Patent Grant
7204100
This invention relates to a process for processing natural gas or other methane-rich gas streams to produce a liquefied natural gas (LNG) stream that has a high methane purity and a liquid stream containing predominantly hydrocarbons heavier than methane.
Natural gas is typically recovered from wells drilled into underground reservoirs. It usually has a major proportion of methane, i.e., methane comprises at least 50 mole percent of the gas. Depending on the particular underground reservoir, the natural gas also contains relatively lesser amounts of heavier hydrocarbons such as ethane, propane, butanes, pentanes and the like, as well as water, hydrogen, nitrogen, carbon dioxide, and other gases.
Most natural gas is handled in gaseous form. The most common means for transporting natural gas from the wellhead to gas processing plants and thence to the natural gas consumers is in high pressure gas transmission pipelines. In a number of circumstances, however, it has been found necessary and/or desirable to liquefy the natural gas either for transport or for use. In remote locations, for instance, there is often no pipeline infrastructure that would allow for convenient transportation of the natural gas to market. In such cases, the much lower specific volume of LNG relative to natural gas in the gaseous state can greatly reduce transportation costs by allowing delivery of the LNG using cargo ships and transport trucks.
Another circumstance that favors the liquefaction of natural gas is for its use as a motor vehicle fuel. In large metropolitan areas, there are fleets of buses, taxi cabs, and trucks that could be powered by LNG if there were an economic source of LNG available. Such LNG-fueled vehicles produce considerably less air pollution due to the clean-burning nature of natural gas when compared to similar vehicles powered by gasoline and diesel engines which combust higher molecular weight hydrocarbons. In addition, if the LNG is of high purity (i.e., with a methane purity of 95 mole percent or higher), the amount of carbon dioxide (a “greenhouse gas”) produced is considerably less due to the lower carbon:hydrogen ratio for methane compared to all other hydrocarbon fuels.
The present invention is generally concerned with the liquefaction of natural gas while producing as a co-product a liquid stream consisting primarily of hydrocarbons heavier than methane, such as natural gas liquids (NGL) composed of ethane, propane, butanes, and heavier hydrocarbon components, liquefied petroleum gas (LPG) composed of propane, butanes, and heavier hydrocarbon components, or condensate composed of butanes and heavier hydrocarbon components. Producing the co-product liquid stream has two important benefits: the LNG produced has a high methane purity, and the co-product liquid is a valuable product that may be used for many other purposes. A typical analysis of a natural gas stream to be processed in accordance with this invention would be, in approximate mole percent, 84.2% methane, 7.9% ethane and other C2 components, 4.9% propane and other C3 components, 1.0% 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.
There are a number of methods known for liquefying natural gas. For instance, see Finn, Adrian J., Grant L. Johnson, and Terry R. Tomlinson, “LNG Technology for Offshore and Mid-Scale Plants”, Proceedings of the Seventy-Ninth Annual Convention of the Gas Processors Association, pp. 429–450, Atlanta, Ga., Mar. 13–15, 2000 and Kikkawa, Yoshitsugi, Masaaki Ohishi, and Noriyoshi Nozawa, “Optimize the Power System of Baseload LNG Plant”, Proceedings of the Eightieth Annual Convention of the Gas Processors Association, San Antonio, Tex., Mar. 12–14, 2001 for surveys of a number of such processes. U.S. Pat. Nos. 4,445,917; 4,525,185; 4,545,795; 4,755,200; 5,291,736; 5,363,655; 5,365,740; 5,600,969; 5,615,561; 5,651,269; 5,755,114; 5,893,274; 6,014,869; 6,053,007; 6,062,041; 6,119,479; 6,125,653; 6,250,105 B1; 6,269,655 B1; 6,272,882 B1; 6,308,531 B1; 6,324,867 B1; 6,347,532 B1; PCT Patent Application No. WO 01/88447; and our co-pending U.S. patent application Ser. Nos. 10/161,780 filed Jun. 4, 2002 and Ser. No. 10/278,610 filed Oct. 23, 2002 also describe relevant processes. These methods generally include steps in which the natural gas is purified (by removing water and troublesome compounds such as carbon dioxide and sulfur compounds), cooled, condensed, and expanded. Cooling and condensation of the natural gas can be accomplished in many different manners. “Cascade refrigeration” employs heat exchange of the natural gas with several refrigerants having successively lower boiling points, such as propane, ethane, and methane. As an alternative, this heat exchange can be accomplished using a single refrigerant by evaporating the refrigerant at several different pressure levels. “Multi-component refrigeration” employs heat exchange of the natural gas with one or more refrigerant fluids composed of several refrigerant components in lieu of multiple single-component refrigerants. Expansion of the natural gas can be accomplished both isenthalpically (using Joule-Thomson expansion, for instance) and isentropically (using a work-expansion turbine, for instance).
