Increased CO2 hydrogenation to liquid products using promoted iron catalysts

Journal of Catalysis 369 (2019) 239–248

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Increased CO2 hydrogenation to liquid products using promoted iron catalysts Wilson D. Shafer a,⇑, Gary Jacobs b, Uschi M. Graham c, Hussein H. Hamdeh d, Burtron H. Davis e,1 a

Asbury University, 1 Macklem Drive, Wilmore, KY 40390, USA University of Texas at San Antonio, Chemical Engineering Program – Dept. of Biomedical Engineering, and Dept. of Mechanical Engineering, 1 UTSA Circle, San Antonio, TX 78249, USA c Topasol, 1525 Bull Lea Road Suite # 5, Lexington, KY 40511, USA d Department of Physics, Wichita State University, 1845 Fairmount, Wichita, KS 67260, USA e Center for Applied Energy Research, University of Kentucky, 2540 Research Park Drive, Lexington, KY 40511, USA b

a r t i c l e

i n f o

Article history: Received 6 September 2018 Revised 30 October 2018 Accepted 1 November 2018

Keywords: CO2 hydrogenation Alkali promoter Iron-based catalysts Activity Product selectivity

a b s t r a c t The effect of alkali promoter (K, Rb and Cs) on the performance of precipitated iron-based catalysts was investigated for carbon dioxide (CO2) hydrogenation. Characterization by temperature-programmed reduction with CO, Mössbauer spectroscopy, and transmission electron microscopy were used to study the effect of alkali promoter interactions on the carburization and phase transformation behavior of the catalysts. Under similar reaction conditions, cesium (Cs) and rubidium (Rb) promoted catalysts exhibited the highest initial CO2 conversions to higher hydrocarbons. CO2 conversions then decreased to reach steady state conversions around 170 h on stream. At steady state conversion, all three catalysts exhibited similar CO2 conversions and selectivities. For comparison, a lower loaded Cs (1.5 Cs) promoted iron-based catalyst was prepared. It exhibited slightly lower initial conversion than the higher loaded Cs catalyst, but remained very stable. Among all the catalysts at steady state conversion, the 1.5 Cs promoted catalyst exhibited the highest stability. Results indicate a synergistic effect brought on by these promoters that, if balanced, could potentially yield superior CO2 hydrogenation catalysts. Ó 2018 Elsevier Inc. All rights reserved.

1. Introduction The steady increase of carbon dioxide emissions into the atmosphere, caused by human activities (i.e., electricity, transportation, buildings and deforestation), contributes to harmful global warming and climate change [1]. Two different technologies have been proposed to mitigate the carbon dioxide concentration in the atmosphere. These are carbon dioxide capture and storage (CCS), [2,3] and carbon dioxide capture and utilization (CCU) in chemical synthesis [4,5]. Carbon dioxide storage in a geological reservoir has significant drawbacks including possible leakage, questionable long-term stability, and availability of adequate storage capacity in many regions of the world. Thus, attention has been shifting to utilization (CCU) from CCS. The utilization of CO2 as a feedstock for producing chemicals not only contributes to reducing global climate change caused by increasing CO2 emissions, but also provides a grand challenge in exploring new concepts and opportunities for ⇑ Corresponding author. 1

E-mail address: [email protected] (W.D. Shafer). Deceased.

https://doi.org/10.1016/j.jcat.2018.11.001 0021-9517/Ó 2018 Elsevier Inc. All rights reserved.

catalytic and industrial development [6]. CO2 hydrogenation to hydrocarbons is a modification of Fischer–Tropsch synthesis (FTS), where CO2 is a reactant instead of carbon monoxide (CO). FTS is a promising technology to catalytically convert synthesis gas (a mixture of CO and H2) for the production of clean transportation fuels and chemicals via a surface polymerization reaction [7,8]. Hydrogenation of carbon dioxide is carried out with catalysts that have been demonstrated to be active for the Fischer-Tropsch synthesis (FTS) reaction. Several group VIII metals such as Fe, Co and Ru are known to be active for Fischer-Tropsch synthesis [7– 13] and these catalysts have also been tested for CO2 hydrogenation to hydrocarbons [14–20]. Supported Co, Ni and Ru catalysts yielded mainly methane from CO2 hydrogenation and only small amounts of higher hydrocarbons were observed [14–20]. Iron catalysts are more suitable for the CO2 hydrogenation reaction, because they possess both higher intrinsic reverse water-gas shift (RWGS) and FTS activity [20,21]. Varying the Group I alkali metal promoters altered adsorption of syngas [22–24]. These studies describe the ability of the larger, more basic alkali to suppress methane by suppressing hydrogen dissociation and promoting

