EA040662B1 - METHOD FOR PRODUCING A CATALYST FOR CATALYTIC OXIDATION CRACKING OF A GAS FLOW CONTAINING H2S - Google Patents

Description

The field of technology to which the invention belongs

The invention relates to a process for preparing a catalyst suitable for use in catalytic oxidative cracking of a gas stream containing H2S.

State of the art

Sulfur recovery plants are designed to remove H2S from acid gases containing H2S, from amine recovery systems and from acid water strippers, producing sulfur, which is a non-toxic product that can be stored and sold in liquid or solid form to various users for many different applications. areas of industrial application. Acid gases from amine recovery systems and sour water strippers containing variable amounts of H2S are treated in a sulfur recovery unit (SRU), generally based on a modified Claus process, for bulk sulfur recovery and subsequently in a tailings treatment section. gases (TGT) for deep sulfur recovery. Other impurities contained in acid gases, including ammonia and hydrocarbons, are destroyed in the Claus section.

The modified Claus process alone recovers approximately 94-96% (2 catalytic stages) or 95-98% (3 stages) of sulfur in the feedstock. Therefore, if a higher sulfur recovery (SRE) is required, additional Claus tail gas treatment will be required.

The modified Claus process involves substoichiometric combustion of the acid gas stream in a thermal reactor (thermal stage) followed by catalytic conversion in Claus reactors (catalytic stage). In the Claus section, one third of the total H2S is oxidized to form SO2, which reacts with the remaining hydrogen sulfide H2S to form sulfur and water according to the following reactions:

H2S + 1.5 O ₂ -> H ₂ O + SO ₂ (oxidation reaction) (1)

H ₂ S + SO2 1.5 S2 + 2 H2O (Claus reaction) (2)

H2S + 1.5 O2 —> 3 H2O + 1.5 S2 (total reaction) (3)

The purpose of this method is to stimulate the passage of the cumulative reaction almost to completion. In the Claus thermal reactor, hydrogen sulfide H2S contained in acid gas is burned using air (or using oxygen-enriched air in some specific cases) in a specific burner, and only one third of the total amount of H2S is oxidized to form SO2, while as the remaining two-thirds do not react. The total amount of air is exactly sufficient to oxidize one third of the total amount of H2S and completely oxidize all hydrocarbons and ammonia contained in the feedstock; therefore, the molar ratio of H2S/O2 in the feedstock is approximately 2:1 in order to achieve an H2S/SO2 ratio in the Claus tail gas of exactly or as close to 2:1 as possible, which is the stoichiometric ratio for the Claus reaction, such thus maximizing the degree of sulfur recovery. During acid gas combustion, a small portion of H2S (typically 5-7%) dissociates to form hydrogen and sulfur according to the following reaction:

H2S <—> H2 + 0.5 S2 (dissociation or cracking reaction) (4)

According to Clark et al., Alberta Sulfur Research Ltd. (ASRL) the formation of hydrogen also occurs according to the following reaction:

H2S + O2 A 2 H2 + 2 H2O + 2 S2 (H2 formation reaction) (5)

Several side reactions also occur, leading to the destruction of ammonia and hydrocarbons and the formation of carbonyl sulfide COS and carbon disulfide CS2. In order to complete the Claus reactions, a suitable residence time in a thermal reactor at high temperature is required.

The Claus thermal reactor is usually followed by a waste heat boiler where the furnace effluent is cooled to approximately 300°C and heat is recovered by generating high pressure steam and using a sulfur condenser where the process gas is cooled to the sulfur condensing temperature by generating low pressure steam. pressure, and liquid sulfur is separated.

The thermal Claus stage is generally followed by two or three catalytic stages, with each stage consisting of a gas reheater to bring the gas to the optimum reaction temperature, a catalytic reactor where the Claus reaction takes place, and a sulfur condenser where the gas is cooled and liquid sulfur condense and separate. The Claus reaction is an exothermic equilibrium reaction thermodynamically enhanced by low temperatures. The first catalytic Claus reactor is partially filled with a Claus catalyst (based on alumina) to improve the Claus reaction and partially filled with a specific high conversion catalyst (based on titanium dioxide) to improve the hydrolysis of COS and CS2. The second and third catalytic Claus reactors, if any, are generally filled with a Claus catalyst (based on alumina) to improve the progress of the Claus reaction.