Regardless of the method used to liquefy the natural gas stream, it is common to require removal of a significant fraction of the hydrocarbons heavier than methane before the methane-rich stream is liquefied. The reasons for this hydrocarbon removal step are numerous, including the need to control the heating value of the LNG stream, and the value of these heavier hydrocarbon components as products in their own right. Unfortunately, little attention has been focused heretofore on the efficiency of the hydrocarbon removal step.
In accordance with the present invention, it has been found that careful integration of the hydrocarbon removal step into the LNG liquefaction process can produce both LNG and a separate heavier hydrocarbon liquid product using significantly less energy than prior art processes. The present invention, although applicable at lower pressures, is particularly advantageous when processing feed gases in the range of 400 to 1500 psia [2,758 to 10,342 kPa(a)] or higher.
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. The production rates reported as pounds per hour (Lb/Hr) correspond to the stated molar flow rates in pound moles per hour. The production rates reported as kilograms per hour (kg/Hr) correspond to the stated molar flow rates in kilogram moles per hour.
Referring now to
The feed stream 31 is cooled in heat exchanger 10 by heat exchange with refrigerant streams and flashed separator liquids at −44° F. [−42° C.] (stream 39a). Note that in all cases heat exchanger 10 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 cooling services will depend on a number of factors including, but not limited to, inlet gas flow rate, heat exchanger size, stream temperatures, etc.) The cooled stream 31a enters separator 11 at 0° F. [−18° C.] and 1278 psia [8,812 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33).
The vapor (stream 32) from separator 11 is divided into two streams, 34 and 36, with stream 34 containing about 15% of the total vapor. Some circumstances may favor combining stream 34 with some portion of the condensed liquid (stream 38) to form combined stream 35, but in this simulation there is no flow in stream 38. Stream 35 passes through heat exchanger 13 in heat exchange relation with refrigerant stream 71e and liquid distillation stream 40, resulting in cooling and substantial condensation of stream 35a. The substantially condensed stream 35a at −109° F. [−78° C.] is then flash expanded through an appropriate expansion device, such as expansion valve 14, to the operating pressure (approximately 465 psia [3,206 kPa(a)]) of fractionation tower 19. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in
The remaining 85% of the vapor from separator 11 (stream 36) enters a work expansion machine 15 in which mechanical, energy is extracted from this portion of the high pressure feed. The machine 15 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 36a to a temperature of approximately −76° F. [−60° C.]. The typical commercially available expanders are capable of recovering on the order of 80–85% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item 16) that can be used to re-compress the tower overhead gas (stream 49), for example. The expanded and partially condensed stream 36a is supplied as feed to absorbing section 19a in distillation column 19 at a lower mid-column feed point. Stream 39, the remaining portion of the separator liquid (stream 33) is flash expanded to slightly above the operating pressure of demethanizer 19 by expansion valve 12, cooling stream 39 to −44° F. [−42° C.] (stream 39a) before it provides cooling to the incoming feed gas as described earlier. Stream 39b, now at 85° F. [29° C.], then enters stripping section 19b in demethanizer 19 at a second lower mid-column feed point.
