Patent Application
20060272981
This invention relates to a process and an installation for the treatment of heavy hydrocarbon feedstocks that contain sulfur-containing impurities. It relates to a process that makes it possible to convert at least in part such a hydrocarbon feedstock, for example a vacuum distillate that is obtained by direct distillation of a crude oil, into a gas oil that meets the 2005 sulfur specifications, i.e., that has less than 50 ppm of sulfur, and into a heavier product that can advantageously be used as a feedstock for catalytic cracking (such as the fluidized-bed catalytic cracking).
Until 2000, the sulfur content allowed in diesel fuel was 350 ppm. Drastically more restricting values are expected for 2005, however, since this maximum content will be reduced to 50 ppm.
The inventors therefore sought a process that makes it possible to achieve this goal. In providing such a process the goal was to a large extent exceeded since contents of less than 20 ppm and even 10 ppm were generally obtained.
More specifically, the invention relates to a process for treating petroleum feedstocks of which at least 80% by weight boils above 340° C. and which contains at least 0.05% by weight of sulfur for producing at least one gas oil fraction with a sulfur content of at most 50 ppm by weight, whereby said process comprises the following stages:
The treated feedstocks are heavy, i.e., 80% by weight boils above 340° C. Their initial boiling point is generally established at at least 340° C., often at at least 370° C. and even at least 400° C. Very advantageously, the process makes it possible to treat feedstocks that have a final boiling temperature of at least 450° C. and that can even go beyond 700° C.
The sulfur content is at least 0.05% by weight, often at least 1% and very often at least 2%, and even at least 2.5% by weight. Feedstocks with 3% sulfur or more are very suitable in this process.
The feedstocks that can be treated within the framework of this invention are vacuum distillates of direct distillation, vacuum distillates that are obtained from a conversion process such as, for example, those that are obtained from coking, a fixed-bed hydroconversion (such as those that are obtained from the HYVAHL(R) processes for treatment of heavy products developed by Institut Francais du Petrole) or processes for hydrotreatment of heavy products in a boiling bed (such as those that are obtained from H-OIL(R) processes) or else oils that are deasphalted with solvent (for example with propane, butane or pentane) that are obtained from deasphalting of direct distillation vacuum residue or residues that are obtained from HYVAHL(R) and H-OIL(R) processes. The feedstocks can also be formed by mixing these various fractions. They can also contain gas oil fractions and heavy gas oils that are obtained from catalytic cracking that have in general a distillation range of from about 150° C. to about 370° C. They can also contain aromatic extracts and paraffins that are obtained within the framework of the production of lubricating oils. According to this invention, the feedstocks that are treated are preferably vacuum distillates, DAO-type feedstocks, i.e., that contain metals and/or asphaltenes, and for example, more than 10 ppm of metals and more than 1000 ppm of asphaltenes.
Stage a) for hydroconversion where the feedstock that is described below is reacted in an ebulliated-bed reactor.
Said hydrocarbon feedstock is treated in a treatment section in the presence of hydrogen, whereby said section comprises at least one three-phase reactor that contains at least one hydroconversion catalyst, whose mineral support is at least in part amorphous, in an ebulliated bed, operating with an upward flow of liquid and gas, and whereby said reactor comprises at least one means for drawing off the catalyst beyond said reactor that is located near the bottom of the reactor and at least one means for adding fresh catalyst into said reactor that is located close to the top of said reactor.
The operation is usually carried out under an absolute pressure of 2 to 35 MPa, often 4 to 20 MPa, and most often 6 to 20 MPa at a temperature of from about 300 to about 550° C. and often from about 350 to about 470° C. The hourly space velocity (VVH) relative to the volume of catalyst and the partial pressure of hydrogen are important factors that are selected based on characteristics of the product to be treated and the desired conversion. Most often, the VVH relative to the volume of catalyst lies in a range of from about 0.1 h−1 to about 10 h−1 and preferably about 0.5 h−1 to about 5 h−1. The amount of hydrogen that is mixed with the feedstock is usually from about 50 to about 5000 normal cubic meters (Nm3) per cubic meter (m3) of liquid feedstock and most often from about 100 to about 1500 Nm3/m3 and preferably from about 200 to about 500 Nm3/m3.
The conversion of the feedstock into lighter fractions than 360° C. is usually between 10-80% by weight, most often from 25 to 60%.
