R&D on Hydrogen Production by Autothermal Reforming

Transcription

1 sin 2.3 R&D on Hydrogen Production by Autothermal Reforming (Shinnen Satte Laboratory) Takashi Suzuki, Katsumi Miyamoto, Shuichi Kobayashi, Noriyuki Aratani, Tomoyuki Yogo 1. R&D Objectives The purpose of the present R&D is to develop high-efficiency hydrogen production technology through reforming of naphtha fraction to kerosene fraction, including GTL oil, etc. Another aim is to investigate the constituents that make up the raw material hydrocarbons and reforming characteristics. The R&D is scheduled to take place over 5 years from JFY2000 to JFY2004. As part of this R&D program, the Shinnen Satte Laboratory will endeavor to develop technological processes involving the autothermal reforming method, to investigate the reforming characteristics of naphtha fraction to kerosene fraction, including GTL oil, etc., and to develop desulfurization catalyst and processes essential for removing sulfur compounds in everything from naphtha to kerosene fraction. 2. R&D Contents The contents of R&D conducted are summarized below. Design of autothermal reforming process Basic establishment of Autothermal Reforming - Fuel Processing System Development of autothermal reforming catalyst Investigation of hydrocarbon constituents in fuel and of autothermal reforming characteristics Development of catalytic desulfurization process 3. R&D Results 3.1 Design of Autothermal Reforming Process Basic Establishment of Autothermal Reforming - Fuel Processing System An autothermal reforming - fuel processing system (hereinafter ATR-FPS) was investigated for the purpose of evaluating autothermal reforming reactions. A flow diagram of the basic process in autothermal reforming is presented in Figure The ATR-FPS is comprised of a reform reactor + CO shift reactor + CO selective remover. Air equivalent to 3 times the amount of oxygen required for CO combustion after CO shift reaction (O 2 /CO = 1.5) is supplied at the inlet to the CO selective remover. In addition, heat balance was determined from changes in enthalpy arising from compositional changes, on the assumption that in reactions after CO shift reaction, methanation reaction does not advance, and that discharge heat loss from the system is zero. 1

2 Raw material Reformer Shift CO removal Figure 3.1.1: Flow Diagram of Autothermal Reforming Process Establishment of Oxygen/Carbon Ratio and Steam/Carbon Ratio In determining the oxygen/carbon (hereinafter O 2 /C) ratio and the steam/carbon (hereinafter S/C) ratio supplied to ATR-FPS, O 2 /C and S/C were adjusted by equilibrium calculation so that the total of heat absorption/generation from each reactor and of heat exchange required for reaching each reaction temperature becomes 0-1 [kj/mol-crude oil] in the entire ATR-FPS with autothermal reforming + CO shift + CO selective removal. System internal pressure is kg/cm 2 and each reactor is isothermal. Taking normal decane as an example, the relationship between S/C and O 2 /C obtained by equilibrium calculation when the autothermal reforming reaction temperature was set at 650 C is shown in Figure In the figure, when O 2 /C is 0.45, if S/C is set to 2.37, the heat balance of ATR-FPS as a whole becomes 0.36 kj/mol-nc 10. In evaluating catalytic activity or comparing the autothermal reforming reactivity of each fuel oil, standard conditions of O 2 /C and S/C must be set up. In making these settings, the following two setting standards must be satisfied because with the current power generation system, with PEMFC built in, unless the CO concentration in gas at the CO shift reactor outlet is less than 0.5 vol%, it will be difficult to have the CO concentration reach under 10 ppm at the CO selective remover at a later stage. Another reason is that it is important to have the concentration of hydrogen in generated gas as high as possible. (1) CO concentration in CO shift outlet gas must be less than 0.5 vol%. (2) O 2 /C must be as low as possible. Taking normal decane as an example, Figure shows the relationship between O 2 /C obtained by equilibrium calculation and CO concentration at shift reaction outlet when the autothermal reforming reaction temperature was set at 650 C. It was clarified that when O 2 /C was set at 0.45, the two aforesaid standards are satisfied. It was also confirmed that similar results obtain with raw material hydrocarbon other than normal decane. 2

