Effects of Carbon Formation on Catalytic Performance for CO2

Aug 26, 2013 - Fax: + 82-31-290-7272. ... Magnetic field acceleration of CO 2 reforming of methane over novel hierarchical Co/MgO catalyst in fluidize...
0 downloads 0 Views 3MB Size
Download PDF
Article pubs.acs.org/IECR

Effects of Carbon Formation on Catalytic Performance for CO2 Reforming with Methane on Ni/Al2O3 Catalyst: Comparison of FixedBed with Fluidized-Bed Reactors Young Kyu Han, Chang-Il Ahn, Jong-Wook Bae,* A Rong Kim, and Gui Young Han* School of Chemical Engineering, Sungkyunkwan University (SKKU), Suwon, Kyonggi-do 440-746, Republic of Korea S Supporting Information *

ABSTRACT: The amount of carbon formed and the H2/CO molar ratio for the carbon dioxide reforming (CDR) reaction with methane were investigated on Ni/Al2O3 catalyst using a laboratory-scale fixed-bed reactor and a bench-scale fluidized-bed reactor. A significant suppression of carbon deposition in the fluidized-bed reactor compared with the fixed-bed reactor can be mainly induced from different product gas flow patterns by the continuous circulation of catalysts in oxidizing and reducing regions. This approach also enhances the gasification rate of deposited carbon in an expanded catalyst bed by increasing the amount of water adsorbed. The higher H2/CO ratio above 1.0 in the fluidized-bed reactor is also attributed to the enhanced gasification rate of deposited carbon precursors. The differences in the conversions of CH4 and CO2 and the H2/CO ratios in the two reactors are responsible for the different competitive rates of the reverse water−gas shift (RWGS) reaction, the Boudouard reaction, and the gasification of carbon precursors.

1. INTRODUCTION The production of synthesis gas (syngas) through various reforming processes such as steam reforming, carbon dioxide reforming, and autothermal reforming of hydrocarbons has been largely investigated because of increased demands for the production of hydrogen fuel and clean synthetic fuels.1,2 The steam reforming of hydrocarbons has been commercially established to produce hydrogen-rich syngas that can be utilized for hydrogen fuel through fuel processing technologies or for chemical syntheses such as those of ammonia and methanol.1 However, carbon dioxide reforming (CDR) with methane, which is also called dry reforming of CH4, on nickel-, noble-metal-, and mixed-oxide-based catalysts is attractive because it effectively transforms the greenhouse gas (GHG) CO2 into useful chemicals.3−5 Use of the CDR reaction for syngas production has been largely reported in recent years to produce useful chemicals and clean fuels such as gasoline, methanol, dimethyl ether, and Fischer−Tropsch hydrocarbons. In addition, economical methods of using low-grade small-scale stranded gases containing large amounts of CO2 have been largely investigated so that these methods can be used properly.1 However, industrial implementation of the CDR reaction has been restricted because of significant coke formation and a strongly endothermic nature. The main reactions related to the CDR reaction are categorized as the dry reforming reaction of CH4 and CO2 (eq 1); the simultaneous reverse water−gas shift (RWGS) reaction (eq 2), which is responsible for a H2/CO ratio in the product gas of less than unity for the CDR reaction; and the Boudouard reaction (eq 3) for coke formation as follows CDR reaction

RWGS reaction CO2 + H 2 ↔ H 2O + CO

(2)

Boudouard reaction 2CO ↔ CO2 + C

© 2013 American Chemical Society

ΔH 0(298 K) = −171 kJ/mol

(3)

The general reforming catalysts can be categorized into two groups, namely, noble-metal-based catalysts such as Pt, Rh, and Ru on porous supports, and nickel-based supported catalysts on Al2O3, MgO, TiO2, CeO2, ZrO2, or La2O3 metal oxides with incorporation of an appropriate promoter that can enhance the dispersion of the active metal and/or suppress the formation of coke.1,6−10 For example, hydrotalcite-like MgAl2O4 supports have been abundantly investigated to design proper catalysts that suppress catalytic deactivation with the help of their basic nature. These basic supports are beneficial in suppressing coke formation by decreasing acidic sites on the surface of reforming catalysts.11−13 The types of reforming catalysts and deactivation mechanisms are similar for steam reforming and CO2 reforming with methane, and the coke formation rates through filamentous carbon deposition on the active sites are generally known to be a crucial factor in catalyst deactivation for various reforming reactions.1 In the present investigation, the simplest prototype Ni/Al2O3 catalyst, which has a high mechanical strength and a high surface area of the Al2O3 support, was studied to elucidate the different characteristics of filamentous carbon formation on Ni/ Al2O3 catalyst in a laboratory-scale fixed-bed reactor and a Received: Revised: Accepted: Published:

CH4 + CO2 ↔ 2H 2 + 2CO ΔH 0(298 K) = 247 kJ/mol

ΔH 0(298 K) = 41 kJ/mol

(1) 13288

October 19, 2012 August 21, 2013 August 26, 2013 August 26, 2013 dx.doi.org/10.1021/ie3028558 | Ind. Eng. Chem. Res. 2013, 52, 13288−13296

Industrial & Engineering Chemistry Research

Article

bench-scale fluidized-bed reactor by eliminating the effects of complicated contributions from structural and chemical promoters on the modified reforming catalysts, which is the general approach of methods for developing highly stable reforming catalysts by improving the lower sintering properties of nickel particles with higher resistance to nickel aluminate formation.6−10 A fluidized-bed reactor has the advantages of easy scaleup because of the facile control and even temperature distribution in the catalyst bed with significant coke reduction during the CDR reaction. However, the different conversions of CH4 and CO2 and the H2/CO molar ratio in two different reactors have not been sufficiently exploited by considering the main side reactions, such as the RWGS reaction, Boudouard reaction, and gasification of deposited carbons with water generated simultaneously by the RWGS reaction. Therefore, the coke formation properties and these side reactions on Ni/ Al2O3 catalysts during the CDR reaction were investigated to explain the different H2/CO molar ratios in the product gas and the conversions of CH4 and CO2 in two different reactors by means of X-ray diffraction (XRD), scanning electron microscopy (SEM), Raman spectroscopy, and thermogravimetric analysis (TGA).