It is possible to use a standard granular catalyst for hydroconversion that comprises, on an amorphous support, at least one metal or metal compound that has a hydro-dehydrogenating function. This catalyst can be a catalyst that comprises metals of group VIII, for example nickel and/or cobalt most often combined with at least one metal of group IB, for example molybdenum and/or tungsten. It is possible, for example, to use a catalyst that comprises 0.5 to 10% by weight of nickel and preferably 1 to 5% by weight of nickel (expressed in nickel oxide NiO) and 1 to 30% by weight of molybdenum, preferably 5 to 20% by weight of molybdenum (expressed in molybdenum oxide MoO3) on an amorphous mineral support. This support will be selected, for example, from the group that is formed by alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. This support can also contain other compounds and, for example, oxides that are selected from the group that is formed by boron oxide, zirconia, titanium oxide, and phosphoric anhydride. Most often an alumina support, and, very often an alumina subtrate that is doped with phosphorus and optionally boron is used. The concentration of phosphoric anhydride P2O5 is usually less than about 20% by weight and most often less than about 10% by weight. This concentration of P2O5 is usually from at least 0.001% by weight. The concentration of boron trioxide B2O3 is usually from about 0 to about 10% by weight. The alumina that is used is usually a γ-alumina or η-alumina. This catalyst is most often in extrudate form. The total content of metal oxides of groups VI and VIII is often from about 5 to about 40% by weight and in general from about 7 to 30% by weight, and the ratio by weight that is expressed in metallic oxide between metal (or metals) of group VI to metal (or metals) of group VIII is in general from about 20 to about 1 and most often from about 10 to about 2.
The used catalyst is replaced in part by fresh catalyst by drawing off fresh catalyst or new catalyst at the bottom of the reactor and introducing it at the top of the reactor at regular intervals, i.e., for example, in bursts or almost continuously. It is possible, for example, to introduce fresh catalyst every day. The rate of replacement of the waste catalyst by fresh catalyst can be, for example, from about 0.05 kilogram to about 10 kilograms per cubic meter of feedstock. This draw-off and this replacement are carried out with devices that make possible the continuous operation of this hydroconversion stage. The unit usually comprises a recirculation pump that makes it possible to maintain the catalyst in an ebulliated bed by continuous recycling of at least a portion of liquid that is drawn off from stage a) and reinjected in the bottom of the zone of stage a). It is also possible to send the used catalyst that is drawn off from the reactor into a regeneration zone in which the carbon and the sulfur that it contains are eliminated and then to send this regenerated catalyst into hydroconversion stage b).
Stage b) in which said hydroconverted effluent is subjected at least in part, and preferably completely, to one or more separations.
The object of this stage is to separate the gases from the liquid, and, in particular, to recover the hydrogen and the bulk of hydrogen sulfide H2S that is formed in stage a), then to obtain a liquid effluent that is free of dissolved H2S.
During the separation of H2S from the liquid, a portion of naphtha can be separated. This portion is then stabilized (H2S is removed).
The liquid effluent that is depleted in H2S and optionally treated with stabilized naphtha is distilled to obtain at least one distillate fraction that includes a gas oil fraction and at least one fraction that is heavier than the gas oil.
The distillate fraction can be a gas oil fraction or a gas oil fraction that is mixed with naphtha. It feeds stage c).
The liquid fraction that is heavier than the gas oil type fraction optionally can be sent into a catalytic cracking process in which it is advantageously treated under conditions that make it possible to produce a gas fraction, a gasoline fraction, a gas oil fraction and a fraction that is heavier than the gas oil fraction and is often called slurry fraction by one skilled in the art.
In other cases, this liquid fraction that is heavier than the gas oil fraction can be used as an industrial fuel with a low sulfur content or as a thermal cracking feedstock.
When the naphtha is not sent into the mixture with the gas oil in stage c), it is distilled. The naphtha fraction that is obtained can advantageously be separated into heavy gasoline, which preferably will be a feedstock for a reforming process, and into light gasoline, which preferably will be subjected to a process for isomerization of paraffins.
At the output of stage b), the gas oil fraction most often has a sulfur content of between 100 and 500 ppm by weight, and the gasoline fraction most often has a sulfur content of at most 200 ppm by weight. The gas oil fraction therefore does not meet 2005 sulfur specifications. The other characteristics of the gas oil are also at a low level; for example, the cetane is on the order of 45 and the aromatic content is greater than 20% by weight.
