Patent Application
20100326134
Hydrocarbons are found in a variety of gases, such as natural gas, refinery gas, and synthetic gas streams obtained from other hydrocarbon materials such as coal, crude oil, naphtha, oil shale, tar sands, and lignite. The hydrocarbon constituents of natural gas usually have a major proportion of methane and ethane, i.e., methane and ethane together comprise at least 50 mole percent of the total hydrocarbon components. The gas also usually contains relatively lesser amounts of heavier hydrocarbons such as propane, butanes, pentanes, and the like, as well as carbon dioxide, hydrogen, nitrogen, and other gases.
In many cases, the gas streams from these sources are contaminated with high concentrations of carbon dioxide, making the gas streams unsuitable for use as fuel, chemical plant feedstock, or other purposes. There are a variety of processes that have been developed to remove the carbon dioxide using chemical, physical, and hybrid solvents. Other processes have been developed that use a refrigerated absorbent stream composed of heavy (C4-C10 typically) hydrocarbons to remove the carbon dioxide in a distillation column, such as the process described in U.S. Pat. No. 4,318,723. All of these processes have increasingly higher capital cost and operating cost as the carbon dioxide concentration in the gas stream increases, which often makes processing of such gas streams uneconomical.
One method for improving the economics of processing gas streams containing high concentrations of carbon dioxide is to provide bulk separation of the carbon dioxide from the gas stream prior to processing with solvents or absorbents, so that only a minor fraction of the carbon dioxide must then be removed from the gas stream. For example, semi-permeable membranes have often been used for bulk removal of carbon dioxide. However, a significant fraction of the lighter hydrocarbons in the gas stream are often “lost” in the carbon dioxide stream that is separated by bulk removal processes of this type.
A better alternative for bulk removal of carbon dioxide is to use distillation to fractionate the gas stream into a light hydrocarbon stream and a carbon dioxide stream, so that removal of the residual carbon dioxide from the light hydrocarbon stream is all that is required to produce pipeline-quality gas for use as fuel, chemical plant feedstock, etc. The majority of the carbon dioxide that is removed is recovered as a liquid rather than a vapor, allowing the carbon dioxide to be pumped (rather than compressed) for subsequent use in tertiary oil recovery operations or for other purposes, resulting in substantial reductions in capital cost and operating cost.
The present invention is generally concerned with the removal of the majority of the carbon dioxide from such gas streams. A typical analysis of a gas stream to be processed in accordance with this invention would be, in approximate mole percent, 31.0% methane, 0.4% ethane and other C2 components, 0.1% propane and other C3 components, 0.1% butanes plus, and 67.7% carbon dioxide, with the balance made up of nitrogen. Sulfur containing gases are also sometimes present.
In a typical distillation process for removing carbon dioxide, a feed gas stream under pressure is cooled by heat exchange with other streams of the process and/or external sources of refrigeration such as a Freon compression-refrigeration system. The gas is condensed as it is cooled, and the high-pressure liquid is expanded to an intermediate pressure, resulting in further cooling of the stream due to the vaporization occurring during expansion of the liquids. The expanded stream, comprising a mixture of liquid and vapor, is fractionated in a distillation column to separate residual methane, nitrogen, and other volatile gases as overhead vapor from the carbon dioxide and the heavier hydrocarbon components as bottom liquid product. A portion of the liquid carbon dioxide can be flash expanded to lower pressure and thereafter used to provide low level refrigeration to the process streams if desired.
The present invention employs a novel means of performing the various steps described above more efficiently and using fewer pieces of equipment. This is accomplished by combining what heretofore have been individual equipment items into a common housing, thereby reducing the plot space required for the processing plant and reducing the capital cost of the facility. Surprisingly, applicants have found that the more compact arrangement also reduces the power consumption required to achieve a given level of carbon dioxide removal, thereby increasing the process efficiency and reducing the operating cost of the facility. In addition, the more compact arrangement also eliminates much of the piping used to interconnect the individual equipment items in traditional plant designs, further reducing capital cost and also eliminating the associated flanged piping connections. Since piping flanges are a potential leak source for carbon dioxide (a greenhouse gas) and for hydrocarbons (which are volatile organic compounds, VOCs, that contribute to greenhouse gases and may also be precursors to atmospheric ozone formation), eliminating these flanges reduces the potential for atmospheric emissions that can damage the environment.
