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, to provide a volatile methane-rich residue 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 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 heavier hydrocarbons from LNG, these processes often must compromise between high recovery, low utility costs, and process simplicity (and hence low capital investment). In U.S. Pat. No. 2,952,984 Marshall describes an LNG process capable of very high ethane recovery via the use of a refluxed distillation column. Markbreiter describes in U.S. Pat. No. 3,837,172 a simpler process using a non-refluxed fractionation column, limited to lower ethane or propane recoveries. Rambo et al describe in U.S. Pat. No. 5,114,451 an LNG process capable of very high ethane or very high propane recovery using a compressor to provide reflux for the distillation column.
The present invention is generally concerned with the recovery of ethylene, ethane, propylene, propane, and heavier hydrocarbons from such LNG streams. It uses a novel process arrangement to allow high ethane or high propane 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 to give lower operating cost than the prior art processes. A typical analysis of an LNG stream to be processed in accordance with this invention would be, in approximate mole percent, 86.7% methane, 8.9% ethane and other C2 components, 2.9% propane and other C3 components, and 1.0% butanes plus, with the balance made up of nitrogen.
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 International System of Units (SI). The molar flow rates given in the tables may be interpreted as either pound moles per hour or kilogram moles per hour. The energy consumptions reported as horsepower (HP) and/or thousand British Thermal Units per hour (MBTU/Hr) correspond to the stated molar flow rates in pound moles per hour. The energy consumptions reported as kilowatts (kW) correspond to the stated molar flow rates in kilogram moles per hour.
Referring now to
The second portion, stream 43, is heated prior to entering fractionation tower 16 so that all or a portion of it is vaporized, reducing the amount of liquid flowing down fractionation tower 16 and allowing the use of a smaller diameter column. In the example shown in
Fractionation tower 16, commonly referred to as a demethanizer, 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 trays and/or packing provide the necessary contact between the liquids falling downward in the column and the vapors rising upward. As shown in
The demethanizer overhead vapor, stream 46, is the methane-rich residue gas, leaving the column at −141° F. [−96° C.]. After being heated to −40° F. [−40° C.] in cross exchanger 29 so that conventional metallurgy may be used in compressor 28, stream 46a enters compressor 28 (driven by a supplemental power source) and is compressed to sales line pressure (stream 46b). Following cooling to 50° F. [10° C.] in cross exchanger 29, the residue gas product (stream 46c) flows to the sales gas pipeline at 1315 psia [9,067 kPa(a)] for subsequent distribution.
The relative split of the LNG into streams 42 and 43 is typically adjusted to maintain the desired recovery level of the desired C2 components and heavier hydrocarbon components in the bottom liquid product (stream 47). Increasing the split to stream 42 feeding the top of fractionation tower 16 will increase the recovery level, until a point is reached where the composition of demethanizer overhead vapor (stream 46) is in equilibrium with the composition of the LNG (i.e., the composition of the liquid in stream 42a). Once this point has been reached, further increasing the split to stream 42 will not raise the recovery any further, but will simply increase the amount of high level utility heat required in reboiler 22 because less of the LNG is split to stream 43 and heated with low level utility heat in heat exchanger 14. (High level utility heat is normally more expensive than low level utility heat, so lower operating cost is usually achieved when the use of low level heat is maximized and the use of high level heat is minimized.) For the process conditions shown in
A summary of stream flow rates and energy consumption for the process illustrated in
This prior art process can also be adapted to produce an LPG product containing the majority of the C3 components and heavier hydrocarbon components present in the feed stream as shown in
The liquid product stream 47 exits the bottom of fractionation tower 16 (commonly referred to as a deethanizer when producing an LPG product) at 189° F. [87° C.], based on a typical specification of an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product. After cooling to 125° F. [52° C.] in heat exchanger 13, the liquid product (stream 47a) flows to storage or further processing.
The deethanizer overhead vapor (stream 46) leaves the column at −90° F. [−68° C.], is heated to −40° F. [−40° C.] in cross exchanger 29 (stream 46a), and is compressed by compressor 28 to sales line pressure (stream 46b). Following cooling to 83° F. [28° C.] in cross exchanger 29, the residue gas product (stream 46c) flows to the sales gas pipeline at 1315 psia [9,067 kPa(a)] for subsequent distribution.