3 Total heat balance (kj/mol-nc10) CO concentration (vol%) after shift (250 C) CO concentration upper limit: Figure 3.1.2: O 2 /C Ratio vs S/C Ratio Obtained by Equilibrium Calculation (n-decane, 650 C) Figure 3.1.3: O 2 /C Obtained by Equilibrium Calculation vs CO Concentration at Shift Reaction Outlet (n-decane, 650 C): CO Concentration Upper Limit at CO Selective Remover Inlet After Shift (estimated value) Investigation of Hydrogen Production on 1 Nm 3 /hr Scale With normal decane equivalent to kerosene fraction taken as the raw material, with throughput at 240 ml/hr, and with O 2 /C set at 0.45 and S/C at 2.37 as indicated in Section 3.1.2, hydrogen production at 1 Nm 3 /hr was implemented using a fixed-bed, flow-type unit for evaluating the activity of heat-resistant reactor equipment. Used as the catalyst was autothermal reforming catalyst A (hereinafter ATR catalyst A), obtained as a result of investigation of partial oxidation reaction and preliminary investigation of autothermal reforming conducted in the previous fiscal year. Using for reactions the gas generated after autothermal reforming reaction, reforming reactivity was evaluated from the gasification rate that indicates the percentage of conversion of normal decane to CO, CO 2 and C 4 hydrocarbon or below, and from the C 2 -C 4 hydrocarbon selection rate that indicates the percentage of conversion to hydrocarbon of C 4 or below. Because CO is denatured to H 2 of equivalent mol volume in the post-stage shift reactor of an autothermal reformer, H 2 + CO yield and H 2 + CO production volume were used to evaluate hydrogen production. Shown in Figure is the temperature distribution of catalytic layer in autothermal reform reaction. Heat generation can be seen at the upper catalytic layer, and it was confirmed that oxidation reaction progresses on the catalyst. Next, the impact of reaction temperature was investigated. Gasification rate and C 2 -C 4 hydrocarbon selection rate are shown in Figure 3.1.5, while hydrogen yield and hydrogen production volume are shown in Figure Gasification rate increases as temperature rises, and with normal decane at around 700 C, gasification of 97% or more took place. It was also recognized that when the reaction temperature rises, the types of hydrocarbon produced without complete reforming tend to decrease. In this way it was confirmed that as the reaction temperature rises, hydrogen yield and hydrogen production volume increase, gradually approaching the equilibrium value. 3

4 Temperature ( C) Upper catalyst layer Middle catalyst layer Lower catalyst layer Figure 3.1.4: Time (hr) Catalyst Layer Temperature Distribution (650 C) in Autothermal Reforming Reaction at 1 Nm 3 /hr Gasification rate (%) Gasification rate (%) C 2-C 4 selection rate C2-C4 selection rate (%) (H2 + CO) yield (%) Equilibrium value Equilibrium value (H2 + CO) Production Volume (Nm 3 /hr) Figure 3.1.5: Reaction temperature ( C) Impact of Reaction Temperature on Gasification Rate and C 2 -C 4 Hydrocarbon Selection Rate Figure 3.1.6: Reaction temperature ( C) Impact of Reaction Temperature on Hydrogen Yield and Hydrogen Production Volume 3.2 Basic Design of Autothermal Reforming - Fuel Processing System ATR-FPS Basic Design (1) ATR-FPS outline and design conditions ATR-FPS is a hydrogen production process that uses the autothermal reforming method. In the latest investigation, two units each of shift and selective oxidation reactor were installed so that CO shift or selective removal would be accomplished more securely. The ATR-FPS flow process is shown in Figure

