DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

[0025] Many refineries face an abundant supply of lower alkanes, i.e., C ₁ -C ₄ alkanes such as methane, and relatively few means of converting them to more valuable products. Much research has been devoted to investigating the conversion of methane to more easily transportable products. One technique that has been developed entails the partial oxidation of light hydrocarbons in the presence of a catalyst. This technique results in the production of synthesis gas, i.e., “syngas”, a mixture of CO and H ₂ . The catalytic partial oxidation of methane can be represented by the following reaction scheme:

CH ₄ +1/20 ₂ CO+2H ₂ . (4)

[0026] Such catalytic oxidation reactions are exothermic and require good composition control in order to avoid over-oxidation resulting in too high a reaction temperature.

[0027] Several schemes for carrying out such partial oxidation are known in the art. One scheme for carrying out the exothermic oxidation reaction entails a brief exposure of the methane feed to a hot catalyst followed by cooling the resultant gas stream. A catalyst is positioned in the flow path of the feed gas. The catalyst comprises a wire gauze, several layers of wire gauze, or a porous ceramic impregnated with a catalyst.

[0028] A new system according to the present invention for carrying out catalytic partial oxidation of methane or other light hydrocarbons replaces the burner of a Claus process. In addition to H ₂ S, the feed stream includes methane (or a similar light hydrocarbon) and air, oxygen, or a mixture of both. Thus, while sulfur is produced according to Equation (3) above, additional hydrogen is generated, which allows the hydrogen originally consumed in the desulfurization process to be recovered.

[0029] Referring initially to FIG. 1, a preferred embodiment of the present system includes a Claus reactor 10 that includes feed injection openings 12 , 14 , and 16 , a mixing zone 19 , a reaction zone 20 and a cooling zone 30 . Reaction zone 20 preferably includes a thermal radiation barrier 22 positioned immediately upstream of a catalytic device 24 . Radiation barrier 22 is preferably a porous ceramic or refractory material that is suited to withstand operating temperatures and provide sufficient thermal insulation, such as are described in U.S. Pat. 4,038,036 (Beavon) which is incorporated herein by reference in its entirety.

[0030] Catalytic device 24 is preferably a layer or layers of wire gauze 25 or a porous ceramic monolith (not shown) having a suitable catalyst supported on its surface. Gauze 25 is preferably one or more layers of a substantially planar, flexible woven metal-containing or metal-coated screen or gauze having about 20-120 mesh. More preferably, it is a gauze of metal wires about 25 micrometers to about 2.5 millimeters in diameter, which are made of about 87-93% by weight (wt-%) Pt and about 7-13 wt-% Rh. Alternative catalyst structures could include a disk with multiple perforations formed therethrough, a honeycomb-like structure, an etched foil and any other structure that provides the desired amount of transparency to effect the desired partial oxidation. A detailed discussion of the catalyst structure and composition can be found in U.S. Pat. No. 5,654,491 to Goetsch et al., which is incorporated herein in its entirety.

[0031] Examples of suitable catalysts that can be included in the metal of the gauze or incorporated at its surface include, but are not limited to, platinum, rhodium, nickel, palladium, iridium, Pt/ZrO ₂ , Pt/Al ₂ O _3.

[0032] In operation, H ₂ S is fed into one of the feed injection openings 12 . A light hydrocarbon, such as methane, is fed into a second feed injection opening 14 . Air or oxygen is fed into the third feed injection opening 16 . It will be understood that the feed injection openings can be configured differently from the configuration shown without affecting the principles or operation of the present system.

[0033] As the feed gases from feed injection openings 12 , 14 , 16 flow toward catalytic device 24 , they are preferably subjected to thorough mixing by static mixer 18 . During mixing, they are shielded by radiation barrier 22 from radiant heat that is generated downstream in the process. It is preferred that the temperature on the upstream side of barrier 22 be in the range of about 20° C. to about 300° C. The feed gas stream is preferably at ambient temperature prior to contact with the catalyst. Preheating the feed gas stream is not desired, as it can cause homogeneous reactions and reduce the selectivity of the process of the present invention for the desired compounds. Therefore, preheating the feed gas mixture is typically avoided, although in some applications feed gas temperatures up to about 300° C. can be tolerated.

[0034] After the gases pass barrier 22 , they flow past catalytic device 24 and are simultaneously heated to an oxidation temperature in the range of from about 900° C. to about 1500° C. The gas flow rate is preferably maintained such that the contact time for the portion of the gas that contacts the catalyst is between about 0.00001 to 0.01 seconds and more preferably between about 0.001 to 0.005 seconds.

[0035] This degree of contact produces a favorable balance between competing reactions and produces sufficient heat to maintain the catalyst at approximately 900-1500° C. Specifically, sulfur is produced by catalyzed partial oxidation according to the equation:

H ₂ S+1/2O ₂ 1/xS _x +H ₂ O (5)

[0036] where x equals 2, 6, or 8, with x =2 being the most likely. At the same time, exposure to the hot catalyst partially oxidizes the hydrocarbons in the feed, according to the equation:

CH ₄ +1/2O ₂ CO+2H ₂ . (6)

[0037] Oxygen for these reactions comes from the air, oxygen, or air/oxygen mix that is fed into the system with the H ₂ S and hydrocarbon feed gases.

[0038] Typically, the catalyst structure is heated as a result of the exothermic chemical reactions occurring at its surface; however, it can additionally or alternatively be heated by external means, such as electrical resistance, magnetic induction, RF, etc. Heating by external means can allow for increases in the rate at which feed gas can be passed through the catalyst structure while still obtaining desirable reaction products. In many cases it is helpful to heat the catalytic device 24 with external means at least at the start of the process, so as to initiate the exothermic reactions on the catalyst structure. This initial heating can be accomplished in any suitable manner including electrical resistance, magnetic induction, RF, or the like. Once the system is running, it is preferably run adiabatically or nearly adiabatically (i.e., without the loss of heat aside from convective losses in the exiting gas), so as to reduce the formation of solid carbon (e.g., coke) on the surface of the gauze catalyst.