2. EXPERIMENTAL SECTION 2.1. Catalyst Preparation and Activity Test. The prototype Ni/Al2O3 catalyst was prepared by the impregnation method using γ-Al2O3 support having a specific surface area of 161.2 m2/g (supplied by Aldrich) with the nickel precursor Ni(NO3)2·6H2O (98%, Junsei Chemical Co., Ltd.) in deionized water. The CDR catalysts were dried overnight at 110 °C and subsequently calcined at 500 °C for 3 h at a heating rate of 10 °C/min. The weight ratio of nickel was varied from 10 to 20 wt % on the Al2O3 support, and the catalysts are denoted as NiA(x), where NiA represents the Ni/Al2O3 catalyst and x is the weight percentage of metallic nickel on the Al2O3 support. The performances of the catalysts were tested separately in a laboratory-scale fixed-bed reactor and a bench-scale fluidizedbed reactor. For the CDR reaction in the laboratory-scale fixed-bed reactor, Ni/Al2O3 catalysts were tested in a fixed-bed quartz reactor with an outer diameter of 8 mm and a height of 500 mm, as shown in Figure 1A. Before each activity test, the CDR catalyst was reduced at 500 °C for 1 h under a flow of 5 vol % H2 balanced with N2. The reaction conditions for the CDR reaction in the fixed-bed reactor were as follows: catalyst loading = 0.3 g, temperature = 550−850 °C, pressure = 0.1 MPa, residence time = 1.748 s (based on the total amount of reactants CH4 and CO2), feed-gas CH4/CO2/N2 molar ratio = 1/1/1, and space velocity = 4000 mL (total gas)/(gcat·h). The reaction temperature was monitored with a thermocouple placed in the middle of the catalyst bed. For the CDR reaction in the bench-scale fluidized-bed reactor, Ni/Al2O3 catalyst was tested in a fluidized-bed quartz reactor with an outer diameter of 55 mm and a height of 1000 mm, as shown in Figure 1B. The selected representative catalyst NiA(15) had a bulk density of 1020 kg/m3 with a particle size distribution of 50−120 μm, and the activity was measured at a minimum fluidization velocity (Umf) of ∼1.14 cm/s. Before each activity test, the CDR catalyst was reduced at 500 °C for 2 h under a flow of 5 vol % H2 balanced with nitrogen. The CDR reaction conditions in the fluidized-bed reactor were as follows: catalyst loading = 100 g, temperature = 550−850 °C, pressure = 0.1 MPa, residence time = 1.809 s at 2Umf (based on the total

Figure 1. Schematic diagrams of the reaction apparatus: (A) laboratory-scale fixed-bed reactor, (B) bench-scale fluidized-bed reactor.

amount of reactants CH4 and CO2), and feed-gas CH4/CO2/ N2 molar ratio = 1/1/1. All product gases were analyzed by an online gas chromatograph (Young Lin Acme 6100) equipped with a thermal conductivity detector (TCD) connected to a Carbosphere packed column. 2.2. Catalyst Characterization. The surface areas, pore volumes, and average pore diameters of the NiA catalysts were 13289

dx.doi.org/10.1021/ie3028558 | Ind. Eng. Chem. Res. 2013, 52, 13288−13296

Industrial & Engineering Chemistry Research

Article

determined by the N2 sorption method at −196 °C using a constant-volume adsorption apparatus (Micromeritics, ASAP2400). The pore volumes were determined at a relative pressure (P/P0) of 0.99. The pore size distributions of the NiA catalysts was calculated using the Barrett−Joyner−Halenda (BJH) model from the data of the desorption branch of the N2 isotherm. The calcined NiA catalysts were degassed at 300 °C in a He flow for 2 h before measurement to remove any contaminant and physisorbed water. To identify crystalline phases of nickel particles and the average sizes of Ni and NiO, powder X-ray diffraction (XRD) patterns of the fresh (prereduced at 500 °C for 1 h) and used NiA catalysts were obtained using a Bruker AXS D8 X-ray diffractometer equipped with Cu Kα radiation operating at 40 kV and 40 mA at a scanning rate of 5°/min from 20° to 80°. The particle sizes of nickel species were calculated by using the values of the full width at half-maximum (fwhm) with the Scherrer equation at the diffraction peaks of 2θ = 43° for NiO and 44.5° for Ni particles. Identification of the different crystalline phases in the fresh and used NiA catalysts was performed by comparing collected spectra with the Joint Committee on Powder Diffraction Standards (JCPDS) files for Ni, NiO, and Al2O3 species. The morphologies of the filamentous carbon after the CDR reaction for 4 h at 650 °C in the bench-scale fluidized-bed reactor were characterized by scanning electron microscopy (SEM; JEOL, JSM7000F). Thermogravimetric analysis (TGA) on the used NiA catalysts (after 4 h on stream) was carried out using a DMA instrument. The used NiA catalyst (30 mg) was heated from 50 to 1000 °C at a rate of 10 °C/min under an air flow of 30 mL/min. Raman spectroscopy analysis was also carried out under ambient conditions using a Bruker FRA106/S spectrometer with a Nd:YAG excitation laser (wavelength of 1064 nm) and an optical power of 300 mW at the sample position.