5 Autothermal reform type hydrogen production Reactor system Flow sheet Figure 3.2.1: ATR-FPS Flow Diagram (2) Raw material specifications In the latest investigation, the crude oil composition ratio used was n-decane:n-butyl-cyclohexane:diethyl benzene (= 64:16:20 mol%); crude oil molecular formula was C 11.3 H 22.3 ; crude oil molecular volume was ; crude oil density (25 C) was kg/m 3, and crude oil flow volume was 380 ml/hr (= g/hr = gmol/hr). Steam and air supplied for reforming were each set at O 2 /C = 0.45 and S/C = 2.44, the conditions for attaining heat balance. In addition, as opposed to CO concentration at the shift reactor outlet, O 2 /CO was set to 1.5 (3 times the theoretical ratio) at the first CO selective remover and to 0.15 at the second reactor (1/10 the volume at the first reactor). (3) Substance balance Table presents substance balance as calculated based on the conditions given in Section (2). The following was also clarified. Hydrogen was produced at approximately 950 Nl/hr from raw material kerosene at 380 ml/hr (approx. 290 g/hr). The composition of hydrogen-rich gas product was H 2 (42.83 mol%): CH 4 (0.015 mol%): CO 2 (20.66 mol%): N 2 (36.49 mol%), with the ratios of H 2, CO 2 and N 2 at roughly 6:3:5. 5

6 Table 3.2.1: ATR-FPS Substance Balance Kerosene Kerosene + Steam + Air (4) Heat Balance (a) Heat recovery system Optimization of ATR-FPS heat recovery was attempted using pinch technology. It was confirmed that all the heat required (e.g., raw material vaporization, steam vaporization) could not be replenished by means of waste heat recovery. The heat balance obtained is shown in Table and heat recovery for ATR-FPS is given in Figure As Table indicates, whereas the total of received heat is kcal/hr, that of heat given off is kcal/hr, and since the volume of heat given off is greater than that of heat received, it is apparent that heating by an outside source is not required. According to pinch technology shown in Figure in 3.2.2, however, when a temperature differential (10 C) has been secured at pinch point (60 C), some 34.0 kcal/hr of heat, equivalent to roughly 4% of the total heat volume from heating (962.2 kcal/hr), is required from an external source. 6

7 Table 3.2.2: Heat Balance Received heat Emitted heat Heat exchanger Inlet temperature ( C) Outlet temperature ( C) Heat volume (kcal/hr) Crude oil vaporization Steam vaporization Fuel oil + steam + heated air Total ATR to LTS LTS-1 to LTS LTS-2 to PROX PROX-1 to PROX From PROX Total 1,140.7 Figure 3.2.2: ATR-FPS Heat Recovery 3.3 Development of Autothermal Reforming Catalyst Experiment Method A fixed-bed, flow-type micro reactor was used for reactions. Deep desulfurized kerosene (Sf < 0.5 ppm) was used as raw material oil. At prescribed volumes, calculated in accordance with Section 3.1.2, deep desulfurized kerosene and water were supplied via fluid feed pump, and air was supplied by thermal mass flow. Used as reactor was an electric oven heated to a prescribed temperature. ATR catalyst A served as the standard catalyst. In the current fiscal year, plans call for preparation of ATR catalysts B-E, to which tertiary ingredients have been added to ATR catalyst A in order to have catalyst that offers greater steam reforming activity, in an effort to achieve autothermal reforming at high activity. Carrier was obtained by kneading two types of oxide with binder, and then after molding, by sintering at a prescribed temperature under air flow. The carriers obtained were globular in shape, measuring 2-4 mm in diameter and m 2 /g in surface area. For active metal retention, metallic salt water solution was impregnated in obtained carrier, dried and adjusted. Prior to reaction, pretreatment took place for 3 hrs at 700 C under hydrogen current. Used in evaluating the reaction was CO + CO 2 selection rate, which gives the percentage of carbon in raw material hydrocarbons that has been transformed into CO and CO 2. 7

