3.2.1 Brunauer–Emmett–Teller (BET) surface area measurement results

The total surface area, the pore volume and the average pore diameter of the support; fresh catalysts were measured using a BELSORP-mini instrument. The samples were degassed at 350 °C for 4 h using N₂ physisorption. The results of BET and porosity tests over the ruthenium particle catalyst are summarized in Table 2. The specific surface area of alumina was found to be 252 m² g⁻¹ while the pore volume was 3.18 cm³ g⁻¹, and the pore diameter was 0.201 nm.

3.2.2 H₂-Temperature Programed Reduction (TPR) results

The catalyst samples were initially purged in a flow of argon at 200 °C for 30 min to remove traces of water and then cooled to 40 °C. The TPR experiments were carried out in a stainless steel reactor loaded with 20 mg of each sample catalyst. Once the thermal conductivity detector (TCD) signal was stable, the gas stream was switched to 5% hydrogen in an argon gas mixture with a flow rate of 30 mL min⁻¹ via a rota-meter. The sample temperature was increased from 40 to 800 °C at a heating rate of 10 °C min⁻¹. Fig. 3 presents the hydrogen profiles resulting from TPR experiments which were performed with intensities in arbitrary units (a.u.). Our 10% Ru/γ-Al₂O₃ catalyst showed one hydrogen reduction with a maximum peak at nearly 200 °C. This evidence might be assigned to the single reduction step from RuO₂ to Ru⁰. Hence, it can be concluded that the presence of low temperature peaks in the TPR profiles of monometallic catalysts did not complicate the reducibility of our catalysts. In addition, similar reduction behavior over γ-Al₂O₃ catalyst-supported catalysts can be observed from Nurunnabi et al. (2007; 2008).

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Figure 3 TPR profile of the 10% Ru/γ-Al₂O₃ catalyst.

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3.2.3 Temperature Programed Surface Reaction (TPSR) results

The TPSR experiment is essentially a temperature programed desorption experiment carried out under reactive conditions using a gas chromatograph (GC) fitted with flame ionization detector (FID) to detect the species evolving as a function of temperature. In the TPSR experiment, samples (20 mg) were reduced in flowing H₂ at 400 °C for 2 h to reduce the surface oxide and remove impurities. Subsequently, the catalyst samples were cooled in flowing argon at room temperature. CO injection was continuous at room temperature for 0.5 h, and then the argon stream was replaced to ensure that physisorption on the surface catalyst was eliminated. Then, the catalyst was swept with pure H₂ for 30 min, the temperature was increased linearly in flowing H₂, and the samples were heated to 850 °C at a rate of 10 °C min⁻¹. The signal of the methanation reaction of CH₄ and CO was recorded using FID as a function of temperature. The profiles resulting from the TPSR experiments for 10% Ru/γ-Al₂O₃ catalysts were investigated; the hydrogenation of the adsorbed CO leads to the formation of single CH₄ peak and can be observed in Fig. 4. Because the hydrogenation of the carbon atom takes place very quickly on an active catalyst surface, the assumption of the production of CH₄ means that the dissociation of CO can be detected by inspecting the formation of CH₄. The maximal CH₄ peak temperature was observed at approximately 180 °C for the 10% Ru/γ-Al₂O₃ catalyst, which can be verified by the FT synthesis starting at 180 °C. However, the result obtained from the TPSR experiment under flowing H₂/CO with a ratio equal to 1 may not be sufficient to claim the perfection of catalyst modification because of the transient experiment. However, it was a good sign for reacting under mild conditions; this observation may be a reason why the Ru-based catalyst prepared using a sol–gel technique showed better catalytic performance (Mirzaei et al., 2009b), which is tested within our catalyst using steady state experiments and the various effects of BTL conditions.

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Figure 4 TPSR profile obtained from hydrogenation of pre-adsorbed CO of the 10% Ru/γ-Al₂O₃ catalyst.