decreased from 0.23 to 0.17 cm3/g. The unimodal pore size distribution on the NiA catalysts, shown in Figure 2, suggests

Figure 2. Pore size distribution of γ-Al2O3 support and NiA catalysts.

the uniform distribution of NiO particles inside the Al2O3 pores. In addition, a slight increase in average pore diameter in the range of 6.5−6.7 nm on the NiA catalysts was observed with increasing nickel content, compared to the value of 5.85 nm for the Al2O3 support alone, and this is possibly due to the blockage of small pores by the deposition of NiO particles on the Al2O3 surface.3,14 This conclusion is also supported by the observation that the peak intensity of pores at ∼3.5 nm decreased significantly upon the addition of nickel species to the γ-Al2O3 support (Figure 2). As summarized in Table 1, the average particle sizes of NiO on the reduced fresh NiA catalysts were found to be in the range of 4.4−8.1 nm by XRD analysis, and their sizes increased with increasing nickel content on the γ-Al2O3 support because of the increased aggregation of nickel particles at higher concentrations of nickel interacting with the tetrahedral and octahedral vacant sites of γ-Al2O3, as reported previously for the impregnation step.10,15 Furthermore, well-defined NiO particles of the NiO(111) and NiO(200) phases and Al2O3 in the form of the γ phase assigned to the peaks at 2θ = 43° and 67°, respectively, were also observed on the reduced NiA catalysts, as shown in Figure 3a−c. The intensities of the diffraction peaks of NiO on the reduced NiA catalysts also increased significantly for catalyst NiA(20), which is responsible for the larger NiO particle formation on this catalyst and is in line with the calculated particle size of NiO species from the XRD analysis. 3.2. Catalytic Activities in Two Reactors at Different Temperatures. To clearly differentiate the effects of the two different type of reactors on the characteristics of filamentous carbon formation, catalyst NiA(15), which has a medium surface area of 115.4 m2/g and an average pore diameter of 6.59 nm, was selected to elucidate the effects of temperature from 550 to 850 °C. Even though catalyst deactivation is significantly affected by the initial metal particle size, sintering degree, and metal−support interaction for Ni-based reforming catalysts,6−10 the intrinsic effects of the reactor type on coke formation were verified by using a simply prepared prototype Ni/Al2O3 catalyst. The CDR reaction was carried out at a pressure of 0.1 MPa and a CH4/CO2 molar ratio of 1.0 for 4 h at each temperature in the fixed-bed and fluidized-bed reactors. As

3. RESULTS AND DISCUSSION 3.1. Physicochemical Properties of Ni/Al2O3 Catalysts. The physical properties of NiA catalysts such as surface area, pore volume, and average pore diameter are summarized in Table 1. With an increase in the Ni content on Al2O3, the surface area gradually decreased from 161 to 102 m2/(g of γAl2O3 support) for catalyst NiA(20), and the pore volume also Table 1. Characteristics of Ni/Al2O3 Catalysts for the CDR Reaction particle sizea (nm)

N2 sorption method

notation

Ni content (wt%)

surface area (m2/g)

pore volume (cm3/g)

average pore size (nm)

NiO (before)

Ni (after)

NiA(10) NiA(15) NiA(20) Al2O3

10 15 20 −

123 115 102 161

0.20 0.19 0.17 0.23

6.7 6.6 6.5 5.8

4.4 6.6 8.1 −

14.9 25.7 55.0 −

b

a

Particle sizes of NiO and Ni calculated from fwhm values at the diffraction peaks of 2θ = 43° for NiO for the reduced fresh catalysts and 2θ = 44.5° for metallic nickel after the CDR reaction for 4 h in a fluidized-bed reactor. bNi/Al2O3 catalysts for the CDR reaction denoted as NiA(x), where Ni represents nickel, A represents Al2O3, and x is the weight percentage of metallic nickel on the γ-Al2O3 support. 13290

