PROCESS FOR REMOVAL OF HYDROGEN CYANIDE FROM SYNTHESIS GAS

Abstract: Hydrogen cyanide is removed from an HCN containing synthesis gas, CO + H2, by contact with a metal oxide catalyst comprised of the oxides of molybdenum, titanium and aluminum in the presence of water vapor, and subsequently water washing the resulting synthesis gas.

Description: PROCESS FOR REMOVAL OF HYDROGEN CYANIDE FROM SYNTHESIS GAS FIELD OF THE INVENTION This invention relates to the removal of hydrogen cyanide, HCN, from synthesis gas streams. More particularly, this invention relates to the use of a composite catalyst for HCN removal from essentially sulfur free streams containing hydrogen and carbon monoxide. BACKGROUND OF THE INVENTION Synthesis gas, hydrogen and carbon monoxide, is used in a variety of hydrocarbon synthesis processes, e.g., Fischer-Tropsch. However, trace components that find their way into the synthesis gas are often poisons for hydrocarbon synthesis catalysts. An example of a trace component that poisons Fischer-Tropsch catalysts is hydrogen cyanide. This component is difficult to remove from synthesis gas because of its low solubility in common solvents, e.g., water, and because of its low concentration, usually less than about 100 ppm, removal by adsoφtion is difficult Also, chemical removal by, for example, alkaline scrubbing, is hampered by the presence of other acidic materials, e.g., CO2- Consequently, a need exists for the economic removal of HCN from synthesis gas at levels of at least about 95%, and particularly at temperatures similar to those employed in the hydrocarbon synthesis step. SUMMARY OF THE INVENTION In accordance with this invention hydrogen cyanide is substantially removed from an HCN containing synthesis gas by contacting the synthesis gas with a composite metal oxide catalyst containing the oxides of a Group VI metal, e.g , molybdenum, a Group IVB metal, e.g. titanium, and aluminum in the presence of water vapor. A synthesis gas of reduced HCN content is then produced for use in subsequent hydrocarbon synthesis reactions, such as the Fischer-Tropsch reaction Reaction conditions include elevated temperatures and elevated pressures, and at these conditions at least about 95% ofthe HCN contained in the synthesis gas feed stream is removed, preferably at least about 98%, more preferably at least about 99%, removal of HCN is obtained by this process. Subsequent to the contacting step, the synthesis gas is scrubbed with water to remove NH3 originally present or as hydrolyzed from HCN. Preferably, the catalyst is characterized by the absence or substantial absence of Group VIII metals or the oxides thereof. BRIEF DESCRIPTION OF THE DRAWINGS Figure 1 shows the performance of a molybdenum-alumina catalyst, without hydrogen pre-treatment. Figure 2 shows the performance of a molybdenum-alumina catalyst with hydrogen pre-treatment. Figure 3 shows the performance of a molybdenum-alumina catalyst with hydrogen pre-treatment at a higher temperature. Figure 4 shows the performance of a molybdenum-titania-alumina catalyst with hydrogen pre-treatment. Figure 5 shows the performance of a titania-alumina catalyst without hydrogen pretreatment. Figure 6 shows the performance of a titania-alumina catalyst with hydrogen pre-treatment. Curve A is always 100 (HCN out/HCN in), or % HCN conversion, Curve B is always temperature, Curve C is always ppm HCN, and Curve D is always ppm CH4. In all figures the left ordinate is 100 (HCN out/HCN in), the abscissa is days on stream, and the right ordinate is ppm methane or ppm HCN in effluent. The HCN removal process proceeds via the reaction of hydrogen cyanide with water: HCN + H2O -> NH3 + CO and whereas NH3, ammonia, is also a Fischer-Tropsch catalyst porous, NH3 can be readily removed, e.g., by a washing ofthe treated synthesis gas. Feed gases for the process generally and primarily contain hydrogen, carbon monoxide and water vapor, while small amounts of CO2 may also be present. Usually, the feed gas is sulfur free, e.g., less than 1 wppm sulfur, preferably less than 500 wppb. In a preferred embodiment, the catalyst is treated with hydrogen at elevated pressures, thereby causing an increase in catalyst activity. While the metal oxides used in this invention, molybdenum, titanium, and aluminum, are known as difficulty reducible oxides, there is evidence that the