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In order to meet the requirement to achieve >99% sulfur recovery normally required in a sulfur recovery unit, the Claus section is generally followed by a tail gas treatment section. Over the years, several different alternative methods for increasing sulfur recovery have been proposed, such as the SCOT process, the RAR process, the CBA process, the CLINSULF/DEGSULF process, or the BSR Selectox process. In a traditional tail gas reduction section, process gas from the Claus section is preheated and combined with hydrogen from an external source before being fed to the hydrogenation reactor where all sulfur compounds are converted to H2S on a specific reduction catalyst (based on Co and Mo oxides) that performs the functions both hydrogenation and hydrolysis. The reactor effluent is cooled in a quench tower using circulating steam condensate. Hydrogen sulfide H2S produced in the hydrogenation reactor is recovered in the amine absorber using a specific aqueous amine solution and recycled to the Claus section from the top of the amine regenerator where the enriched solution is desorbed.

The tail gases from the amine absorber are sent to a thermal afterburner to oxidize the residual hydrogen sulfide H ₂ S and other sulfur compounds such as COS and CS ₂ to form SO ₂ before being discharged to the atmosphere through an appropriate stack.

The main disadvantages of the traditional Claus plant are the need for process gas recovery in the TGT section in large and expensive equipment with very low economic value of sulfur, continuous emissions of SO _x (SO2 and SO ₃ ), CO, CO2, NO _x plus trace amounts of H ₂ S into the atmosphere and a continuous supply of hydrogen from the network.

In some applications where hydrogen is not available, such as in gas fields, the reducing gas mixture is generated in a reducing gas generator by substoichiometric combustion of the fuel gas.

The main disadvantage of this alternative configuration is the larger size of the equipment compared to the traditional Claus installation. This is due to a large 1015% consumption of process gas due to significant amounts of inert compounds coming from the fuel gas introduced into the combustion process line (mainly nitrogen from air and water and carbon dioxide from combustion).

As mentioned in Clark, Catalysis Communications, 5 (2004), 743-747, the recovery of H ₂ from H ₂ S is a long standing goal in the industry. Clark refers to it when using the partial oxidation of H ₂ S on catalysts based on alumina. The key to this process is said to be to promote the reaction between H2S and O2 to form hydrogen, water and sulfur at a controlled temperature using an external furnace. Remediation of air emissions is not addressed.

Over the years, several alternative processes have been proposed that address thermal or catalytic partial oxidation of H2S.

U.S. Patent Nos. 6,946,111 and 6,800,269 to Conoco Inc. describes methods for removing H2S from a gas stream containing H ₂ S - in the first patent - and from an exhaust gas stream enriched in H ₂ S in the second patent, including a flameless short-term contact reactor filled with a suitable catalyst for the partial oxidation reaction of H ₂ S with the formation of sulfur and water using air or enriched air or pure oxygen at a ratio of H ₂ S/O ₂ in the feedstock of approximately 2:1, followed by a cooling zone and a sulfur condenser. The main purpose of the first patent is to desulphurize the gas stream, while the main purpose of the second patent is to offer an alternative solution for the traditional Claus thermal reactor. Both patents are based on the catalytic partial oxidation of hydrogen sulfide using oxygen to produce sulfur and water.

U.S. Patent No. 7,560,088 to the Conoco Phillips Company describes a process for removing sulfur from a gas stream containing H2S using a compact system comprising a catalytic partial oxidation reaction zone in flameless brief contact with the following temperature control zone, first Claus catalytic reaction zone, second zone temperature control, a first outlet for liquid sulfur and a first outlet for an exhaust gaseous stream. The main purpose of this patent is to propose an alternative solution to the traditional Claus plant based on the catalytic partial oxidation of hydrogen sulfide to produce sulfur and water.