In accordance with the present invention, it has been found that more than 87% of the carbon dioxide can be removed while leaving more than 99% of the methane in the residue gas stream. The present invention, although applicable at lower pressures and warmer temperatures, is particularly advantageous when processing feed gases in the range of 400 to 1500 psia [2,758 to 10,342 kPa(a)] or higher under conditions requiring distillation column overhead temperatures of −50° F. [−46° C.] or colder.
For a better understanding of the present invention, reference is made to the following examples and drawings. Referring to the drawings:
In the following explanation of the above figures, tables are provided summarizing flow rates calculated for representative process conditions. In the tables appearing herein, the values for flow rates (in moles per hour) have been rounded to the nearest whole number for convenience. The total stream rates shown in the tables include all non-hydrocarbon components and hence are generally larger than the sum of the stream flow rates for the hydrocarbon components. Temperatures indicated are approximate values rounded to the nearest degree. It should also be noted that the process design calculations performed for the purpose of comparing the processes depicted in the figures are based on the assumption of no heat leak from (or to) the surroundings to (or from) the process. The quality of commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art.
For convenience, process parameters are reported in both the traditional British units and in the units of the Système International d'Unités (SI). The molar flow rates given in the tables may be interpreted as either pound moles per hour or kilogram moles per hour. The energy consumptions reported as horsepower (HP) and/or thousand British Thermal Units per hour (MBTU/Hr) correspond to the stated molar flow rates in pound moles per hour. The energy consumptions reported as kilowatts (kW) correspond to the stated molar flow rates in kilogram moles per hour.
The feed stream 31 is cooled to 1° F. [−17° C.] in heat exchanger 10 by heat exchange with column reboiler liquids at 43° F. [6° C.] (stream 36), column side reboiler liquids at 30° F. [−1° C.] (stream 35), and Freon R134a refrigerant. Stream 31a is further cooled in heat exchanger 11 by heat exchange with cool carbon dioxide vapor at −49° F. [−45° C.] (stream 39), cold residue gas at −60° F. [−51° C.] (stream 33), and pumped liquid at −60° F. [−51° C.] (stream 34a). The further cooled stream 31b at −14° F. [−25° C.] and 1060 psia [7,308 kPa(a)] is expanded to the operating pressure (approximately 617 psia [4,254 kPa(a)]) of fractionation tower 13 by expansion valve 12, cooling stream 31c to −35° F. [−37° C.] before it is supplied to fractionation tower 13 at its top column feed point.
Overhead vapor stream 32 leaves fractionation tower 13 at −35° F. [−37° C.] and is cooled and partially condensed in heat exchanger 14. The partially condensed stream 32a enters separator 15 at −60° F. [−51° C.] and 610 psia [4,206 kPa(a)] where the vapor (cold residue gas stream 33) is separated from the condensed liquid (stream 34). Liquid stream 34 is pumped to slightly above the operating pressure of fractionation tower 13 by pump 19 before stream 34a enters heat exchanger 11 and is heated to −6° F. [−21° C.] and partially vaporized by heat exchange with the feed gas as described previously. The partially vaporized stream 34b is thereafter supplied as feed to fractionation tower 13 at a mid-column feed point.
Fractionation tower 13 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. It also includes reboilers (such as the reboiler and the side reboiler described previously) 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 column bottom liquid product (stream 37) of methane and lighter components. The trays and/or packing provide the necessary contact between the stripping vapors rising upward and cold liquid falling downward, so that the bottom product stream 37 exits the bottom of the tower at 45° F. [7° C.], based on reducing the methane concentration in the bottom product to 0.50% on a molar basis.
Column bottom product stream 37 is predominantly liquid carbon dioxide. A small portion (stream 38) is subcooled in heat exchanger 16 by cool residue gas stream 33a. The subcooled liquid (stream 38a) at 0° F. [−18° C.] is expanded to lower pressure by expansion valve 20 and partially vaporized, further cooling stream 38b to −63° F. [−53° C.] before it enters heat exchanger 14. The residual liquid in stream 38b functions as refrigerant in heat exchanger 14 to provide cooling of stream 32 as described previously, with the resulting carbon dioxide vapor leaving at −49° F. [−45° C.] as stream 39. Since stream 38b contains a small amount of heavier hydrocarbons, a small liquid purge (stream 40) is drawn off from heat exchanger 14 to prevent an accumulation of heavier hydrocarbons in the refrigerant liquid that could elevate its boiling point and reduce the cooling efficiency of heat exchanger 14.