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CO dissociation, thus promoting the formation of longer hydrocarbon chained products (e.g., wax). It was proposed that this ability could extend to CO2 utilization. In essence, these catalysts could have the potential of turning CO2 into liquids and waxes. Trends from the data indicate that CO2 hydrogenation proceeds via a two-step reaction mechanism [25]. In the first step, carbon dioxide is converted into carbon monoxide through RWGS (Eq. (1)). The carbon monoxide produced in this step subsequently reacts with hydrogen by FTS, producing mainly hydrocarbons, as described in the general FTS equation (2):

RWGS: CO2 + H2 CO + H2 O DHR;573K = +38 kJ/mol

ð1Þ

FTS: nCO + (2n + 1) H2 ! Cn H2nþ2 + nH2 O DHR;573K =  166 kJ/mol ð2Þ

The first equation describes RWGS, yet the norm for iron catalysts is water-gas shift (WGS). However, at these lower temperatures the equilibrium of the RWGS/WGS will swing based on the ratio of the reactants and the temperature. As WGS is exothermic, higher equilibrium conversion is favored at lower temperatures, which are also necessary for FTS (based on Eq. (1)). Yet, in the current work, the concentration of CO is essentially 0; thus, based on Le Chatelier’s principle, RWGS will be preferred [26]. Some standard equilibrium calculations for the RWGS reaction describe an equilibrium constant of 0.017, indicating only a slight shift, but predictions yielded a molar amount of CO per hour at 0.06 mol/h. This data agreed with equilibrium data from WGS at these lower temperatures, where Keq = 0.018 for the RWGS is equal to 1/Keq for WGS [27]. The second equation is a general equation for FTS, which produces a vast array of hydrocarbons. A calculated equilibrium constant for FTS was significantly higher where Keq = 5.56e+2, however this too agreed with previous research on iron FTS catalysts [28]. It is generally suggested that RWGS is the slower of the two reactions. The product distribution for iron catalysts is more diverse than the other active metals and, when promoted with Group I alkali metals, such as potassium, more olefins (and specifically, 1-olefins) are produced. Iron-based catalysts often contain small amounts of structural promoters (e.g., Al2O3, SiO2) along with chemical promoters (e.g., Cu, Mn, and alkali elements) to improve both activity and selectivity [29]. Alkali metals alter the catalyst basicity and, in turn, influence the adsorption of reactants (CO and H2) on the active sites, which is suggested to cause an increase in the heat of CO adsorption on the catalytic surface [30], thereby increasing the surface coverage of carbon species due to increased CO bond scission rates. This leads to some effects on the FTS activity: an enhancement in the selectivity to olefins, a suppression in the formation of methane, and a selectivity shift to higher molecular weight products [31]. Choi et al. [32] studied the effect of potassium loading on an iron-based catalyst for the CO2 hydrogenation reaction. They reported that CO2 conversion and yield of hydrocarbons increased with increasing the potassium promoter up to K/Fe = 0.5 and then slightly decreased with further increases in potassium loading. Wang et al. [33] investigated the effect of alkali promoters (Li, Na, K, Rb, and Cs) over zirconia-supported iron catalysts for CO2 hydrogenation. These authors reported that, with the exception of Li, other alkali-promoted catalysts enhance the activity, decrease the methane selectivity, and increase the higher hydrocarbon selectivity compared to the un-promoted Fe/ZrO2 catalyst. In this contribution, the goal is to investigate the influence of alkali promoters (K, Rb, and Cs) over a traditional FTS precipitated ironbased catalyst during CO2 hydrogenation and to compare the activity, selectivity, and stability of the catalysts.