U.S. Patent No. 4,481,181 to GA Technologies Inc. describes a process for removing sulfur and recovering hydrogen from a gas stream containing H2S, combining the thermal partial oxidation of H2S to form sulfur and water and the thermal dissociation of H2S to form hydrogen and sulfur in the same reaction zone with the preceding feedstock heating section and the following zone cooling and sulfur condenser using pure oxygen and a significant proportion of nitrogen at a ratio of H2S/O2 in the feedstock in the range from 10:1 to 25:1. The main purpose of this patent is the thermal decomposition resulting from the incomplete oxidation and dissociation of hydrogen sulfide to form sulfur and hydrogen.

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Chevron USA Inc. WO 2010/036941. and Drexel University describes a method for conducting thermal dissociation of H ₂ S at a temperature less than 1600°C, based on H and SH radicals, in one embodiment, on a suitable plasma catalyst.

In addition, Siirtec-Nigi Italian Patent 1203898 describes a process referred to as HCR based on the operation of a conventional Claus thermal reactor with a slightly increased H ₂ S/O ₂ ratio in the feedstock in order to maintain the H ₂ S/SO ₂ ratio in the tailings. gases of the Claus method, significantly greater than 2:1. The main purpose of this method is to increase the production of hydrogen in the thermal reactor and avoid the import of external hydrogen into the TGT section. In addition, when such a method is used, emissions from the sulfur recovery unit are not avoided.

From the foregoing discussion, it is clear that some effort has been made in the past in an attempt to offer a workable alternative to the traditional Klaus setup. In particular, some processes that have been proposed over the years are based on thermal or catalytic partial oxidation of H2S, while some others focus on thermal or catalytic cracking of H2S. None of the proposed methods is conceived or designed to carry out the conversion of H ₂ S with the formation of hydrogen and sulfur on a suitable catalyst capable of favoring the passage of both reactions at the same time.

Applicants' co-pending application PCT/NL2012/050308 describes a process for producing hydrogen from a gas stream containing H2S, in which the gas stream is catalytically oxidatively cracked to produce H2 and S2. The invention described in this document is used to address the problem of gas emissions into the atmosphere and at the same time obtain a valuable exportable stream of hydrogen.

One problem with gas streams containing H2S that are typically fed to sulfur recovery units is the co-presence of methane. Methane is susceptible to conversion to sulfur compounds such as CS2 or COS, which is undesirable. Another problem with gas streams containing H2S, which are usually fed to sulfur recovery units, is the co-presence of ammonia. Ammonia is usually converted in the thermal stage of the Claus plant. However, thermal conversion of ammonia involves the risk of formation of solid salts such as ammonium sulfides or sulfites arising from incomplete combustion. These salts cause blocking in the coldest sections of the Claus plant, in particular in the sulfur condensers. Proper combustion of ammonia requires a homogeneous mixture of ammonia and air and a high flame temperature. However, the formation of nitrogen oxides promotes the oxidation of sulfur dioxide SO2 to form sulfur trioxide SO3. Thereafter, the Claus catalyst becomes sulfided and corrosion is observed for the cold parts of the plant.

It would be desirable to provide a catalyst that is active and selective in the oxidative cracking of H2S, in particular in the case of the simultaneous presence of NH3 and/or CH ₄ , and more generally of carbonaceous compounds.

Disclosure of invention

In order to better address one or more of the aforementioned desires, the invention relates in one aspect to a process for preparing a catalyst suitable for use in the catalytic oxidative cracking of an H2S containing gas stream that contains iron and molybdenum supported on an aluminum containing support, in which an aqueous iron and molybdenum precursor solution selected from the group consisting of iron tetracarbonyl, iron pentacarbonyl, iron nonacarbonyl, iron nitrates, iron bromides, iron chlorides, iron fluorides, iron phosphates, iron sulfates, iron acetylacetonates, iron acetates, iron fumarates, gluconates iron, iron citrates, iron benzoates, iron maleates, iron oxalates, iron oleates and iron stearates, iron-ammonium complexes and similar molybdenum precursors, and said solution is dispersed on the carrier, and before the solution is dispersed on the carrier, the carrier is subjected to calcination at a temperature erature from 750 to 1000°C, while the stage of calcination is carried out in the presence of air.

Implementation of the invention

The invention in its broadest sense is based on identifying a reasonable combination of catalytically active metals and using them in the form of a supported type catalyst. The metals chosen are iron and molybdenum and the support is a carrier containing aluminum.