The cool carbon dioxide vapor from heat exchanger 14 (stream 39) is heated to −6° F. [−21° C.] in heat exchanger 11 by heat exchange with the feed gas as described previously. The warm carbon dioxide vapor (stream 39a) at 76 psia [524 kPa(a)] is then compressed to high pressure in four stages by compressors 21, 23, 25, and 27, with cooling to 120° F. [49° C.] after each stage of compression by discharge coolers 22, 24, 26, and 28. The remaining portion (stream 41) of column bottom product stream 37 is pumped to high pressure by pump 29 so that stream 41a can combine with the dense phase fluid (stream 39i) leaving discharge cooler 28, forming high pressure carbon dioxide stream 42 which then flows to reinjection at 79° F. [26° C.] and 2015 psia [13,893 kPa(a)].
The cool residue gas (stream 33a) leaves heat exchanger 11 at −6° F. [−21 ° C.] after heat exchange with the feed gas as described previously, and is further heated to 27° F. [−3° C.] in heat exchanger 16 by heat exchange with liquid carbon dioxide stream 38 as described previously. The warm residue gas stream 33b is then re-compressed by compressor 17 and cooled to 120° F. [49° C.] in discharge cooler 18 before it flows to treating at 1215 psia [8,377 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
In the simulation of the
The partially cooled stream 31a enters a heat exchange means in feed cooling section 113b inside processing assembly 113. This heat exchange means may also be comprised of a fin and tube type heat exchanger, a plate type heat exchanger, a brazed aluminum type heat exchanger, or other type of heat transfer device, including multi-pass and/or multi-service heat exchangers. The heat exchange means is configured to provide heat exchange between stream 31a flowing through one pass of the heat exchange means, Freon R134a refrigerant streams, a residue gas stream arising from rectifying section 113c inside processing assembly 113, and cool carbon dioxide stream 38c (which is described further below in paragraph [0028]). Stream 31a is further cooled while heating these other streams, whereupon the further cooled stream 31b at −14° F. [−25° C.] and 1076 psia [7,419 kPa(a)] is expanded to the operating pressure (approximately 618 psia [4,264 kPa(a)]) of stripping section 113d by expansion valve 12, cooling stream 31c to −35° F. [−37° C.] before it is supplied to processing assembly 113 above the mass transfer means in the upper region of stripping section 113d.
Stripping section 113d inside processing assembly 113 contains a mass transfer means consisting of 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 vapors rising upward and cold liquid falling downward. The liquid portion of expanded stream 31c commingles with liquids falling downward from rectifying section 113c and the combined liquid continues downward to the mass transfer means in the upper region of stripping section 113d. The stripping vapors arising from the heat and mass transfer means in the lower region of stripping section 113d rise upward through the mass transfer means in the upper region of stripping section 113d to partially strip the combined liquid of methane and lighter components, with the remaining stripping vapors leaving the mass transfer means and combining with the vapor portion of expanded stream 31c to form a distillation vapor stream that continues upward to rectifying section 113c. The partially stripped distillation liquid leaving the mass transfer means in the upper region of stripping section 113d continues downward into the heat and mass transfer means in the lower region of stripping section 113d to continue the stripping of methane and lighter components from the liquid. The resulting liquid product (stream 37) exits the lower region of stripping section 113d and leaves processing assembly 113 at 45° F. [7° C.], based on reducing the methane concentration in the bottom product to 0.50% on a molar basis.