8 The gas obtained was analyzed using TCD or FID gas chromatogram; an oxygen combustion type carbon analyzer was used to analyze the volume of carbon accumulated on catalyst after the reaction Autothermal Reforming Reaction with Each Type of Catalyst Figure presents the results of a comparison of the dependency on reaction temperature of autothermal reforming reaction activity with each type of catalyst. The catalysts can be arranged in sequence as follows. Catalyst B > Catalyst C > Catalyst A (standard) > Catalyst D > Catalyst E At 750 C or above, a CO + CO 2 selection rate of virtually 100% was exhibited with all the catalysts. The temperature distribution of each catalytic layer at this time is shown in Figure 3.3.2, with 600 C taken as sample reaction temperature. With all the catalysts, it was confirmed that the temperature rises sharply near the inlet and declines as you go to the lower layers. Dissanayake et al. 1) and Groote et al. 2) report that in autothermal reforming reaction with methane, a complete oxidation reaction of methane takes place first, followed by water vapor reforming, CO 2 reforming, and aqueous gas shift reaction. It is believed that in the latest results as well, oxidation reaction advanced over the upper catalyst layer. When the catalysts are arranged in order of highest temperature at the catalyst layer inlet, the following sequence obtains. Catalyst E > Catalyst D > Catalyst C > Catalyst A (standard) > Catalyst B This sequence manifests a trend virtually opposite that of the activity sequence. It shows that when the catalyst inlet temperature is high (that is when the heat given off is great), the complete oxidation reaction takes precedence over the reforming reaction (steam reforming reaction or CO 2 reforming reaction). In catalyst with low autothermal reforming activity, the reforming reaction tends to advance less easily than the complete oxidation reaction. Conversely, in catalyst with high autothermal reforming activity, the complete oxidation reaction and reforming reaction both advance easily. The fact that the complete oxidation reaction and reforming reaction take place at nearby locations on the catalyst suggests that the temperature at catalyst inlet is kept relatively low. With catalyst A (standard) and catalyst C, however, this trend is reversed, and the factors determining catalyst layer temperature distribution are not just the catalyst s complete oxidation activity or reforming activity. Such things as thermal conductivity and filling density are also contributing factors. In the development of autothermal reforming catalyst, it is not only high activity that is important but also curtailment of carbon precipitation. Of the catalysts considered in Figure 3.3.1, ATR catalyst A and ATR catalysts B and C, which manifest higher activity than ATR catalyst A, were used in reactions that took place for 16 hrs at 600 C and the volumes of precipitated carbon thereafter were compared as shown in 8

9 Figure It was recognized that the volume of precipitated carbon with ATR catalyst B, which exhibited the highest activity among the catalysts compared, was the smallest, and it became evident that the volume of precipitated carbon with ATR catalyst A can be reduced by about 35%. With ATR catalyst C, which manifested higher activity than ART catalyst A, approximately 1.2 times greater carbon precipitation was noted as compared to ATR catalyst A. This shows that escalation of activity by the addition of tertiary constituents does not always match with curtailment of carbon precipitation. It is conjectured that in the design of autothermal reforming catalyst, escalation of activity and curtailment of carbon precipitation must be considered from separate standpoints. CO + CO2 selection rate (%) Catalyst A Catalyst B Catalyst C Catalyst D Catalyst E Catalyst layer temperature ( C) Outlet Inlet Catalyst A Catalyst B Catalyst C Catalyst D Catalyst E Catalyst layer outlet temperature ( C) (Deep desulfurized kerosene, LHSV = 1,O 2/C = 0.45, S/C = 2.437) Distance (cm) from catalyst layer outlet Figure 3.3.1: Comparison of Autothermal Reforming Reactivity with Each Type of Catalyst Figure 3.3.2: Catalyst Layer Temperature Distribution in Autothermal Reforming Reaction with Each Type of Catalyst Precipitated carbon volume (mass%) Catalyst A Catalyst B Catalyst C Figure 3.3.3: Comparison of Precipitated Carbon Volume after Reaction with Each Type of Catalyst (600 C, 16 hrs) 9