Supported catalyst will be understood to refer to a catalyst composition comprising a catalytically active moiety (i.e., particles in the established order that are either active or converted into an active phase in situ) and a catalytically inactive moiety, while being catalytically inactive part (substrate) generally forms the majority of the catalyst. This distinguishes a supported catalyst from a bulk catalyst in which the catalytically inactive portion is generally a smaller portion. Thus, in a supported catalyst, the catalytically inactive part in the general case is more than

50% (wt.) of the composition of the catalyst. Preferably the support forms more than 60% (wt.), and more preferably more than 80% (wt.) of the total composition of the catalyst.

The catalytically active part of the catalyst composition contains iron and molybdenum. These metals are usually present in the form of particles dispersed on the carrier in an amount of at least 1% (wt.) and in General up to 50% (wt.) based on the weight of the catalyst composition. The substrate contains aluminum and will generally be aluminum oxide (Al ₂ O ₃ ). The alumina may be, for example, alpha or theta or gamma alumina.

The iron in the catalyst composition is more preferably present in an amount ranging from 1 to 50% (w/w), most preferably 2% (w/w), of the catalyst composition. Molybdenum is present in the catalyst composition in an amount ranging from 1 to 50% (wt.), most preferably 6% (wt.), of the catalyst composition.

The catalyst may consist essentially of said catalyst composition, ie. an aluminum-containing carrier and iron and molybdenum contained thereon. In such a case, the catalyst will usually be in a suitable shaped form, for example in the form of a powder or granules. The catalyst, in addition to the catalyst composition containing the carrier and active metals, may also contain the mechanical structure of the support, i.e. substrate.

It is to be understood that such a substrate is not part of the above catalyst composition, but is present in addition to it. The substrate may be any structure known in the art as a substrate for catalysts. In one embodiment of the present invention, the substrate may be in the form of beads, beads, monolithic honeycomb structures, or open cell foams. The substrate may be formed from alumina, silica-alumina, silica, titanium dioxide, mixtures thereof, or any other suitable material available in the catalyst substrate art.

Where the catalyst contains a substrate, it will typically be coated with a supported alumina-iron-molybdenum oxide catalyst composition.

In one preferred embodiment, the catalytically active metals are in the form of their sulfides. That is, the iron is preferably in the form of iron sulfide, the molybdenum is preferably in the form of molybdenum sulfide, and most preferably both are in the form of sulfide.

The catalyst composition of the invention may be prepared in a manner known to those skilled in the art. Reference is made, for example, to Catalyst Handbook, M.V. Twigg (Ed.), Wolfe Publishing Ltd, 1989 and Structured Catalysts And Reactors, A. Cybulski and J.A. Moulijn (Eds.), Marcel Dekker Inc., 1998 - Chapter 21 (Transformation of a structured carrier into structured catalyst), p. 599-615.

In one particularly well-suited method, an aqueous solution is obtained from the precursor and said solution is dispersed on a carrier material in accordance with the definition in paragraphs. 1 or 4 claims. Examples of iron precursors are inorganic and organic iron salts, iron chelates, iron clusters, iron hydroxides and oxyhydroxides, and iron organometallic complexes. Representatives of these compounds are iron tetracarbonyl, iron pentacarbonyl, iron nonacarbonyl, nitrates, bromides, chlorides, fluorides, phosphates, sulfates, acetylacetonates, acetates, fumarates, gluconates, citrates, benzoates, maleates, oxalates, oleates, iron stearates and the like. Iron precursors may provide iron in ferrous form, ferric form, or a combination thereof. The catalyst precursors preferably contain Fe(II) or Fe(III) in combination with organic ligands or anions such as acetate, citrate, EDTA (ethylenediaminetetraacetate) or NTA (nitrilotriacetate). The same is true for molybdenum with appropriate corrections. Iron acetate and ammonium heptamolybdate are preferred.