dx.doi.org/10.1021/ie3028558 | Ind. Eng. Chem. Res. 2013, 52, 13288−13296

Industrial & Engineering Chemistry Research

Article

competitive adsorption, with a lower reaction rate for CO2 compared to CH4.16,17 These different catalytic properties were found to be more significant in the fluidized-bed reactor because of the high heat- and mass-transfer rates of the product gases through the circulation of catalyst particles by backmixing in the expanded catalyst bed.1,18,19 For reforming reactions, use of a fluidized-bed reactor is generally known to be beneficial because of the properties of fluidization with a suitable reducibility and a lower coke formation because the fluidized catalyst particles are continuously circulated between the oxidizing and reducing regions with facile carbon gasification with water generated by the RWGS reaction, which results in a lower deactivation.18,19 This phenomenon was also supported by the results for the catalytic performances of the NiA catalysts in the laboratory-scale fixedbed and bench-scale fluidized-bed reactors. In general, a H2/ CO molar ratio of less than 1.0, which is the stoichiometric molar ratio of the intrinsic CDR reaction, could be obtained because of the possible simultaneous RWGS reaction.17 Lower H2/CO ratios in the range of 0.91−0.98 at all tested temperatures were observed in the fixed-bed reactor, as reported in Table 2; however, somewhat higher H2/CO ratios above 1.10 were observed in the fluidized-bed reactor, except at the lowest temperature of 550 °C. The H2/CO ratios were also found to be much lower than the equilibrium values, especially at lower temperatures. The ratios decreased from 1.76 to 1.07 as the reaction temperature increased from 550 to 850 °C in the fluidized-bed reactor. The higher H2/CO ratios are responsible for the suppressed CO2 conversion resulting from competitive adsorption with the product gases, especially with H2O in the expanded circulating bed in the fluidized-bed reactor. To verify the effects of water generated by the RWGS reaction in the fluidized-bed reactor for the suppressed carbon formation and lower CO2 conversion, additional CDR experiments were carried out by adding 3 vol % water based on the CH4 molar flow rate during the CDR reaction, and the results are summarized in Table 2. As shown in Table 2, the CO2 conversion decreased to 86.1% from 94.5% at 850 °C upon addition of a small quantity of water as a result of the possible competitive adsorption of water on the NiA(15) surface, and no significant coke formation was observed upon the addition of water during the CDR reaction. The H2/CO molar ratios also increased above 1.0 because of the suppressed CO2 conversion. This observation also supports the possible gasification of coke precursors with water generated by suppressing CO2 conversion through the competitive adsorption of water without a significant variation in the CH4 conversions from the equilibrium conversion values. Because the conversions of CH4 and CO2 approached the equilibrium values at 850 °C and the extents of coke formation at high temperatures were similar irrespective of variations in the reaction conditions, as shown in Supplementary Figure 1 of the Supporting Information, the reaction temperature of 650 °C was selected to properly compare the activities of filamentous carbon formation in the fixed-bed and fluidized-bed reactors, even though the equilibrium conversions were somewhat lower at that temperature. In addition, the temperature of 650 °C was found to provide a medium degree of nickel aggregation and appropriate CH4 and CO2 conversions with a larger difference in H2/CO ratios in the two different reactors. The different conversions with different H2/CO molar ratios in the two reactors can be explained quantitatively by considering the rates

Figure 3. XRD patterns of NiA catalysts that had been (a−c) reduced and (d−f) used for 4 h of reaction in a fluidized-bed reactor: (a) NiA(10) before, (b) NiA(15) before, (c) NiA(20) before, (d) NiA(10) after, (e) NiA(15) after, (f) NiA(20) after.

shown in Table 2, the CH4 and CO2 conversions approached the equilibrium conversions at the highest temperature of 850 Table 2. Catalytic Performances on Catalyst NiA(15) at Different Reaction Temperatures in Fixed-Bed and Fluidized-Bed Reactorsa equilibrium values

fixed-bed reactor

fluidized-bed reactor

temp (°C)

conversion (CH4/ CO2)

H2/ CO molar ratio

conversion (CH4/ CO2)

H2/ CO molar ratio

conversion (CH4/ CO2)

H2/ CO molar ratio

550 650 750 850 850b

75.7/67.5 88.0/80.6 96.6/91.2 99.2/96.1 99.6/96.2

1.76 1.31 1.14 1.07 1.11

41.5/30.9 83.0/72.0 95.2/92.6 98.1/94.5 96.7/86.1

0.91 0.98 0.93 0.94 1.07

25.8/18.8 65.2/57.3 79.7/74.9 97.5/87.9 −

0.80 1.15 1.10 1.12 −

a

Catalytic performances measured at P = 0.1 MPa and CH4/CO2 molar ratio = 1.0 for 4 h at each temperature. Catalyst NiA(15) was tested using the fixed-bed reactor with 0.3 g of catalyst at a space velocity of 4000 mL (CH4)/(gcat·h) and using the fluidized-bed reactor with 100 g of catalyst at 2Umf (i.e., twice the minimum fluidization velocity). bCDR reaction carried out in the fixed-bed reactor at the same reaction conditions but with the addition of a small quantity of H2O (H2O/CH4 molar ratio = 0.03) to verify the suppressed CO2 conversion and coke formation.

°C (where the equilibrium conversions were calculated using HSC Chemistry by minimizing the Gibbs free energy), and the differences between equilibrium and experimental conversions became significant as the reaction temperature decreased. Interestingly, the catalytic activities were found to be somewhat higher in the fixed-bed reactor than in the fluidized-bed reactor at all tested temperatures. The observed lower CO2 conversions in the fluidized-bed reactor compared to the fixed-bed reactor, especially at lower temperatures, can possibly be attributed to the different gas flow patterns in the two reactors. For example, the CO2 conversion at 650 °C was found to be around 80.6% at the equilibrium point, and different CO2 conversions of 72.0% and 57.3% were observed in the fixed-bed and fluidized-bed reactors, respectively. The product gases of CO, H2, and H2O formed by the RWGS reaction can inhibit the adsorption properties of CO 2 molecules on active sites through 13291

dx.doi.org/10.1021/ie3028558 | Ind. Eng. Chem. Res. 2013, 52, 13288−13296

Industrial & Engineering Chemistry Research

13292

a

Catalytic performances of and coke formation on NiA catalysts measured at T = 650 °C, P = 0.1 MPa, and CH4/CO2 molar ratio = 1.0 for a reaction duration of 4 h. bReaction rates of CH4 and CO2 calculated using the flow rate of each reactant in units of number of moles of CH4 (or CO2) reacted at standard state per unit weight of catalyst per second [mol/(gcat·s)].