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3.3.1 Small scale fixed bed reactor system

An experimental system used in the first approach was of laboratory scale. The simple criteria for system design were the ease of construction, low cost, ease of sampling and product composition analysis, and the effectiveness of the contact between the catalyst and the gaseous reactants. A Tubular Fixed-Bed Reactor (TFBR) with a length of 300 mm and an internal diameter (ID) of approximately 8 mm that was made from stainless steel (SS316) was designed for this work. A Ru-based catalyst loading of ∼10–15 g was used to maintain the residence time through the bed, which depends on the reactant flowrate used in each experimental run. A diagram of the FT synthesis system is shown in Fig. 5, and photographs of the bench-scale FT synthesis apparatus can be referenced from Narataruksa et al. (2012). The FT reaction experiments were conducted in a TFBR, which had been packed with 10% Ru/γ-Al₂O₃ catalyst particles. The meshed catalyst was held in the middle of the reactor using quartz wool. The TFBR was heated using an electric tube furnace (Model RS485) of 220 V, 1500 W, and equipped with constant thermocouple (K-Type) indicators located inside and outside the catalytic bed. The reactive gases were supplied from three cylinders consisting of hydrogen (H₂), carbon monoxide (CO) and nitrogen (N₂), and the mass flow rates of each component were controlled by three separate mass electronic mass flow controllers (Aalborg model GFC-17), enabling the desired H₂/CO ratio to be obtained. System pressure was regulated by a spring-loaded back pressure regulator located at the bottom of TFBR. Two pressure transducers were also installed at the top and bottom of TFBR to monitor the pressure difference across the bed.

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Figure 5 A schematic diagram of the small FT synthesis’s experimental apparatus.

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After the product gas had been discharged from the TFBR, it was passed through a condensing section to separate the heavy components before entering the gas detector. The liquid hydrocarbon products were collected and the temperature was maintained at approximately −5 °C by silicon oil, and then the non-condensable hydrocarbons were analyzed to determine the compositions using GC. The gas flow rate of the effluent stream was measured using a bubble gas meter.

3.3.2 Pilot scale fixed bed reactor system

In 2010, our research group was supported by NSTDA to construct a pilot plant for FT synthesis. The aim of the second approach was to design and commission a catalytic rig, which can be used to test the reaction and system performance of the pilot-scale fixed-bed reactor. The design of the pilot-scale fixed-bed reactor and supporting systems followed the development of the reactor design, efficient catalyst regeneration methods and the integrated reactive system to achieve the optimum efficiency, i.e., a simple structure, the ability to be fabricated within the country, controllability, easy maintenance and operation, detachability and endurance (Hill, 1977; Froment and Bischoff, 1990; Hewitt et al., 1994; Robinson, 1996; Donati and Paludetto, 1997; Levenspiel, 1999; Fogler, 2000; Davis and Davis, 2003; Bartholomew and Farrauto, 2005). In detail, the design of the pilot system was different from the small-scale system in such a way that the expanded pre-heater section, product separation/product collection section and more precise devices for manipulating process variables were integrated. The pilot system was designed to overcome operability problems and to be operated continuously with the use of system monitoring and safety devices. The diameters of the tubes (D) and particles (D_p) used for the pilot reactor module were in agreement with the value of D/D_p applied in the small-scale reactor units. The geometry of the multi-tube reactor was similar to shell-and-tube heat exchangers. The reactor configuration was a 10-cm external diameter reactor fitted with a three-zone thermo-well inside the catalytic bed. The internal reactor contained 24 tubes that were 90 cm in height, 10 mm in internal diameter and 1 mm in thickness. The reactor tubes are filled with ∼2 kg of the Ru-based catalyst. A three-zone electric heating system was used to provide the thermal energy. Several temperature sensors (K-Type) were mounted along the reactor to measure the axial temperature profiles to ensure isothermal conditions. In addition to the temperature control equipment, a pressure control device was installed after the reactor. All pressures and pressure differences across the bed were monitored and controlled. These installations enabled the successful operation of the reactive system. A schematic of the pilot reactor is shown in Figs. 6 and 7. To monitor and control the pilot system, the commercial software LabVIEW™ was adapted to communicate between an operator with a personal computer and all of monitoring and control equipment. Interface control screens were utilized to adjust and/or monitor the process parameters, such as the temperatures, gas flow rates, pressures, and on/off status of certain valves. Important operating parameters were recorded and displayed on data logging screens. The data history can also be viewed via data logging functions.

