section of deethanizer 62 at −46° F. [−43° C.]. This stream is then cooled and partially condensed (stream 73a) in exchanger 52 by heat exchange with LNG stream 71a and distillation liquid stream 72 as described previously. The partially condensed stream 73a flows to reflux separator 64 at −104° F. [−76° C.]. The operating pressure of reflux separator 64 (452 psia [3,113 kPa(a)]) is slightly below the operating pressure of deethanizer 62 to provide the driving force which causes distillation vapor stream 73 to flow through heat exchanger 52 and into reflux separator 64, where the condensed liquid (stream 75) is separated from the uncondensed vapor (stream 74).
The liquid stream 75 from reflux separator 64 is pumped by pump 65 to a pressure slightly above the operating pressure of deethanizer 62, and the pumped stream 75a is then divided into two portions. One portion, stream 76, is supplied as top column feed (reflux) to deethanizer 62. This cold liquid reflux absorbs and condenses the C3 components and heavier components rising in the upper rectification region of the absorbing section of deethanizer 62. The other portion, stream 77, is supplied to deethanizer 62 at a mid-column feed position located in the upper region of the stripping section in substantially the same region where distillation vapor stream 73 is withdrawn, to provide partial rectification of stream 73. The deethanizer overhead vapor (stream 79) exits the top of deethanizer 62 at −105° F. [−76° C.] and is combined with the uncondensed vapor (stream 74) to form cold vapor stream 83 at −105° F. [−76° C.]. The liquid product stream 80 exits the bottom of the tower at 174° F. [79° C.], based on a typical specification of an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product.
Cold vapor stream 83 flows to compressor 56 driven by expansion machine 55 to increase the pressure of stream 83a sufficiently so that it can be totally condensed in heat exchanger 52. Stream 83a exits the compressor at −58° F. [−50° C.] and 669 psia [4,611 kPa(a)] and is cooled to −114° F. [−81° C.] (stream 83b) by heat exchange with the high pressure LNG feed stream 71a and distillation liquid stream 72 as discussed previously. Condensed stream 83b is pumped by pump 63 to a pressure slightly above the sales gas delivery pressure for subsequent vaporization in heat exchangers 23 and 12, heating stream 83c from −94° F. [−70° C.] to 40° F. [4° C.] as described in paragraphs [0033] and [0037] below to produce warm lean LNG stream 83e.
In the simulation of the
The vapor from separator 13 (stream 34) enters a work expansion machine 10 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 10 expands the vapor substantially isentropically to the operating pressure of fractionation tower 20 (approximately 441 psia [3,039 kPa(a)]), with the work expansion cooling the expanded stream 34a to a temperature of approximately −73° F. [−58° C.]. The partially condensed expanded stream 34a is then supplied as feed to fractionation tower 20 at an upper mid-column feed point. The liquid portion of stream 34a commingles with the liquids falling downward from the absorbing section and the combined liquid proceeds downward into the stripping section of deethanizer 20 (which includes reboiler 19). The vapor portion of expanded stream 34a rises upward through the absorbing section and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier components.
Liquid stream 35 is flash expanded through an appropriate expansion device, such as expansion valve 17, to slightly above the operating pressure of fractionation tower 20. The expanded stream 35a leaving expansion valve 17 reaches a temperature of −62° F. [−52° C.] before it provides cooling to the incoming feed gas in heat exchanger 12 as described previously. The heated stream 35b at 82° F. [28° C.] then enters fractionation tower 20 at a lower mid-column feed point to be stripped of its methane and C2 components.
A distillation liquid stream 36 is withdrawn from the lower region of the absorbing section in deethanizer 20 and is routed to heat exchanger 23. The distillation liquid stream is heated from −86° F. [−66° C.] to −12° F. [−24° C.], partially vaporizing stream 36a before it is returned as a lower mid-column feed to deethanizer 20, in the middle region of the stripping section.
