# Patent

## Title
Liquefied natural gas and hydrocarbon gas processing

## Abstract

A process for recovering heavier hydrocarbons from a liquefied natural gas (LNG) stream and a hydrocarbon gas stream is disclosed. The LNG stream is heated to vaporize at least part of it, expanded, and supplied to a fractionation column at a first mid-column feed position. The gas stream is expanded, cooled, and supplied to the column at a second mid-column feed position. A distillation vapor stream is withdrawn from the column below the mid-column feed positions and cooled by the LNG stream sufficiently to condense at least a part of it, with at least a portion of the condensed stream directed to the column at an upper mid-column feed position. A portion of the column overhead stream is cooled by the LNG feed stream to condense it and form both a “lean” LNG stream and a reflux stream that is supplied to the column at a top column feed position.

## Background
This application is a continuation of U.S. patent application Ser. No. 12/466,661, filed May 15, 2009.


BACKGROUND OF THE INVENTION
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 ethylene, ethane, 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 C2 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 U.S. Pat. No. 7,216,507 has been used to recover C2 components and heavier hydrocarbon components in plants processing LNG, while assignee's co-pending application Ser. No. 11/430,412 could be used to recover C2 components and heavier hydrocarbon components in plants processing natural gas. Surprisingly, applicants have found that by integrating certain features of the assignee's U.S. Pat. No. 7,216,507 invention with certain features of the assignee's co-pending application Ser. No. 11/430,412, extremely high C2 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:
FIG. 1 is a flow diagram of a base case natural gas processing plant using LNG to provide its refrigeration;
FIG. 2 is a flow diagram of base case LNG and natural gas processing plants in accordance with U.S. Pat. No. 7,216,507 and co-pending application Ser. No. 11/430,412, respectively;
FIG. 3 is a flow diagram of an LNG and natural gas processing plant in accordance with the present invention; and
FIGS. 4 through 8 are flow diagrams illustrating alternative means of application of the present invention to LNG and natural gas streams.



FIGS. 1 and 2 are provided to quantify the advantages of the present invention.
In the following explanation of the above figures, tables are provided summarizing flow rates calculated for representative process conditions. In the tables appearing herein, the values for flow rates (in moles per hour) have been rounded to the nearest whole number for convenience. The total stream rates shown in the tables include all non-hydrocarbon components and hence are generally larger than the sum of the stream flow rates for the hydrocarbon components. Temperatures indicated are approximate values rounded to the nearest degree. It should also be noted that the process design calculations performed for the purpose of comparing the processes depicted in the figures are based on the assumption of no heat leak from (or to) the surroundings to (or from) the process. The quality of commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art.
For convenience, process parameters are reported in both the traditional British units and in the units of the 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.
FIG. 1 is a flow diagram showing the design of a processing plant to recover C2+ components from natural gas using an LNG stream to provide refrigeration. In the simulation of the FIG. 1 process, inlet gas enters the plant at 126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. If the inlet gas contains a concentration of sulfur compounds which would prevent the product streams from meeting specifications, the sulfur compounds are removed by appropriate pretreatment of the feed gas (not illustrated). In addition, the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccant has typically been used for this purpose.
The inlet gas stream 31 is cooled in heat exchanger 12 by heat exchange with a portion (stream 72a) of partially warmed LNG at −174° F. [−114° C.] and cool distillation stream 38a at −107° F. [−77° C.]. The cooled stream 31a enters separator 13 at −79° F. [−62° 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 430 psia [2,965 kPa(a)]) of fractionation tower 20. The expanded stream 35a leaving expansion valve 17 reaches a temperature of −93° F. [−70° 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 slightly above the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −101° F. [−74° 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 distillation stream (stream 38b), for example. The expanded stream 34a is further cooled to −124° F. [−87° C.] in heat exchanger 14 by heat exchange with cold distillation stream 38 at −143° F. [−97° C.], whereupon the partially condensed expanded stream 34b is thereafter supplied to fractionation tower 20 at a second mid-column feed point.
The demethanizer 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 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 and lighter components. Liquid product stream 41 exits the bottom of the tower at 99° F. [37° C.], based on a typical specification of a methane to ethane 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 −143° F. [−97° 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 demethanizer 20 and stream 46a is then supplied as cold top column feed (reflux) to demethanizer 20. This cold liquid reflux absorbs and condenses the C2 components and heavier hydrocarbon components from the vapors rising in the upper section of demethanizer 20.
The second portion (stream 47) of overhead vapor stream 43 combines with any uncondensed vapor (stream 45) from reflux separator 24 to form cold distillation stream 38 at −143° F. [−97° C.]. Distillation stream 38 passes countercurrently to expanded stream 34a in heat exchanger 14 where it is heated to −107° F. [−77° C.] (stream 38a), and countercurrently to inlet gas in heat exchanger 12 where it is heated to 47° F. [8° C.] (stream 38b). The distillation 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 38c to sales line pressure (stream 38d). After cooling to 126° F. [52° C.] in discharge cooler 22, stream 38e 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 −174° 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 43° F. [6° C.] in heat exchanger 12 as it provides cooling to the inlet gas. The second portion, stream 73, is heated to 35° F. [2° C.] in heat exchanger 53 using low level utility heat. The heated streams 72b and 73a recombine to form warm LNG stream 71b at 40° F. [4° C.], which thereafter combines with distillation stream 38e to form residue gas stream 42 as described previously.
A summary of stream flow rates and energy consumption for the process illustrated in FIG. 1 is set forth in the following table:







TABLE I





(FIG. 1)


Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]




















Stream
Methane
Ethane
Propane
Butanes+
Total





31
42,545
5,048
2,972
1,658
53,145


34
33,481
1,606
279
39
36,221


35
9,064
3,442
2,693
1,619
16,924


43
50,499
25
0
0
51,534


44
8,055
4
0
0
8,221


45
0
0
0
0
0


46
8,055
4
0
0
8,221


47
42,444
21
0
0
43,313


38
42,444
21
0
0
43,313


71
40,293
2,642
491
3
43,689


72
27,601
1,810
336
2
29,927


73
12,692
832
155
1
13,762


42
82,737
2,663
491
3
87,002


41
101
5,027
2,972
1,658
9,832














Recoveries*





Ethane
65.37%



Propane
85.83%



Butanes+
99.83%



Power













LNG Feed Pump
3,561
HP
[5,854
kW]



Reflux Pump
23
HP
[38
kW]



Residue Gas Compressor
24,612
HP
[40,462
kW]



Totals
28,196
HP
[46,354
kW]



Low Level Utility Heat



LNG Heater
68,990
MBTU/Hr
[44,564
kW]



High Level Utility Heat



Demethanizer Reboiler
80,020
MBTU/Hr
[51,689
kW]











Specific Power





HP-Hr/Lb. Mole
2.868



[kW-Hr/kg mole]

[4.715]







*(Based on un-rounded flow rates)