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Figure 6 3D schematic view of the FT pilot plant mounted on the movable carriage; (1)-Control panel, (2)-Pre-heater, (3)-FT reactor, (4)-Pre-condenser, (5)-Chiller unit, (6)-Condenser, (7)-Separator, (8)-Cooling Tower and (9)-Gas cylinders.

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Figure 7 Photographs of a FT synthesis pilot scale. (a) Reactor system and apparatus. (b) Control screens interface of pilot reacting system.

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Initially, the FT pilot system was purged by a nitrogen flow for 30 min to ensure oxygen removal and an inert atmosphere in the reactor, followed by the pre-reducing stage with hydrogen at 400 °C for 12 h. The syngas was fed into the pilot system until the desired pressure was reached; afterward, the reaction condition was subsequently adapted to achieve the desired temperature and pressure. Syngas passed through the multi-tubular reactor (R-101) and was converted into hydrocarbon products. The vapor phase composed of mixed products (∼C1–C24) was cooled and condensed using a condenser (C-101) that was maintained at −5 °C. Water as a by-product was roughly separated from the oil phase (C5+) in a flash drum (S-101), while the non-condensable gas (C1–C4) was vacuumed to analyze online using a GC analyzer and a mini-compressor.

3.4.1 Small FT fixed bed scale

In addition to the two major gaseous products (mainly H₂ and CO) obtained through the gasification of biomass, other mixture gases, including CO₂ and CH₄, were produced simultaneously after the wet and dry cleaning process. However, to avoid possible disturbances during the ongoing performance of each activity, the H₂/CO ratios and the reaction temperatures delivered from the biomass gasification were the two key factors that were varied during the FT synthesis to study the carbon monoxide conversion and selectivity of the hydrocarbon products. Additionally, these two parameters were chosen to represent the effects of the reaction conditions on the performance of a Ru-based catalyst.

The total number of experimental runs was 24. The experimental sets for run number 1 through run number 24 were arranged to determine the percent of CO and hydrocarbon conversion, while the reaction pressure and space velocity were kept constant. In the experimental method, a stabilization period with a CO conversion % as a function of reaction time was maintained to ensure that stable catalytic reactions were established. For each operating condition, at least 36 h was necessary to ensure that the steady state behavior of the catalyst did not change after a change in the reaction conditions. Descriptive samples were cumulatively collected during a typical period of 20–36 h. The experimental data obtained in this investigation are listed in Tables 3 and 4.

3.4.2 Pilot FT fixed bed scale

Three experimental runs were conducted in the FT pilot scale system to further clarify the results. These tests also helped research the effects of various process parameters on an appropriate pilot scale. Three different continuous feeds of syngas were passed onto a ruthenium supported alumina bed. The influence of different operating parameters, such as the total pressure up to approximately 4 atm and GHSV (200–1000 h⁻¹), on the behavior of the FT process was studied. The CO conversion and hydrocarbon distribution were measured by maintaining a temperature of 180 °C. Mass balance and carbon balance were calculated based on the inlet/outlet flows, gas composition analysis and liquid product weight and composition.

5.1.1 Effect of reaction temperatures

The effects of the reaction temperatures were studied at temperatures of 160, 180, 200 and 220 °C over the catalytic performance of the 10% Ru/γ-Al₂O₃ catalyst prepared using sol–gel procedures. The molar H₂/CO feed ratios were set at 1/1, 2/1 and 3/1. The reduced catalyst was tested to compare the CO conversion, C1–C4 and C5+ selectivity for 36 h. The CO conversion was found to increase with increasing reaction temperature from 160 to 220 °C during the time-on-stream tests, which can be further seen in Fig. 8 for each H₂/CO ratio.