In one particularly preferred embodiment, before impregnation with the precursor solution, the carrier is first subjected to calcination. The calcination is preferably carried out at a temperature greater than 700°C, more preferably at least 750°C, most preferably in the range of 750-1000°C. Preferably the calcination is carried out in the presence of oxygen, more preferably in air. After calcination, the catalyst active metal is applied to the calcined support described above. For example, wet impregnation or catalytic metal deposition can be used. Without wishing to be bound by theory, it is believed that calcination as described above makes it possible to stabilize the structure of the catalyst. Thus, the resulting catalyst is particularly well suited for use at high temperatures, such as 1100°C, occurring in catalytic oxidative cracking. In this embodiment, it is important to first carry out the calcination step to obtain a stabilized substrate, and then apply

- 4 040662 catalytic metal. In the case of carrying out the process in the reverse order, it is believed that the structure of the catalyst may change in such a way that the catalytic metal would become unavailable for the catalytic reaction.

The catalyst according to the invention may find use in particular in a process for producing hydrogen from a gas stream containing H2S, in which said gas stream is subjected to catalytic oxidative cracking to produce H2 and S2 using the above catalyst.

It should be emphasized that catalytic oxidative cracking is a fundamentally different process compared to both the thermal stage and the catalytic stage in the existing Claus type process. As can be seen from reference to the above reaction equations (1) to (5), the Claus methods are directed towards causing the above reaction (3) to proceed almost to completion. The above method basically has a reasonable disclosure of the essential causes and relationships to obtain a method based on side reactions (4) and (5) and promote the passage of these reactions to obtain from the gas stream containing H ₂ S, both hydrogen and sulfur.

The above method is also fundamentally different from the recent proposals by Clark et al.. The references they authored are based on the theory of direct oxidation of H2S to form hydrogen, water and sulfur. The resulting conversion, despite avoiding SO2 formation, is to be improved in terms of H2S conversion and sulfur production simultaneously with H ₂ .

In the above process, a catalytic oxidative cracking (COC) stage replaces the Claus thermal stage. The above method thus favors the dissociation and partial oxidation of H2S instead of the complete oxidation and Claus reaction. However, this method does not preclude the addition of a Claus thermal stage after the COC stage.

Catalytic oxidative cracking is carried out in one or more reaction zones, preferably provided in one reaction chamber. Throughout this text, the term chamber may refer to one or more reaction zones. The reaction chamber is defined as the reactor volume including optionally containing a catalyst bed. There is only one type of catalyst in one reaction chamber. Typically, the reaction chamber is essentially cylindrical, and the flow of the reactants is in the axial direction. If the reaction chamber contains a catalyst bed, then one or more reactions can take place in the axial direction of the gas flow. In one embodiment in which more than one reaction takes place, the reaction conversion profile for one reaction may be different from that for another reaction. In other words, one reaction may take place, for example, mainly at the beginning of the catalyst bed, while another reaction may take place, for example, throughout the entire length of the catalyst bed.

The invention provides those skilled in the art with the disclosure of the essential causes and relationships for promoting the aforementioned reactions (4) and (5). The fact that for this the gas stream must be subjected to catalytic oxidative cracking, suggests a clear message to experts in the relevant field of information on how to do this.

In general, the presence of a catalyst will normally be provided on the catalyst bed through which the gas stream to be treated is passed. The choice of types of layers and their volumes is well within the limits of the usual possibilities of specialists in the relevant field of technology.

Oxygen is supplied to the catalytic oxidative cracking reaction zone or zones. The oxygen is preferably supplied as a gas enriched in oxygen relative to air. Preferably it is an oxygen-containing gas stream containing at least 40% (vol.) oxygen, preferably at least 60% (vol.) oxygen. More preferably, this oxygen is supplied as substantially pure oxygen, namely at 90-99% (vol.) oxygen, or as close to 100% as possible.

The use of an oxygen-enriched gas, and preferably pure oxygen, is not only about optimizing the catalytic oxidative cracking process, it also shows advantages such as avoiding unnecessarily large equipment that would be required due to the presence of large volumes of inert gas (nitrogen). In addition, while addressing the object of the invention in the form of hydrogen production in addition to sulfur recovery and at the same time with reduced emissions, it will be beneficial to reduce, and preferably avoid, the presence of nitrogen in the tail gases of the process.