[0039] The rapid heating of the feed gases as a result of contact with the hot catalyst promotes fast reaction rates. In accordance with the present invention, the feed gas stream velocity past catalyst structure 24 is preferably at least about 0.1 meter/second, often as high as 4-5 meters/second, and even as high as 70 meters/second. The maximum velocity will generally determined by the specific equipment used; however, the theoretical limit is that velocity at which the reaction would be extinguished. If an external means of heating the catalytic device 24 is used, this theoretical limit is significantly large.

[0040] According to one preferred embodiment, the feed gas stream velocity is between about 0.1 and 100 meters/second. As a result, the superficial contact time of the feed gas stream with a preferred embodiment of gauze catalytic device 24 is less than about 10,000 microseconds, and typically within a range of about 1,000-5,000 microseconds. When used in the present invention, it is preferred that the superficial contact time of the feed gas stream with the catalyst be less than about 5,000 microseconds, more preferably less than about 2,000 microseconds. As used herein, “superficial contact time” is calculated as the wire diameter divided by the feed gas stream velocity at inlet conditions (i.e., temperature and pressure at the inlet to the reactor). Superficial contact time is inversely proportional to the term “space velocity” that is used in many chemical process descriptions.

[0041] Although for ease in comparison with prior art, space velocities at standard conditions have been used to describe the present invention, it is well recognized in the art that residence i time is the inverse of space velocity and that the disclosure of high space velocities equates to low residence times.

[0042] From reaction zone 20 , the reacted gases enter a firetube boiler 40 , where they are cooled to below 425° C. and preferably to below 3400° C. As shown, it is preferred that heat removed from the partially oxidized gases can be recaptured by boiling water to make steam or the like. The rapid cooling that occurs in the boiler drops the temperature to below about 425° C. and thus ceases the above reactions. A detailed description of the considerations involved in operating a reactor using extremely small contact times is given in U.S. Pat. No. 5,654,491, which is incorporated herein by reference in its entirety.

[0043] Referring now to FIG. 2 , the present system preferably includes the reactor 10 , firetube boiler 40 , a condenser 50 , heater 55 , one or more tailgas converter units 60 , a quench tower 80 , an amine absorber/contactor 90 , and an amine regenerator 100 . The cooled, partially oxidized gases flow from boiler 40 into condenser 50 , where they are cooled further until the dew point of the elemental sulfur is reached. This allows for the removal of elemental sulfur, as desired, from the process. Once almost all of the elemental sulfur is removed, the partially oxidized gases are reheated in heat exchanger 55 and passed through one or more tailgas converter units 60 . Each tailgas converter unit 60 includes at least a catalyst bed 56 in contact with the fluid and a quench device 58 . More specifically, in each converter unit 60 , the hot gas stream is passed over a bed of conventional cobalt-molybdenum based Claus tail gas treating unit hydrogenation catalyst. In this catalyst bed, any elemental sulfur is converted to H ₂ S. The CO in the hot gas reacts with water generated in the short contact time reactor (equation (5)) to form CO ₂ and hydrogen according to the following equations:

H ₂ +1/xS _x H ₂ S (7)

CO+H ₂ O CO ₂ +H ₂ (8)

[0044] If any additional water vapor is required for the water gas shift (Equation (8)), it can be added after the sulfur condensation stage. It is desirable to carry out the water gas shift reaction, as CO will require incineration to CO ₂ before it can be emitted from the stack. Since the water gas shift reactor forms the CO ₂ , anyway, it is much more valuable to generate hydrogen from the CO that to simply incinerate it to CO ₂ . The effluent from the water gas shift reactor(s) is then preferably cooled sufficiently to condense the bulk of any remaining water from the gas stream and to adjust the temperature of the gas to the proper level for alkanolamine treating.

[0045] Following the final quenching by counter-current flow through quench tower 80 , the partially oxidized gases, including any hydrogen gas, are fed into an alkanolamine absorber 90 , where H ₂ S is removed. In absorber 90 , an alkanolamine absorber, preferably based on methyl diethanolamine or diisopropanolamine, is used to remove any H ₂ S that may be present in the product gas from the water condensation stage. The treated gases, which comprise hydrogen, nitrogen, and some CO ₂ , with trace amounts of H ₂ S, are then compressed and purified using Pressure Swing Absorption (PSA), membranes, or cryogenic separation. From this process, purified hydrogen is made available for use in the hydrogen consuming processes. The waste gas from the purification process is preferably sent to the refinery fuel system. Hence, there is no direct stack emission from the sulfur recovery unit. H ₂ S and CO ₂ removed from the hydrogen-rich product gas in the alkanolamine absorber go to the alkanolamine regenerator 100 , where they are boiled out of the alkanolamine solution and recycled to the front of the sulfur recovery unit.

[0046] While a preferred embodiment of the present invention has been shown and described, it will be understood that variations can be to the preferred embodiment, without departing from the scope of the present invention. For example, the mixing process can be altered or replaced with an active mixer, the thermal barrier can be modified, the structure and composition of the catalyst can be varied, and the tail gas treatment steps can be modified.

[0047] The complete disclosure of all patents, patent documents, and publications cited herein are incorporated by reference. The foregoing detailed description and examples have been given for clarity of understanding only. No unnecessary limitations are to be understood therefrom. The invention is not limited to the exact details shown and described, for variations obvious to one skilled in the art will be included within the invention defined by the claims.