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Figure 8 Carbon monoxide conversion (%) against time on stream for each H₂/CO ratio; (a) H₂/CO ratio = 3/1, (b) H₂/CO ratio = 2/1 and (c) H₂/CO ratio = 1/1 for the FT synthesis by considering the effect of different temperatures at 160 °C, 180 °C, 200 °C and 220 °C over the 10% Ru/γ-Al₂O₃ catalyst under atmospheric pressure and with space velocity = 1036 h⁻¹.

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When the inlet H₂/CO was maintained at 3/1 (Fig. 8a) and a temperature of 160°C, the corresponding CO conversion was found to be 31.1%, which slightly increased to 41.2%, 42.4% and 43.0% as the temperature was raised to 180, 200 and 220 °C, respectively. In a similar fashion for the inlet H₂/CO of 2/1 (Fig. 8b), the CO conversion trend was not distinguishable for the three different temperatures of 180, 200 and 220 °C with CO conversions of 31.5%, 33.4% and 34.8%, except for the temperature of 160 °C, which showed the lowest of the conversion level of 25.0%. Consequently, at a H₂/CO ratio of 1/1 (Fig. 8c), an increase in the reaction temperature from 180 to 200 °C resulted in an increase in CO conversion from 15.9% to 17.9%, unless the temperature was 160 °C, for which the CO conversion cannot be measured. This result might not be appropriate for any purpose if it is compared to other temperatures. Based on the data above, the temperature of 160 °C showed the lowest conversion level and availability of any hydrocarbon product for a H₂/CO ratio of 1/1. This observation may be possible because the catalyst active site structure of 10% Ru/γ-Al₂O₃ was not amenable to CO hydrogenation and chain growth under these conditions. The performance of catalyst design depends not only on the pretreatment handling but also on the values of the thermodynamic conditions.

A detailed comparison of the hydrocarbon products obtained at four different temperatures is presented in Fig. 9, which illustrated the effect of the reaction temperature on the liquid hydrocarbon distribution (C5+) of three different H₂/CO ratios in units of μmole per gram of catalyst per unit of time (second) against carbon number (n). Fig. 9a contains the highest H₂/CO ratio (3/1) at temperatures of 200 and 220 °C, which favored the C7–C24 distribution as the temperature increased, while the hydrocarbon distribution of C9–C22 was observed at a temperature of 180 °C, and the range of C10-C15 appeared at a temperature of 160 °C. Moreover, the liquid hydrocarbon distribution obtained with H₂/CO ratios of 2/1 and 1/1 (as depicted in Fig. 9b and c) showed different quantities at temperatures of 200 and 220 °C. On the contrary, it did not show such differences between reaction temperatures of 160 °C and 180 °C. When the inlet H₂/CO was carefully set to 1/1, as shown in Fig. 9c, our catalyst showed very high activities at a reaction temperature of 180 °C (as cross-checked from the TPSR analysis) with small amounts of liquid hydrocarbon products. The 10% Ru/γ-Al₂O₃ catalyst was active for the hydrogenation of carbon monoxide, which is related to its high natural activity. It was also the important reaction at low temperatures for the desired hydrocarbon products (Pichler, 1952; Vannice, 1975; Schulz, 1977).

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Figure 9 Liquid hydrocarbon distribution (C5+) against time on stream for each H₂/CO ratio; (a) H₂/CO ratio = 3/1, (b) H₂/CO ratio = 2/1 and (c) H₂/CO ratio = 1/1 for the FT synthesis by considering the effect of different temperatures at 160 °C, 180 °C, 200 °C and 220 °C over the 10% Ru/γ-Al₂O₃ catalyst under atmospheric pressure and with space velocity = 1036 h⁻¹.