In distillation, the conditions are generally selected such that the initial boiling point of the heavy fraction is from about 340° C. to about 400° C. and preferably from about 350° C. to about 380° C., and, for example, about 360° C.
For naphtha, the final boiling point is between about 120° C. and 180° C.
The gas oil boiling point range is between naphtha and the heavy fraction.
The boiling point ranges that are provided here are given only by way of examples but the user will select the boiling point range based on the quality and the quantity of the desired products, as is generally done.
Stage c) in which at least a portion, and preferably all, of the distillate fraction undergoes hydrotreatment so as to reduce the sulfur content below 50 ppm by weight, and most often below 10 ppm.
With said distillate fraction, it is possible to treat a fraction that is produced external to the process according to the invention, which normally cannot be incorporated directly into the gas oil pool. This hydrocarbon fraction can be selected from, for example, the group that is formed by the LCO (light cycle oils that are obtained from fluidized-bed catalytic cracking).
The operation is usually carried out under an absolute pressure of about 2 to 12 MPa, often from about 2 to 10 MPa and most often from about 4 to 9 MPa; it is also possible to work under 3 to 7 MPa. The temperature in this stage is usually from about 300 to about 500° C., often from about 300° C. to about 450° C. and very often from about 350 to about 420° C. This temperature is usually adjusted based on the desired level of hydrodesulfurization and/or saturation of aromatic compounds and should be compatible with the desired cycle duration. The hourly space velocity (VVH) and the partial hydrogen pressure are selected based on the characteristics of the product that is to be treated and the desired conversion.
The VVH most often lies in a range that goes from about 0.1 h−1 to about 10 h−1 and preferably 0.1 h−1-5 h−1 and advantageously from about 0.2 h−1 to about 2 h−1.
The total amount of hydrogen mixed with the feedstock is usually from about 200 to about 5000 normal cubic meters (Nm3) per cubic meter (m3) of liquid feedstock and most often from about 250 to 2000 Nm3/m3 and preferably from about 300 to 1500 Nm3/m3.
The operation is also usefully carried out with a reduced partial pressure of hydrogen sulfide that is compatible with the stability of the sulfide catalysts. In the preferred case of this invention, the partial pressure of the hydrogen sulfide is preferably less than 0.05 MPa, preferably 0.03 MPa, or better yet 0.01 MPa.
In the hydrodesulfurization zone, the ideal catalyst should have a strong hydrogenating power so as to carry out a deep refining of the products and to obtain a significant reduction of sulfur. In the preferred embodiment, the hydrotreatment zone operates at a relatively low temperature which tends to produce an intense hydrogenation therefore with a lowered aromatic compound content of the product, an improvement in its cetane and less coking. The scope of this invention would not be exceeded by using a single catalyst or several different catalysts in the hydrotreatment zone in a simultaneous manner or in a successive manner. Usually, this stage is carried out industrially in one or more reactors with one or more catalytic beds and with liquid downflow.
In the hydrotreatment zone, at least one fixed bed of hydrotreatment catalyst that comprises a hydro-dehydrogenating function and an amorphous support is used. A catalyst whose support is selected from, for example, the group that is formed by alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals will preferably be used. This support can also contain other compounds and, for example, oxides that are selected from the group that is formed by boron oxide, zirconia, titanium oxide, and phosphoric anhydride. Most often an alumina support, and, better, η-alumina or γ-alumina is used.
The hydrogenating function is ensured by at least one metal of group VIII and/or group VIB.
In an advantageous case, the total content of metal oxides of groups VI and VIII is often from about 5 to about 40% by weight and in general from about 7 to 30% by weight, and the ratio by weight that is expressed by metal oxide between metal (metals) of group VI to metal (or metals) of group VIII is in general from about 20 to about 1 and most often from about 10 to about 2.