A summary of stream flow rates and energy consumption for the process illustrated in
If a slightly lower recovery level is acceptable, this prior art process can produce an LPG product using less power and high level utility heat as shown in
A summary of stream flow rates and energy consumption for the process illustrated in
In the simulation of the
Overhead stream 46 leaves the upper section of fractionation tower 16 at −146° F. [−99° C.] and flows to reflux condenser 17 where it is cooled to −147° F. [−99° C.] and partially condensed by heat exchange with the cold LNG (stream 41a) as described previously. The partially condensed stream 46a enters reflux separator 18 wherein the condensed liquid (stream 49) is separated from the uncondensed vapor (stream 48). The liquid stream 49 from reflux separator 18 is pumped by reflux pump 19 to a pressure slightly above the operating pressure of demethanizer 16 and stream 49a is then supplied as cold top column feed (reflux) to demethanizer 16. 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 16.
The liquid product stream 47 exits the bottom of fractionation tower 16 at 71° F. [22° C.], based on a methane to ethane ratio of 0.005:1 on a volume basis in the bottom product. After cooling to 18° F. [−8° C.] in heat exchanger 13 as described previously, the liquid product (stream 47a) flows to storage or further processing. The residue gas (stream 48) leaves reflux separator 18 at −147° F. [−99° C.], is heated to −40° F. [−40° C.] in cross exchanger 29 (stream 48a), and is compressed by compressor 28 to sales line pressure (stream 48b). Following cooling to 43° F. [6° C.] in cross exchanger 29, the residue gas product (stream 48c) flows to the sales gas pipeline at 1315 psia [9,067 kPa(a)] for subsequent distribution.
A summary of stream flow rates and energy consumption for the process illustrated in
Comparing the recovery levels displayed in Table IV above for the
This prior art process can also be adapted to produce an LPG product containing the majority of the C3 components and heavier hydrocarbon components present in the feed stream as shown in
The liquid product stream 47 exits the bottom of deethanizer 16 at 190° F. [88° C.], based on an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product. After cooling to 125° F. [52° C.] in heat exchanger 13, the liquid product (stream 47a) flows to storage or further processing. The residue gas (stream 48) leaves reflux separator 18 at −94° F. [−70° C.], is heated to −40° F. [−40° C.] in cross exchanger 29 (stream 48a), and is compressed by compressor 28 to sales line pressure (stream 48b). Following cooling to 79° F. [26° C.] in cross exchanger 29, the residue gas product (stream 48c) flows to the sales gas pipeline at 1315 psia [9,067 kPa(a)] for subsequent distribution.
A summary of stream flow rates and energy consumption for the process illustrated in
If a slightly lower recovery level is acceptable, this prior art process can produce an LPG product using less power and high level utility heat as shown in
A summary of stream flow rates and energy consumption for the process illustrated in
In the simulation of the
The proportion of the total feed in stream 41a flowing to the column as stream 42 is controlled by valve 12, and is typically 50% or less of the total feed. Stream 42a flows from valve 12 to heat exchanger 17 where it is heated as it cools, substantially condenses, and subcools stream 49a. The heated stream 42b then flows to demethanizer 16 at an upper mid-column feed point at −160° F. [−107° C.].
Tower overhead stream 46 leaves demethanizer 16 at −147° F. [−99° C.] and is divided into two portions. The major portion, stream 48, is the methane-rich residue gas. It is heated to −40° F. [−40° C.] in cross exchanger 29 (stream 48a) and compressed by compressor 28 to sales line pressure (stream 48b). Following cooling to 43° F. [6° C.] in cross exchanger 29, the residue gas product (stream 48c) flows to the sales gas pipeline at 1315 psia [9,067 kPa(a)] for subsequent distribution.