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At a constant H₂/CO feed ratio, it is clear that an increase in the temperature of the FT reaction led to an improvement in the CO conversion and yielded greater formation of methane and other light gas components. In agreement with previous results reported by Chernobaev et al. (1997), Schulz (1999), and Escalona et al. (2009), an increase in the reaction temperature would favor a shift to the formation of lower-molecular-weight hydrocarbons on the catalysts. At the same time, an increase in the reaction temperature also led to a shift in the distribution of the liquid hydrocarbon with higher chain lengths. From our experimental data, by increasing the reaction temperature to 220 °C, an increase in higher chain lengths than in the experiment with a reaction temperature of 160 °C can be observed. This observation could be correlated to the product desorption behavior in two hydrocarbon product groups (gaseous and liquid hydrocarbon) on 10% Ru/γ-Al₂O₃. First, it is well known that hydrogen diffuses faster on the catalyst surface than does carbon monoxide (Yu et al., 2002; McDaniel et al., 2007) and carbon atoms in the carbide phase react with hydrogenate and provide more CH, CH₂ and CH₃ species; these species are intermediates in the formation of methane and other light hydrocarbon gases. More unstable carbides for further hydrogenation to CH₄ and light gas hydrocarbons may be released. Therefore, the effects of increasing the reaction temperature could not be explained solely by observing the light hydrocarbon gases. However, the effects of increasing the reaction temperature could be clarified by considering the liquid hydrocarbon products as shown in Fig. 9a–c. Apparently, increasing the temperature causes a shift to products with higher carbon number (C5+). The reason is that the desorption term (chain termination) increases in this case. Moreover, at higher temperatures, the process of heavier hydrocarbon production in this reactive system accelerates.

The above results can be used to estimate the chain growth probability (α). Fig. 10 displays plots of the Anderson–Schulz–Flory (ASF) distribution as a function of reaction temperature. As seen in this figure, the combination of the gaseous (C1–C4) and the liquid hydrocarbon was used to calculate the values of W_n/n. As observed, there was a spread deviation in the ASF distribution curves between the light gas (C1–C4) and the liquid product (C5+) because of the separate GC analyses. According to the shapes of the ASF distribution curves, an H₂/CO ratio of 3/1 was chosen as the inlet ratio value for further experiments. All product distributions can approximately be classified into three ranges: C1–C4, C5–C9 and C10+. To compare on a quantitative basis, the temperature effect on the α value and the slope of the best line for the range of C10+ were fitted to evaluate the value of α with an R² value of almost 95%. The ASF distributions have a slight declining tendency with an increase in the carbon number, while the lower temperature led to more heavy hydrocarbons (C10+) and fewer light ones (C1–C9). Moreover, for the chain growth probability, the value of α decreased from 0.88 to 0.85 when the reaction temperature was increased from 160 to 220 °C. These results are in the agreement with the early results reported by Puskas and Hurlbut (2003), Liu et al. (2007), and Dasgupta and Wiltowski (2011), which improve as the catalytic basicity increases. Higher temperatures tend to favor the formation of light hydrocarbons, while lower temperatures were favorable for the production of heavy hydrocarbons. Therefore, the selectivity to C5+ slightly decreased with increasing temperature.

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Figure 10 Effect of reaction temperature on the ASF distribution of hydrocarbon against carbon number (H₂/CO ratio = 3/1, atmospheric pressure, space velocity = 1036 h⁻¹).

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The selectivity of C1–C4 (light gas hydrocarbons) gradually increased from 0.01% when the reaction temperature was 160 °C to 14.55% as the reaction temperature was increased to 220 °C. However, in agreement with our expectations, the selectivity of C5+ (liquid hydrocarbons) was observed to be opposite to that of the light gas hydrocarbons; namely, it was diminished by approximately 12%. From these trends, it can be concluded that the CO conversion increases with increasing reaction temperature, while the selectivity to C1–C4 (light hydrocarbon) increases, except for the C5+ selectivity, which gradually decreases over the four different temperatures.