The demethanizer in fractionation tower 19 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 is often the case in natural gas processing plants, the fractionation tower may consist of two sections. The upper absorbing (rectification) section 19a contains the trays and/or packing to provide the necessary contact between the vapor portion of the expanded stream 36a rising upward and cold liquid falling downward to condense and absorb the ethane, propane, and heavier components; and the lower, stripping section 19b contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The stripping section also includes one or more reboilers (such as reboiler 20) 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. The liquid product stream 41 exits the bottom of demethanizer 19 at 150° F. [66° C.], based on a typical specification of a methane to ethane ratio of 0.020:1 on a molar basis in the bottom product. The overhead distillation vapor stream 37, containing predominantly methane and lighter components, leaves the top of demethanizer 19 at −108° F. [−78° C.].
A portion of the distillation vapor (stream 42) is withdrawn from the upper region of stripping section 19b. This stream is cooled from −58° F. [−50° C.] to −109° F. [−78° C.]and partially condensed (stream 42a) in heat exchanger 13 by heat exchange with refrigerant stream 71e and liquid distillation stream 40. The operating pressure in reflux separator 22 (461 psia [3,182 kPa(a)]) is maintained slightly below the operating pressure of demethanizer 19. This provides the driving force which causes distillation vapor stream 42 to flow through heat exchanger 13 and thence into the reflux separator 22 wherein the condensed liquid (stream 44) is separated from any uncondensed vapor (stream 43). Stream 43 combines with the distillation vapor stream (stream 37) leaving the upper region of absorbing section 19a of demethanizer 19 to form cold residue gas stream 47 at −108° F. [−78° C.].
The condensed liquid (stream 44) is pumped to higher pressure by pump 23, whereupon stream 44a at −109° F. [−78° C.] is divided into two portions. One portion, stream 45, is routed to the upper region of absorbing section 19a of demethanizer 19 to serve as the cold liquid that contacts the vapors rising upward through the absorbing section. The other portion is supplied to the upper region of stripping section 19b of demethanizer 19 as reflux stream 46.
Liquid distillation stream 40 is withdrawn from a lower region of absorbing section 19a of demethanizer 19 and is routed to heat exchanger 13 where it is heated as it provides cooling of distillation vapor stream 42, combined stream 35, and refrigerant (stream 71a). The liquid distillation stream is heated from −79° F. [−62° C.] to −20° F. [−29° C.], partially vaporizing stream 40a before it is supplied as a mid-column feed to stripping section 19b in demethanizer 19.
The cold residue gas (stream 47) is warmed to 94° F. [34° C.] in heat exchanger 24, and a portion (stream 48) is then withdrawn to serve as fuel gas for the plant. (The amount of fuel gas that must be withdrawn is largely determined by the fuel required for the engines and/or turbines driving the gas compressors in the plant, such as refrigerant compressors 64, 66, and 68 in this example.) The remainder of the warmed residue gas (stream 49) is compressed by compressor 16 driven by expansion machines 15, 61, and 63. After cooling to 100° F. [38° C.] in discharge cooler 25, stream 49b is further cooled to −93° F. [−69° C.] (stream 49c) in heat exchanger 24 by cross exchange with cold residue gas stream 47.
Stream 49c then enters heat exchanger 60 and is further cooled by expanded refrigerant stream 71d to −256° F. [−160° C.] to condense and subcool it, whereupon it enters a work expansion machine 61 in which mechanical energy is extracted from the stream. The machine 61 expands liquid stream 49d substantially isentropically from a pressure of about 638 psia [4,399 kPa(a)] to the LNG storage pressure (15.5 psia [107 kPa(a)]), slightly above atmospheric pressure. The work expansion cools the expanded stream 49e to a temperature of approximately −257° F. [−160° C.], whereupon it is then directed to the LNG storage tank 62 which holds the LNG product (stream 50).