The amount of oxygen supplied to the reactor is chosen so as to obtain a ratio of H2S/O2 in the feedstock greater than the typical value of approximately 2:1. Preferably, the ratio of H2S/O2 in the feedstock should be in the range of 2:16:1, or preferably more than 2:1 and up to 6:1, more preferably in the range of 3:1-5:1, even more preferably in range 4:1-4.5:1.

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In a preferred embodiment, when conducting catalytic oxidative cracking based on the ratio of H2S/O2 in the range from 4:1 to 4.5:1, most preferably from

4.1:1 to 4.5:1, the preferred reaction temperatures for simultaneously obtaining cracking and partial oxidation of H2S are in the range of 900-1500°C, preferably in the range of 900-1200°C.

More preferably, a temperature of approximately 1100°C is obtained.

In one embodiment, the feedstock for the catalytic oxidative cracking reaction zone or zones (acid gas containing H2S and oxygen-containing gas) is preheated to increase the reaction temperature, increase hydrogen production, and suppress SO2 formation.

In one embodiment, the H2S-containing acid gas and the oxygen-containing gas are mixed in a static mixer just prior to entering the catalytic bed of the catalytic oxidative cracking reaction zone or zones.

In one embodiment, the hydrogen concentration in the reaction chamber effluent (after quenching) is at least 3% (vol.), preferably at least 5% (vol.), most preferably at least 7% (vol.).

It should be noted that the reaction is preferably carried out autothermally. This refers to the fact that, despite the preferred adiabaticity of the process, heat transfer actually takes place, since the oxidation reaction is exothermic and the cracking reaction is endothermic, whereby the heat made available due to the passage of the exothermic reaction is used in the endothermic reaction.

Summarizing the above, it should be noted that the above method, due to a reasonable choice of catalyst, presumably favors reactions (4) and (5) in comparison with reactions (1) and (2), which leads to a lower degree of conversion of H2S, but, on the other hand, , to much more H2 production and to much less SO2 production. As a consequence of the lower H2S conversion compared to the conventional Claus plant, a higher acid gas recycle is obtained from the _H2S containing gas source (eg amine regenerator) to the reaction chamber.

The catalytic oxidative cracking process is used to reduce the temperature in order to obtain the desired reaction equilibrium. This results in an increase in hydrogen yield and minimization of SO2 generation, which in turn is used to minimize hydrogen consumption in the tail gas treatment section to reduce _SO2 to form H2S.

Preferably, the acid gas containing H2S and the oxygen-containing gas are fed separately to the reaction zone, and these gases are mixed before entering the catalyst bed.

The effluent from the reaction chamber is preferably quenched to avoid recombination of H2 and S2 to form H2S, namely the reverse reaction for (4), which would make the process sub-optimal in terms of overall conversion. Preferably, this quenching is carried out substantially instantaneously. Hardening is preferably carried out to a temperature of less than 950°C, preferably in the range of 850-750°C. The residence time in the quench zone is preferably as short as possible, typically in the range of 10 to 300 ms, preferably 10 to 100 ms, more preferably 10 to 50 ms.

The quench zone (which could be the reaction chamber zone) is preferably followed by a waste heat boiler and a sulfur condenser to cool the process gas and recover liquid sulfur. The latter is preferably carried out by generating high or medium pressure steam in the waste heat boiler or low or medium pressure steam in the sulfur condenser.

In one embodiment, quenching of the effluent gas stream from the reaction chamber is achieved by agitation with water at the end of the reaction chamber, said gas-water agitation being carried out using a water sprayer in a suitable mixing chamber immediately below the catalyst bed. In the most preferred embodiment, the quenching (in the sense of quenching) is carried out in the first part of the two-zone recovery boiler. In this zone containing short pipes, the gas reaches a temperature of approximately 700°C, and in the second zone containing ordinary pipes, the gas reaches a temperature of 350300°C.

Despite the significant reduction in SO2 formation in the above process, some SO2 formation will be unavoidable. In order to remove such SO2, the catalytic oxidative cracking stage is preferably followed by a tail gas treatment section. Here, a portion (eg, approximately 10-15% (v/v)) of the produced hydrogen is consumed to reduce residual SO2 to form H2S in the hydrogenation reactor. Due to the much higher hydrogen content and much lower SO ₂ content in the tail gas compared to a conventional Claus plant, the tail gas treatment section recovery step can be carried out in the absence of any hydrogen import.