2. Experimental 2.1. Catalyst preparation Precipitated iron catalysts were prepared using a ferric nitrate solution obtained by dissolving iron (III) nitrate nonahydrate (1.17 M) in deionized water, and then tetraethylorthosilicate (TEOS) was added to provide the desired Fe:Si (100:4.6) ratio. The mixture was agitated vigorously until the TEOS had hydrolyzed. A flow of the TEOS and iron nitrate mixture was added to a continuously stirred tank reactor (CSTR) precipitation vessel together with a stream of ammonium hydroxide (14.8 M) that was added at a rate to maintain a pH of 9.0. By maintaining the slurry pH at 9.0 and an average residence time of six minutes, a base catalyst material containing iron and silicon was obtained. The slurry from the CSTR was filtered with a vacuum drum filter and the solid was washed twice with deionized water. The final filter cake was dried for 24 h in an oven at 110 °C with flowing air. For this study, the Fe:Si catalyst base powder was then impregnated with the proper amount of aqueous alkali nitrate (e.g., KNO3 pH  6.2, RbNO3 pH  6.0 and CsNO3 pH  6.0) and Cu (NO3)2 (pH  2.0) solution to produce the desired composition of Fe:Si:alkali:Cu = 100:4.6:3.0:2.0 (atomic ratios). For the purpose of comparison, a 100Fe:4.6Si:1.5Cs:2.0Cu catalyst was also prepared in a similar manner. Finally, the catalysts were dried at 110 °C overnight followed by impregnation, drying, and calcination in a muffle furnace at 350 °C in an air flow for 4 h. 2.2. Catalyst characterization 2.2.1. Surface area and pore size distribution BET surface area and porosity measurements on the calcined catalysts were conducted using a Micromeritics 3-Flex unit. To degas the samples, the temperature was ramped to 160 °C followed by the chamber being evacuated for at least 12 h to approximately 50 mTorr. The BET surface area, pore volume, and average pore radius were obtained for each sample. 2.2.2. Temperature programmed reduction (TPR) of CO Temperature programmed reduction (TPR) of CO profiles for the calcined catalysts were obtained using a Zeton-Altamira AMI-200 unit furnished with a thermal conductivity detector (TCD). The TPR was achieved by a 10% CO/He carrier gas mixture (referenced to helium) with a flow rate of 30 cm3/min. The catalyst samples were heated from 50 to 525 °C using a ramp rate of 10 °C/min then held for 1 h. A liquid nitrogen trap was utilized to constantly remove the CO2 produced. 2.2.3. Morphology and phase transformation by HRTEM and STEM The morphology and phase transformations of distinct catalyst particles were investigated by TEM. Specimens were arranged on copper grids (200 mesh) from representative catalyst powders after FTS. TEM imaging was accomplished using a JEOL 2010F field-emission gun transmission electron microscope (accelerating voltage of 200 keV and magnification ranging from 50 to 1000 K). Images were analyzed using a Gatan Ultrascan 4 k  4 k CCD camera, and all data-managing and examination were completed by applying the Gatan Digital Micrograph software. Furthermore, STEM imaging was achieved with a high-angle annular dark-field (HAADF) detector that further utilized a Gatan imaging filter (GIF). The GIF permitted imaging of solely the catalyst particles, excluding the carbon deposits because the light elements were not detected in HAADF. This provided an excellent contrast in the TEM and STEM images that expose the carbon residues.