oxidation state of at least the molybdenum and titanium is reduced by virtue ofthe hydrogen treatment. It is unlikely, however, that the hydrogen treatment leads to reduction ofthe oxides to the elemental state; nevertheless, there is some reduction of at least some ofthe oxides. The composite oxide catalyst generally has >0 to 30 wt% molybdenum as the oxide, preferably 10 to 20 wt%; the titania concentration is about >0 to 30 wt%, preferably 4 to 20 wt%, more preferably 8-16 wt%, e.g., 8 wt%; the remainder being alumina. The catalyst may be treated with hydrogen, or a hydrogen containing stream, although the effects of hydrogen treatment are not always manifest. Because the materials ofthe composite catalyst are essentially difficulty reducible oxides, we believe that there is virtually no metal present in the zero valence state, e.g., less than 0.1 wt% zero valence metal. However, some ofthe molybdenum, in particular, and perhaps some ofthe titanium, may be in a reduced valence state after hydrogen treatment. When hydrogen treatment is effected, temperatures may range from about 200-600°C, preferably 230-550°C for periods of about 1-24 hours. The catalyst ofthe present invention may be readily prepared by depositing suitable sources of molybdenum and titanium on an alumina support. After deposition, for example, by impregnation or incipient wetness techniques, the material is dried, e.g., overnight at about 100-150°C, followed by calcination at temperatures of about 250-500°C, preferably 350-450°C to produce the oxide form. Suitable sources of molybdenum include ammonium hepta molybdate, ammonium molybdate, molybdenum trioxide; while suitable titanium sources include titanium isopropoxide, titanium oxychloride, titanium sulfate, titanium chloride, potassium titanium oxalate and other similar sources well known to those skilled in the art. The alumina may be any alumina useful as a catalyst support, including eta and gamma forms of alumina, and may have surface areas ranging from about 100- 400 m 2 /gm. HCN removal can be carried out over a relatively wide temperature range, e.g., 150-350°C. However, the preferred temperature range is that compatible with the subsequent process step, e.g., Fischer-Tropsch processing at 150°-250°C, more preferably about 160-235°C. Pressures are similarly wide ranging, e.g., 1-100 bar, although preferred pressures are in the range of 10-50 bar, more preferably 15-30 bar. The hydrolyzed HCN and product NH3 can then be removed from the feed synthesis gas by any applicable method well known to those skilled in the art, for example, water scrubbing or absoφtion onto a solid absorbent. The synthesis gas of relatively low HCN concentration may then be employed in a Fischer-Tropsch hydrocarbon synthesis process using shifting or non-shifting catalysts. Suitable Fischer- Tropsch catalysts include non-shifting Group VIII metals, preferably cobalt or ruthenium in bulk or supported form. In supported form, supports may be silica, alumina, silica-alumina, or titania. Promoters may also be employed, e.g., zirconium, rhenium, hafnium, etc. The process of this invention will be more fully appreciated by the following examples which serve to illustrate, but not limit, the invention. CATALYST TESTING The following general procedure was used for testing Catalysts A and C: A weighed amount of catalyst was mixed with 2 cc (about 2.7 gms) of 14 x 35 mesh crushed inert material (Denstone). This was placed in a 0.4 inch I.D. tubular stainless steel reactor. The catalyst/diluent was supported by a plug of Pyrex wool at the bottom ofthe reactor. On top of this bed, 4 cc (about 5.4 gms) of 14 x 35 mesh crushed Denstone was placed. A thermocouple was inserted into the bed for temperature control. The charged reactor was then placed into a vertically mounted infrared furnace and connected to supply and withdrawal tubing. The flow path was downflow, first through the diluent layer and then to the catalyst/diluent bed. Product analysis was by gas chromatography. HCN conversion was determined by comparison vs. an internal standard. Gases were supplied by electronic mass flow controllers. Water, HCN, and internal standard were supplied via a liquid feed pump. The liquid feed was mixed with the gas feed and vaporized in the upper section ofthe reactor. All ofthe catalysts described below were prepared using