The tail gases are preferably preheated and fed to the hydrogenation reactor. In it, SO2, as well as other residual sulfur compounds such as COS and CS2, are converted to hydrogen sulfide.

H2S, which is then removed. This removal can be carried out in a conventional manner, for example by gas scrubbing using a lean amine solution in the absorber.

In one embodiment, the catalytic oxidative cracking stage is followed by a single Claus catalytic stage containing a gas reheater, a catalytic Claus reactor, and a sulfur condenser to convert most SO2 to sulfur, thereby minimizing H2 consumption to reduce SO2 to tail gas treatment sections.

In one embodiment, the hydrogen stream obtained from the TGT absorber is sent to end users such as hydroprocessing units, hydrocracking units, or hydrodesulfurization units. It should be noted that the composition of the hydrogen-rich stream from the top of the TGT absorber may vary depending on variables such as SRU feedstock quality, plant configuration and operating conditions, and may contain trace amounts or percentage levels of H2O, N2 , CO, CO2, H2S, COS and CS2.

In one preferred embodiment, the hydrogen stream obtained from the TGT absorber is further purified in a hydrogen purification section (eg, a pressure swing absorber). It should be noted that prior to purification, the composition of the hydrogen-rich stream from the top of the TGT absorber may vary depending on variables such as SRU feedstock quality, plant configuration and operating conditions, and may contain trace amounts or percentage levels of H2O. , N2, CO, CO2, H2S, COS and CS2.

The purified hydrogen is sent to end users such as hydroprocessing units, hydrocracking units or hydrodesulfurization units.

The invention will be illustrated with reference to the following non-limiting examples.

Example 1 Wet impregnation/precipitation catalyst preparation procedure.

Al ₂ O ₃ with a spherical shape with a diameter of 3 mm was used as a substrate after preliminary calcination in air at 900°C for 12 h and grinding to obtain the desired particle size distribution (355-710 μm).

As salt precursors of active substances, (CH ₃ COO) ₂ Fe and ammonium heptamolybdate tetrahydrate (NH4) ₆ Mo ₇ O ₂₄ -4H ₂ O were chosen.

The preparation procedure includes obtaining an aqueous solution of the salt of the precursor, in which the support is also dispersed.

The precursor salt, the amount of which is calculated on the basis of the loading of the active substance, is dissolved in doubly distilled water. After complete dissolution, the substrate is added to the solution while stirring and heating on a hot plate until the water is almost completely evaporated. The sample is subsequently dried at 120° C. until complete evaporation of the water is achieved. After the drying procedure was completed, the catalyst was pre-treated using 5% H2S in nitrogen and heated up to 1000° C. (heating rate: 10° C./min and isothermal for 60 min).

To obtain bimetallic samples, salt precursors of active substances are stirred and then dissolved in 1.5 dm ³ of water. In particular, 0.685 and 0.452 g of ammonium heptamolybdate tetrahydrate and iron acetate, respectively, and 5 g of alumina were used. Thus, the resulting catalyst is characterized by a nominal loading of 10% (wt.) MoS ₂ and 5% (wt.) FeS ₂ where the loading is calculated as follows:

MeS ₂ ,% (wt.) \u003d MeS ₂ , grams * 100 / (MeS ₂ , grams + Al ₂ O ₃ , grams).

Example 2 Catalytic Activity Tests

Conducted two types of tests for catalytic activity; the main differences are the composition of the stream fed to the oxidative cracking reactor and the contact time:

40% H2S, 0.7% CH ₄ , 8% O ₂ , nitrogen balance (H ₂ S/O ₂ =5).

Contact time selectable as 20 or 40 ms.

Temperature: 1100°C.

15% H ₂ S, 10% NH ₃ , 11% O ₂ , nitrogen balance.

Contact time selectable as 1 or 2 s.

Temperature: 1100°C.

It should be understood that the actual contact time will vary depending on the form of the catalyst (eg, powder or structured catalyst). The contact time, as defined herein, is the ratio between the volume of catalytic material and the total volume flow rate of the feed stream (calculated at 0° C. and 1 atm). The catalytic material herein is the active metals (iron and molybdenum) plus an aluminum-containing support. The contact time thus determined is in the range of 0.05 to 5 s, preferably 0.1 to 1 s.