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2.2.4. Mössbauer spectroscopy Mössbauer spectra were collected in transmission mode by a constant acceleration spectrometer (MS-1200, Ranger Scientific). A radiation source of 30 mCi 57Co in Rh matrix was used and spectra were obtained using a gas detector. The catalyst samples collected from the CSTR were dispersed in the wax to become a solid phase at room temperature. For the low temperature measurements, the samples were placed inside a vibration-free closed cycle cryostat (Cryo-Industries of America). Structural analysis of the samples was performed by least-squares fitting of the Mössbauer spectra to a summation of hyperfine sextets. The leastsquares fitting procedure employed user defined functions within the PeakFit program. The parameters for each sextet in the fit consisted of the position, width, and height of the first peak; the hyperfine magnetic field; and the quadrupole electric field. These parameters were allowed to vary freely to obtain the best fit of the experimental data. Errors in the determined percentages for the Fe values are about ±3% for well resolved spectra; in those that contain several iron oxide and carbide phases, the uncertainty increased with the complexity of the representative spectrum (i.e., depending on the degree of overlap and the weakness of the signal). However, these complex spectra were obtained during the course of transformation from a predominantly iron oxide form to an iron carbide/oxide mixture and conform to a general trend. 2.2.5. Temperature programmed decarburization of hydrogen (TPDec) Samples were positioned into a 0.500 O.D stainless steel reactor between two separate beds of glass wool. A H2/He mixture was introduced at 50 standard cubic centimeters per second (SCCM). The temperature was increased around 2 °C/min. and injections transpired every 5 min. The temperature program commenced at room temperature and was ramped to 850 °C. The effluent was delivered into an SRI-GC 8610 capable of Flame Ionization Detection (FID) and TCD evaluation; thus, the analysis was performed on the amount of methane that evolved from this method.

2.4. Analytical procedures for product analysis The reaction products were collected with a post reactor distillation unit that housed three pressure vessels retained at separate temperatures: a hot trap (200 °C), a warm trap (100 °C) and a cold trap (0 °C). These FTS products were divided into distinctive fractions (wax, oil and aqueous phase) for identification and quantification. Liquids were extracted online daily from the warm trap which and was separated into the light wax and water mixture. Alongside the warm trap, the cold trap also was pulled daily and separated into the oil plus water phase. Gas phase CO2 and H2 were delivered by in-house calibrated mass flow controllers, and the vapor phase components (e.g. CO, CO2, N2, C1, C2 – C6 (olefin/paraffin) were passed into an online HP Quad Series Micro GC. This instrument was calibrated externally for each vapor phase component and the results were continually coupled with online flow measurements. By doing this, accurate measurements of the mole % of all the reactants and products could be assessed online. Aqueous Phase - The analysis of the aqueous phase used an SRI 8610C GC with a thermal conductivity detector (TCD). Necessary external calibrations of these materials were performed to quantify the materials in this phase. This is especially important in the case of iron catalysts, which normally display a much more diverse array of products, including oxygenates that enter the aqueous phase. Oil + Wax Phase – No sample preparation for this phase was necessary before analysis. The organic products were directly analyzed using an Agilent 6890 GC with a flame ionization detector (FID) and a 60 m DB-5 column. All of these data were then combined and converted into carbon mole % to relate the diverse product distribution to CO consumption.

3. Results and discussion 2.3. Catalyst assessment The CO2 hydrogenation reactions were performed in a 1 L CSTR furnished with a magnetically propelled agitator containing a turbine impeller, a gas-inlet and vapor exhaust line, and a stainless steel fritted filter (2 mm) outside the reactor. A tube fitted with a stainless steel frit filter (0.5 mm opening) was set underneath the liquid layer inside the reactor. This filter was utilized to extract reactor wax to conserve a uniform liquid level in the reactor. Another stainless steel dip tube (1/8” OD), set with pressure vessels, extended to the side of the FTS system and was utilized to extract aliquots of the catalyst/wax slurry at different times on stream (TOS). Three individual mass flow controllers (MFC) were utilized for flow regulation of hydrogen and carbon dioxide and carbon monoxide, individually. The reactant gases were premixed in a 500 cc pressure vessel prior to entering the reactor. After being mixed, the reactant gases entered the CSTR beneath the stirrer which was controlled at 750 rpm. The CSTR slurry temperature was preserved at a constant temperature (±1 °C) by a Model 3254 Omega temperature controller. Normally, 20 g of an iron catalyst was added to C30 oil (310 g) in the CSTR to generate a slurry to be comprised of roughly 6% iron oxide. The CSTR temperature was then increased to 270 °C at a rate of 1 °C/min. Next, the catalyst was activated by CO at a space velocity of 3.0 sl/h/gcat at 270 °C and 0.1 MPa for 24 h. Then, reaction temperature was kept constant so that the pressure could be increased to 1.3 MPa. Lastly, the CO2 hydrogenation reaction commenced by adding the H2 and CO2 gas mixture to the reactor at a space velocity of 2.0 sl/h/gcat and a H2/CO2 ratio of 3.0.