A portion of the distillation vapor (stream 37) is withdrawn from the upper region of the stripping section of deethanizer 20 at −9° F. [−23° C.]. This stream is then cooled and partially condensed (stream 37a) in exchanger 23 by heat exchange with cold lean LNG stream 83c and with distillation liquid stream 36 as described previously. The partially condensed stream 37a flows to reflux separator 24 at −86° F. [−65° C.]. The operating pressure of reflux separator 24 (437 psia [3,012 kPa(a)]) is slightly below the operating pressure of deethanizer 20 to provide the driving force which causes distillation vapor stream 37 to flow through heat exchanger 23 and into reflux separator 24, where the condensed liquid (stream 45) is separated from the uncondensed vapor (stream 44).
The liquid stream 45 from reflux separator 24 is pumped by pump 25 to a pressure slightly above the operating pressure of deethanizer 20, and the pumped stream 45a is then divided into two portions. One portion, stream 46, is supplied as top column feed (reflux) to deethanizer 20. This cold liquid reflux absorbs and condenses the C3 components and heavier components rising in the upper rectification region of the absorbing section of deethanizer 20. The other portion, stream 47, is supplied to deethanizer 20 at a mid-column feed position located in the upper region of the stripping section in substantially the same region where distillation vapor stream 37 is withdrawn, to provide partial rectification of stream 37.
The deethanizer overhead vapor (stream 43) exits the top of deethanizer 20 at −88° F. [−67° C.] and is directed into heat exchanger 23 to provide cooling to distillation vapor stream 36 as described previously. The heated overhead vapor stream 43a at −56° F. [−49° C.] is combined with the uncondensed vapor (stream 44) to form cool residue vapor stream 38 at −58° F. [−50° C.]. The liquid product stream 40 exits the bottom of the tower at 208° F. [98° C.], based on a typical specification of an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product.
Cool residue vapor stream 38 passes countercurrently to inlet gas stream 31 in heat exchanger 12 where it is heated to 8° F. [−13° C.] (stream 38a). The heated residue vapor stream is then re-compressed in two stages. The first stage is compressor 11 driven by expansion machine 10. The second stage is compressor 21 driven by a supplemental power source which compresses stream 38b to sales line pressure (stream 38c). After cooling to 126° F. [52° C.] in discharge cooler 22, stream 38d combines with warm lean LNG stream 83e to form the residue gas product (stream 42). Residue gas stream 42 flows to the sales gas pipeline at 1262 psia [8,701 kPa(a)], sufficient to meet line requirements.
A summary of stream flow rates and energy consumption for the process illustrated in
Comparison of the recovery levels displayed in Tables I and II shows that the liquids recovery of the
In the simulation of the
The heated stream 71c enters separator 54 at −12° F. [−24° C.] and 1339 psia [9,232 kPa(a)] where the vapor (stream 77) is separated from any remaining liquid (stream 78). Vapor stream 77 enters a work expansion machine 55 in which mechanical energy is extracted from the high pressure feed. The machine 55 expands the vapor substantially isentropically to the tower operating pressure (455 psia [3,135 kPa(a)]), with the work expansion cooling the expanded stream 77a to a temperature of approximately −105° F. [−76° C.]. The work recovered is often used to drive a centrifugal compressor (such as item 56) that can be used to re-compress the cold second overhead vapor portion (stream 83), for example. The partially condensed expanded stream 77a is thereafter supplied as feed to fractionation column 20 at a first mid-column feed point. The separator liquid (stream 78), if any, is expanded to the operating pressure of fractionation column 20 by expansion valve 59 before expanded stream 78a is supplied to fractionation tower 20 at a first lower mid-column feed point.