Column bottom product stream 37 is predominantly liquid carbon dioxide. A small portion (stream 38) enters a heat exchange means in refrigerant subcooling section 113a inside processing assembly 113. This heat exchange means may likewise be comprised of a fin and tube type heat exchanger, a plate type heat exchanger, a brazed aluminum type heat exchanger, or other type of heat transfer device, including multi-pass and/or multi-service heat exchangers. The heat exchange means is configured to provide heat exchange between stream 38 flowing through one pass of the heat exchange means and the residue gas stream from rectifying section 113c inside processing assembly 113 that has been heated in the heat exchange means in feed cooling section 113b, so that stream 38 is subcooled while further heating the residue gas stream. The resulting subcooled liquid (stream 38a) at 0° F. [−18° C.] is expanded to lower pressure by expansion valve 20 and partially vaporized, further cooling stream 38b to −64° F. [−54° C.] before it enters a heat and mass transfer means in rectifying section 113c inside processing assembly 113. This heat and mass transfer means may likewise be comprised of a fin and tube type heat exchanger, a plate type heat exchanger, a brazed aluminum type heat exchanger, or other type of heat transfer device, including multi-pass and/or multi-service heat exchangers. The heat and mass transfer means is configured to provide heat exchange between the distillation vapor stream arising from stripping section 113d flowing upward through one pass of the heat and mass transfer means and the expanded stream 38b flowing downward, so that the distillation vapor is cooled while heating the expanded stream. As the distillation vapor stream is cooled, a portion of it is condensed and falls downward while the remaining distillation vapor continues flowing upward through the heat and mass transfer means. The heat and mass transfer means provides continuous contact between the condensed liquid and the distillation vapor so that it also functions to provide mass transfer between the vapor and liquid phases, thereby providing rectification of the distillation vapor. The condensed liquid is collected from the bottom of the heat and mass transfer means, combined with the liquid portion of expanded stream 31c, and directed to the mass transfer means in the upper region of stripping section 113d as described previously. The residual liquid in stream 38b functions as refrigerant in the heat and mass transfer means to provide the cooling and rectification of the distillation vapor stream, with the resulting mostly vaporized carbon dioxide leaving the heat and mass transfer means in rectifying section 113c at −48° F. [−44° C.] as stream 38c.
The cool carbon dioxide from the heat and mass transfer means in rectifying section 113c inside processing assembly 113 (stream 38c) is heated to −10° F. [−23° C.] in the heat exchange means in feed cooling section 113b inside processing assembly 113 by heat exchange with the feed gas as described previously. The warm carbon dioxide vapor (stream 38d) at 76 psia [527 kPa(a)] is then compressed to high pressure in four stages by compressors 21, 23, 25, and 27, with cooling to 120° F. [49° C.] after each stage of compression by discharge coolers 22, 24, 26, and 28. The remaining portion (stream 41) of column bottom product stream 37 is pumped to high pressure by pump 29 so that stream 41a can combine with the dense phase fluid (stream 38l) leaving discharge cooler 28, forming high pressure carbon dioxide stream 42 which then flows to reinjection at 78° F. [26° C.] and 2015 psia [13,893 kPa(a)].
The vapor arising from rectifying section 113c is the residue gas. It is warmed in feed cooling section 113b as it provides cooling to the feed gas as described previously. The residue gas stream is further heated in refrigerant subcooling section 113a as it provides cooling to stream 38 as described previously, whereupon residue gas stream 33 leaves processing assembly 113 at 20° F. [−7° C.]. The residue gas stream is then re-compressed by compressor 17 and cooled to 120° F. [49° C.] in discharge cooler 18 before it flows to treating at 1215 psia [8,377 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in
A comparison of Tables I and II shows that, compared to the prior art, the present invention maintains essentially the same methane recovery (99.03% versus 99.04% for the prior art), carbon dioxide removal (87.36% versus 87.37% for the prior art), carbon dioxide concentration in the residue gas (21.37% versus 21.37% for the prior art), and carbon dioxide purity (98.69% versus 98.70% for the prior art). However, further comparison of Tables I and II shows that the process performance was achieved using less power than the prior art. In terms of the total power consumption, the present invention represents a 0.4% improvement over the prior art of the
The improvement in energy efficiency provided by the present invention over that of the prior art of the
Second, using the heat and mass transfer means in the lower region of stripping section 113d to simultaneously heat the distillation liquid leaving the mass transfer means in the upper region of stripping section 113d while allowing the resulting vapors to contact the liquid and strip its volatile components is more efficient than using a conventional distillation column with external reboilers. The volatile components are stripped out of the liquid continuously, reducing the concentration of the volatile components in the stripping vapors more quickly and thereby improving the stripping efficiency for the present invention.