3.1. Catalyst characterization BET surface area and pore size distribution results of unpromoted and various alkali-promoted iron-based catalysts are shown in Table 1. The surface area of the base catalyst (100Fe:4.6Si) was found to be 113.4 m2/g. After the addition of the alkali (K, Rb, and Cs) and Cu promoters, the surface areas were found to be lower than that of the base catalyst. The surface areas of various alkali-promoted catalysts were similar (±2%, within experimental error). For the alkali and Cu promoted catalysts, if the promoter oxides (alkali and Cu) do not contribute to the surface area, the corrected surface area would have been slightly higher than the obtained values listed in Table 1. These values indicate that there is some pore blockage in the promoted catalysts and this is likely due to the promoter oxide clusters blocking a fraction of pores from the adsorbing gas or in decreasing the porosity during the second drying step.

Table 1 BET surface area and PSD results of the various iron-based catalysts. Catalyst

BET surface area (m2/g)

Single point pore volume (cm3/g)

Single point pore diameter (nm)

100Fe:4.6Si 100Fe:4.6Si:3.0K:2.0Cu 100Fe:4.6Si:3.0Rb:2.0Cu 100Fe:4.6Si:3.0Cs:2.0Cu 100Fe:4.6Si:1.5Cs:2.0Cu

113.4 97.5 103.3 98.7 103.1

0.218 0.199 0.179 0.167 0.202

7.7 8.2 6.9 6.8 7.8

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Temperature programmed reduction of carbon monoxide (COTPR) was used to investigate the reduction and carburization behavior of iron-based catalysts in CO atmosphere. The CO-TPR profiles of 100Fe:4.6Si base catalyst and various alkali and Cu promoted iron-based catalysts are shown in Fig. 1. The base catalyst showed only three reduction peaks, whereas the promoted catalyst profiles show four apparent reduction/carburization peaks. Actually, a very weak peak around 200 °C is also detected, which could be ascribed to the reduction of hematite (a-Fe2O3) to magnetite (Fe3O4) and the first major peak is located in the temperature range of 220–240 °C for the promoted catalysts, but for the base catalyst this peak appeared at slightly higher temperature (275 °C). These peaks are associated with reduction of Fe2O3 to lower oxides (i.e., Fe3O(4-x) and a defect-laden form of this oxide similar to FeO) prior to carburization. These peaks are fully consistent with previous CO-TPR XANES/EXAFS studies [29–31]. The second peak for the base catalyst is located in the temperature range of 350–525 °C, whereas for the promoted catalysts, the second peak is located in the temperature range of 300–400 °C. This could be ascribed to the carburization of partially reduced iron oxides [7,12,15]. The alkali and copper promoters significantly affect the second main peak in that the carburization rate is more rapid and the peak maximum is shifted to lower temperature. These results clearly reveal that the addition of promoters (alkali and copper) has a significant effect on the carburization rate of the iron catalyst. The second peak temperature was not significantly influenced by the basicity of the alkali promoter (i.e., all three catalysts showed similar Tmax). These results are consistent with earlier CO-TPR XANES/EXAFS studies [34] (i.e., the carburization rates of potassium, rubidium and cesium promoted iron catalysts are similar). Mössbauer spectroscopy followed the dynamic phase transformation between iron oxide and iron carbide after carburization and CO2 hydrogenation. Representative Mössbauer spectra for iron-based catalysts, including the catalysts following CO activation, as well as used catalysts after running the CO2 hydrogenation reaction, are shown in Tables 2 and 3, respectively. Compositional changes in terms of percentage peak areas of the different iron species were determined from curve fitting of the Mössbauer measurements for all samples and are shown in Table 2 for various alkali loaded copper-promoted iron-based catalysts. During carburization with CO, the as-synthesized catalyst first transformed from hematite (Fe2O3) to magnetite (Fe3O4) and then further reduced to iron carbide phases, which was confirmed with CO-

Table 2 Summary of phase identification of iron from Mössbauer spectroscopy analysis of freshly carburized various alkali loaded iron-based catalysts. Catalyst