The ideal catalyst is to have a strong hydrogenating power so as to carry out a deep refining of the products and to obtain a significant reduction of sulfur. This catalyst can be a catalyst that comprises metals of group VIII, for example nickel and/or cobalt most often combined with at least one metal of group VIB, for example molybdenum and/or tungsten. A catalyst with an NiMo base will preferably be used. For the gas oils that are difficult to hydrotreat and for very high hydrodesulfurization rates, it is known to one skilled in the art that desulfurization with a catalyst having an NiMo base is higher than that of a CoMo catalyst because the first shows a more significant hydrogenating function than the second. It is possible, for example, to use a catalyst that comprises 0.5 to 10% by weight of nickel and preferably 1 to 5% by weight of nickel (expressed in terms of nickel oxide NiO) and 1 to 30% by weight of molybdenum and preferably 5 to 20% by weight of molybdenum (expressed in terms of molybdenum oxide (MoO3) on an amorphous mineral support.
The catalyst can also contain an element such as phosphorus and/or boron. This element may have been introduced into the matrix or have been deposited on the support. It is also possible to deposit the silicon on the support, alone or with phosphorus and/or boron.
The concentration of said element is usually less than about 20% by weight (theoretical oxide) and most often less than about 10% by weight, and it is usually at least 0.001% by weight. The concentration of boron trioxide B2O3 is usually from about 0 to about 10% by weight.
Preferred catalysts contain silicon that is deposited on a support (such as alumina), optionally with P and/or B also deposited and also containing at least one metal of GVIII (Ni, Co) and at least one metal of GVIB (W, Mo).
In the process according to the invention, the gasolines and the gas oils that are obtained from the conversion process, such as, for example, hydroconversion, are very refractory in the hydrotreatment if they are compared to gas oils that are obtained directly from the atmospheric distillation of crude oils. To obtain very low sulfur contents, the critical point is the conversion of the most refractory radicals, particularly the dibenzothiophenes that are di- and trialkylated or more for which the access of the sulfur atom to the catalyst is limited by the alkyl groups. For this family of compounds, the path of the hydrogenation of an aromatic ring before desulfurization by rupture of the Csp3-S bond is faster than the direct desulfurization by rupture of the Csp2-S bond.
The conversion gas oils therefore require very strict operating conditions to reach future sulfur specifications. If it is desired to hydrotreat these conversion gas oils under operating conditions that make it possible to maintain moderate investment with a reasonable cycle length of the hydrotreatment catalyst, an optimization of the integration of the equipment of the process is necessary.
We discovered that it is possible to obtain good quality gas oils while reducing the investment costs by maximizing the partial pressure of hydrogen. To do this, according to this invention, make-up hydrogen is introduced into hydrotreatment stage c).
The amount of make-up hydrogen introduced in this stage c) is preferably larger than the chemical consumption of hydrogen that is necessary to obtain fixed performance levels under operating conditions that are fixed for this stage c).
This means that this amount is greater than necessary for the desired hydrogenation level of the compounds that can be hydrogenated.
If a hydrogen material balance is carried out between the input corresponding to the hydrocarbon feedstock and the output corresponding to the liquid and gaseous effluents beyond separated hydrogen, the amount of make-up hydrogen is at least equal to the difference of the material balance; the difference that is found corresponds approximately to the chemical consumption of hydrogen.
A suitable means for measuring the hydrogen content in the feedstock or the liquid effluent is the RMN-1H measurement. For the gaseous effluent, chromatographic analysis is suitable.
In a preferred embodiment, all of the make-up hydrogen that is necessary to the process is introduced into stage c).
Accordingly, the amount that is provided will also take into account the chemical consumption of hydrogen in stage a) so as to provide the hydrogen that is necessary for the hydrogenation that is also desired in stage a).
Thus, in the process, the make-up hydrogen can therefore be introduced:
Another consequence is that it is possible to optimize the addition of hydrogen in stage c) according to the refractory level of the gas oils to be treated.
This advantageous arrangement of the invention thus makes it possible to improve considerably the performance levels of the hydrotreatment catalyst and in particular the hydrodesulfurization for conditions of temperature and total pressure that are provided and that correspond to values that can be practiced industrially.
Actually, it makes it possible to maximize the partial hydrogen pressure, and therefore the performance level, in stage c), while maintaining an almost identical total pressure of stages a) and c) (and therefore their investment cost).
It is thus possible to reduce the residual sulfur content of the gas oil on the order of 30% relative to a process where all of the make-up hydrogen would be brought to stage a) or else the make-up hydrogen brought to stage c) would be just equal to the chemical consumption of hydrogen in stage c).