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Claims (6)

In the case of a feed stream containing ammonia, more oxygen will be added as the ammonia and H2S compete with each other for oxygen consumption. This additional amount of oxygen is preferably such as to obtain an NH3/O2 molar ratio in the range of 0.9 to 1.5. The NH3 concentration in the feed stream ranges from 0 to 45% (vol.), while the relative concentration of O2 ranges from 0 to 33% (vol.). In particular, the amount of oxygen was modulated in such a way as to obtain the same H2S/O2 ratio typical for oxidative cracking, (5/1), so in general Applicants can say about the consumption of 30% O2 by hydrogen sulfide H2S and the consumption of the remaining 70% ammonia NH3. In addition, longer contact time is required because ammonia is much less reactive than H2S, while H2S and methane show identical reactivity. The main results are shown in the following tables. Table 1 Tests for catalytic activity for oxidative cracking in the presence of methane. Parameter Experiment Equilibrium The degree of conversion of CH ₄ ,% 100 100 Degree of conversion of H ₂ S,% 46 54 Yield H ₂ ,% 9 19 Selectivity for SO ₂ ,% 0-0.3 0.4 Selectivity for CO ₂ ,% 70 45 The results were observed to be independent of contact time and the formation of sulfur compounds such as CS2 and COS was not detected over the entire range of contact time studied. table 2 Tests for catalytic activity for oxidative cracking in the presence of ammonia. Parameter Experiment Equilibrium The degree of conversion of NH ₃ ,% 100 100 Degree of conversion of H ₂ S,% 47-64 64 Output H ₂ ,% 4-6 12 Selectivity for SO ₂ ,% 0-0.1 6 The results are observed to be slightly dependent on contact time, in particular the H2S conversion decreases with decreasing contact time, while the H2 yield shows the opposite trend with increasing with decreasing contact time (Table 3). No NO _x formation is detected because all of the oxygen that reacts with ammonia is consumed in the ammonia combustion reaction to form nitrogen (2NH3+3/2O ₂ =N ₂ +3H2O). Table 3 Consumption, normal cube. cm/min Contact time, sec H ₂ S conversion, % NH ₃ conversion, % H ₂ yield, % SO ₂ selectivity, % 300 1.2 64 100 4 0.07 900 0.4 55 100 5 0.08 1180 o,h 47 100 6 0.09 CLAIM 1. A method for preparing a catalyst suitable for use in the catalytic oxidative cracking of a gas stream containing H ₂ S, which contains iron and molybdenum, supported on a carrier containing aluminum, in which an aqueous solution of iron and molybdenum precursors selected from the group consisting of from iron tetracarbonyl, iron pentacarbonyl, iron nonacarbonyl, iron nitrates, iron bromides, iron chlorides, iron fluorides, iron phosphates, ferrous sulfates, ferrous acetylacetonates, ferrous acetates, ferrous fumarates, ferrous gluconates, ferrous citrates, ferrous benzoates, ferrous maleates, oxalates iron, iron oleates and iron stearates, iron-ammonium complexes and similar molybdenum precursors, and said solution is dispersed on the carrier, and before dispersing the solution on the carrier, the carrier is subjected to calcination at a temperature of from 750 to 1000 ° C, while the calcination step is carried out in the presence of air . 2. The method of claim 1 wherein the iron in the catalyst is in the form of iron sulfide. 3. Process according to claim 1 or 2, wherein the molybdenum in the catalyst is in the form of molybdenum sulfide. 4. Process according to any one of claims 1 to 3, wherein the catalyst support is formed from Al ₂ O ₃ . - 8 040662 5. Process according to any one of claims 1 to 4, wherein the iron in the catalyst is present in an amount ranging from 1 to 50% (w/w), preferably 2% (w/w), of the catalyst. 6. The method according to any one of claims 1 to 5, wherein the molybdenum in the catalyst is present in an amount ranging from 1 to 50% (w/w), preferably 6% (w/w), of the catalyst. Eurasian Patent Organization, EAPO Russia, 109012, Moscow, Maly Cherkassky per., 2