10 3.4 Investigation of Hydrocarbon Constituents in Fuel and of Autothermal Reforming Characteristics Experiment Method A fixed-bed, flow-type micro reactor was used for reactions. The hydrocarbon compounds shown in Table were used as raw material, together with standard ATR catalyst A. So as to compare the autothermal reforming reaction characteristics among each hydrocarbon, the most ideal reaction temperature was one at which side reactions such as combustion or decomposition could be suppressed as much as possible. Representative of each hydrocarbon, deep desulfurized kerosene and normal hexane were taken as raw materials. Reaction was initiated under the same conditions as for autothermal reforming reaction, using a micro-reactor without catalyst, and the presence of side reactions in the reactor was confirmed. As shown in Figure 3.4.1, because the C 2 -C 4 hydrocarbon selection rate is zero at 600 C or below, it was confirmed that no side reactions take place inside the reactor. Consequently, autothermal reforming reaction was implemented at 600 C, where there would be no side reactions. Taking the carbon mol flow volume included in the raw material hydrocarbons as standard, the rate was 0.71 mol/hr-c. O 2 /C and S/C were determined in accordance with Section Reforming reactivity was evaluated using a virtual speed constant k (CO + CO 2 ) covering CO + CO 2 production, determined on the assumption of a primary reaction and CO + CO 2 selection rate that correlates with the hydrogen selection rate targeted. Table 3.4.1: Model Hydrocarbon Compounds Equivalent to Naphtha - Kerosene Fractions Used Hydrocarbon n-p i-p O N A count LN 6 n-hexane 2,2 dimethylbutane 1-hexane cyclohexane benzene HN 8 n-octane 2,2,4-trimethylp 1-octane ethylcyclohexane m-xylene entane ethylbenzene KERO 10 n-decane n-butylcyclohexane diethylbenzene Other n-dodecane, n-hexadecane 1,2,4-trimethylbenzene Desulfurized kerosene C2-C4 selection rate (%) Reaction temperature ( C) (Reaction conditions: O 2 /C = 0.45, S/C heat balance conditions, LHSV = 1.0 to 1.6) Figure 3.4.1: Temperature Dependency vs Side Reaction in Reactor 10

11 3.4.2 Reforming Reactivity of Normal Paraffin Due to Differences in Carbon Number Figure presents a comparison of relative reforming activities for normal paraffin due to differences in carbon number. It can be seen that as carbon number increases, relative reforming activity declines. A comparison of precipitated carbon volume at this time is shown in Figure Because the precipitated carbon volume increased together with an increase in carbon number, relative reforming activity and precipitated carbon volume exhibited similar trends with ATR catalyst A. Figure gives the selection rates of C 2 -C 4 hydrocarbons produced at this time. It can be seen that the C 2 -C 4 hydrocarbon selection rates become greater as the carbon number increases. This fact suggests that when the carbon chain becomes large, even though it decomposes midway, unreformed hydrocarbon increases. What is more, the bulk of unreformed hydrocarbons are olefins. That precipitated carbon volume increases together with an increase in carbon number can be ascribed to the fact that unreformed olefins condense on catalyst surface, making it easy for carbons to be formed. Relative activity k (CO + CO2) /- C deposition weight (mass%) C number C number Figure 3.4.2: Comparison of Autothermal Reforming Reactivity of Normal Paraffin Due to Differences in Carbon Number Figure 3.4.3: Comparison of Precipitated Carbon Magnitude on Catalyst Due to Differences in Carbon Number C2-C4 selection rate (%) C number Figure 3.4.4: Selection Rate of C 2 -C 4 Hydrocarbon in Normal Paraffin Due to Differences in Carbon Number 11

12 3.4.3 Reforming Reactivity of Model Hydrocarbons Equivalent to Gasoline Fraction Figure presents a comparison of relative reforming activity for each model hydrocarbon compound of normal paraffin, isoparaffin, olefin, naphthene and aroma; the carbon number of model hydrocarbon equivalent to gasoline fraction was taken as 8. A comparison of precipitated carbon volume at this time is shown in Figure In each case, the sequence of relative reforming activity and precipitated carbon volume becomes as follows. Isoparaffin = Naphthene = Normal paraffin > Aromatic > Olefin It is believed that factors due to hydrocarbon structure play a large role in these sequences, but the details will have to be further investigated in the future. Relative activity k (CO + CO2) /- C deposition weight (mass%) Figure 3.4.5: Comparison of Autothermal Reforming Reactivity Due to Differences in Hydrocarbon Type (Carbon No.: 8) Figure 3.4.6: Comparison of Precipitated Carbon Volume on Catalyst Due to Differences in Hydrocarbon Type (Carbon No.: 8) 12