6.65/6.77 7.89/6.93 7.86/6.99 90 136 137 1.02 1.15 1.17 55.0/56.0 65.2/57.3 65.0/57.8 9.59/8.69 10.39/9.01 9.89/8.66 204 339 355 76.6/69.4 83.0/72.0 79.0/69.2 NiA(10) NiA(15) NiA(20)

0.88 0.98 0.93

coke (mg/gcat) conversion (CH4/CO2) reaction rate (CH4/CO2)b [mol/(gcat·s) × 106] coke (mg/gcat) H2/CO molar ratio conversion (CH4/CO2) notation

fixed-bed reactor

Table 3. Catalytic Performances and Extents of Coke Formation in Fixed-Bed and Fluidized-Bed Reactorsa

H2/CO molar ratio

fluidized-bed reactor

of the RWGS reaction, the Boudouard reaction, and the gasification of deposited carbons with water generated by the RWGS reaction, as explained below. 3.3. Characteristics of Carbon Formation in Fixed-Bed and Fluidized-Bed Reactors. The formation of filamentous carbon on the NiA catalysts was investigated at 650 °C for 4 h in the two reactors, and the results are summarized in Table 3. The catalytic performance and stability for 4 h in the fixed-bed and fluidized-bed reactors are also displayed in Figure 4. Even though stable catalytic activities were observed in both reactors for this short reaction period, a very small decrease in CO2 conversion was observed on catalyst NiA(15) in the fluidizedbed reactor. To verify the catalyst stability in the fixed-bed reactor, the activity variation for a longer time of around 48 h of catalyst NiA(15) at 750 °C was also tested, as shown in Supplementary Figure 2 (Supporting Information). Slow catalyst deactivation was observed on catalyst NiA(15) at 750 °C, which suggests that the initial deactivation rate was more significant than indicated by comparing the extent of coke formation as shown in Table 3. Therefore, the short reaction time of 4 h can also represent deactivation in the two different reactors. To clearly compare the amounts of carbon deposited in the two different reactors, CDR reactions were also carried out at similar reaction rates, as summarized in Table 3. The residence times, based on the total amount of reactants of CH4 and CO2, were selected to be similar at around 1.748 and 1.809 s in the fixed-bed and fluidized-bed reactors, respectively. With respect to nickel contents of 10−20 wt % on alumina, the reaction rates [defined as moles of CH4 (or CO2) reacted per gram of catalyst per second, times 106] in the fixed-bed reactor were found to be in the ranges of 9.58−10.39 for CH4 and 8.66−9.01 for CO2, and they were observed to be in the ranges of 6.65−7.86 for CH4 and 6.77−6.99 for CO2 in the fluidizedbed reactor without significant deviation. Even though the observed slightly lower reaction rates in the fluidized-bed reactor could be responsible for the expanded catalyst bed with a uniform temperature profile and this could also be responsible for the lower coke formation as shown in Table 3, the significantly different amounts of filamentous carbon formed are mainly attributed to the different types of reactors instead of the different extents of conversion of CH4 and CO2 during the CDR reaction. The CH4 and CO2 conversions and H2/CO molar ratio in the fixed-bed reactor increased with increasing Ni content; however, decreased conversions and H2/CO molar ratio on catalyst NiA(20) were observed with values of 79.0%, 69.2%, and 0.93, respectively, compared to corresponding values for catalyst NiA(15) of 83.0%, 72.0%, and 0.98, respectively. This result is mainly attributed to the severe formation of filamentous carbon at the very beginning of CDR reaction. The amount of deposited coke also increased continuously with increasing nickel content in the fixed-bed reactor, from 204 to 355 mg/gcat. The CH4 and CO2 conversions and H2/CO molar ratio in the fluidized-bed reactor also increased with increasing nickel content; however, the values stabilized on the catalysts with at least 15 wt % Ni on Al2O3, at 65.0%, 57.8%, and 1.17, respectively. Interestingly, the amount of deposited coke was not significantly greater on catalyst NiA(20) compared to catalyst NiA(15) in the fluidized-bed reactor. Less coke formed in the fluidized-bed reactor than in the fixed-bed reactor, and the values were found to be in the range of 90−137 mg/gcat. With respect to the general advantage of fluidized-bed reactors for various reforming reactions,18,19 it can be attributed to the

reaction rate (CH4/CO2)b [mol/(gcat·s) × 106]

Article

dx.doi.org/10.1021/ie3028558 | Ind. Eng. Chem. Res. 2013, 52, 13288−13296

Industrial & Engineering Chemistry Research

Article

action,1,2 which was the intrinsic reason for the formation of larger nickel particles on catalyst NiA(20). Finally, catalyst NiA(15) was selected as an optimum catalyst to be applied in the fluidized-bed reactor for the CDR reaction for robust operation because of the medium degree of nickel aggregation and the lower coke formation rate. Further morphological characterizations of the deposited cokes and the amounts of coke were carried out by using SEM, TGA, and Raman spectroscopy. The particle size of the reduced fresh catalyst NiA(15) was found to be in the range of 50−120 μm, as shown in Figure 5A. In addition, as shown in Figure 5B, the deposited carbons were mainly in the form of filamentous carbons irrespective of the type of reactor, and their morphologies were not significantly altered. In general, the formation of filamentous carbon was more significant on the larger metal particles because of the weak metal−support interaction,20 which is in line with the present results for NiA

Figure 4. Catalytic performances with time on stream (h) of catalyst NiA(15) at T = 650 °C, P = 0.1 MPa, CH4/CO2 molar ratio = 1.0 for a reaction duration of 4 h: (A) laboratory-scale fixed-bed reactor, (B) bench-scale fluidized-bed reactor.