All of the cooling for stream 49c and a portion of the cooling for streams 35 and 42 is provided by a closed cycle refrigeration loop. The working fluid for this refrigeration cycle is a mixture of hydrocarbons and nitrogen, with the composition of the mixture adjusted as needed to provide the required refrigerant temperature while condensing at a reasonable pressure using the available cooling medium. In this case, condensing with cooling water has been assumed, so a refrigerant mixture composed of nitrogen, methane, ethane, propane, and heavier hydrocarbons is used in the simulation of the
The refrigerant stream 71 leaves discharge cooler 69 at 100° F. [38° C.] and 607 psia [4,185 kPa(a)]. It enters heat exchanger 10 and is cooled to −15° F. [−26° C.] and partially condensed by the partially warmed expanded refrigerant stream 71f and by other refrigerant streams. For the
The superheated refrigerant vapor (stream 71g) leaves heat exchanger 10 at 93° F. [34° C.] and is compressed in three stages to 617 psia [4,254 kPa(a)]. Each of the three compression stages (refrigerant compressors 64, 66, and 68) is driven by a supplemental power source and is followed by a cooler (discharge coolers 65, 67, and 69) to remove the heat of compression. The compressed stream 71 from discharge cooler 69 returns to heat exchanger 10 to complete the cycle.
A summary of stream flow rates and energy consumption for the process illustrated in
The efficiency of LNG production processes is typically compared using the “specific power consumption” required, which is the ratio of the total refrigeration compression power to the total liquid production rate. Published information on the specific power consumption for prior art processes for producing LNG indicates a range of 0.168 HP-Hr/Lb [0.276 kW-Hr/kg] to 0.182 HP-Hr/Lb [0.300 kW-Hr/kg], which is believed to be based on an on-stream factor of 340 days per year for the LNG production plant. On this same basis, the specific power consumption for the
There are two primary factors that account for the improved efficiency of the present invention. The first factor can be understood by examining the thermodynamics of the liquefaction process when applied to a high pressure gas stream such as that considered in this example. Since the primary constituent of this stream is methane, the thermodynamic properties of methane can be used for the purposes of comparing the liquefaction cycle employed in the prior art processes versus the cycle used in the present invention.
Contrast this now with the liquefaction cycle of the present invention. After partial cooling at high pressure (path A–A′), the gas stream is work expanded (path A′–A″) to an intermediate pressure. (Again, fully isentropic expansion is displayed in the interest of simplicity.) The remainder of the cooling is accomplished at the intermediate pressure (path A″–B′), and the stream is then expanded (path B′–C) to the pressure of the LNG storage vessel. Since the lines of constant entropy slope less steeply in the vapor region of the phase diagram, a significantly larger enthalpy reduction is provided by the first work expansion step (path A′–A″) of the present invention. Thus, the total amount of cooling required for the present invention (the sum of paths A–A′ and A″–B′) is less than the cooling required for the prior art processes (path A–B), reducing the refrigeration (and hence the refrigeration compression) required to liquefy the gas stream.
The second factor accounting for the improved efficiency of the present invention is the superior performance of hydrocarbon distillation systems at lower operating pressures. The hydrocarbon removal step in most of the prior art processes is performed at high pressure, typically using a scrub column that employs a cold hydrocarbon liquid as the absorbent stream to remove the heavier hydrocarbons from the incoming gas stream. Operating the scrub column at high pressure is not very efficient, as it results in the co-absorption of a significant fraction of the methane from the gas stream, which must subsequently be stripped from the absorbent liquid and cooled to become part of the LNG product. In the present invention, the hydrocarbon removal step is conducted at the intermediate pressure where the vapor-liquid equilibrium is much more favorable, resulting in very efficient recovery of the desired heavier hydrocarbons in the co-product liquid stream.
One skilled in the art will recognize that the present invention can be adapted for use with all types of LNG liquefaction plants to allow co-production of an NGL stream, an LPG stream, or a condensate stream, as best suits the needs at a given plant location. Further, it will be recognized that a variety of process configurations may be employed for recovering the liquid co-product stream. The present invention can be adapted to recover an NGL stream containing a significantly higher fraction of the C2 components present in the feed gas, to recover an LPG stream containing only the C3 and heavier components present in the feed gas, or to recover a condensate stream containing only the C4 and heavier components present in the feed gas, rather than producing an NGL co-product containing only a moderate fraction of the C2 components as described earlier. The present invention is particularly advantageous over the prior art processes when only partial recovery of the C2 components in the feed gas is desired while capturing essentially all of the C3 and heavier components, as the reflux stream 45 in the
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 pumped condensed liquid (stream 44a) leaving reflux separator 22 and all or a part of the expanded substantially condensed stream 35b from expansion valve 14 can be combined (such as in the piping joining the expansion valve 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 the two streams shall be considered for the purposes of this invention as constituting an absorbing section.