In the simulation of the
The vapor from separator 13, stream 34, enters a work expansion machine 10 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 10 expands the vapor substantially isentropically to the operating pressure of fractionation tower 20, with the work expansion cooling the expanded stream 34a to a temperature of approximately −93° F. [−70° C.]. The work recovered is often used to drive a centrifugal compressor (such as item 11) that can be used to re-compress the heated residue vapor stream (stream 38a), for example. The partially condensed expanded stream 34a is then supplied to fractionation tower 20 at a second mid-column feed point.
Liquid stream 35 is flash expanded through an appropriate expansion device, such as expansion valve 17, to slightly above the operating pressure of fractionation tower 20. The expanded stream 35a leaving expansion valve 17 reaches a temperature of −85° F. [−65° C.] before it provides cooling to the incoming feed gas in heat exchanger 12 as described previously. The heated stream 35b at 81 ° F. [27° C.] then enters fractionation tower 20 at a second lower mid-column feed point to be stripped of its methane and C2 components.
The deethanizer in fractionation column 20 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The fractionation tower 20 may consist of two sections. The upper absorbing (rectification) section 20a contains the trays and/or packing to provide the necessary contact between the vapors rising upward and cold liquid falling downward to condense and absorb the C3 components and heavier components; the lower stripping (deethanizing) section 20b contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The deethanizing section also includes one or more reboilers (such as reboiler 19 using high level utility heat) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column. The column liquid stream 41 exits the bottom of the tower at 208° F. [98° C.], based on a typical specification of an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product.
The partially condensed expanded streams 77a and 34a are supplied to fractionation tower 20 in the lower region of absorbing section 20a. The liquid portions of streams 77a and 34a commingle with the liquids falling downward from absorbing section 20a and the combined liquid proceeds downward into stripping section 20b of deethanizer 20. The vapor portions of expanded streams 77a and 34a rise upward through absorbing section 20a and are contacted with cold liquid falling downward to condense and absorb the C3 components and heavier components.
A distillation liquid stream 36 is withdrawn from the lower region of absorbing section 20a in deethanizer 20 and is routed to heat exchanger 23. The distillation liquid stream is heated from −106° F. [−77° C.] to −24° F. [−31° C.], partially vaporizing stream 36a before it is returned to deethanizer 20 at a third lower mid-column feed position in the middle region of stripping section 20b.
A portion of the distillation vapor (stream 37) is withdrawn from the upper region of stripping section 20b in deethanizer 20 at −21 ° F. [−29° C.]. This stream is then cooled and partially condensed (stream 37a) in exchanger 23 by heat exchange with cold LNG stream 71a and distillation liquid stream 36 as described previously, and with cold first overhead vapor portion 43. The partially condensed stream 37a flows to reflux separator 24 at −87° F. [−66° C.]. The operating pressure of reflux separator 24 (452 psia [3,113 kPa(a)]) is slightly below the operating pressure of deethanizer 20 to provide the driving force which causes distillation vapor stream 37 to flow through heat exchanger 23 and into reflux separator 24, where the condensed liquid (stream 45) is separated from the uncondensed vapor (stream 44).
The liquid stream 45 from reflux separator 24 is pumped by pump 25 to a pressure slightly above the operating pressure of deethanizer 20, and the pumped stream 45a is then divided into two portions. One portion, stream 46, is supplied as top column feed (reflux) to deethanizer 20. This cold liquid reflux absorbs and condenses the C3 components and heavier components rising in the upper rectification region of absorbing section 20a of deethanizer 20. The other portion, stream 47, is supplied to deethanizer 20 at a mid-column feed position located in the upper region of stripping section 20b in substantially the same region where distillation vapor stream 37 is withdrawn, to provide partial rectification of stream 37.