The minor portion of the tower overhead, stream 49, enters compressor 26, which supplies a modest boost in pressure to overcome the pressure drops in heat exchanger 17 and control valve 27, as well as the static head due to the height of demethanizer 16. The compressed stream 49a is cooled to −247° F. [−155° C.] to substantially condense and subcool it (stream 49b) by a portion of the LNG feed (stream 42a) in heat exchanger 17 as described previously. Stream 49b flows through valve 27 to lower its pressure to that of fractionation tower 16, and resulting stream 49c flows to the top feed point of demethanizer 16 to serve as reflux for the tower.
The liquid product stream 47 exits the bottom of fractionation tower 16 at 70° F. [21° C.], based on a methane to ethane ratio of 0.005:1 on a volume basis in the bottom product. After cooling to 18° F. [−8° C.] in heat exchanger 13 as described previously, the liquid product (stream 47a) flows to storage or further processing.
A summary of stream flow rates and energy consumption for the process illustrated in
Comparing the recovery levels displayed in Table VII above for the
This prior art process can also be adapted to produce an LPG product containing the majority of the C3 components and heavier hydrocarbon components present in the feed stream as shown in
The liquid product stream 47 exits the bottom of deethanizer 16 at 189° F. [87° C.], based on an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product. After cooling to 124° F. [51° C.] in heat exchanger 13, the liquid product (stream 47a) flows to storage or further processing. The residue gas (stream 48) at −93° F. [−70° C.] is heated to −40° F. [−40° C.] in cross exchanger 29 (stream 48a) and compressed by compressor 28 to sales line pressure (stream 48b). Following cooling to 78° F. [25° C.] in cross exchanger 29, the residue gas product (stream 48c) flows to the sales gas pipeline at 1315 psia [9,067 kPa(a)] for subsequent distribution.
A summary of stream flow rates and energy consumption for the process illustrated in
If a slightly lower recovery level is acceptable, this prior art process can produce an LPG product using less power and high level utility heat as shown in
A summary of stream flow rates and energy consumption for the process illustrated in
In the simulation of the
The second portion, stream 43, is heated prior to entering fractionation tower 16 so that all or a portion of it is vaporized, reducing the amount of liquid flowing down fractionation tower 16 and allowing the use of a smaller diameter column. In the example shown in
The demethanizer in fractionation tower 16 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. As shown in
Overhead distillation stream 46 is withdrawn from the upper section of fractionation tower 16 at −146° F. [−99° C.] and flows to reflux condenser 17 where it is cooled to −147° F. [−99° C.] and partially condensed by heat exchange with the cold LNG (stream 41a) as described previously. The partially condensed stream 46a enters reflux separator 18 wherein the condensed liquid (stream 49) is separated from the uncondensed vapor (stream 48). The liquid stream 49 from reflux separator 18 is pumped by reflux pump 19 to a pressure slightly above the operating pressure of demethanizer 16 and stream 49a is then supplied as cold top column feed (reflux) to demethanizer 16. 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 16.
The residue gas (stream 48) leaves reflux separator 18 at −147° F. [−99° C.], is heated to −40° F. [−40° C.] in cross exchanger 29 (stream 48a), and is compressed by compressor 28 to sales line pressure (stream 48b). Following cooling to 43° F. [6° C.] in cross exchanger 29, the residue gas product (stream 48c) flows to the sales gas pipeline at 1315 psia [9,067 kPa(a)] for subsequent distribution.
A summary of stream flow rates and energy consumption for the process illustrated in
Comparing the recovery levels displayed in Table X above for the
Comparing the recovery levels displayed in Table X with those in Tables IV and VII for the
There are three primary factors that account for the improved efficiency of the present invention. First, compared to the
The present invention can also be adapted to produce an LPG product containing the majority of the C3 components and heavier hydrocarbon components present in the feed stream as shown in
The processing scheme for the
The liquid product stream 47 exits the bottom of deethanizer 16 at 189° F. [87° C.], based on an ethane to propane ratio of 0.020:1 on a molar basis in the bottom product. After cooling to 124° F. [51° C.] in heat exchanger 13, the liquid product (stream 47a) flows to storage or further processing. The residue gas (stream 48) leaves reflux separator 18 at −94° F. [−70° C.], is heated to −40° F. [−40° C.] in cross exchanger 29 (stream 48a), and is compressed by compressor 28 to sales line pressure (stream 48b). Following cooling to 79° F. [26° C.] in cross exchanger 29, the residue gas product (stream 48c) flows to the sales gas pipeline at 1315 psia [9,067 kPa(a)] for subsequent distribution.