an extruded alumina support. The properties of this support were: Surface Area 169 m 2 /gm Pore Volume (H2P) 0.725 ml/gm Compacted Bulk Density 0.50 gm/ml Median Pore Diameter (Hg) 13 nm Shape Asymmetric Quadralobe Major Diameter 1.40 mm Minor Diameter 1.14 mm Average Extrudate Length 4.6 mm Non-volatile matter (%) 90.7 The catalysts were prepared as follows: Catalyst A Ammonium heptamolybdate (22.229 gms) and citric acid (30.5 gms) were dry-mixed in a 300 ml flask. To this was added 60 ml of deionized water. The solution was mixed with mild heating until it was clear. Deionized water was then added to a final volume of 83.8 ml. This solution was added to 110.25 gms ofthe alumina extrudates. The flask was stoppered and allowed to sit overnight. It was then dried at 120°C in flowing air for 4 hours and then calcined in air for 2 hours at 1000°F. Catalyst B Titanium Isopropoxide (42.028 gms) was dissolved with 30 ml of isopropyl alcohol in a 300 ml flask. Isopropyl alcohol was then added to bring the solution volume to 88 ml. Then 110.25 gms ofthe alumina extrudates were added to the solution in the flask. The flask was stoppered and shaken until all ofthe extrudates were wet. The wet extrudates were then put into an evaporation dish and dried in air at room temperature for 2 hours. The air dried extrudates were then further dried in an oven overnight. The extrudates were then calcined in air for 2 hours at 1000°F. Titanium isopropoxide (26.849 gms) was dissolved in isopropyl alcohol to give a total solution volume of 78 ml. To this, 97.2 gms ofthe above calcined extrudates were added and the flask stoppered. The flask was shaken until all ofthe extrudates were wet. The wet extrudates were then put into an evaporation dish and dried in air at room temperature for 2 hours. The air dried extrudates were then further dried in an oven at 120°C in flowing air overnight. The extrudates were then calcined in air for 2 hours at 1000°F to give Catalyst B. Catalyst C Ammonium molybdate (9.704 gms) and citric acid (12.8 gms) were dry- mixed in a 250 ml flask. To this was added 25 ml of deionized water. The solution was stirred with mild heating until it was clear. Deionized water was then added to give a final solution volume of 42 ml. While the solution was still warm, 50.0 gms of dry Catalyst B was added to the flask. The flask was stoppered and shaken until all ofthe catalyst particles were wet. The stoppered flask was allowed to sit overnight. The catalyst then was put in an evaporation dish and dried in flowing air at 120°C for 4 hours. The dried catalyst was then calcined at 1000°F in air for 2 hours to give Catalyst C. Catalyst D - Comparative Example To activated alumina (LaRoche Chemicals A-2, surface area 299 m 2 /gm, 0.65 gm/mJ bulk density, 12 x 32 mesh) sufficient titanium oxychloride was added to give 10% by weight Tiθ2- This material was calcined in air for 3 hours at 751°F. This material was then impregnated to incipient wetness with a water solution of ammonium heptamolybdate to give a loading of 7.5 wt% Mo03- This material was dried in air at 220°F and then calcined in air at 751°F for 3 hours. This material was then crushed to <150 microns to give Catalyst D. Example 1 - Catalyst A Without Prior Reduction Catalyst A (1.08 gms, 2 cc) was charged as described above to the reactor. Gas flows were established and the reactor temperature was set at 450°F. Liquid flow was then established. The operating conditions were: Feed Gas Composition Hydrogen 44.1 mole% Carbon Dioxide 8.3 mole% Water 19.1 mole% Argon 9.1 mole% Carbon Monoxide 18.8 mole% HCN 637 ppm Pyrrole (Internal Std) 50 ppm Pressure 320 psig GHSV 10360 1/hr Results of this operation are shown in the Figure 1. HCN conversion improved with time on stream. Increasing the temperature to 650°F resulted in HCN conversion of >99%. A portion of this activity was retained when the temperature was reduced back to 450°F, suggesting the catalyst activation was not complete at the end ofthe first 450°F period. At 450°F the methane content ofthe product gas was about 5 ppm, comparable to the methane observed at 450°F with only inert Denstone in the reactor (blank run). Methane at 650°F was about 8 ppm, showing a net methane make of 3 ppm. Example 2 - Catalyst A With Prior Reduction bv Hydrogen at 850°F Catalyst A (1.08 gms, 2 cc) was charged as described