The deethanizer overhead vapor (stream 79) exits the top of deethanizer 20 at −97° F. [−71° C.] and is divided into two portions, first overhead vapor portion 43 and second overhead vapor portion 83. First overhead vapor portion 43 is directed into heat exchanger 23 to provide cooling to distillation vapor stream 37 as described previously. The heated first overhead vapor portion 43a at −24° F. [−31 ° C.] is combined with any uncondensed vapor (stream 44) to form cool residue vapor stream 38, which passes countercurrently to inlet gas stream 31 in heat exchanger 12 where it is heated to −24° F. [−31° C.] (stream 38a). The residue vapor stream is then re-compressed in two stages. The first stage is compressor 11 driven by expansion machine 10. The second stage is compressor 21 driven by a supplemental power source which compresses stream 38b to sales line pressure (stream 38c). (Note that discharge cooler 22 is not needed in this example. Some applications may require cooling of compressed residue vapor stream 38c so that the resultant temperature when mixed with warm lean LNG stream 83d is sufficiently cool to comply with the requirements of the sales gas pipeline.)
Second overhead vapor portion 83 flows to compressor 56 driven by expansion machine 55, where it is compressed to 701 psia [4,833 kPa(a)] (stream 83a). At this pressure, the stream is totally condensed as it is cooled to −109° F. [−78° C.] in heat exchanger 23 as described previously. The condensed liquid (stream 83b) is the methane-rich lean LNG stream, which is pumped by pump 63 to 1275 psia [8,791 kPa(a)] for vaporization in heat exchanger 12, heating stream 83c to −25° F. [−32° C.] as described previously to produce warm lean LNG stream 83d which then combines with compressed residue vapor stream 38c/38d to form the residue gas product (stream 42). Residue gas stream 42 flows to the sales gas pipeline at 30° F. [−1° C.] and 1262 psia [8,701 kPa(a)], sufficient to meet line requirements.
A summary of stream flow rates and energy consumption for the process illustrated in
The improvement offered by the
Comparing the recovery levels displayed in Table III for the
There are six primary factors that account for the improved efficiency of the present invention. First, compared to many prior art processes, the present invention does not depend on the LNG feed itself to directly serve as the reflux for fractionation column 20. Rather, the refrigeration inherent in the cold LNG is used in heat exchanger 23 to generate a liquid reflux stream (stream 46) that contains very little of the C3 components and heavier hydrocarbon components that are to be recovered, resulting in efficient rectification in absorbing section 20a of fractionation tower 20 and avoiding the equilibrium limitations of such prior art processes. Second, the partial rectification of distillation vapor stream 37 by reflux stream 47 results in a top reflux stream 46 that is predominantly liquid methane and C2 components and contains very little C3 components and heavier hydrocarbon components. As a result, nearly 100% of the C3 components and substantially all of the heavier hydrocarbon components are recovered in liquid product 41 leaving the bottom of deethanizer 20. Third, the rectification of the column vapors provided by absorbing section 20a allows all of the LNG feed to be vaporized before entering work expansion machine 55 as stream 77, resulting in significant power recovery. This power can then be used to compress second overhead vapor portion 83 to a pressure sufficiently high so that it can be condensed in heat exchanger 23 and thereafter pumped to the pipeline delivery pressure. (Pumping uses significantly less power than compressing.)
Fourth, vaporization of the LNG feed (with part of the vaporization duty provided by low level utility heat in heat exchanger 53) means less total liquid feeding fractionation column 20, so that the high level utility heat consumed by reboiler 19 to meet the specification for the bottom liquid product from the deethanizer is minimized. Fifth, using the cold lean LNG stream 83c to provide “free” refrigeration to inlet gas stream 31 in heat exchanger 12 eliminates the need for a separate vaporization means (such as heat exchanger 53 in the
An alternative method of processing LNG and natural gas is shown in another embodiment of the present invention as illustrated in
In the simulation of the
Vapor stream 77 enters a work expansion machine 55 in which mechanical energy is extracted from the high pressure feed. The machine 55 expands the vapor substantially isentropically to the tower operating pressure (455 psia [3,135 kPa(a)]), with the work expansion cooling the expanded stream 77a to a temperature of approximately −105° F. [−76° C.]. The partially condensed expanded stream 77a is thereafter supplied as feed to fractionation column 20 at a first mid-column feed point. The separator liquid (stream 78), if any, is expanded to the operating pressure of fractionation column 20 by expansion valve 59 before expanded stream 78a is supplied to fractionation tower 20 at a first lower mid-column feed point.