13 3.4.4 Reforming Reactivity of Each Model Hydrocarbon Type Equivalent from Naphtha to Kerosene Fraction Respecting the reforming reactivity of each model hydrocarbon equivalent to fractions from naphtha to kerosene, Figure presents a comparison of relative reforming activity, organized by carbon number, and Figure gives a comparison of carbon deposition weight. The relative reforming activity of normal paraffin is relatively high at low class but at high class, it tends to become lower than that of other hydrocarbons. Regardless of the carbon number, the reforming activity of naphthene was high in comparison to other hydrocarbons. With aromatic compounds, a clear correlation with carbon number could not be confirmed. On the contrary, structural factors such as substituent position or chain length are suspected. For isoparaffin and olefin, the trends in relative reforming activity could not be determined. Carbon deposition weight was small with naphthene and isoparaffin, but large with olefin. The deposited carbon weight with normal paraffin was small in comparison to other hydrocarbons at low class, the same as relative reforming activity, and large at high class. Among aromatic compounds, the deposited carbon weight did not exhibit a clear correlation with carbon number, because of substituent reactivity or structural factors such as electron polarization in aromatic rings. Relative reforming activity also exhibited different trends. These results indicate that among the hydrocarbons equivalent to fractions from naphtha to kerosene, the reforming reactivity is dependent upon carbon number in some cases, as in normal paraffin, but this factor is not adequate for explaining all hydrocarbons. To explore the details in greater depth, such things as hydrocarbon physical properties, chemical properties and interactions with catalyst will have to be further investigated. Relative activity k (CO + CO2) /- n-paraffin Naphthene Aroma Olefin i-paraffin C deposition weight (mass%) n-paraffin Naphthene Aroma Olefin i-paraffin C number C number Figure 3.4.7: Carbon No. vs Relative Reactivity Figure 3.4.8: Carbon No. vs Precipitated Carbon Magnitude 3.5 Development of Catalytic Desulfurization Process Experiment Method Almost all the kerosene in circulation in the market contains a sulfur component at the level of massppm. As a sample for desulfurization reaction, kerosene on the market at 50 massppm was prepared. Properties are indicated in Table

14 At the oil refinery, the hydrodesulfurization reactor operates with hydrogen partial pressure at MPa for light fraction and at a high pressure of 10.0 MPa or greater for heavy fraction. Nevertheless, with the small-scale fuel cell power generation system, including power for household use, atmospheric pressure must be considered in terms of the High-Pressure Gas Control Law, the Electric Utility Law, and so on. In the present research, therefore, the following two points were assumed for desulfurization reaction. 1. The raw material is kerosene on the market (with 50 massppm sulfur component). 2. The reaction pressure is atmospheric pressure. The purpose of the initial investigation, therefore, was to determine desulfurization performance of active metals at atmospheric pressure. The constituents and configurations of catalysts used in evaluating activity are listed in Table Desulfurization catalyst A is a regular extrusion-mold-type hydrodesulfurization catalyst. Desulfurization catalyst B was prepared by having precious metal, the active metal, retained in globular-shaped alumina, the carrier. Desulfurization catalyst C was obtained by molding base metal and zinc oxide in columnar-shaped tablets. A microreactor was used in evaluating the hydrodesulfurization activity of these three catalysts. Prior to the hydrodesulfurization reaction, desulfurization catalyst A filled in the reactor underwent preliminary sulfurization for 2 hrs at 350 C via 5% hydrogen sulfide/hydrogen gas; desulfurization catalysts B and C underwent hydrogenation pretreatment reduction for 2 hrs at 350 C under conditions of hydrogen flow Impact of Catalytically Active Metals Figure presents the results of an evaluation of the hydrodesulfurization activity of each catalyst in kerosene on the market with the H 2 /oil ratio at 100 and the reaction temperature at 250 C. The desulfurization activity of these three catalysts with LHSV = 10 can be expressed in relative terms from reaction speed constant. When the activity of desulfurization catalyst A is taken as 100, that of desulfurization catalyst B becomes 250 and that of desulfurization catalyst C, 773, revealing that the activity of desulfurization catalyst C is the highest. Table 3.5.1: Properties of Raw Material Kerosene Used Distillation ( C) IBP 152 Density g/cm % 170 Composition Aroma % % 177 Olefin % % 185 Saturation % % 192 Sulfur component massppm 50 50% % % % % 248 EP