Fe% as Fe3O4

Fe% as v-Fe5C2

100Fe:4.6si:3.0 K:2.0Cu 100Fe:4.6si:3.0Rb:2.0Cu 100Fe:4.6si:3.0Cs:2.0Cu 100Fe:4.6si:1.5Cs:2.0Cu

20 16 14 21

80 84 86 79

Table 3 Summary of phase identification of iron from Mössbauer spectroscopy analysis of CO2 hydrogenated various alkali loaded iron-based catalysts. Catalyst

Fe% as Fe3O4

Fe% as v-Fe5C2

100Fe:4.6si:3.0K:2.0Cu 100Fe:4.6si:3.0Rb:2.0Cu 100Fe:4.6si:3.0Cs:2.0Cu 100Fe:4.6si:1.5Cs:2.0Cu

43 36 42 44

57 64 58 56

TPR results. For the potassium loaded copper-promoted iron catalyst, at the end of the 24 h carburization period (Table 2), low temperature (20 K) Mössbauer results reveal that the initial hematite converted to a mixture of 80% v-Fe5C2 and 20% Fe3O4. For Rb and Cs promoted iron catalysts, initial hematite converted into mixtures of iron carbide (84 and 86%) and iron oxide (16 and 14%), respectively. The lower loaded Cs-promoted catalyst also exhibited similar fractions of iron phases. After the 24 h CO activation period, with increasing the basicity of the alkali promoter the carburization rate is either similar or slightly increased. These results further confirm the observations from CO-TPR measurements (Fig. 1). Low temperature (20 K) Mössbauer spectroscopy results of various alkali-promoted copper-promoted iron-based catalysts used for CO2 hydrogenation are shown in Table 3. CO2 hydrogenation used catalysts exhibited lower fractions of iron carbide and higher fractions of iron oxide phases compared to their counterpart freshly carburized catalysts. Among all the catalysts, the rubidiumpromoted iron-based catalyst exhibited a higher fraction of iron carbide and a lower fraction of iron oxide relative to the other alkali-promoted iron catalysts. 3.2. CO2 utilization The effect on CO2 conversion against time on stream (TOS) for various alkali-promoted iron-based catalysts is shown in Fig. 2. Figs. 3–10 display selectivity data for the carbon distribution and the selectivity against the TOS. All selectivity data was calculated in terms of mole % of carbon to maintain consistency with the moles of CO2 utilized. To maintain experimental control, activation

% Carbon Dioxide Conversion

30 25 20 15 10

3.0Rb 3.0Cs 3.0K 1.5Cs

5 0 0

Fig. 1. CO-TPR profiles of various alkali loaded iron-based catalysts.

50

100 150 TIme on Stream

200

250

Fig. 2. Effect on CO2 conversion against TOS for various alkali loaded iron-based catalysts.

243

60

70

50

60

Selectivity (% mole of Carbon)

Selectivity (% mole of Carbon)

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40 30

Methane C2-C4

20

C5+

10

40 Methane 30

C2-C4 C5+

20 10

0 0

50

100 150 Time on Stream (h)

200

0

250

0

Fig. 3. Product selectivity results of 100Fe:4.6Si:3.0Cs:2.0Cu catalyst.

50

100 150 Time on Stream (h)

200

250

Fig. 6. Product selectivity results of 100Fe:4.6Si:1.5Cs:2.0Cu catalyst.

70

50 Selectivity (% mole of arbon)

Selectivity (% mole of Carbon)

50

60 50 40 Methane

30

C2-C4 20

C5+

45 40 35

Isomerized Parafin Alcohol 1- Olifen Cis-Olifen Trans-Olifen

30 25 20 15 10 5 0

10

0

50

100 150 Time on Stream (hours)

200

250

0 0

50

100 150 Time on Stream (h)

200

250

Fig. 7. Product selectivity results of 100Fe:4.6Si:3.0Cs:2.0Cu catalyst. Products are totaled between C1 to C25 and compared to TOS.