As described earlier, the distillation vapor stream 42 is partially condensed and the resulting condensate used to absorb valuable C3 components and heavier components from the vapors rising through absorbing section 19a of demethanizer 19 (
In the practice of the present invention, there will necessarily be a slight pressure difference between demethanizer 19 and reflux separator 22 which must be taken into account. If the distillation vapor stream 42 passes through heat exchanger 13 and into reflux separator 22 without any boost in pressure, the reflux separator shall necessarily assume an operating pressure slightly below the operating pressure of demethanizer 19. In this case, the liquid stream withdrawn from the reflux separator can be pumped to its feed position(s) in the demethanizer. An alternative is to provide a booster blower for distillation vapor stream 42 to raise the operating pressure in heat exchanger 13 and reflux separator 22 sufficiently so that the liquid stream 44 can be supplied to demethanizer 19 without pumping.
The high pressure liquid (stream 33 in
In accordance with this invention, the splitting of the vapor feed may be accomplished in several ways. In the processes of
Some circumstances may favor withdrawing all of the cold liquid distillation stream 40 leaving absorbing section 19a in
The disposition of the gas stream remaining after recovery of the liquid co-product stream (stream 47 in
In accordance with the present invention, the cooling of the inlet gas stream and the feed stream to the LNG production section may be accomplished in many ways. In the processes of
Further, the supplemental external refrigeration that is supplied to the inlet gas stream and to the feed stream to the LNG production section may also be accomplished in many different ways. In
Subcooling of the condensed liquid stream leaving heat exchanger 60 (stream 49d in
Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where appropriate. For example, conditions may warrant work expansion of the substantially condensed feed stream (stream 35a in
It will also be recognized that the relative amount of feed found in each branch of the split vapor feed will depend on several factors, including gas pressure, feed gas composition, the amount of heat which can economically be extracted from the feed, the hydrocarbon components to be recovered in the liquid co-product stream, and the quantity of horsepower available. More feed to the top of the column may increase recovery while decreasing power recovered from the expander thereby increasing the recompression horsepower requirements. Increasing feed lower in the column reduces the horsepower consumption but may also reduce product recovery. The relative locations of the mid-column feeds may vary depending on inlet composition or other factors such as desired recovery levels and amount of liquid formed during inlet gas cooling. 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.
| Number | Name | Date | Kind |
|---|---|---|---|
| 2952984 | Marshall | Sep 1960 | A |
| 3292380 | Bucklin | Dec 1966 | A |
| 3837172 | Markbreiter et al. | Sep 1974 | A |
| 4140504 | Campbell et al. | Feb 1979 | A |