The recoveries reported in Table I are computed relative to the total quantities of ethane, 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 (99.58%, 100.00%, and 100.00%, respectively, for ethane, propane, and butanes+), none of the heavier hydrocarbons contained in the LNG stream are captured in the FIG. 1 process. In fact, depending on the composition of LNG stream 71, the residue gas stream 42 produced by the FIG. 1 process may not meet all pipeline specifications. The specific power reported in Table I is the power consumed per unit of liquid product recovered, and is an indicator of the overall process efficiency.
FIG. 2 is a flow diagram showing processes to recover C2+ components from LNG and natural gas in accordance with U.S. Pat. No. 7,216,507 and co-pending application Ser. No. 11/430,412, respectively, with the processed LNG stream used to provide refrigeration for the natural gas plant. The processes of FIG. 2 have been applied to the same LNG stream and inlet gas stream compositions and conditions as described previously for FIG. 1.
In the simulation of the FIG. 2 process, the LNG to be processed (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 expansion machine 55. Stream 71a exits the pump at −242° F. [−152° C.] and 1364 psia [9,404 kPa(a)] and is split into two portions, streams 75 and 76. The first portion, stream 75, is expanded to the operating pressure (approximately 415 psia [2,859 kPa(a)]) of fractionation column 62 by expansion valve 58. The expanded stream 75a leaves expansion valve 58 at −238° F. [−150° C.] and is thereafter supplied to tower 62 at an upper mid-column feed point.
The second portion, stream 76, is heated to −79° F. [−62° C.] in heat exchanger 52 by cooling compressed overhead distillation stream 79a at −70° F. [−57° C.] and reflux stream 82 at −128° F. [−89° C.]. The partially heated stream 76a is further heated and vaporized in heat exchanger 53 using low level utility heat. The heated stream 76b at −5° F. [−20° C.] and 1334 psia [9,198 kPa(a)] enters work expansion machine 55 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 55 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 76c to a temperature of approximately −107° F. [−77° C.] before it is supplied as feed to fractionation column 62 at a lower mid-column feed point.
The demethanizer in fractionation column 62 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing consisting of two sections. The upper absorbing (rectification) section contains the trays and/or packing to provide the necessary contact between the vapors rising upward and cold liquid falling downward to condense and absorb the ethane and heavier components; the lower stripping (demethanizing) section contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section also includes one or more reboilers (such as side reboiler 60 using low level utility heat, and reboiler 61 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 80 exits the bottom of the tower at 54° F. [12° C.], based on a typical specification of a methane to ethane ratio of 0.020:1 on a molar basis in the bottom product.
Overhead distillation stream 79 is withdrawn from the upper section of fractionation tower 62 at −144° F. [−98° C.] and flows to compressor 56 driven by expansion machine 55, where it is compressed to 807 psia [5,567 kPa(a)] (stream 79a). At this pressure, the stream is totally condensed as it is cooled to −128° F. [−89° C.] in heat exchanger 52 as described previously. The condensed liquid (stream 79b) is then divided into two portions, streams 83 and 82. The first portion (stream 83) is the methane-rich lean LNG stream, which is pumped by pump 63 to 1278 psia [8,809 kPa(a)] for subsequent vaporization in heat exchangers 14 and 12, heating stream 83a to −114° F. [−81° C.] and then to 40° F. [4° C.] as described in paragraphs [0036] and [0033] below to produce warm lean LNG stream 83c. 
The remaining portion of condensed liquid stream 79b, reflux stream 82, flows to heat exchanger 52 where it is subcooled to −237° F. [−149° C.] by heat exchange with a portion of the cold LNG (stream 76) as described previously. The subcooled stream 82a is then expanded to the operating pressure of demethanizer 62 by expansion valve 57. The expanded stream 82b at −236° F. [−149° C.] is then supplied as cold top column feed (reflux) to demethanizer 62. This cold liquid reflux absorbs and condenses the C2 components and heavier hydrocarbon components from the vapors rising in the upper rectification section of demethanizer 62.
In the simulation of the FIG. 2 process, inlet gas enters the plant at 126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. The feed stream 31 is cooled in heat exchanger 12 by heat exchange with cool lean LNG (stream 83b), cool overhead distillation stream 38a at −114° F. [−81° C.], and demethanizer liquids (stream 39) at −51° F. [−46° C.]. The cooled stream 31a enters separator 13 at −91° F. [−68° 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 390 psia [2,687 kPa(a)]) of fractionation tower 20. The expanded stream 35a leaving expansion valve 17 reaches a temperature of −111° F. [−80° C.] and is supplied to fractionation tower 20 at a first lower mid-column feed point.
Vapor stream 34 from separator 13 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 −121° F. [−85° C.]. The partially condensed expanded stream 34a is thereafter supplied as feed to fractionation tower 20 at a second lower mid-column feed point.
The demethanizer 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 consisting of two sections. The upper absorbing (rectification) section contains the trays and/or packing to provide the necessary contact between the vapors rising upward and cold liquid falling downward to condense and absorb the ethane and heavier components; the lower stripping (demethanizing) section contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section also includes one or more reboilers (such as the side reboiler in heat exchanger 12 described previously, and 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 40 exits the bottom of the tower at 89° F. [31° C.], based on a typical specification of a methane to ethane ratio of 0.020:1 on a molar basis in the bottom product, and combines with stream 80 to form the liquid product (stream 41).
A portion of the distillation vapor (stream 44) is withdrawn from the upper region of the stripping section of fractionation column 20 at −125° F. [−87° C.] and compressed to 545 psia [3,756 kPa(a)] by compressor 26. The compressed stream 44a is then cooled from −87° F. [−66° C.] to −143° F. [−97° C.] and condensed (stream 44b) in heat exchanger 14 by heat exchange with cold overhead distillation stream 38 exiting the top of demethanizer 20 and cold lean LNG (stream 83a) at −116° F. [−82° C.]. Condensed liquid stream 44b is expanded by expansion valve 16 to a pressure slightly above the operating pressure of demethanizer 20, and the resulting stream 44c at −146° F. [−99° C.] is then supplied as cold liquid reflux to an intermediate region in the absorbing section of demethanizer 20. This supplemental reflux absorbs and condenses most of the C3 components and heavier components (as well as some of the C2 components) from the vapors rising in the lower rectification region of the absorbing section so that only a small amount of recycle (stream 36) must be cooled, condensed, subcooled, and flash expanded to produce the top reflux stream 36c that provides the final rectification in the upper region of the absorbing section of demethanizer 20. As the cold reflux stream 36c contacts the rising vapors in the upper region of the absorbing section, it condenses and absorbs the C2 components and any remaining C3 components and heavier components from the vapors so that they can be captured in the bottom product (stream 40) from demethanizer 20.
Overhead distillation stream 38 is withdrawn from the upper section of fractionation tower 20 at −148° F. [−100° C.]. It passes countercurrently to compressed distillation vapor stream 44a and recycle stream 36a in heat exchanger 14 where it is heated to −114° F. [−81° C.] (stream 38a), and countercurrently to inlet gas stream 31 and recycle stream 36 in heat exchanger 12 where it is heated to 20° F. [−7° C.] (stream 38b). The distillation 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 38c to sales line pressure (stream 38d). After cooling to 126° F. [52° C.] in discharge cooler 22, stream 38e is divided into two portions, stream 37 and recycle stream 36. Stream 37 combines with warm lean LNG stream 83c 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.
Recycle stream 36 flows to heat exchanger 12 and is cooled to −105° F. [−76° C.] by heat exchange with cool lean LNG (stream 83b), cool overhead distillation stream 38a, and demethanizer liquids (stream 39) as described previously. Stream 36a is further cooled to −143° F. [−97° C.] by heat exchange with cold lean LNG stream 83a and cold overhead distillation stream 38 in heat exchanger 14 as described previously. The substantially condensed stream 36b is then expanded through an appropriate expansion device, such as expansion valve 15, to the demethanizer operating pressure, resulting in cooling of the total stream to −151° F. [−102° C.]. The expanded stream 36c is then supplied to fractionation tower 20 as the top column feed. Any vapor portion of stream 36c combines with the vapors rising from the top fractionation stage of the column to form overhead distillation stream 38, which is withdrawn from an upper region of the tower as described previously.
A summary of stream flow rates and energy consumption for the process illustrated in FIG. 2 is set forth in the following table:







TABLE II





(FIG. 2)


Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]




















Stream
Methane
Ethane
Propane
Butanes+
Total





31
42,545
5,048
2,972
1,658
53,145


34
28,762
1,051
163
22
30,759


35
13,783
3,997
2,809
1,636
22,386


44
6,746
195
3
0
7,000


38
49,040
39
0
0
50,064


36
6,595
5
0
0
6,733


37
42,445
34
0
0
43,331


40
100
5,014
2,972
1,658
9,814


71
40,293
2,642
491
3
43,689


75
4,835
317
59
0
5,243


76
35,458
2,325
432
3
38,446


79
45,588
16
0
0
45,898


82
5,348
2
0
0
5,385


83
40,240
14
0
0
40,513


80
53
2,628
491
3
3,176


42
82,685
48
0
0
83,844


41
153
7,642
3,463
1,661
12,990












Recoveries*




Ethane
 99.38%


Propane
100.00%


Butanes+
100.00%


Power











LNG Feed Pump
3,552
HP
[5,839
kW]


LNG Product Pump
1,774
HP
[2,916
kW]


Residue Gas Compressor
29,272
HP
[48,123
kW]


Reflux Compressor
601
HP
[988
kW]


Totals
35,199
HP
[57,866
kW]


Low Level Utility Heat


Liquid Feed Heater
66,200
MBTU/Hr
[42,762
kW]


Demethanizer Reboiler 60
23,350
MBTU/Hr
[15,083
kW]


Totals
89,550
MBTU/Hr
[57,845
kW]


High Level Utility Heat


Demethanizer Reboiler 19
26,780
MBTU/Hr
[17,298
kW]


Demethanizer Reboiler 61
3,400
MBTU/Hr
[2,196
kW]


Totals
30,180
MBTU/Hr
[19,494
kW]


Specific Power









HP-Hr/Lb. Mole
2.710



[kW-Hr/kg mole]

[4.455]





*(Based on un-rounded flow rates)






Comparison of the recovery levels displayed in Tables I and II shows that the liquids recovery of the FIG. 2 processes is much higher than that of the FIG. 1 process due to the recovery of the heavier hydrocarbon liquids contained in the LNG stream in fractionation tower 62. The ethane recovery improves from 65.37% to 99.38%, the propane recovery improves from 85.83% to 100.00%, and the butanes+recovery improves from 99.83% to 100.00%. In addition, the process efficiency of the FIG. 2 processes is improved by more than 5% in terms of the specific power relative to the FIG. 1 process.
DESCRIPTION OF THE INVENTION
Example 1
FIG. 3 illustrates a flow diagram of a process in accordance with the present invention. The LNG stream and inlet gas stream compositions and conditions considered in the process presented in FIG. 3 are the same as those in the FIG. 1 and FIG. 2 processes. Accordingly, the FIG. 3 process can be compared with the FIG. 1 and FIG. 2 processes to illustrate the advantages of the present invention.
In the simulation of the FIG. 3 process, the LNG to be processed (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 separator 54. Stream 71a exits the pump at −242° F. [−152° C.] and 1364 psia [9,404 kPa(a)] and is heated prior to entering separator 54 so that all or a portion of it is vaporized. In the example shown in FIG. 3, stream 71a is first heated to −54° F. [−48° C.] in heat exchanger 52 by cooling compressed distillation stream 81a at −32° F. [−36° C.], reflux stream 82, and distillation vapor stream 44. The partially heated stream 71b is further heated in heat exchanger 53 using low level utility heat. (High level utility heat, such as the heating medium used in tower reboiler 19, is normally more expensive than low level utility heat, so lower operating cost is usually achieved when use of low level heat, such as sea water, is maximized and the use of high level utility heat is minimized.) Note that in all cases exchangers 52 and 53 are representative of either a multitude of individual heat exchangers or a single multi-pass heat exchanger, or any combination thereof. (The decision as to whether to use more than one heat exchanger for the indicated heating services will depend on a number of factors including, but not limited to, inlet LNG flow rate, heat exchanger size, stream temperatures, etc.)
The heated stream 71c enters separator 54 at 11° F. [−12° C.] and 1334 psia [9,198 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 (approximately 412 psia [2,839 kPa(a)]), with the work expansion cooling the expanded stream 77a to a temperature of approximately −100° F. [−73° C.]. The work recovered is often used to drive a centrifugal compressor (such as item 56) that can be used to re-compress a portion (stream 81) of the column overhead vapor (stream 79), 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 FIG. 3 process, inlet gas enters the plant at 126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. The feed stream 31 is cooled in heat exchanger 12 by heat exchange with cool lean LNG (stream 83a) at −99° F. [−73° C.], cold distillation stream 38, and demethanizer liquids (stream 39) at −57° F. [−50° C.]. The cooled stream 31a enters separator 13 at −82° F. [−63° C.] and 584 psia [4,027 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35). Note that in all cases exchanger 12 is representative of either a multitude of individual heat exchangers or a single multi-pass heat exchanger, or any combination thereof. (The decision as to whether to use more than one heat exchanger for the indicated heating services will depend on a number of factors including, but not limited to, inlet gas flow rate, heat exchanger size, stream temperatures, etc.)
The vapor (stream 34) from separator 13 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 −108° F. [−78° 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 distillation stream (stream 38a), for example. The expanded partially condensed stream 34a is 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 the operating pressure of fractionation tower 20. The expanded stream 35a leaving expansion valve 17 reaches a temperature of −99° F. [−73° C.] and is supplied to fractionation tower 20 at a second lower mid-column feed point.
The demethanizer 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 ethane and heavier components; the lower stripping (demethanizing) section 20b contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. Demethanizing section 20b also includes one or more reboilers (such as the side reboiler in heat exchanger 12 described previously, side reboiler 18 using low level utility heat, and 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 83° F. [28° C.], based on a typical specification of a methane to ethane ratio of 0.020:1 on a molar basis in the bottom product.
A portion of the distillation vapor (stream 44) is withdrawn from the upper region of stripping section 20b of fractionation column 20 at −120° F. [−84° C.] and is cooled to −143° F. [−97° C.] and condensed (stream 44a) in heat exchanger 52 by heat exchange with the cold LNG (stream 71a). Condensed liquid stream 44a is pumped to slightly above the operating pressure of fractionation column 20 by pump 27, whereupon stream 44b at −143° F. [−97° C.] is then supplied as cold liquid reflux to an intermediate region in absorbing section 20a of fractionation column 20. This supplemental reflux absorbs and condenses most of the C3 components and heavier components (as well as some of the C2 components) from the vapors rising in the lower rectification region of absorbing section 20a so that only a small amount of the lean LNG (stream 82) must be subcooled to produce the top reflux stream 82b that provides the final rectification in the upper region of absorbing section 20a of fractionation column 20.
Overhead distillation stream 79 is withdrawn from the upper section of fractionation tower 20 at −145° F. [−98° C.] and is divided into two portions, stream 81 and stream 38. The first portion (stream 81) flows to compressor 56 driven by expansion machine 55, where it is compressed to 1092 psia [7,529 kPa(a)] (stream 81a). At this pressure, the stream is totally condensed as it is cooled to −106° F. [−77° C.] in heat exchanger 52 as described previously. The condensed liquid (stream 81b) is then divided into two portions, streams 83 and 82. The first portion (stream 83) is the methane-rich lean LNG stream, which is pumped by pump 63 to 1273 psia [8,777 kPa(a)] for subsequent vaporization in heat exchanger 12, heating stream 83a to 65° F. [18° C.] as described previously to produce warm lean LNG stream 83b. 
The remaining portion of stream 81b (stream 82) flows to heat exchanger 52 where it is subcooled to −234° F. [−148° C.] by heat exchange with the cold LNG (stream 71a) as described previously. The subcooled stream 82a is expanded to the operating pressure of fractionation column 20 by expansion valve 57. The expanded stream 82b at −232° F. [−146° C.] is then supplied as cold top column feed (reflux) to demethanizer 20. This cold liquid reflux absorbs and condenses the C2 components and heavier hydrocarbon components from the vapors rising in the upper rectification region of absorbing section 20a of demethanizer 20.
The second portion of overhead distillation stream 79 (stream 38) flows countercurrently to inlet gas stream 31 in heat exchanger 12 where it is heated to −62° F. [−52° C.] (stream 38a). The distillation 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 gas line pressure (stream 38c). (Note that discharge cooler 22 is not needed in this example. Some applications may require cooling of compressed distillation stream 38c so that the resultant temperature when mixed with warm lean LNG stream 83b is sufficiently cool to comply with the requirements of the sales gas pipeline.) Stream 38c/38d then combines with warm lean LNG stream 83b to form the residue gas product (stream 42). Residue gas stream 42 at 89° F. [32° C.] 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 FIG. 3 is set forth in the following table:







TABLE III





(FIG. 3)


Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]




















Stream
Methane
Ethane
Propane
Butanes+
Total





31
42,545
5,048
2,972
1,658
53,145


34
32,557
1,468
247
35
35,112


35
9,988
3,580
2,725
1,623
18,033


71
40,293
2,642
491
3
43,689


77
40,293
2,642
491
3
43,689


78
0
0
0
0
0


44
23,473
771
21
0
24,399


79
91,871
58
0
0
93,147


38
55,581
35
0
0
56,354


81
36,290
23
0
0
36,793


82
9,186
6
0
0
9,313


83
27,104
17
0
0
27,480


42
82,685
52
0
0
83,834


41
153
7,638
3,463
1,661
13,000












Recoveries*




Ethane
 99.33%


Propane
100.00%


Butanes+
100.00%


Power











LNG Feed Pump
3,552
HP
[5,839
kW]


LNG Product Pump
569
HP
[935
kW]


Reflux Pump
87
HP
[143
kW]


Residue Gas Compressor
22,960
HP
[37,746
kW]


Totals
27,168
HP
[44,663
kW]


Low Level Utility Heat


Liquid Feed Heater
58,100
MBTU/Hr
[37,530
kW]


Demethanizer Reboiler 18
8,000
MBTU/Hr
[5,167
kW]


Totals
66,100
MBTU/Hr
[42,697
kW]


High Level Utility Heat


Demethanizer Reboiler 19
31,130
MBTU/Hr
[20,108
kW]


Specific Power









HP-Hr/Lb. Mole
2.090



[kW-Hr/kg mole]

[3.436]





*(Based on un-rounded flow rates)