Fig. 4. Product selectivity results of 100Fe:4.6Si:3.0Rb:2.0Cu catalyst. 50 45 Selectivity (% mole of arbon)

Selectivity (% mole of Carbon)

70 60 50 40 30

Methane C2-C4

20

40 35

Isomerized

30

Parafin

25

Alcohol 1- Olifen

20

Cis-Olifen

15

Trans-Olifen

10 5

C5+

0

10

0

50

100 150 Time on Stream (hours)

200

250

0 0

50

100 150 Time on Stream (h)

200

250

Fig. 8. Product selectivity results of 100Fe:4.6Si:3.0Rb:2.0Cu catalyst. Products are totaled between C1 to C25 and compared to TOS.

Fig. 5. Product selectivity results of 100Fe:4.6Si:3.0K: 2.0Cu catalyst.

and reaction conditions (temperature, pressure, and GHSV) were maintained constant for all catalysts. The initial CO2 conversion was improved by increasing the basicity of the promoter. However, the higher basicity promoted catalysts displayed a CO2 conversion decrease with time to finally reach a steady state conversion level at around 170 h TOS. Unlike the Cs and Rb promoted catalysts, the potassium-promoted iron catalyst exhibited very stable CO2 conversion with time. After several hours of TOS a similar steady state conversion level for all three alkali-promoted catalysts (100Fe:4.6Si:2.0Cu:3.0alkali) was reached. For the purpose of comparison, a lower loaded Cs-promoted iron-based catalyst was prepared and tested for CO2 hydrogenation. The lower Cs-promoted

(1.5Cs) catalyst exhibited lower initial CO2 conversion than the higher loaded catalyst (3.0Cs), but at steady state conversion exhibited an elevated activity when compared to the higher Cs-promoted catalyst. The reason for this higher CO2 conversion to longer hydrocarbon chains might be due to the higher basicity of this alkali promoter [24,34], which is suggested to cause an increase in the heat of CO2 adsorption on the catalytic surface, thereby increasing the surface coverage of carbon species due to an increased CO2 bond scission rate. The mechanism (i.e., RWGS followed by FTS) mainly involves breaking the CAO bond regardless of the species involved in each step - CO2 or CO. This is because the mechanistic route described involves RWGS (i.e., with CAO bond scission in converting CO2 to

W.D. Shafer et al. / Journal of Catalysis 369 (2019) 239–248 60

6.00E-03

50

5.00E-03

3.0Rb Rate of FT (mole/h/g of catalyst)

Selectivity (% mole of arbon)

244

40 Isomerized Parafin Alcohol 1- Olifen Cis-Olifen Trans-Olifen

30 20 10

3.0Cs 3.0K

4.00E-03

1.5Cs 3.00E-03 2.00E-03 1.00E-03

0 0

50

100 150 Time on Stream (hours)

200

250

0.00E+00 0

Fig. 9. Product selectivity results of 100Fe:4.6Si:3.0K: 2.0Cu catalyst. Products are totaled between C1 to C25 and compared to TOS.

50

100 150 Time on Stream (h)

200

250

Fig. 11. Rate of the FTS (mole/hour/gram of catalyst) with TOS i.e. the products that were formed from CO2 – CO - FTS.

1.40E-03

60 Isomerized Parafin Alcohol 1- Olifen Cis-Olifen Trans-Olifen

50 40 30 20 10 0 0

50

100 150 Time on Stream (hours)

200

250

Rate of CO (mole/h/g of catalyst)

Selectivity (% mole of arbon)

70

Fig. 10. Product selectivity results of 100Fe:4.6Si:1.5Cs:2.0Cu catalyst. Products are totaled between C1 to C25 and compared to TOS.

1.20E-03 1.00E-03 8.00E-04

3.0Rb 3.0Cs

6.00E-04

3.0K 4.00E-04

1.5Cs

2.00E-04 0.00E+00 0

100 150 Time on Stream (h)

200

250

Fig. 12. Rate of the CO (mole/hour/gram of catalyst) with TOS, i.e. the amount of CO that was produced from CO2, but not converted to FTS.