above to the reactor. Hydrogen flow was established (9435 GHSV) and the temperature increased to 850°F. The reactor was held for two hours at this condition. The pressure averaged 158 psig. At the end of this period, the reactor was cooled under flowing hydrogen to 450°F and the pressure increased to 320 psig. The other gas flows and liquid flow were then established. The operating conditions were: Feed Gas Composition Hydrogen 43.0 mole% Carbon Dioxide 8.5 mole% Water 19.0 mole% Argon 9.5 mole% Carbon Monoxide 19.3 mole% HCN 640 ppm Pyrrole (Internal Std) 50 ppm Pressure 320 psig GHSV 10590 1/hr The catalyst showed very high HCN conversion immediately, in contrast with the previous example where a "break-in" period was apparent. Furthermore, the activity at 450°F was significantly improved. In Example 1 the best HCN conversion at 450°F was 91.8%. In this example HCN conversion at 450°F was 99.1%. Activity maintenance was also good. The final test temperature was 450°F and HCN conversion was as high as during the initial part ofthe test at 450°F. Example 3 - Catalyst A With Prior Reduction bv Hydrogen at 1000°F Catalyst A (1.08 gms, 2cc) was charged as described above to the reactor. Hydrogen flow was established (9330 GHSV) and the temperature increased to 1000°F. The reactor was held for 2 hours at this condition. The pressure averaged 152 psig. At the end of this period, the reactor was cooled under flowing hydrogen to 400°F and the pressure increased to 320 psig. The other gas flows and liquid flow were then established. The operating conditions were: Feed Gas Composition Hydrogen 43.8 mole% Carbon Dioxide 8.5 mole% Water 19.1 mole% Argon 9.4 mole% Carbon Monoxide 19.3 mole% HCN 670 ppm Pyrrole (Internal Std) 50 ppm Pressure 320 psig GHSV 10540 1/hr The results are shown in the figure. Similar to Example 2, the catalyst immediately showed good HCN conversion. The activities at all temperatures were marginally higher than in Example 2. Consistent with Examples 1 and 2, net methane make was very low. A comparison of Example 1 with Examples 2 and 3 (in Table I) clearly shows prior reduction with hydrogen improves catalyst activity. A comparison of Examples 2 and 3 shows a small benefit for increasing reduction temperature to 1000°F from 850°F. Example 4 - Catalyst C With Prior Reduction bv Hydrogen at 850°F Catalyst C (1.14 gms, 2 cc) was charged as described above to the reactor. Hydrogen flow was established (9460 GHSV) and the temperature increased to 850°F. The reactor was held for 2 hours at this condition. The pressure averaged 155 psig. At the end of this period, the reactor was cooled under flowing hydrogen to 400°F and the pressure increased to 320 psig. The other gas flows and liquid flow were then established. The operating conditions were: Feed Gas Composition Hydrogen 43.8 mole% Carbon Dioxide 8.5 mole% Water 19.1 mole% Argon 9.4 mole% Carbon Monoxide 19.3 mole% HCN 670 ppm Pyrrole (Internal Std) 50 ppm Pressure 320 psig GHSV 10550 1/hr Similar to Example 2, the catalyst immediately showed good HCN conversion. The activities at all temperatures were significantly higher than in Example 2. Consistent with Examples 1 and 2, net methane make was very low. This example shows the combination of titania and molybdenum oxide on an alumina support provides a particularly active catalyst for HCN removal from syngas. Comparison ofthe first and second periods at 400°F shows good activity maintenance, with only a slight loss occurring over the test period. Example 5 - Catalyst D Without Prior Reduction Catalyst D (0.4625 gms, 0.5 cc) was mixed with crushed high purity tubular alpha alumina (10.72 gms, 5.5 cc). This was placed in a 0.4 inch I.D. tubular stainless steel reactor. The catalyst/diluent was supported by a plug of Pyrex wool at the bottom ofthe reactor. A thermocouple was inserted into the bed for temperature control. The charged reactor was then placed into a vertically mounted infrared furnace and connected to supply and withdraw tubing. The flow path was downflow through the catalyst/diluent bed. Product analysis was by gas chromatography. HCN conversion was determined by comparison vs. an internal standard. Gases were supplied by electronic massflow controllers. Water, HCN, and internal standard were supplied via a liquid feed pump. The liquid feed was mixed with the gas feed and vaporized in