continuous circulation of reforming catalyst powders in oxidizing and reducing regions and the easy gasification of deposited carbons with the water generated by the RWGS reaction. Furthermore, H2/CO molar ratios in the fluidized-bed reactor, which were in the range of 1.02−1.17, were much higher than those in the fixed-bed reactor. This finding can be attributed to the decreased CO2 conversion because of the competitive adsorption properties of product gases on the catalyst surfaces in the expanded circulating-bed region. Severe aggregation of the nickel particles after 4 h of the CDR reaction at 650 °C in the fluidized-bed reactor was observed on NiA catalysts. This was verified by XRD analyses, as shown in the diffraction peaks of Ni particles with the Ni(111), Ni(200), and Ni(220) phases in Figure 3d−f, and supported by the observation of a larger intensity of the peak for metallic nickel at 2θ = 44.5° and the abundant presence of coke precursors assigned to 2θ = 26.2° on the catalyst NiA(20) surface. The calculated particle sizes of metallic nickel after the CDR reaction in the fluidized-bed reactor are summarized in Table 1. The particle sizes of NiO on reduced NiA catalysts were significantly increased after the CDR reaction because of the easy aggregation at high temperature, and their sizes were found to be in the range of 14.9−55.0 nm for metallic nickel. Increased aggregation of nickel particles after the CDR reaction was also observed with an increase in the Ni content on the Al2O3 support because of the weak metal−support inter-

Figure 5. SEM images of catalyst NiA(15) before and after CDR reaction for 4 h at T = 650 °C, CH4/CO2 = 1: (A) prereduced fresh catalyst NiA(15), (B-1) used catalyst NiA(15) in a fixed-bed reactor, and (B-2) used catalyst NiA(15) in a fluidized-bed reactor. 13293

dx.doi.org/10.1021/ie3028558 | Ind. Eng. Chem. Res. 2013, 52, 13288−13296

Industrial & Engineering Chemistry Research

Article

catalysts. The amounts of carbon formed on the NiA catalysts were further characterized by TGA, and the weight losses on the used NiA catalysts in the fluidized-bed reactor operated at 650 °C are displayed in Figure 6. All used NiA catalysts showed

Figure 7. Raman spectroscopy of the used NiA catalysts obtained from a fixed-bed reactor [NiA(x)-fixed] and from a fluidized-bed reactor [NiA(x)-fluidized].

coke formation, supported by the observed peak of G band at 1580 cm−1 with a lower ID/IG ratio. Smaller peak intensities were observed on the used NiA catalysts in the fluidized-bed reactor, and the larger ID/IG ratios are responsible for low coke formation and for the lower formation of ordered filamentous carbon. The amounts and types of coke in the two different reactors could be induced from the different extents of the Boudouard and gasification reactions with water generated by the RWGS reaction for the removal of deposited carbon precursors. The different coke species can generally be categorized as polymeric, filamentous, and graphitic carbon23 during a reforming reaction, and the degree of removal activity by gasification becomes difficult with increasing extent of graphitization. The facile removal capacities of polymeric carbon species (assigned D band with a Raman shift of around 1330 cm−1 in Figure 7) by gasification with water generated by the RWGS reaction are responsible for the decreased carbon formation in the fluidized-bed reactor by retaining the filamentous carbon species on the Ni/Al2O3 catalyst surface. 3.4. Effects of Side Reactions during CDR to Coke Formation and CO2 Conversion. Based on the experimental results, we propose possible mechanisms to explain the different activities of filamentous carbon formation at different conversions of CH4 and CO2 and different H2/CO molar ratios below the stoichiometric ratio of 1.0, which corresponds to the ratio of an ideal CDR reaction, according to the type of reactor. In general, H2/CO molar ratios could be obtained below 1.0 because of the concomitant RWGS reaction consuming CO2 and H2. However, as summarized in Figure 8 (at reaction conditions of T = 650 °C and CH4/CO2 = 1.0), a higher H2/ CO molar ratio could also be obtained by the Boudouard reaction (assigned as R1 in Figure 8) consuming CO molecules with the concomitant formation of carbon species on the catalyst surfaces.16,17 The observed higher H2/CO ratios with lower CO2 conversion and the suppressed formation of filamentous carbons in the fluidized-bed reactor, compared to the fixed-bed reactor, can possibly be attributed to the degree of gasification (assigned as R2 in Figure 8) of deposited carbon precursors (Cx). The degree of gasification is responsible for the formation of filamentous carbon, through reaction with H2O generated by the RWGS reaction. The expanded catalyst bed region in the fluidized-bed reactor could change the

Figure 6. TGA of the used NiA catalysts obtained from a fluidized-bed reactor.

a significant weight loss of around 8 wt %. The observation of a weight gain below 600 °C can be assigned to the oxidation of the reduced metallic nickel particles (Ni + 1/2O2 → NiO), and the weight loss above that temperature can be assigned to coke elimination by gasification with oxygen.1,3 The weight losses in the TGA of the used NiA catalysts increased with increasing Ni content and were found to be 6.32%, 6.48%, and 6.57% on catalysts NiA(10), NiA(15), and NiA(20), respectively. These results are also in line with those for the coke contents on NiA catalysts after the CDR reaction, as summarized in Table 3. With an increase in the nickel content on the Al2O3 support, the rate of formation of filamentous carbon increased continuously because of the increased nickel particle size,20 and they stabilized at some extent above 15 wt % Ni on the Al2O3 support in the fluidized-bed reactor. However, the coke formation rates steadily increased without an optimum nickel content on the NiA catalysts in the fixed-bed reactor with rapid catalyst deactivation. As shown in Table 3, the catalytic performances were significantly altered by the type of reactor and operating conditions, rather than by the type of catalyst. Raman spectroscopy is useful for verifying the types of graphite, and the results for the NiA catalysts after the CDR reaction in fixed-bed and fluidized-bed reactors are shown in Figure 7. The observed Raman shifts at around 1330, 1580, and 2720 cm−1 are assigned to the D, G, and 2D bands, respectively.21,22 The peaks at around 1580 and 2720 cm−1 (G and 2D bands) can be assigned to hexagonal crystal graphite, which are the characteristic peaks for carbon nanotubes. The peak at around 1330 cm−1 is known to decrease after the graphitization process;22 therefore, the ID/IG intensity ratio of the Raman shift peaks can be related to the extent of formation of ordered filamentous carbon. On the used NiA catalysts, the peak intensities and the ratios of the D and G bands were strongly affected by the typ of reactor. The observed larger peak intensities on the used NiA catalysts in the fixed-bed reactor were responsible for significant 13294