| 4157904 | Campbell et al. | Jun 1979 | A |
| 4171964 | Campbell et al. | Oct 1979 | A |
| 4185978 | McGalliard et al. | Jan 1980 | A |
| 4251249 | Gulsby | Feb 1981 | A |
| 4278457 | Campbell et al. | Jul 1981 | A |
| 4445917 | Chiu | May 1984 | A |
| 4519824 | Huebel | May 1985 | A |
| 4525185 | Newton | Jun 1985 | A |
| 4545795 | Liu et al. | Oct 1985 | A |
| 4600421 | Kummann | Jul 1986 | A |
| 4617039 | Buck | Oct 1986 | A |
| 4687499 | Aghili | Aug 1987 | A |
| 4689063 | Paradowski et al. | Aug 1987 | A |
| 4690702 | Paradowski et al. | Sep 1987 | A |
| 4707170 | Ayres et al. | Nov 1987 | A |
| 4710214 | Sharma et al. | Dec 1987 | A |
| 4755200 | Liu et al. | Jul 1988 | A |
| 4851020 | Montgomery, IV | Jul 1989 | A |
| 4854955 | Campbell et al. | Aug 1989 | A |
| 4869740 | Campbell et al. | Sep 1989 | A |
| 4889545 | Campbell et al. | Dec 1989 | A |
| 4895584 | Buck et al. | Jan 1990 | A |
| RE33408 | Khan | Oct 1990 | E |
| 5114451 | Rambo et al. | May 1992 | A |
| 5275005 | Campbell et al. | Jan 1994 | A |
| 5291736 | Paradowski | Mar 1994 | A |
| 5363655 | Kikkawa et al. | Nov 1994 | A |
| 5365740 | Kikkawa et al. | Nov 1994 | A |
| 5555748 | Campbell et al. | Sep 1996 | A |
| 5566554 | Vijayaraghavan et al. | Oct 1996 | A |
| 5568737 | Campbell et al. | Oct 1996 | A |
| 5600969 | Low | Feb 1997 | A |
| 5615561 | Houshmand et al. | Apr 1997 | A |
| 5651269 | Prevost et al. | Jul 1997 | A |
| 5755114 | Foglietta | May 1998 | A |
| 5755115 | Manley | May 1998 | A |
| 5771712 | Campbell et al. | Jun 1998 | A |
| 5799507 | Wilkinson et al. | Sep 1998 | A |
| 5881569 | Campbell et al. | Mar 1999 | A |
| 5890378 | Rambo et al. | Apr 1999 | A |
| 5893274 | Nagelvoort et al. | Apr 1999 | A |
| 5983664 | Campbell et al. | Nov 1999 | A |
| 6014869 | Elion et al. | Jan 2000 | A |
| 6023942 | Thomas et al. | Feb 2000 | A |
| 6053007 | Victory et al. | Apr 2000 | A |
| 6062041 | Kikkawa et al. | May 2000 | A |
| 6116050 | Yao et al. | Sep 2000 | A |
| 6119479 | Roberts et al. | Sep 2000 | A |
| 6125653 | Shu et al. | Oct 2000 | A |
| 6182469 | Campbell et al. | Feb 2001 | B1 |
| 6250105 | Kimble | Jun 2001 | B1 |
| 6269655 | Roberts et al. | Aug 2001 | B1 |
| 6272882 | Hodges et al. | Aug 2001 | B1 |
| 6308531 | Roberts et al. | Oct 2001 | B1 |
| 6324867 | Fanning et al. | Dec 2001 | B1 |
| 6336344 | O'Brien | Jan 2002 | B1 |
| 6347532 | Agrawal et al. | Feb 2002 | B1 |
| 6363744 | Finn et al. | Apr 2002 | B2 |
| 6367286 | Price | Apr 2002 | B1 |
| 6526777 | Campbell et al. | Mar 2003 | B1 |
| 6604380 | Reddick et al. | Aug 2003 | B1 |
| 6742358 | Wilkinson et al. | Jun 2004 | B2 |
| 6907752 | Schroeder et al. | Jun 2005 | B2 |
| 6941771 | Reddick et al. | Sep 2005 | B2 |
| 20030005722 | Wilkinson et al. | Jan 2003 | A1 |
| 20030158458 | Prim | Aug 2003 | A1 |
| 20040079107 | Wilkinson et al. | Apr 2004 | A1 |
| 20050061029 | Narinsky | Mar 2005 | A1 |
| 20050155381 | Yang et al. | Jul 2005 | A1 |
| Number | Date | Country |
|---|---|---|
| 1535846 | Aug 1968 | FR |
| 0188447 | Nov 2001 | WO |
| 2004109180 | Dec 2004 | WO |
| 2005015100 | Feb 2005 | WO |
| 2005035692 | Apr 2005 | WO |
| Number | Date | Country | |
|---|---|---|---|
| 20050247078 A1 | Nov 2005 | US |