the upper section ofthe reactor. The operating conditions were: Feed Gas Composition Hydrogen 53.7 mole% Carbon Dioxide 6.4 mole% Water 18.2 mole% Argon 8.4 mole% Carbon Monoxide 13.3 mole% HCN 200 ppm Pyrrole (Internal Std) 500 ppm Pressure 320 psig GHSV 41000 1/hr The results are shown in the table Very good HCN removal activity was achieved. HCN conversions were lower than those in Examples 1-4, but this is due to the much higher space velocity (41000 vs. 10500) in this Example This example shows the benefit of higher alumina support surface area (299 m- * Vgm vs 169 m-^/gm for Catalysts A-C) and smaller particle size (<150 micro vs 1/20" extrudates for Catalysts A-C). Example 6 - Catalyst B With prior Reduction by Hydrogen at 850°F Catalyst B (0.98 gm, 2 cc) was mixed with 2 cc (about 27 gms) of 14 x 35 mesh crushed inert material (Denstone) This was placed in a 0 4 inch I.D tubular stainless steel reactor The catalyst/diluent was supported by a plug of Pyrex wool at the bottom ofthe reactor On top of this bed, 4 cc (about 5 4 gms) of 14 x 35 mesh crushed Denstone was placed. A thermocouple was inserted into the bed for temperature control. The charged reactor was then placed into a vertically mounted infrared furnace and connected to supply and withdraw tubing The flow path was downflow, first through the diluent layer and then to the catalyst/diluent bed Product analysis was by gas chromatography HCN conversion was determined by comparison vs. an internal standard. Gases were supplied by electronic mass flow controllers Water, HCN, and internal standard were supplied via a liquid feed pump The liquid feed was mixed with the gas feed and vaporized in the upper section ofthe reactor Hydrogen flow was established (9830 GHSV) and the temperature increased to 850°F. The reactor was held for 2 hours at this condition The pressure averaged 90 psig. At the end of this period, the reactor was cooled under flowing hydrogen to 450°F and the pressure increased to 320 psig. The other gas flows and liquid flow were then established. The operating conditions were: Feed Gas Composition Hydrogen 43.8 mole% Carbon Dioxide 8.5 mole% Water 18.6 mole% Argon 9.7 mole% Carbon Monoxide 19.4 mole% HCN 670 ppm Pyrrole (Internal Std) 50 ppm Pressure 320 psig GHSV 10830 1/hr The results are shown in the figure. As can be seen, HCN conversion initially decreased with time, finally stabilizing at about 70%. Table I below shows a compilation ofthe results of Examples 1-6. Example 4, using a molybdenum-titania on alumina catalyst showed the best HCN reduction at the lowest temperature, i.e., the highest activity catalyst. TABLE I

Claims: CLAIMS: 1. A process for reducing HCN concentrations of HCN containing synthesis gas which comprises treating the synthesis gas, in the presence of water vapor and elevated temperatures and pressures with a composite metal oxide containing catalyst comprising the oxides of molybdenum, titanium and aluminum, and in the substantial absence of Group VIII metals or compounds. 2. The process of claim 1 wherein the gas is essentially free of sulfur. 3. The process of claim 1 wherein a synthesis gas of reduced HCN concentration is recovered. 4. The process of claim 3 wherein the recovered synthesis gas is water washed. 5. The process of claim 1 wherein the catalyst is pre-treated with a hydrogen containing stream. 6. The process of claim 1 wherein the molybdenum and titanium oxides are present in amounts of >0 to 30 wt% each. 7. The process of claim 1 wherein the treated synthesis gas is water washed and converted to hydrocarbons in the presence of shifting or non-shifting Fischer-Tropsch catalyst. 8. The process of claim 7 wherein the Fischer-Tropsch catalyst comprises supported or bulk cobalt or ruthenium. 9. The process of claim 1 wherein the temperature ranges from about 150-350°C and pressures range from about 10-50 bar. 10. The process of claim 1 wherein the HCN containing synthesis gas is essentially sulfur free. 11. The process of claim 10 wherein the HCN containing synthesis gas contains less than 1 wppm sulfur.

Inventors: CORKERN KEITH E, KOVEAL RUSSELL J, RILEY KENNETH L, CORKERN, KEITH, E., KOVEAL, RUSSELL, J., RILEY, KENNETH, L.
Applicants: EXXON RESEARCH ENGINEERING CO, EXXON RESEARCH AND ENGINEERING COMPANY
Application year: 1997
Main classification: C01B3
IPC codes: B01D53, C01B3, C10K1