A summary of stream flow rates and energy consumption for the process illustrated in
Comparing the recovery levels displayed in Table XI above for the
Comparing the recovery levels displayed in Table XI with those in Tables V and VIII for the
If a slightly lower recovery level is acceptable, another embodiment of the present invention may be employed to produce an LPG product using much less power and high level utility heat.
In the simulation of the
The combined liquid stream 44 from the bottom of the absorber column 16 is flash expanded to slightly above the operating pressure (430 psia [2,965 kPa(a)]) of stripper column 21 by expansion valve 20, cooling stream 44 to −11° F. [−24° C.] (stream 44a) before it enters fractionation stripper column 21 at a top column feed point. In the stripper column 21, stream 44a is stripped of its methane and C2 components by the vapors generated in reboiler 22 to meet the specification of an ethane to propane ratio of 0.020:1 on a molar basis. The resulting liquid product stream 47 exits the bottom of stripper column 21 at 191° F. [88° C.] and is cooled to 126° F. [52° C.] in heat exchanger 13 (stream 47a) before flowing to storage or further processing.
The overhead vapor (stream 45) from stripper column 21 exits the column at 52° F. [11° C.] and enters overhead compressor 23 (driven by a supplemental power source). Overhead compressor 23 elevates the pressure of stream 45a to slightly above the operating pressure of absorber column 16 so that stream 45a can be supplied to absorber column 16 at a lower column feed point. Stream 45a enters absorber column 16 at 144° F. [62° C.], whereupon it rises upward through absorber column 16 and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier hydrocarbon components.
Overhead distillation stream 46 is withdrawn from contacting device absorber column 16 at −63° F. [−53° C.] and flows to reflux condenser 17 where it is cooled to −78° F. [−61° C.] and partially condensed by heat exchange with the cold LNG (stream 41a) as described previously. The partially condensed stream 46a enters reflux separator 18 wherein the condensed liquid (stream 49) is separated from the uncondensed vapor (stream 48). The liquid stream 49 from reflux separator 18 is pumped by reflux pump 19 to a pressure slightly above the operating pressure of absorber column 16 and stream 49a is then supplied as cold top column feed (reflux) to absorber column 16. This cold liquid reflux absorbs and condenses the C3 components and heavier hydrocarbon components from the vapors rising in absorber column 16.
The residue gas (stream 48) leaves reflux separator 18 at −78° F. [−61° C.], is heated to −40° F. [−40° C.] in cross exchanger 29 (stream 48a), and is compressed by compressor 28 to sales line pressure (stream 48b). Following cooling to −37° F. [−38° C.] in cross exchanger 29, stream 48c is heated to 30° F. [−1° C.] using low level utility heat in heat exchanger 30 and the residue gas product (stream 48d) flows to the sales gas pipeline at 1315 psia [9,067 kPa(a)] for subsequent distribution.
A summary of stream flow rates and energy consumption for the process illustrated in
Comparing Table XII above for the
Comparing the recovery levels displayed in Table XII with those in Tables III, VI, and IX for the
Comparing this embodiment of the present invention to the prior art process displayed in
Since the prior art processes perform rectification and stripping in the same tower (i.e., absorbing section 16a and stripping section 16b contained in fractionation tower 16 in
With overhead compressor 23 supplying the motive force to cause the overhead from stripper column 21 (stream 45 in
The dramatic reduction in the duty of reboiler 22 for the
A slightly more complex design that maintains the same C3 component recovery with lower power consumption can be achieved using another embodiment of the present invention as illustrated in the
In the simulation of the
The combined liquid stream 44 from the bottom of contacting device absorber column 16 is flash expanded to slightly above the operating pressure (430 psia [2,965 kPa(a)]) of stripper column 21 by expansion valve 20, cooling stream 44 to −24° F. [−31° C.] (stream 44a) before it enters fractionation stripper column 21 at a top column feed point. In the stripper column 21, stream 44a is stripped of its methane and C2 components by the vapors generated in reboiler 22 to meet the specification of an ethane to propane ratio of 0.020:1 on a molar basis. The resulting liquid product stream 47 exits the bottom of stripper column 21 at 191° F. [88° C.] and is cooled to 126° F. [52° C.] in heat exchanger 13 (stream 47a) before flowing to storage or further processing.