Third, using the heat and mass transfer means in rectifying section 113c to simultaneously cool the distillation vapor stream arising from stripping section 113d while condensing the carbon dioxide and the heavier hydrocarbon components from the distillation vapor stream provides more efficient rectification than using reflux in a conventional distillation column. As a result, less of the liquid carbon dioxide in bottom product stream 37 is needed to provide low level refrigeration to the process compared to the prior art of the
The present invention offers two other advantages over the prior art in addition to the increase in processing efficiency. First, the compact arrangement of processing assembly 113 of the present invention replaces seven separate equipment items in the prior art (heat exchangers 10, 11, 14, and 16, separator 15, pump 19, and fractionation tower 13 in
As described earlier for the embodiment of the present invention shown in
In some circumstances, it may be advantageous to use an external separator vessel to separate cooled feed stream 31b, rather than including separator section 113e in processing assembly 113. As shown in
Feed gas conditions, plant size, available equipment, or other factors may indicate that elimination of work expansion machine 50, or replacement with an alternate expansion device (such as an expansion valve), is feasible in the
In accordance with the present invention, the use of external refrigeration to supplement the cooling available to the inlet gas from the distillation vapor and liquid streams may be employed. In such cases, the heat exchange means in feed cooling section 113b may include one or more passes for the refrigerant as shown by the dashed lines in
Depending on the temperature and richness of the feed gas and the amount of methane allowed in liquid product stream 37, there may not be sufficient heating available from feed stream 31 to cause the liquid leaving stripping section 113d to meet the product specifications. In such cases, the heat and mass transfer means in stripping section 113d may include provisions for providing supplemental heating with heating medium as shown by the dashed lines in
Depending on the type of heat transfer devices selected for the heat exchange means in refrigerant subcooling section 113a and feed cooling section 113b, it may be possible to combine these heat exchange means in a single multi-pass and/or multi-service heat transfer device. In such cases, the multi-pass and/or multi-service heat transfer device will include appropriate means for distributing, segregating, and collecting stream 31a, stream 38, stream 38c, and stream 33 in order to accomplish the desired cooling and heating. Likewise, the type of heat and mass transfer device selected for the heat and mass transfer means in rectifying section 113c may allow combining it with the heat exchange means in feed cooling section 113b (and possibly with the heat exchange means in refrigerant subcooling section 113a as well) in a single multi-pass and/or multi-service heat and mass transfer device. In such cases, the multi-pass and/or multi-service heat and mass transfer device will include appropriate means for distributing, segregating, and collecting stream 31a, stream 38, stream 38b, stream 38c, and stream 33 in order to accomplish the desired cooling and heating.
In some circumstances, the portion (stream 38) of column bottom product stream 37 that is used to provide refrigeration in
The present invention provides improved separation of carbon dioxide from hydrocarbon gas streams 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 re-compression, reduced power requirements for pumping, reduced power requirements for external refrigeration, reduced energy requirements for supplemental heating, reduced energy requirements for tower reboiling, or a combination thereof.
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.
This invention relates to a process and apparatus for the separation of a gas containing hydrocarbons and carbon dioxide. The applicants claim the benefits under Title 35, United States Code, Section 119(e) of prior U.S. Provisional Application No. 61/186,361 which was filed on Jun. 11, 2009. The applicants also claim the benefits under Title 35, United States Code, Section 120 as a continuation-in-part of U.S. patent application Ser. No. 12/781,259 which was filed on May 17, 2010, and as a continuation-in-part of U.S. patent application Ser. No. 12/772,472 which was filed on May 3, 2010, and as a continuation-in-part of U.S. patent application Ser. No. 12/750,862 which was filed on Mar. 31, 2010, and as a continuation-in-part of U.S. patent application Ser. No. 12/717,394 which was filed on Mar. 4, 2010, and as a continuation-in-part of U.S. patent application Ser. No. 12/689,616 which was filed on Jan. 19, 2010, and as a continuation-in-part of U.S. patent application Ser. No. 12/372,604 which was filed on Feb. 17, 2009. Assignees S.M.E. Products LP and Ortloff Engineers, Ltd. were parties to a joint research agreement that was in effect before the invention of this application was made.
| Number | Date | Country | |
|---|---|---|---|
| 61186361 | Jun 2009 | US |
| Number | Date | Country | |
|---|---|---|---|
| Parent | 12781259 | May 2010 | US |
| Child | 12792136 | US | |
| Parent | 12772472 | May 2010 | US |
| Child | 12781259 | US | |
| Parent | 12750862 | Mar 2010 | US |
| Child | 12772472 | US | |
| Parent | 12717394 | Mar 2010 | US |
| Child | 12750862 | US | |
| Parent | 12689616 | Jan 2010 | US |
| Child | 12717394 | US | |
| Parent | 12372604 | Feb 2009 | US |
| Child | 12689616 | US |