The improvement offered by the FIG. 3 embodiment of the present invention is astonishing compared to the FIG. 1 and FIG. 2 processes. Comparing the recovery levels displayed in Table III above for the FIG. 3 embodiment with those in Table I for the FIG. 1 process shows that the FIG. 3 embodiment of the present invention improves the ethane recovery from 65.37% to 99.33%, the propane recovery from 85.83% to 100.00%, and the butanes+recovery from 99.83% to 100.00%. Further, comparing the utilities consumptions in Table III with those in Table I shows that the power required for the FIG. 3 embodiment of the present invention is nearly 4% lower than the FIG. 1 process, meaning that the process efficiency of the FIG. 3 embodiment of the present invention is significantly better than that of the FIG. 1 process. The gain in process efficiency is clearly seen in the drop in the specific power, from 2.868 HP-Hr/Lb. Mole [4.715 kW-Hr/kg mole] for the FIG. 1 process to 2.090 HP-Hr/Lb. Mole [3.436 kW-Hr/kg mole] for the FIG. 3 embodiment of the present invention, an increase of more than 27% in the production efficiency. In addition, the high level utility heat requirement for the FIG. 3 embodiment of the present invention is only 39% of the requirement for the FIG. 1 process.
Comparing the recovery levels displayed in Table III for the FIG. 3 embodiment with those in Table II for the FIG. 2 processes shows that the liquids recovery levels are essentially the same. However, comparing the utilities consumptions in Table III with those in Table II shows that the power required for the FIG. 3 embodiment of the present invention is nearly 23% lower than the FIG. 2 processes. This results in reducing the specific power from 2.710 HP-Hr/Lb. Mole [4.455 kW-Hr/kg mole] for the FIG. 2 processes to 2.090 HP-Hr/Lb. Mole [3.436 kW-Hr/kg mole] for the FIG. 3 embodiment of the present invention, an improvement of nearly 23% in the production efficiency.
There are five 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 52 to generate a liquid reflux stream (stream 82) that contains very little of the C2 components and heavier hydrocarbon components that are to be recovered, resulting in efficient rectification in the upper region of absorbing section 20a in fractionation tower 20 and avoiding the equilibrium limitations of such prior art processes. Second, using distillation vapor stream 44 to produce supplemental reflux for the lower region of absorbing section 20a in fractionation column 20 allows using less top reflux (stream 82b) for fractionation tower 20. The lower top reflux flow, plus the greater degree of heating using low level utility heat in heat exchanger 53, results in less total liquid feeding fractionation column 20, reducing the duty required in reboiler 19 and minimizing the amount of high level utility heat needed to meet the specification for the bottom liquid product from demethanizer 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 the first portion (stream 81) of distillation overhead stream 79 to a pressure sufficiently high so that it can be condensed in heat exchanger 52 and so that the resulting lean LNG (stream 83) can then be pumped to the pipeline delivery pressure. (Pumping uses significantly less power than compressing.)
Fourth, using the cold lean LNG stream 83a to provide “free” refrigeration to the gas stream in heat exchanger 12 eliminates the need for a separate vaporization means (such as heat exchanger 53 in the FIG. 1 process) to re-vaporize the LNG prior to delivery to the sales gas pipeline. Fifth, this “free” refrigeration of inlet gas stream 31 means less of the cooling duty in heat exchanger 12 must be supplied by distillation vapor stream 38, so that stream 38a is cooler and less compression power is needed to raise its pressure to the pipeline delivery condition.
Example 2
An alternative method of processing LNG and natural gas is shown in another embodiment of the present invention as illustrated in FIG. 4. The LNG stream and inlet gas stream compositions and conditions considered in the process presented in FIG. 4 are the same as those in FIGS. 1 through 3. Accordingly, the FIG. 4 process can be compared with the FIGS. 1 and 2 processes to illustrate the advantages of the present invention, and can likewise be compared to the embodiment displayed in FIG. 3.
In the simulation of the FIG. 4 process, the LNG to be processed (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 separator 54. Stream 71a exits the pump at −242° F. [−152° C.] and 1364 psia [9,404 kPa(a)] and is heated prior to entering separator 54 so that all or a portion of it is vaporized. In the example shown in FIG. 4, stream 71a is first heated to −66° F. [−54° C.] in heat exchanger 52 by cooling compressed distillation stream 81a at −54° F. [−48° C.], reflux stream 82, and distillation vapor stream 44. The partially heated stream 71b is further heated in heat exchanger 53 using low level utility heat.
The heated stream 71c enters separator 54 at 3° F. [−16° C.] and 1334 psia [9,198 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 (approximately 420 psia [2,896 kPa(a)]), with the work expansion cooling the expanded stream 77a to a temperature of approximately −102° F. [−75° 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 FIG. 4 process, inlet gas enters the plant at 126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. The feed stream 31 enters a work expansion machine 10 in which mechanical energy is extracted from the high pressure feed. The machine 10 expands the vapor substantially isentropically to a pressure slightly above the operating pressure of fractionation tower 20, with the work expansion cooling the expanded stream 31a to a temperature of approximately 93° F. [34° C.]. The expanded stream 31a is further cooled in heat exchanger 12 by heat exchange with cool lean LNG (stream 83a) at −93° F. [−69° C.], cool distillation stream 38a, and demethanizer liquids (stream 39) at −76° F. [−60° C.].
The cooled stream 31b enters separator 13 at −81° F. [−63° C.] and 428 psia [2,949 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35). Vapor stream 34 is cooled to −122° F. [−86° C.] in heat exchanger 14 by heat exchange with cold distillation stream 38, and the partially condensed stream 34a is then supplied to fractionation tower 20 at a second mid-column feed point. Liquid stream 35 is directed through valve 17 and is supplied to fractionation tower 20 at a second lower mid-column feed point.
A portion of the distillation vapor (stream 44) is withdrawn from the upper region of the stripping section of fractionation column 20 at −119° F. [−84° C.] and is cooled to −145° F. [−98° C.] and condensed (stream 44a) in heat exchanger 52 by heat exchange with the cold LNG (stream 71a). Condensed liquid stream 44a is pumped to slightly above the operating pressure of fractionation column 20 by pump 27, whereupon stream 44b at −144° F. [−98° C.] is then supplied as cold liquid reflux to an intermediate region in the absorbing section of fractionation column 20. This supplemental reflux absorbs and condenses most of the C3 components and heavier components (as well as some of the C2 components) from the vapors rising in the lower rectification region of the absorbing section of fractionation column 20.
The column liquid stream 41 exits the bottom of the tower at 85° F. [29° C.], based on a typical specification of a methane to ethane ratio of 0.020:1 on a molar basis in the bottom product. Overhead distillation stream 79 is withdrawn from the upper section of fractionation tower 20 at −144° F. [−98° C.] and is divided into two portions, stream 81 and stream 38. The first portion (stream 81) flows to compressor 56 driven by expansion machine 55, where it is compressed to 929 psia [6,405 kPa(a)] (stream 81a). At this pressure, the stream is totally condensed as it is cooled to −108° F. [−78° C.] in heat exchanger 52 as described previously. The condensed liquid (stream 81b) is then divided into two portions, streams 83 and 82. The first portion (stream 83) is the methane-rich lean LNG stream, which is pumped by pump 63 to 1273 psia [8,777 kPa(a)] for subsequent vaporization in heat exchanger 12, heating stream 83a to 65° F. [18° C.] as described previously to produce warm lean LNG stream 83b. 
The remaining portion of stream 81b (stream 82) flows to heat exchanger 52 where it is subcooled to −235° F. [−148° C.] by heat exchange with the cold LNG (stream 71a) as described previously. The subcooled stream 82a is expanded to the operating pressure of fractionation column 20 by expansion valve 57. The expanded stream 82b at −233° F. [−147° C.] is then supplied as cold top column feed (reflux) to demethanizer 20. This cold liquid reflux absorbs and condenses the C2 components and heavier hydrocarbon components from the vapors rising in the upper rectification region of the absorbing section of demethanizer 20.
The second portion of overhead distillation stream 79 (stream 38) flows countercurrently to separator vapor stream 34 in heat exchanger 14 where it is heated to −87° F. [−66° C.] (stream 38a), and to expanded inlet gas stream 31a in heat exchanger 12 where it is heated to −47° F. [−44° C.] (stream 38b). The distillation 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 38c to sales gas line pressure (stream 38d). Stream 38d/38e then combines with warm lean LNG stream 83b to form the residue gas product (stream 42). Residue gas stream 42 at 99° F. [37° C.] 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 FIG. 4 is set forth in the following table:







TABLE IV





(FIG. 4)


Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]




