For feedstocks treated in stage a) that have a large amount of sulfur (for example that have at least 1% by weight of sulfur or at least 2%) and that produce refractory and sulfur-containing conversion gas oils, it has thus become possible to obtain good quality middle distillates in particular with a low sulfur content under conditions in particular of relatively low pressure and thus to limit the cost of necessary investments.
Stage d) of final separation on at least a portion, and preferably all, of the hydrotreated effluent that is obtained in stage c).
Hydrogen is separated from the effluent. It contains small amounts of hydrogen sulfide and usually does not require treatment.
The hydrogen sulfide is also separated from the liquid effluent and thus a gas oil is obtained with at most 50 ppm by weight of sulfur, and most often with less than 10 ppm by weight of sulfur. Naphtha is also obtained in general.
Treatment and Recycling of Hydrogen
The gas that contains hydrogen that was separated in stage b) is, if necessary, treated at least in part to reduce its H2S content (preferably by scrubbing with at least one amine) before recycling it in stage a) and optionally in stage c).
The recycle gas preferably contains an amount of H2S that is higher than 0 mol % and up to 1 mol %. Advantageously, this amount is at least 15 ppm, preferably at least 0.1%, and even at least 0.2 mol %.
Thus, for example, at least a portion of the gaseous fraction can be sent into an amine scrubbing section where H2S is completely removed; the other portion can bypass the amine scrubbing section and be sent directly to recycling after compression.
The presence of H2S is useful for keeping the catalysts in the sulfurated state in stages a) and c), but excess H2S could reduce the hydrodesulfurization.
The hydrogen that is separated in stage d) is added to the optionally purified hydrogen that is obtained from stage b). The mixture is re-compressed and then recycled to stage a) and optionally to stage c).
Actually, in the case where the make-up hydrogen is introduced into stage c), the recycling to stage c) may not be necessary, in particular when all of the make-up hydrogen is introduced in stage c).
It is advantageously possible to introduce the recycling hydrogen with the feedstock that enters stage a) and/or in quench form between the catalyst beds.
The gas oil that is obtained has a sulfur content of less than 50 ppm by weight, generally less than 20 ppm, and most often less than 10 pm.
Furthermore, the cetane is improved by 1 to 12 points, generally from 1 to 7, or else 1 to 5 points relative to the gas oil that goes into hydrotreatment.
Its total amount of aromatic compounds is also reduced by at least 10%, and the reduction can go even up to 90%.
The amount of polyaromatic compounds in the final gas oil is at most 11% by weight.
Installation
The invention also relates to an installation for treatment of petroleum feedstocks of which at least 80% by weight boils above 340° C. and which contains at least 0.05% of sulfur comprising:
To facilitate a better understanding of the installation, as well as the process,
The feedstock that is to be treated (as defined above) enters via a pipe (1) into a zone (I) for ebulliated-bed hydroconversion of a hydroconversion catalyst. The effluent that is obtained in pipe (2) is sent into separation zone (II).
Zone (I) also comprises at least one pipe (31) for drawing off catalyst and at least one pipe (32) for providing fresh catalyst.
The effluent first passes into a separator (3) that separates, on the one hand, a gas that contains hydrogen (gaseous phase) into pipe (4) and, on the other hand, a liquid effluent into pipe (5). It is possible to use a hot separator that is followed by a cold separator (preferred) or a cold separator only.
A portion of the liquid effluent that is obtained advantageously can be extracted to be recycled via pipe (33) at the bottom of the boiling bed of stage a) to keep the catalyst in a boiling bed.
The liquid effluent is sent into a separator (6), which is preferably a vapor stripper, to separate the hydrogen sulfide from the hydrocarbon effluent. In the same step, at least a portion of the naphtha fraction can be separated with the hydrogen sulfide. The hydrogen sulfide with said naphtha exits via pipe (7) while the hydrocarbon effluent is obtained in pipe (8).
The hydrocarbon effluent then passes into a distillation column (9), and at least one distillate fraction that includes a gas oil fraction and withdrawn via pipe (11), is separated, and a heavier fraction than the gas oil is withdrawn in pipe (10).
In general, the naphtha that is separated at separator (6) is stabilized (H2S is eliminated). In an advantageous arrangement, the stabilized naphtha is injected into the effluent that enters column (9).