The overhead vapor (stream 45) from stripper column 21 exits the column at 43° F. [6° C.] and flows to cross exchanger 24 where it is cooled to −47° F. [−44° C.] and partially condensed. Partially condensed stream 45a is further cooled to −99° F. [−73° C.] in heat exchanger 13 as previously described, condensing the remainder of the stream. Condensed liquid stream 45b then enters overhead pump 25, which elevates the pressure of stream 45c to slightly above the operating pressure of absorber column 16. Stream 45c returns to cross exchanger 24 and is heated to 38° F. [3° C.] and partially vaporized as it provides cooling to stream 45. Partially vaporized stream 45d is then supplied to absorber column 16 at a lower column feed point, whereupon its vapor portion rises upward through absorber column 16 and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier hydrocarbon components. The liquid portion of stream 45d commingles with liquids falling downward from the upper section of absorber column 16 and becomes part of combined liquid stream 44 leaving the bottom of absorber column 16.
Overhead distillation stream 46 is withdrawn from contacting device absorber column 16 at −64° F. [−53° C.] and flows to reflux condenser 17 where it is cooled to −78° F. [−61° C.] and partially condensed by heat exchange with the cold LNG (stream 41a) as described previously. The partially condensed stream 46a enters reflux separator 18 wherein the condensed liquid (stream 49) is separated from the uncondensed vapor (stream 48). The liquid stream 49 from reflux separator 18 is pumped by reflux pump 19 to a pressure slightly above the operating pressure of absorber column 16 and stream 49a is then supplied as cold top column feed (reflux) to absorber column 16. This cold liquid reflux absorbs and condenses the C3 components and heavier hydrocarbon components from the vapors rising in absorber column 16.
The residue gas (stream 48) leaves reflux separator 18 at −78° F. [−61° C.], is heated to −40° F. [−40° C.] in cross exchanger 29 (stream 48a), and is compressed by compressor 28 to sales line pressure (stream 48b). Following cooling to −37° F. [−38° C.] in cross exchanger 29, stream 48c is heated to 30° F. [−1° C.] using low level utility heat in heat exchanger 30 and the residue gas product (stream 48d) flows to the sales gas pipeline at 1315 psia [9,067 kPa(a)] for subsequent distribution.
A summary of stream flow rates and energy consumption for the process illustrated in
Comparing Table XIII above for the
In the
Some circumstances may favor cooling the high pressure stream leaving overhead compressor 23, such as with dashed heat exchanger 24 in
Some circumstances may favor using a split feed configuration for the LNG feed (as disclosed previously in
In the
In the
Reflux condenser 17 may be located inside the tower above the rectification section of fractionation tower 16 or absorber column 16 as shown in
It also should be noted that valves 12 and/or 15 could be replaced with expansion engines (turboexpanders) whereby work could be extracted from the pressure reduction of stream 42 in
In
It will be recognized that the relative amount of feed found in each branch of the split LNG feed to fractionation tower 16 or absorber column 16 will depend on several factors, including LNG composition, the amount of heat which can economically be extracted from the feed, residue gas delivery pressure, and the quantity of horsepower available. More feed to the top of the column may increase recovery while increasing the duty in reboiler 22 and thereby increasing the high level utility heat requirements. Increasing feed lower in the column reduces the high level utility heat consumption but may also reduce product recovery. The relative locations of the mid-column feeds may vary depending on LNG composition or other factors such as the desired recovery level and the amount of vapor formed during heating of the feed streams. 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.
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.
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