Stream
Methane
Ethane
Propane
Butanes+
Total





31
42,545
5,048
2,972
1,658
53,145


34
37,612
2,081
327
39
40,922


35
4,933
2,967
2,645
1,619
12,223


71
40,293
2,642
491
3
43,689


77
40,293
2,642
491
3
43,689


78
0
0
0
0
0


44
15,646
515
14
0
16,250


79
92,556
62
0
0
93,856


38
48,684
32
0
0
49,369


81
43,872
30
0
0
44,487


82
9,871
7
0
0
10,010


83
34,001
23
0
0
34,477


42
82,685
55
0
0
83,846


41
153
7,635
3,463
1,661
12,988












Recoveries*




Ethane
 99.29%


Propane
100.00%


Butanes+
100.00%


Power











LNG Feed Pump
3,552
HP
[5,839
kW]


LNG Product Pump
1,437
HP
[2,363
kW]


Reflux Pump
58
HP
[95
kW]


Residue Gas Compressor
18,325
HP
[30,126
kW]


Totals
23,372
HP
[38,423
kW]


Low Level Utility Heat


Liquid Feed Heater
66,000
MBTU/Hr
[42,632
kW]


Demethanizer Reboiler 18
17,300
MBTU/Hr
[11,175
kW]


Totals
83,300
MBTU/Hr
[53,807
kW]


High Level Utility Heat


Demethanizer Reboiler 19
32,940
MBTU/Hr
[21,278
kW]


Specific Power









HP-Hr/Lb. Mole
1.800



[kW-Hr/kg mole]

[2.958]





*(Based on un-rounded flow rates)






A comparison of Tables III and IV shows that the FIG. 4 embodiment of the present invention achieves essentially the same liquids recovery as the FIG. 3 embodiment. However, the FIG. 4 embodiment uses less power than the FIG. 3 embodiment, improving the specific power by nearly 14%. However, the high level utility heat required for the FIG. 4 embodiment of the present invention is slightly higher (about 6%) than that of the FIG. 3 embodiment.
Example 3
Another alternative method of processing LNG and natural gas is shown in the embodiment of the present invention as illustrated in FIG. 5. The LNG stream and inlet gas stream compositions and conditions considered in the process presented in FIG. 5 are the same as those in FIGS. 1 through 4. Accordingly, the FIG. 5 process can be compared with the FIGS. 1 and 2 processes to illustrate the advantages of the present invention, and can likewise be compared to the embodiments displayed in FIGS. 3 and 4.
In the simulation of the FIG. 5 process, the LNG to be processed (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 separator 54. Stream 71a exits the pump at −242° F. [−152° C.] and 1364 psia [9,404 kPa(a)] and is heated prior to entering separator 54 so that all or a portion of it is vaporized. In the example shown in FIG. 5, stream 71a is first heated to −71° F. [−57° C.] in heat exchanger 52 by cooling compressed distillation stream 81a at −25° F. [−32° C.], reflux stream 82, distillation vapor stream 44, and separator vapor stream 34. The partially heated stream 71b is further heated in heat exchanger 53 using low level utility heat.
The heated stream 71c enters separator 54 at 1° F. [−17° C.] and 1334 psia [9,198 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 (approximately 395 psia [2,721 kPa(a)]), with the work expansion cooling the expanded stream 77a to a temperature of approximately −107° F. [−77° 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 FIG. 5 process, inlet gas enters the plant at 126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. The feed stream 31 enters a work expansion machine 10 in which mechanical energy is extracted from the high pressure feed. The machine 10 expands the vapor substantially isentropically to a pressure slightly above the operating pressure of fractionation tower 20, with the work expansion cooling the expanded stream 31a to a temperature of approximately 87° F. [30° C.]. The expanded stream 31a is further cooled in heat exchanger 12 by heat exchange with cool lean LNG (stream 83a) at −97° F. [−72° C.], cool distillation stream 38b, and demethanizer liquids (stream 39) at −81° F. [−63° C.].
The cooled stream 31b enters separator 13 at −81° F. [−63° C.] and 403 psia [2,777 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35). Vapor stream 34 is cooled to −117° F. [−83° C.] in heat exchanger 52 by heat exchange with cold LNG stream 71a and compressed distillation stream 38a, and the partially condensed stream 34a is then supplied to fractionation tower 20 at a second mid-column feed point. Liquid stream 35 is directed through valve 17 and is supplied to fractionation tower 20 at a second lower mid-column feed point.
A portion of the distillation vapor (stream 44) is withdrawn from the upper region of the stripping section of fractionation column 20 at −119° F. [−84° C.] and is cooled to −145° F. [−98° C.] and condensed (stream 44a) in heat exchanger 52 by heat exchange with the cold LNG (stream 71a). Condensed liquid stream 44a is pumped to slightly above the operating pressure of fractionation column 20 by pump 27, whereupon stream 44b at −144° F. [−98° C.] is then supplied as cold liquid reflux to an intermediate region in the absorbing section of fractionation column 20. This supplemental reflux absorbs and condenses most of the C3 components and heavier components (as well as some of the C2 components) from the vapors rising in the lower rectification region of the absorbing section of fractionation column 20.
The column liquid stream 41 exits the bottom of the tower at 79° F. [26° C.], based on a typical specification of a methane to ethane ratio of 0.020:1 on a molar basis in the bottom product. Overhead distillation stream 79 is withdrawn from the upper section of fractionation tower 20 at −147° F. [−99° C.] and is divided into two portions, stream 81 and stream 38. The first portion (stream 81) flows to compressor 56 driven by expansion machine 55, where it is compressed to 1124 psia [7,750 kPa(a)] (stream 81a). At this pressure, the stream is totally condensed as it is cooled to −103° F. [−75° C.] in heat exchanger 52 as described previously. The condensed liquid (stream 81b) is then divided into two portions, streams 83 and 82. The first portion (stream 83) is the methane-rich lean LNG stream, which is pumped by pump 63 to 1273 psia [8,777 kPa(a)] for subsequent vaporization in heat exchanger 12, heating stream 83a to 65° F. [18° C.] as described previously to produce warm lean LNG stream 83b. 
The remaining portion of stream 81b (stream 82) flows to heat exchanger 52 where it is subcooled to −236° F. [−149° C.] by heat exchange with the cold LNG (stream 71a) as described previously. The subcooled stream 82a is expanded to the operating pressure of fractionation column 20 by expansion valve 57. The expanded stream 82b at −233° F. [−147° C.] is then supplied as cold top column feed (reflux) to demethanizer 20. This cold liquid reflux absorbs and condenses the C2 components and heavier hydrocarbon components from the vapors rising in the upper rectification region of the absorbing section of demethanizer 20.
The second portion of overhead distillation stream 79 (stream 38) is compressed to 625 psia [4,309 kPa(a)] by compressor 11 driven by expansion machine 10. It then flows countercurrently to separator vapor stream 34 in heat exchanger 52 where it is heated from −97° F. [−72° C.] to −65° F. [−53° C.] (stream 38b), and to expanded inlet gas stream 31a in heat exchanger 12 where it is heated to 12° F. [−11° C.] (stream 38c). The distillation stream is then further compressed to sales gas line pressure (stream 38d) in compressor 21 driven by a supplemental power source, and stream 38d/38e then combines with warm lean LNG stream 83b to form the residue gas product (stream 42). Residue gas stream 42 at 107° F. [42° C.] 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 FIG. 5 is set forth in the following table:







TABLE V





(FIG. 5)


Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]




















Stream
Methane
Ethane
Propane
Butanes+
Total





31
42,545
5,048
2,972
1,658
53,145


34
38,194
2,203
348
40
41,654


35
4,351
2,845
2,624
1,618
11,491


71
40,293
2,642
491
3
43,689


77
40,293
2,642
491
3
43,689


78
0
0
0
0
0


44
17,004
614
16
0
17,715


79
91,637
60
0
0
92,925


38
59,566
39
0
0
60,403


81
32,071
21
0
0
32,522


82
8,952
6
0
0
9,078


83
23,119
15
0
0
23,444


42
82,685
54
0
0
83,847


41
153
7,636
3,463
1,661
12,987












Recoveries*




Ethane
 99.30%


Propane
100.00%


Butanes+
100.00%


Power











LNG Feed Pump
3,552
HP
[5,839
kW]


LNG Product Pump
418
HP
[687
kW]


Reflux Pump
63
HP
[104
kW]


Residue Gas Compressor
19,274
HP
[31,686
kW]


Totals
23,307
HP
[38,316
kW]


Low Level Utility Heat


Liquid Feed Heater
70,480
MBTU/Hr
[45,526
kW]


Demethanizer Reboiler 18
24,500
MBTU/Hr
[15,826
kW]


Totals
94,980
MBTU/Hr
[61,352
kW]


High Level Utility Heat


Demethanizer Reboiler 19
27,230
MBTU/Hr
[17,589
kW]


Specific Power









HP-Hr/Lb. Mole
1.795



[kW-Hr/kg mole

[2.950]]





*(Based on un-rounded flow rates)






A comparison of Tables III, IV, and V shows that the FIG. 5 embodiment of the present invention achieves essentially the same liquids recovery as the FIG. 3 and FIG. 4 embodiments. The FIG. 5 embodiment uses significantly less power than the FIG. 3 embodiment (improving the specific power by over 14%) and slightly less than the FIG. 4 embodiment. However, the high level utility heat required for the FIG. 5 embodiment of the present invention is considerably lower than that of the FIG. 3 and FIG. 4 embodiments (by about 13% and 17%, respectively). The choice of which embodiment to use for a particular application will generally be dictated by the relative costs of power and high level utility heat and the relative capital costs of pumps, heat exchangers, and compressors.
Other Embodiments
FIGS. 3 through 5 depict fractionation towers constructed in a single vessel. FIGS. 6 through 8 depict fractionation towers constructed in two vessels, absorber (rectifier) column 66 (a contacting and separating device) and stripper (distillation) column 20. In such cases, the overhead vapor (stream 43) from stripper column 20 is split into two portions. One portion (stream 44) is routed to heat exchanger 52 to generate supplemental reflux for absorber column 66. The remaining portion (stream 47) flows to the lower section of absorber column 66 to be contacted by the cold reflux (stream 82b) and the supplemental reflux (condensed liquid stream 44b). Pump 67 is used to route the liquids (stream 46) from the bottom of absorber column 66 to the top of stripper column 20 so that the two towers effectively function as one distillation system. The decision whether to construct the fractionation tower as a single vessel (such as demethanizer 20 in FIGS. 3 through 5) or multiple vessels will depend on a number of factors such as plant size, the distance to fabrication facilities, etc.
In accordance with this invention, it is generally advantageous to design the absorbing (rectification) section of the demethanizer 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 cold reflux (stream 82b), all or a part of the condensed liquid (stream 44b), and all or a part of streams 77a and 34a can be combined (such as in the piping to the demethanizer) 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.
In the examples shown, total condensation of streams 44a and 81b is illustrated in FIGS. 3 through 8. Some circumstances may favor subcooling these streams, while other circumstances may favor only partial condensation. Should partial condensation of either or both of these streams be achieved, processing of the uncondensed vapor may be necessary, using a compressor or other means to elevate the pressure of the vapor so that it can join the pumped condensed liquid. Alternatively, the uncondensed vapor could be routed to the plant fuel system or other such use.
When the inlet gas is leaner, separator 13 in FIGS. 3 through 8 may not be needed. Depending on the quantity of heavier hydrocarbons in the feed gas and the feed gas pressure, the cooled stream 31a (FIGS. 3 and 6) or expanded cooled stream 31b (FIGS. 4, 5, 7, and 8) leaving heat exchanger 12 may not contain any liquid (because it is above its dewpoint, or because it is above its cricondenbar), so that separator 13 may not be justified. In such cases, separator 13 and expansion valve 17 may be eliminated as shown by the dashed lines. When the LNG to be processed is lean or when complete vaporization of the LNG in heat exchangers 52 and 53 is contemplated, separator 54 in FIGS. 3 through 8 may not be justified. Depending on the quantity of heavier hydrocarbons in the inlet LNG and the pressure of the LNG stream leaving feed pump 51, the heated LNG stream leaving heat exchanger 53 may not contain any liquid (because it is above its dewpoint, or because it is above its cricondenbar). In such cases, separator 54 and expansion valve 59 may be eliminated as shown by the dashed lines.
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 FIGS. 3 through 8, individual heat exchangers have been shown for most services. However, it is possible to combine two or more heat exchange services into a common heat exchanger, such as combining heat exchangers 52 and 53 in FIGS. 3 through 8 into a common heat exchanger. In some cases, circumstances may favor splitting a heat exchange service into multiple exchangers. The decision as to whether to combine heat exchange services or to use more than one heat exchanger for the indicated service will depend on a number of factors including, but not limited to, inlet gas flow rate, LNG flow rate, heat exchanger size, stream temperatures, etc. In accordance with the present invention, the use and distribution of the methane-rich lean LNG and distillation vapor streams for process heat exchange, and the particular arrangement of heat exchangers for heating the LNG streams and cooling the feed gas stream, must be evaluated for each particular application, as well as the choice of process streams for specific heat exchange services.
In the embodiments of the present invention illustrated in FIGS. 3 through 8, lean LNG stream 83a is used directly to provide cooling in heat exchanger 12. However, some circumstances may favor using the lean LNG to cool an intermediate heat transfer fluid, such as propane or other suitable fluid, whereupon the cooled heat transfer fluid is then used to provide cooling in heat exchanger 12. This alternative means of indirectly using the refrigeration available in lean LNG stream 83a accomplishes the same process objectives as the direct use of stream 83a for cooling in the FIGS. 3 through 8 embodiments of the present invention. The choice of how best to use the lean LNG stream for refrigeration will depend mainly on the composition of the inlet gas, but other factors may affect the choice as well.
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 C2 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 FIGS. 3 through 5 embodiments, recovery of C2 components and heavier hydrocarbon components is illustrated. However, it is believed that the FIGS. 3 through 8 embodiments are also advantageous when recovery of C3 components and heavier hydrocarbon components is desired.
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.