At column (9), a naphtha fraction can be separated and withdrawn via into an additional pipe not shown in
According to
The naphtha that is obtained separately, optionally treated with naphtha that is separated in zone (IV), is advantageously separated into heavy and light gasolines, whereby the heavy gasoline is sent into a reforming zone, and the light gasoline is sent into a zone where the isomerization of paraffins is carried out.
In
The distillate fraction is then sent (alone or optionally treated with a naphtha fraction and/or gas oil fraction that is external to the process) into a hydrotreatment zone (III) that is provided with at least one fixed bed of hydrotreatment catalyst.
The hydrotreated effluent that is obtained exits via pipe (12) to be sent into separation zone (IV) that is circumscribed by dotted lines in
The liquid phase is sent into a separator (16), preferably a stripper, to remove the hydrogen sulfide that exits into pipe (17), most often mixed with naphtha. A gas oil fraction is drawn off via pipe (18); a fraction that is in compliance with sulfur specifications, i.e., having less than 50 ppm by weight of sulfur, generally less than 10 ppm. The H2S-naphtha mixture is then optionally treated to recover the purified naphtha fraction.
The process and the installation according to the invention also advantageously comprise a hydrogen recycling loop for two zones (I) and (II). Thus, the gas that contains hydrogen (gaseous phase in pipe (4) separated in zone (II)) is treated to reduce its sulfur content and optionally to eliminate the hydrocarbon compounds that have been able to pass during the separation.
Advantageously and according to
A portion of the hydrocarbon fraction of pipe (21) is sent into separation zone (II) and advantageously into pipe (5).
A particular embodiment for separating the entrained hydrocarbon compounds will now be described; however, any other method that is known to one skilled in the art is suitable.
The gaseous phase that is obtained in pipe (24) from which hydrocarbon compounds have been removed is, if necessary, sent into a treatment unit (25) for reducing the sulfur content.
Advantageously, this is a treatment with at least one amine.
In some cases, it is sufficient that only a portion of the gaseous phase be treated. In other cases, the entire gaseous phase should be treated, which is what is illustrated in
The gas that contains hydrogen that is thus optionally purified is then re-compressed in compressor (27).
The hydrogen that is separated in pipe (14) is preferably added before compression.
The compressed mixture is then recycled in part to hydrotreatment zone (III) (stage c) and in part to hydroconversion zone (I) (stage a) by pipes (28) and (29) respectively.
In the preferred embodiment of
In another embodiment, it is possible to provide a pipe that brings some make-up hydrogen into zone (I). Thus, an advantageous embodiment comprises, for some make-up hydrogen, a pipe at zone (I) and a pipe at zone (II).
As
The invention thus described offers numerous advantages. In addition to those already described, in the preferred embodiment where the pressures are identical for stages a) and c) because of the unique gas recirculation system, it is possible to use only a single recycling compressor for the two reaction zones, thus also reducing the investment costs. Likewise, when the invention operates at moderate pressures, the operating expenses are reduced. Furthermore, a very good quality feedstock for catalytic cracking (low contents of sulfur and nitrogen, moderate enrichment of hydrogen) is produced.
The preceding specific embodiments are to be construed as merely illustrative, and not limitative of the remainder of the disclosure in any way whatsoever.
The entire disclosure of all applications, patents and publications, cited above and below, and of corresponding French Application No. 01/14.594, are hereby incorporated by reference.
From the foregoing description, one skilled in the art can easily ascertain the essential characteristics of this invention, and without departing from the spirit and scope thereof, can make various changes and modifications of the invention to adapt it to various usages and conditions.
| Number | Date | Country | Kind |
|---|---|---|---|
| 01/14.594 | Nov 2001 | FR | national |
This application claims the priority of Provisional Application Ser. No. 60/357,629 filed Feb. 20, 2002 and is related to Applicants' concurrently filed application Attorney Docket No. PET-1992V1 entitled, “Process For Converting Heavy Petroleum Fractions For Producing A Catalytic Cracking Feedstock And Middle Distillates With A Low Sulfur Content”, based on French Application No. 01/14.531, filed Nov. 9, 2001.
| Number | Date | Country | |
|---|---|---|---|
| 60357629 | Feb 2002 | US |
| Number | Date | Country | |
|---|---|---|---|
| Parent | 10152988 | May 2002 | US |
| Child | 11505285 | Aug 2006 | US |