dx.doi.org/10.1021/ie3028558 | Ind. Eng. Chem. Res. 2013, 52, 13288−13296

Industrial & Engineering Chemistry Research

Article

formation was found to be one of the main factors in catalyst deactivation, rather than the aggregation of nickel particles. The phenomenon of suppressed formation of filamentous carbon in the fluidized-bed reactor compared to that in the fixed-bed reactor was proposed by the intrinsic properties of the fluidization of catalyst particles with a suitable reducibility in a reducing region. The lower coke formation on catalyst surfaces was found by the facile gasification of carbon precursors with water produced from the RWGS reaction. The strong competitive adsorption properties of product gases such as CO2, H2, CO, and H2O could also reduce the CO2 conversion in the fluidized-bed reactor, and the different rates of gasification of deposited carbon precursors and the Boudouard reaction also altered the H2/CO molar ratio according to the type of reactors. The suppressed coke formation and higher H2/CO molar ratios in the fluidizedbed reactor are also proposed to be due to the enhancement of gasification of coke precursors with water generated during the CDR reaction compared with that in a fixed-bed reactor. The simplest prototype catalyst NiA(15) was also found to be a proper catalytic system in a fluidized-bed reactor because it showed equilibrium conversions of CH4 and CO2 at 850 °C with lower aggregation of nickel particles and lower coke formation in the form of filamentous carbon.

Figure 8. Proposed coke formation and gasification reactions resulting in different H2/CO ratios in a fixed-bed reactor and a fluidized-bed reactor during the CDR reaction.

adsorption properties of CO2 with the competitive adsorption properties of the H2O, H2, and CO produced. The water generated by the RWGS reaction could gasify the deposited carbon precursors efficiently in a fluidized-bed reactor with a change in negligible reaction rate for methane formation through the hydrogenation of carbon precursors.17 The effects of water were found to decrease the CO2 conversion through the competitive adsorption of water on the NiA catalyst surface and to suppress the carbon deposition, as confirmed by the experiment of adding water during the CDR reaction reported in Table 2. The extent of gasification of deposited carbon precursors is relatively larger than the extent of the Boudouard reaction in a fluidized-bed reactor, and the opposite is true in a fixed-bed reactor. The rate of gasification of deposited carbon precursors and the rate of the Boudouard reaction finally reached different equilibrium values by forming a hardly removable filamentous carbon in both reactors. These phenomena were confirmed by the observed lower intensity of the D and G bands in the Raman spectra with a lower ID/IG ratio on the NiA catalysts reacted in the fluidized-bed reactor compared to the fixed-bed reactor. This can also be supported by observing a lower amount of filamentous carbon formed in the fluidized-bed reactor at higher H2/CO molar ratios. However, the effects of the Boudouard reaction and gasification on carbon formation cannot be significant at temperatures much higher than 850 °C because of the decreased activity of the Boudouard reaction. In summary, a stable CDR reaction can be obtained in a fluidizedbed reactor by using prototype catalyst NiA(15) to obtain equilibrium conversions of CH4 and CO2 with an appropriate H2/CO molar ratio of around 1.0. These advantages of a fluidized-bed CDR reaction can possibly be attributed to the intrinsic properties of fluidized-bed reactors such as the continuous circulation of catalyst particles in a reducing region with facile gasification of the deposited carbon precursors with water generated by the RWGS on the active catalyst surface.

■

ASSOCIATED CONTENT

S Supporting Information *

Amount of coke deposition at 550−850 °C and different CO2/ CH4 ratios and catalyst stability at 750 °C in a fixed-bed reactor. This material is available free of charge via the Internet at http://pubs.acs.org.

■

AUTHOR INFORMATION

Corresponding Authors

*Tel.: +82-31-290-7347. Fax: + 82-31-290-7272. E-mail: fi[email protected]. *Tel.: +82-31-290-7248. Fax: + 82-31-290-7272. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

■

ACKNOWLEDGMENTS This work was financially supported by a grant from the Energy Technology Development Programs (2012-2010200040) of the Ministry of Knowledge Economy (MKE) of Korea. The authors acknowledge the financial support of a National Research Foundation (NRF) of Korea Grant funded by the Korean government (MEST; 2011-0009003). This work was also supported by a grant from the Industrial Source Technology Development Programs (2012-10042712) of the Korean government Ministry of Trade, Industry and Energy.

■

REFERENCES

(1) Fan, M. S.; Abdullah, A. Z.; Bhatia, S. Catalytic technology for carbon dioxide reforming of methane to synthesis gas. ChemCatChem 2009, 1, 192−208. (2) Jun, K. W.; Roh, H. S.; Chary, K. V. R. Structure and catalytic properties of ceria-based nickel catalysts for CO2 reforming of methane. Catal. Surv. Asia 2007, 11, 97−113. (3) Baek, S. C.; Bae, J. W.; Cheon, J. Y.; Jun, K. W.; Lee, K. Y. Combined steam and carbon dioxide reforming of methane on Ni/ MgAl2O4: Effect of CeO2 promoter to catalytic performance. Catal. Lett. 2011, 141, 224−234.