## Claims

1. A process for the separation of liquefied natural gas containing methane and heavier hydrocarbon components and a gas stream containing methane and heavier hydrocarbon components into a volatile residue gas fraction containing a major portion of said methane and a relatively less volatile liquid fraction containing a major portion of said heavier hydrocarbon components wherein
(a) said liquefied natural gas is heated sufficiently to vaporize it, thereby forming a vapor stream;
(b) said vapor stream is expanded to lower pressure and is thereafter supplied to a distillation column at a first mid-column feed position;
(c) said gas stream is expanded to said lower pressure, is cooled, and is thereafter supplied to said distillation column at a second mid-column feed position;
(d) a distillation vapor stream is withdrawn from a region of said distillation column below said expanded vapor stream and said expanded cooled gas stream, whereupon said distillation vapor stream is cooled sufficiently to at least partially condense it and form thereby a first condensed stream, with said cooling supplying at least a portion of said heating of said liquefied natural gas;
(e) at least a portion of said first condensed stream is supplied to said distillation column at an upper mid-column feed position;
(f) an overhead distillation stream is withdrawn from an upper region of said distillation column and divided into at least a first portion and a second portion, whereupon said first portion is compressed to higher pressure;
(g) said compressed first portion is cooled sufficiently to at least partially condense it and form thereby a second condensed stream, with said cooling supplying at least a portion of said heating of said liquefied natural gas;
(h) said second condensed stream is divided into at least a volatile liquid stream and a reflux stream;
(i) said reflux stream is further cooled, with said cooling supplying at least a portion of said heating of said liquefied natural gas;
(j) said further cooled reflux stream is supplied to said distillation column at a top column feed position;
(k) said volatile liquid stream is heated sufficiently to vaporize it, with said heating supplying at least a portion of said cooling of said expanded gas stream;
(l) said second portion is heated, with said heating supplying at least a portion of said cooling of said expanded gas stream;
(m) said vaporized volatile liquid stream and said heated second portion are combined to form said volatile residue gas fraction containing a major portion of said methane; and
(n) the quantity and temperature of said reflux stream and the temperatures of said feeds to said distillation column are effective to maintain the overhead temperature of said distillation column at a temperature whereby the major portion of said heavier hydrocarbon components is recovered in said relatively less volatile liquid fraction by fractionation in said distillation column.

2. The process according to claim 1, wherein
(a) said liquefied natural gas is heated sufficiently to partially vaporize it;
(b) said partially vaporized liquefied natural gas is separated thereby to provide said vapor stream and a liquid stream; and
(c) said liquid stream is expanded to said lower pressure and thereafter-supplied to said distillation column at a lower mid-column feed position.

3. The process according to claim 1, wherein
(a) said gas stream is expanded to said lower pressure and is thereafter cooled sufficiently to partially condense it;
(b) said partially condensed gas stream is separated thereby to provide a further vapor stream and a liquid stream;
(c) said further vapor stream is cooled and thereafter supplied to said distillation column at said second mid-column feed position;
(d) said liquid stream is supplied to said distillation column at a lower mid-column feed position; and
(e) said distillation vapor stream is withdrawn from a region of said distillation column below said expanded ̃vapor stream and said cooled further vapor stream.

4. The process according to claim 1, wherein
(a) said liquefied natural gas is heated sufficiently to partially vaporize it;
(b) said partially vaporized liquefied natural gas is separated thereby to provide said vapor stream and a first liquid stream;
(c) said first liquid stream is expanded to said lower pressure and thereafter supplied to said distillation column at a first lower mid-column feed position;
(d) said gas stream is expanded to said lower pressure and is thereafter cooled sufficiently to partially condense it;
(e) said partially condensed gas stream is separated thereby to provide a further vapor stream and a second liquid stream;
(f) said further vapor stream is cooled and thereafter supplied to said distillation column at said second mid-column feed position;
(g) said second liquid stream is supplied to said distillation column at a second lower mid-column feed position;
(h) said distillation vapor stream is withdrawn from a region of said distillation column below said expanded vapor stream and said cooled further vapor stream.

5. The process according to claim 1 or 2 wherein
(a) said gas stream is cooled, is expanded to said lower pressure, and is thereafter supplied to said distillation column at said second mid-column feed position;
(b) said distillation vapor stream is withdrawn from a region of said distillation column below said expanded vapor stream and said cooled expanded gas stream;
(c) said volatile liquid stream is heated sufficiently to vaporize it, with said heating supplying at least a portion of said cooling of said gas stream; and
(d) said second portion is heated, with said heating supplying at least a portion of said cooling of said gas stream.

6. The process according to claim 3 wherein
(a) said gas stream is cooled sufficiently to partially condense it and then separated; thereby forming said further vapor stream and said liquid stream;
(b) said further vapor stream is expanded to said lower pressure and is thereafter supplied to said distillation column at said second mid-column feed position;
(c) said liquid stream is expanded to said lower pressure and is thereafter supplied to said distillation column at said lower mid-column feed position;
(d) said distillation vapor stream is withdrawn from a region of said distillation column below said expanded vapor stream and said expanded further vapor stream;
(e) said volatile liquid stream is heated sufficiently to vaporize it, with said heating supplying at least a portion of said cooling of said gas stream; and
(f) said second portion is heated, with said heating supplying at least a portion of said cooling of said gas stream.

7. The process according to claim 4 wherein
(a) said gas stream is cooled sufficiently to partially condense it and then separated; thereby forming said further vapor stream and said second liquid stream;
(b) said further vapor stream is expanded to said lower pressure and is thereafter supplied to said distillation column at said second mid-column feed position;
(c) said second liquid stream is expanded to said lower pressure and is thereafter supplied to said distillation column at said second lower mid-column feed position;
(d) said distillation vapor stream is withdrawn from a region of said distillation column below said expanded first vapor stream and said expanded further vapor stream;
(e) said volatile liquid stream is heated sufficiently to vaporize it, with said heating supplying at least a portion of said cooling of said gas stream; and
(f) said second portion is heated, with said heating supplying at least a portion of said cooling of said gas stream.

8. The process according to claim 1, 2, 3, or 4 wherein
(a) said second portion is compressed to higher pressure;
(b) said compressed second portion is heated, with said heating supplying at least a portion of said cooling of said expanded gas stream; and
(c) said vaporized volatile liquid stream and said heated compressed second portion are combined to form said volatile residue gas fraction.

9. The process according to claim 1, 2, 3, 4, 6, or 7 wherein said volatile residue gas fraction contains a major portion of said methane and C2 components.

10. The process according to claim 5 wherein said volatile residue gas fraction contains a major portion of said methane and C2 components.

11. The process according to claim 8 wherein said volatile residue gas fraction contains a major portion of said methane and C2 components.