15 Table 3.5.2: List of Catalysts for Investigation of Kerosene Hydrodesulfurization Catalyst Constituent Configuration Catalyst A Hydrodesulfuization catalyst Extrusion molded products on market Catalyst B Precious metals Globular shape Catalyst C Base metals Columnar shaped tablets Relative activity value Catalyst A Catalyst B Catalyst C Figure 3.5.1: Evaluation Results for Hydrodesulfurization Activity Impact of Catalyst Adjustment Conditions Desulfurization catalyst C is catalyst in which base metal components and metal oxides have been molded into columnar tablets. In order to investigate the impact of catalyst preparation conditions on hydrodesulfurization activity, four types of catalyst (Table 3.5.3) with different active metal load weight were prepared by means of the impregnation and co-precipitation methods. Using these catalysts, tests to evaluate activity were conducted with LHSV = 0.25 and reaction temperature at 300 C so as to clarify initial deterioration in activity. The results are shown in Figure The performance of catalyst prepared by the impregnation method was such that the Sp value exceeded 0.2 massppm in catalyst with low metal retention magnitude for about 100 hrs and in catalyst with high metal load weight, for about 450 hrs. In catalyst prepared by the co-precipitation method, on the other hand, the Sp value did not exceed 0.2 massppm for up to 500 hrs irrespective of the active metal magnitude. In order to determine the amount of hydrogen required for reaction in the catalyst system, an investigation was made in which the hydrogen supply volume was modified over 100 hrs after reaction startup. The results are shown in Figure With the H 2 /oil ratio at 50 or above, the Sp value was low irrespective of catalyst preparation method, but when the ratio fell below 50, it was found that the catalyst s desulfurization activity drops sharply. Given this fact, it is conjectured that the H 2 /oil ratio must be at least 50 in this catalyst system. Table 3.5.3: List of C-type Catalysts Catalyst name Preparation method Metal load weight Catalyst C-1 Impregnation Low Catalyst C-2 Impregnation High Catalyst C-3 Co-precipitation Low Catalyst C-4 Co-precipitation High 15

16 Catalyst C-1 Catalyst C-2 Catalyst C-3 Catalyst C-4 Figure 3.5.2: Evaluative Tests of Hydrodesulfurization Reaction Life Catalyst C-2 Catalyst C-4 Figure 3.5.3: Impact of H 2 /Oil Ratio on Hydrogenation Desulfurization Reaction 4. Synopsis 4.1 Design of Autothermal Reforming Process (1) Basic Establishment of Autothermal Reforming Type Hydrogen Production System An autothermal reforming - fuel processing system (ATR-FPS) was set up from supply of raw materials to after selective removal of CO. On the assumption that there is no discharge heat loss, a method was established for calculating reaction conditions so that the system as a whole achieves heat balance. (2) Investigation of Hydrogen Production on 1 Nm 3 /hr Scale Hydrogen production on the 1 Nm 3 /hr scale by the autothermal reforming method was confirmed, as was the impact of reaction temperature on the autothermal reforming reaction. 4.2 Basic Design of Autothermal Reforming - Fuel Processing System The basic design of the autothermal reforming - fuel processing system (ATR-FPS) was completed in consideration of raw material specifications, substance balance and heat balance. 16

17 4.3 Development of Autothermal Reforming Catalyst Autothermal reforming catalyst B, superior to the current autothermal reforming catalyst A in activity and in coking resistance, was discovered. 4.4 Investigation of Hydrocarbon Constituents in Fuel and of Autothermal Reforming Characteristics Comparisons were made of the autothermal reforming characteristics of typical constituents comprised of naphtha to kerosene fraction. It was confirmed that the reforming reactivity is higher, the lower the class of hydrocarbon. In addition, with naphtha fraction, the reforming reactivity of saturated hydrocarbon is the highest, followed in sequence by that of aroma and olefin. 4.5 Development of Catalytic Desulfurization Process The reaction conditions required for hydrodesulfurization were established, and desulfurization catalyst B-2 was developed. This catalyst exhibits high activity such that the sulfur concentration in produced oil is 0.2 massppm or less. Service life evaluation tests confirmed that the durability of catalyst B-2 is 500 hrs. 5. Bibliography 1. D. Dissanayake et al., J. Catal., 132 (1991) A. M. D. Groote et al., Appl. Catal. A, 138 (1996) 245 Copyright 2002 Petroleum Energy Center. All rights reserved. 17