In the simulation of the
The further cooled stream 31b enters separator 13 at −76° F. [−60° C.] and 458 psia [3,156 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35) and thereafter supplied to fractionation tower 20 at a second mid-column feed point. Liquid stream 35 is directed through valve 17 and then to heat exchanger 12 where it provides cooling to the incoming feed gas as described previously. The heated stream 35b at 65° F. [18° C.] then enters fractionation tower 20 at a second lower mid-column feed point to be stripped of its methane and C2 components.
A distillation liquid stream 36 is withdrawn from the lower region of the absorbing section in deethanizer 20 and is routed to heat exchanger 23. The distillation liquid stream is heated from −100° F. [−73° C.] to −17° F. [−27° C.], partially vaporizing stream 36a before it is returned to deethanizer 20 at a third lower mid-column feed position in the middle region of the stripping section.
A portion of the distillation vapor (stream 37) is withdrawn from the upper region of the stripping section in deethanizer 20 at −14° F. [−26° C.]. This stream is then cooled and partially condensed (stream 37a) in exchanger 23 by heat exchange with cold LNG stream 71a and distillation liquid stream 36 as described previously, and with cold first overhead vapor portion 43. The partially condensed stream 37a flows to reflux separator 24 at −84° F. [−64° C.]and 452 psia [3,113 kPa(a)] where the condensed liquid (stream 45) is separated from the uncondensed vapor (stream 44).
The liquid stream 45 from reflux separator 24 is pumped by pump 25 to a pressure slightly above the operating pressure of deethanizer 20, and the pumped stream 45a is then divided into two portions. One portion, stream 46, is supplied as top column feed (reflux) to deethanizer 20. The other portion, stream 47, is supplied to deethanizer 20 at a mid-column feed position located in the upper region of the stripping section in substantially the same region where distillation vapor stream 37 is withdrawn.
The column liquid stream 41 exits the bottom of the tower at 208° F. [98° C.], based on a typical specification of an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product. The deethanizer overhead vapor (stream 79) exits the top of deethanizer 20 at −96° F. [−71 ° C.] and is divided into two portions, first overhead vapor portion 43 and second overhead vapor portion 83. First overhead vapor portion 43 is directed into heat exchanger 23 to provide cooling to distillation vapor stream 37 as described previously. The heated first overhead vapor portion 43a at −17° F. [−27° C.] is combined with any uncondensed vapor (stream 44) to form cool residue vapor stream 38, which passes countercurrently to expanded inlet gas stream 31 in heat exchanger 12 where it is heated to −26° F. [−32° C.] (stream 38a). The residue vapor stream is then re-compressed in two stages. The first stage is compressor 11 driven by expansion machine 10. The second stage is compressor 21 driven by a supplemental power source which compresses stream 38b to sales line pressure (stream 38c).
Second overhead vapor portion 83 flows to compressor 56 driven by expansion machine 55, where it is compressed to 686 psia [4,729 kPa(a)] (stream 83a). At this pressure, the stream is totally condensed as it is cooled to −113° F. [−81° C.] in heat exchanger 23 as described previously. The condensed liquid (stream 83b) is the methane-rich lean LNG stream, which is pumped by pump 63 to 1275 psia [8,791 kPa(a)] for vaporization in heat exchanger 12, heating stream 83c to −27° F. [−33° C.] as described previously to produce warm lean LNG stream 83d which then combines with compressed residue vapor stream 38c/38d to form the residue gas product (stream 42). Residue gas stream 42 flows to the sales gas pipeline at 23° F. [−5° C.] and 1262 psia [8,701 kPa(a)], sufficient to meet line requirements.