5.1.2 The effect of the inlet H₂/CO molar ratios

The H₂/CO ratios were varied from 1/1 to 3/1, while the temperature was fixed for each H₂/CO ratio at 160, 180, 200, and 220 °C. The gas hour space velocity for all these experiments was maintained at a value of 1036 h⁻¹. To easily comprehend and conclusively interpret these effects, the performance of the catalyst CO conversion is shown in Fig. 11. It can be observed that CO conversion against time-on-stream was continuously enhanced by increasing the value of the H₂/CO molar ratio. Fig. 11a illustrates that at the reaction temperature that was imposed, namely, 220 °C, the CO conversion remained constant at approximately 43% for H₂/CO molar ratios as high as 3/1 and then decreased to ∼35% and 18% when the hydrogen partial pressure was diminished. A similar pattern of curves between Fig. 11b and c indicate that a further increase in the H₂/CO ratio resulted in a significant increase in the CO conversion because the conversion decreased dramatically when the H₂/CO ratio fell to unity. Furthermore, a temperature of 160 °C (Fig. 11d) had the lowest conversion of all of the experiments. CO conversions of 31% and 25% were obtained for H₂/CO ratios of 3/1 and 2/1; however, for an H₂/CO ratio of 1/1, the CO conversion could not be measured, which could be correlated with the deficient partial pressures of hydrogen present at the reaction conditions.

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Figure 11 Carbon monoxide conversion (%) against time on stream for each temperature; (a) 160 °C, (b) 180 °C, (c) 200 °C and (d) 220 °C for the FT synthesis by considering the effect of different H₂/CO ratios at H₂/CO ratio = 1/1, H₂/CO ratio = 2/1 and H₂/CO ratio = 3/1 over the 10% Ru/γ-Al₂O₃ catalyst under atmospheric pressure; space velocity = 1036 h⁻¹.

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Compared with the above-mentioned experimental runs and the differences in steady-state CO conversion levels, the effect of the H₂/CO feed ratios on the FT reaction has significant influence on the conversion. Clearly, the conversion of the higher inlet H₂/CO ratio was performed over a 10% Ru/γ-Al₂O₃ catalyst for each reaction temperature, but there is a clear difference in the overall conversion when compared to the lower inlet H₂/CO ratio. The experiments at lower H₂/CO ratios (decreased hydrogen partial pressure) led to lower overall CO conversions over FT catalysts. This finding can be attributed mostly to the relative partial pressures of hydrogen present in the feed stream, which is similar to the previously reported results of Riedel et al. (1999), Tavasoli et al. (2010), and De la Osa et al. (2011), which showed that higher hydrogen partial pressure leads to increase in CO conversions over ruthenium catalyst.

The hydrocarbon distributions at three different H₂/CO ratios are illustrated in Fig. 12. It seems that the product distributions were raised due to the increase in the hydrogen partial pressure. At the same time, a higher H₂/CO ratio in the feed leads to higher-molecular-weight hydrocarbons compared with the deficient ratios. At a H₂/CO ratio equal to 3/1 and temperatures of 220 °C and 200 °C (Fig. 12a and b), C7–C24 groups were mainly obtained. Fig. 12c did not show any difference in the hydrocarbon product distributions for varying inlet H₂/CO ratios, except for the decreasing production rate. Additionally, Fig. 12d shows the lowest product distributions and production rate of all subsequent experiments. These results reflect that the H₂/CO ratio plays a very important role in the FT reaction and directly influences the product selectivity and quantity.

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Figure 12 Liquid hydrocarbon distribution (C5+) against time on stream for each temperature; (a) 160 °C, (b) 180 °C, (c) 200 °C and (d) 220 °C for the FT synthesis by considering the effect of different H₂/CO ratios at H₂/CO ratio = 1/1, H₂/CO ratio = 2/1 and H₂/CO ratio = 3/1 over the 10% Ru/γ-Al₂O₃ catalyst under atmospheric pressure; space velocity = 1036 h⁻¹.