4. CONCLUSIONS The degree of filamentous carbon formation and its influence on the conversions of CH4 and CO2 at different H2/CO molar ratios in a laboratory-scale fixed-bed reactor and a bench-scale fluidized-bed reactor were investigated on a prototype Ni/ Al2O3 catalyst during the CDR reaction. The extent of coke 13295

dx.doi.org/10.1021/ie3028558 | Ind. Eng. Chem. Res. 2013, 52, 13288−13296

Industrial & Engineering Chemistry Research

Article

(4) Kambolis, A.; Matralis, H.; Trovarelli, A.; Papadopoulou, Ch. Ni/ CeO2−ZrO2 catalysts for the dry reforming of methane. Appl. Catal. A: Gen. 2010, 377, 16−26. (5) Khalesi, A.; Arandiyan, H. R.; Parvari, M. Production of Syngas by CO2 Reforming on MxLa1−xNi0.3Al0.7O3−d (M = Li, Na, K) Catalysts. Ind. Eng. Chem. Res. 2008, 47, 5892−5898. (6) Rostrup-Nielsen, J. R. Production of synthesis gas. Catal. Today 1993, 18, 305−324. (7) Nakamura, J.; Aikawa, K.; Sato, K.; Uchijima, T. Role of support in reforming of CH4 with CO2 over Rh catalysts. Catal. Lett. 1994, 25, 265−270. (8) Bitter, J. H.; Seshan, K.; Lercher, J. A. The state of zirconia supported platinum catalysts for CO2/CH4 reforming. J. Catal. 1997, 171, 279−286. (9) Stagg-Williams, S. M.; Noronha, F. B.; Fendley, G.; Resasco, D. E. CO2 reforming of CH4 over Pt/ZrO2 catalysts promoted with La and Ce oxides. J. Catal. 2000, 194, 240−249. (10) Ozkara-Aydınoglu, S.; Aksoylu, A. E. CO2 reforming of methane over Pt−Ni/Al2O3 catalysts: Effects of catalyst composition, and water and oxygen addition to the feed. Int. J. Hydrogen Energy 2011, 36, 2950−2959. (11) Efstathiou, A. M.; Kladi, A.; Tsipouriari, V. A.; Verykios, X. E. Reforming of methane with carbon dioxide to synthesis gas over supported rhodium catalysts. J. Catal. 1996, 158, 64−75. (12) Damyanova, S.; Pawelec, B.; Arishtirova, K.; Huerta, M. V. M; Fierro, J. L. G. The effect of CeO2 on the surface and catalytic properties of Pt/CeO2−ZrO2 catalysts for methane dry reforming. Appl. Catal. B: Environ. 2009, 89, 149−159. (13) Bae, J. W.; Kim, A. R.; Baek, S. C.; Jun, K. W. The role of CeO2−ZrO2 distribution on the Ni/MgAl2O4 catalyst during the combined steam and CO2 reforming of methane. Reac. Kinet. Mech. Catal. 2011, 104, 377−388. (14) Koo, K. Y.; Roh, H. S.; Seo, Y. T.; Seo, D. J.; Yoon, W. L.; Park, S. B. Coke study on MgO-promoted Ni/Al2O3 catalyst in combined H2O and CO2 reforming of methane for gas to liquid (GTL) process. Appl. Catal. A: Gen. 2008, 340, 183−190. (15) Wu, G.; Zhang, C.; Li, S.; Han, Z.; Wang, T.; Ma, X.; Gong, J. Hydrogen production via glycerol steam reforming over Ni/Al2O3: Influence of Nickel Precursors. ACS Sustainable Chem. Eng. 2013, 1, 1052−1062. (16) Topalidis, A.; Petrakis, D. E.; Ladavos, A.; Loukatzikou, L.; Pomonis, P. J. A kinetic study of methane and carbon dioxide interconversion over 0.5%Pt/SrTiO3 catalysts. Catal. Today 2007, 127, 238−245. (17) Bradford, M. C. J.; Vannice, M. A. CO2 Reforming of CH4 over supported Pt catalysts. J. Catal. 1998, 173, 157−171. (18) Tomishige, K.; Matsuo, Y.; Yoshinaga, Y.; Sekine, Y.; Asadullah, M.; Fujimoto, K. Comparative study between fluidized bed and fixed bed reactors in methane reforming combined with methane combustion for the internal heat supply under pressurized condition. Appl. Catal. A: Gen. 2002, 223, 225−238. (19) Santos, A.; Menendez, M.; Monzen, A.; Santamaria, J.; Miro, E. E.; Lombardo, E. A. Oxidation of methane to synthesis gas in a fluidized bed reactor using MgO-based catalysts. J. Catal. 1996, 158, 83−91. (20) Wang, H. Y.; Ruckenstein, E. Formation of filamentous carbon during methane decomposition over Co−MgO catalysts. Carbon 2002, 40, 1911−1917. (21) Reshetenko, T. V.; Avdeeva, L. B.; Ismagilov, Z. R.; Pushkarev, V. V.; Cherepanov, S. V.; Chuvilin, A. L.; Likholobov, V. A. Catalytic filamentous carbon: Structural and textural properties. Carbon 2003, 41, 1605−1615. (22) Delhaes, P.; Couzi, M.; Trinquecoste, M.; Dentzer, J.; Hamidou, H.; Vix-Guterl, C. A comparison between Raman spectroscopy and surface characterizations of multiwall carbon nanotubes. Carbon 2006, 44, 3005−3013. (23) Guo, J.; Lou, H.; Zheng, X. The deposition of coke from methane on a Ni/MgAl2O4 catalyst. Carbon 2007, 45, 1314−1321.

13296

dx.doi.org/10.1021/ie3028558 | Ind. Eng. Chem. Res. 2013, 52, 13288−13296