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables III and IV shows that the
Another alternative method of processing LNG and natural gas is shown in the embodiment of the present invention as illustrated in
In the simulation of the
Vapor stream 77 enters a work expansion machine 55 in which mechanical energy is extracted from the high pressure feed. The machine 55 expands the vapor substantially isentropically to the tower operating pressure (455 psia [3,135 kPa(a)]), with the work expansion cooling the expanded stream 77a to a temperature of approximately −101° F. [−74° C.]. The partially condensed expanded stream 77a is thereafter supplied as feed to fractionation column 20 at a first mid-column feed point. The separator liquid (stream 78), if any, is expanded to the operating pressure of fractionation column 20 by expansion valve 59 before expanded stream 78a is supplied to fractionation tower 20 at a first lower mid-column feed point.
In the simulation of the
The further cooled stream 31b enters separator 13 at −72° F. [−58° C.] and 458 psia [3,156 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35) and thereafter supplied to fractionation tower 20 at a second mid-column feed point. Liquid stream 35 is directed through valve 17 and then to heat exchanger 12 where it provides cooling to the incoming feed gas as described previously. The heated stream 35b at 66° F. [19° C.] then enters fractionation tower 20 at a second lower mid-column feed point to be stripped of its methane and C2 components.
A distillation liquid stream 36 is withdrawn from the lower region of the absorbing section in deethanizer 20 and is routed to heat exchanger 23. The distillation liquid stream is heated from −96° F. [−71° C.] to −16° F. [−27° C.], partially vaporizing stream 36a before it is returned to deethanizer 20 at a third lower mid-column feed position in the middle region of the stripping section.
A portion of the distillation vapor (stream 37) is withdrawn from the upper region of the stripping section in deethanizer 20 at −13° F. [−25° C.]. This stream is then cooled and partially condensed (stream 37a) in exchanger 23 by heat exchange with cold LNG stream 71a and distillation liquid stream 36 as described previously, and with cold first overhead vapor portion 43. The partially condensed stream 37a flows to reflux separator 24 at −87° F. [−66° C.]and 452 psia [3,113 kPa(a)] where the condensed liquid (stream 45) is separated from the uncondensed vapor (stream 44).
The liquid stream 45 from reflux separator 24 is pumped by pump 25 to a pressure slightly above the operating pressure of deethanizer 20, and the pumped stream 45a is then divided into two portions. One portion, stream 46, is supplied as top column feed (reflux) to deethanizer 20. The other portion, stream 47, is supplied to deethanizer 20 at a mid-column feed position located in the upper region of the stripping section in substantially the same region where distillation vapor stream 37 is withdrawn.
The column liquid stream 41 exits the bottom of the tower at 208° F. [98° C.], based on a typical specification of an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product. The deethanizer overhead vapor (stream 79) exits the top of deethanizer 20 at −95° F. [−71° C.] and is divided into two portions, first overhead vapor portion 43 and second overhead vapor portion 83. First overhead vapor portion 43 is directed into heat exchanger 23 to provide cooling to distillation vapor stream 37 as described previously. The heated first overhead vapor portion 43a at −16° F. [−27° C.] is combined with any uncondensed vapor (stream 44) to form cool residue vapor stream 38 at −30° F. [−34° C.], which is partially re-compressed by compressor 11 driven by expansion machine 10. Because of the efficiency of the
Second overhead vapor portion 83 flows to compressor 56 driven by expansion machine 55, where it is compressed to 693 psia [4,781 kPa(a)] (stream 83a). At this pressure, the stream is totally condensed as it is cooled to −109° F. [−78° C.] in heat exchanger 23 as described previously. The condensed liquid (stream 83b) is the methane-rich lean LNG stream, which is pumped by pump 63 to 1275 psia [8,791 kPa(a)] for vaporization in heat exchanger 12, heating stream 83c to −11° F. [−24° C.] as described previously to produce warm lean LNG stream 83d which then combines with compressed residue vapor stream 38c/38d to form the residue gas product (stream 42). Residue gas stream 42 flows to the sales gas pipeline at 23° F. [−5° C.] and 1262 psia [8,701 kPa(a)], sufficient to meet line requirements.