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The detailed product distribution of hydrocarbons under three different H₂/CO ratios is depicted in Fig. 13. According to the shapes of the curves for ASF distribution, a reaction temperature of 220 °C was chosen to study the effect of the H₂/CO ratio. As a result, the value of α decreased from 0.92 to 0.89, which occurred when H₂/CO ratio was increased from 1/1 to 3/1. Our data clearly showed that a higher H₂/CO ratio was preferential for the formation of light hydrocarbons, while a lower H₂/CO ratio was favorable for the production of heavy hydrocarbons. This phenomenon may be attributed to the increased H₂ enrichment (high hydrogen partial pressure) corresponding to the increasing H₂/CO ratio in the feed, thus increasing the probability of the CO reacting on the catalyst surface through the dissociation of C-atoms and participating in the chain propagation with H₂ species. Consequently, the monomeric hydrocarbon reacted with incoming H₂ and a new monomeric free radical species, leading to incorporation of hydrocarbons in the chain polymerization and the chain termination step on the active surface areas of the ruthenium catalyst, which promotes formation of the heavier hydrocarbon product compared with the deficient H₂/CO ratio. At the same time, the monomer hydrocarbon may be reacted with incoming H₂ for alternate ways to produce methane and other chemicals. De la Osa et al. (2011) have reported that an increase in the H₂/CO ratio increases the gaseous and liquid hydrocarbon product. Likewise, with increasing H₂/CO partial pressure ratios, the deviation from a normal ASF distribution decreases (Riedel et al., 1999; Jun et al., 2004; Tavasoli et al., 2010). This conclusion has been confirmed in this study.

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Figure 13 Effect of H₂/CO ratio on the distribution of hydrocarbon against carbon number (Temperature = 220 °C, atmospheric pressure, space velocity = 1036 h⁻¹).

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5.1.3 Statistical analysis

A correlation of technical parameters was assembled into a framework that assists in a comprehensive explanation by taking into account all possible variable relationships implied by the experimental data, including the effects of the operating parameters. Within this correlation, we were interested in studying whether the relationships between two variables have an impact on the response variables of interest. Therefore, it was necessary to have some type of the systematic manner, such as statistical tools, to build predictive models relating these parameters and catalyst response variables (which may be used to collect primary data for reactor simulator development). This idea received much attention in the present study, determining the relationships among these parameters and their interactions over Ru catalyst. These relationships were interpreted using the statistical analysis to increase the confidence and indicate a possible direction for an analytical approach. The related variables and data analysis under different conditions were analyzed using the Minitab statistical software (Version 16.1) from Minitab Corporation, which was used under academic license by the Faculty of Engineering, KMUTNB, Thailand.

Our experimental results were used to evaluate the effect of the parameters on the reaction conditions. The relationship of the inlet H₂/CO ratio (X₁) and the reaction temperature (X₂) was determined to verify their response using surface methodology (RSM). Consequently, the relationships between the variables and the response were fitted and statistically derived using the 2nd-order polynomial equation model for each response. This model can be expressed with coded variables (X₁, X₂) as in (Eq. (5)).