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables III, IV, and V shows that the
In accordance with this invention, it is generally advantageous to design the absorbing (rectification) section of the deethanizer to contain multiple theoretical separation stages. However, the benefits of the present invention can be achieved with as few as one theoretical stage, and it is believed that even the equivalent of a fractional theoretical stage may allow achieving these benefits. For instance, all or a part of the condensed liquid (stream 45) leaving reflux separator 24 and all or a part of streams 77a and 34a can be combined (such as in the piping to the deethanizer) and if thoroughly intermingled, the vapors and liquids will mix together and separate in accordance with the relative volatilities of the various components of the total combined streams. Such commingling of these streams shall be considered for the purposes of this invention as constituting an absorbing section.
As described earlier, the distillation vapor stream 37 is partially condensed and the resulting condensate used to absorb valuable C3 components and heavier components from the vapors in streams 77a and 34a. However, the present invention is not limited to this embodiment. It may be advantageous, for instance, to treat only a portion of these vapors in this manner, or to use only a portion of the condensate as an absorbent, in cases where other design considerations indicate portions of the vapors or the condensate should bypass the absorbing section of the deethanizer.
It will also be recognized that the relative amount of feed found in each branch of the condensed liquid contained in stream 45a that is split between the two column feeds in
In the practice of the present invention, there will necessarily be a slight pressure difference between deethanizer 20 and reflux separator 24 which must be taken into account. If the distillation vapor stream 37 passes through heat exchanger 23 and into reflux separator 24 without any boost in pressure, reflux separator 24 shall necessarily assume an operating pressure slightly below the operating pressure of deethanizer 20. In this case, the liquid stream withdrawn from reflux separator 24 can be pumped to its feed position(s) on deethanizer 20. An alternative is to provide a booster blower for distillation vapor stream 37 to raise the operating pressure in heat exchanger 23 and reflux separator 24 sufficiently so that the liquid stream 45 can be supplied to deethanizer 20 without pumping.
When the inlet gas is leaner, separator 13 in
In the examples shown, total condensation of stream 83b in
Feed gas conditions, LNG conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machines 10 and/or 55, or replacement with an alternate expansion device (such as an expansion valve), is feasible. Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where appropriate.
In
Some circumstances may not require using distillation liquid stream 36 to provide cooling in heat exchanger 23, as shown by the dashed lines in
In the embodiments of the present invention illustrated in
The relative locations of the mid-column feeds may vary depending on inlet gas composition, LNG composition, or other factors such as the desired recovery level and the amount of vapor formed during heating of the LNG stream. Moreover, two or more of the feed streams, or portions thereof, may be combined depending on the relative temperatures and quantities of individual streams, and the combined stream then fed to a mid-column feed position.
The present invention provides improved recovery of C3 components and heavier hydrocarbon components per amount of utility consumption required to operate the process. An improvement in utility consumption required for operating the process may appear in the form of reduced power requirements for compression or pumping, reduced energy requirements for tower reboilers, or a combination thereof. Alternatively, the advantages of the present invention may be realized by accomplishing higher recovery levels for a given amount of utility consumption, or through some combination of higher recovery and improvement in utility consumption.
In the examples given for the
While there have been described what are believed to be preferred embodiments of the invention, those skilled in the art will recognize that other and further modifications may be made thereto, e.g. to adapt the invention to various conditions, types of feed, or other requirements without departing from the spirit of the present invention as defined by the following claims.
| Number | Date | Country | |
|---|---|---|---|
| 20100287985 A1 | Nov 2010 | US |