The results of the fitted models are shown in Table 6, where Y₁ is the predicted CO conversion (%), Y₂ is the predicted C1–C4 selectivity (%), and Y₃ is the predicted C5+ selectivity (%). The coefficients of determination (R²) for Y₁, Y₂ and Y₃ were 0.9855, 0.9569 and 0.9358, respectively. Fig. 14 shows the predicted response function and the effects of the independent variables (X₁ and X₂) on the dependent variables (CO conversion, C1–C4 selectivity and C5+ selectivity). The effects of the independent variables on Y₁ are shown in Fig. 14a. An increase in the CO conversion was observed with an increase in the H₂/CO ratio and reaction temperature. Similarly, in Fig. 14b, moving up the curve of C1–C4 selectivity indicates a significant increase between two operating parameters. In contrast, increasing the inlet H₂/CO ratio and reaction temperature tends to decrease the C5+ selectivity (Fig. 14c). These results are consistent with the experimental results and can be explained by the fact that the higher reaction temperature can lead to more hydrogenation of carbon atoms. The increase in the H₂/CO molar ratio led to an enhancement in the CO conversion and FT synthesis rate, thereby producing more light gas and liquid hydrocarbon products.

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Figure 14 Response surface plots showing the interactions among inlet H₂/CO ratio and reaction temperature (°C). (a) CO conversion (%), (b) C1–C4 selectivity (%) and (C) C5+ selectivity (%) over 10% Ru/γ-Al₂O₃.

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5.1.4 Kinetic study

For this study, the Ru-based catalyst was in the form of fine particles; thus, intra-particle diffusion could be ignored, and we could prove this from our experimental results and the plot of GHSV against conversion. These results were not presented in this article. Once GHSV increased to a value between 3000 and 5000 h⁻¹ by varying the catalyst weight, the CO conversion was at almost the same level. The available evidence indicated that external diffusion did not show a significant effect using GHSV. Therefore, our kinetic experiments were conducted free from internal and external mass transfer limitations. Two criteria must be subsequently considered: first, fresh Ru catalyst was used in each experiment to avoid the problem of catalyst deactivation. Second, the catalyst was diluted with an inert material (silicon carbide) to maintain a uniform catalytic bed temperature (isothermal conditions), and the axial temperature distribution was secured using the criterion of Mears (1974) and Mollavali et al. (2008), which should be L/d_p > 50.

The kinetic studies were initially calculated considering the fixed bed as an ideal tubular flow reactor or a plug flow reactor. The experiments were conducted using a mixed gas consisting of H₂, CO and balanced N₂ over a temperature range of 160–220 °C, a feed inlet H₂/CO molar ratio of 1/1–3/1, a total pressure of 1 atm and GHSV of 3600 h⁻¹. The mathematic assumption stated that because the ideal tubular flow reactor was also a steady-state flow device, there was no accumulation term in Eq. (6), and thus, the material balance for a cross-sectional slice of reactor simplifies to

An elementary reaction set on sites for each model is presented in Table 8. Then, all of the models obtained were fitted separately using a multi-variable non-linear regression method to minimize the sum of the least squares for each kinetic model. Afterward, the constants of k_CO and b_CO obtained from the LHHW kinetics, in their logarithmic form, ln k_p and In b_CO were plotted versus 1000/T at different temperatures using the Arrhenius and adsorption equations. As shown in Fig. 15, good straight lines could be drawn. From the slopes of the corresponding lines, the apparent activation energies (E_a) and the enthalpy (ΔH_CO) of adsorption of carbon monoxide were determined and are reported in Table 9.

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Figure 15 Graphical methods (a) Arrhenius plots ln k_CO = f(1000/T) for the estimation of the apparent activation energy and (b) Adsorption plots ln b_CO = f(1000/T) for the estimation of the heat of CO adsorption with respect to two promising models within four different temperatures under P = 1 atm, H₂/CO = 1/1–3/1 and space velocity = 3600 h⁻¹.

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As shown in Table 10, the values of the activation energy of carbon monoxide consumption estimated from the results obtained using the LHHW in Eqs. (11) and (12) were 28.99 kJ/mol and 21.70 kJ/mol, respectively, while the enthalpy of CO adsorption was found to be −12.20 and −17.32 kJ/mol for these equations. The low values of activation energy of carbon dioxide, obtained in the present study, could be attributed to the reaction of CO with Ru/Al₂O₃, which has been characterized as a fast reaction, due to the strong affinity of the syngas reactant with the oxides (van Santen et al., 2011).