2.50E-02 Rate of Hydrogen (mole/h/g of catalyst)

CO) followed by a second CAO scission that provides the active monomer for FTS of hydrocarbon products. The coverage of these molecules (CO2/CO/H2) on the surface will determine how the FTS products are constructed. Thus, the higher the scission rate of CAO, the greater the coverage of C* monomer on the catalyst surface. However, a tradeoff is that if too much of this C* monomer is present on the surface, poisoning by carbon deposition can lead to catalyst instability. Since water-gas shift is a reversible reaction, and the reactants CO2 and H2 provide the driving force, RWGS will occur and tend towards establishing an equilibrium [28]. However, the CAO scission for RWGS/WGS is not well understood as to whether the route involves a redox or an associative mechanism [34]. The second CAO scission step is also not fully understood. Formation of the C* monomer may be direct CO dissociation or involve the assistance of hydrogen [35–45]. The CO2 bond scission rates were observed in the total CO2 conversion by comparing the rates of each catalyst with the time on stream (Figs. 11–14). FTS product formation will most likely occur through dissociated CO, whereas CO2 will remain inert due to the stronger binding of CO to the FTS catalyst surface sites relative to CO2. Thus, considering the mechanistic route of these reactions CO2 ? CO ? FTS products, with all rates being in mole/hour/gram of catalyst, the rate of CO2 is the summation of the rate of FTS product formation and CO formation, where the rate of FTS is the amount of CO2 converted to FTS hydrocarbons, and the rate of CO is the amount of CO2 only converted to CO (presumably via RWGS). The rate of FTS decreases with time, and this decline is more significant for the 3.0 Cs and Rb promoted iron catalysts (Fig. 11). The decline in the FTS hydrocarbon production rate mimics that of the decline in CO2 conversion. However, the rate of CO production increases (Fig. 12), indicating that the RWGS rate is

50

2.00E-02

1.50E-02

1.00E-02

3.0Rb 3.0Cs

5.00E-03

3.0K 1.5Cs

0.00E+00 0

50

100 150 Time on Stream (h)

200

250

Fig. 13. Rate of the H2 (mole/hour/gram of catalyst) with TOS, i.e. the amount of H2 consumed based upon H2in and H2out.

not nearly as affected. Lastly, the rate trends for both the H2 (Fig. 13) and water (Fig. 14) also follow the overall conversion, where the rates are highest for the 3.0Rb and 3.0Cs, but decrease with time. Essentially, group I promoters have been shown to suppress H2 dissociation in conjunction with promoting CO dissociation [46,47]. This trend is observed more so with the larger, more basic metals and is exhibited through higher surface coverages of carbon. However, this carbon coverage is not specific to the C* mono-

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2.00E-02

Decarburization (a.u.)

Rate of Water (mole/h/g of catalyst)

2.50E-02

1.50E-02

1.00E-02

3.0Rb 3.0Cs 3.0K

5.00E-03

3.0 3.0 3.0 1.5

K Rb Cs Cs

1.5Cs 0.00E+00 0

50

100 150 Time on Stream (h)

200

250

0

Fig. 14. Rate of the H2O (mole/hour/gram of catalyst) with TOS, i.e. the amount water produced.

200

400

600

800

Temperature (oC) Fig. 15. Temperature programmed decarburization of freshly carburized various alkali loaded iron-based catalysts.

(1) Active iron carbide phases are gradually oxidized to magnetite (Fe3O4), which is relatively inactive for FTS [9,34,47]; (2) Deposition of inactive carbonaceous compounds takes place on the surface of the catalyst, thereby limiting the contact between reactant gases and the catalytically active phase [17,24,48–52]; (3) Sintering, which is the loss of catalytic surface area due to ripening or migration and coalescence of iron phases [51,53,54]; (4) Poisoning and deactivation by sulfur compounds, which are typically present in most synga feeds, but not here. In the present study, the same reactant gas was used from two pure H2 and CO2 cylinders, so the possibility of poisoning by sulfur compounds was ruled out. Furthermore, provided that iron (unlike Co/Ni/Ru) is the bulk of the catalyst, not needing to be dispersed as fine particles on a support, sintering is also ruled out as the main cause for conversion loss here. Based upon Tables 2 and 3, which display losses in the carbide phase, the results suggest formation of an oxide phase. However, these tables also display a higher carbide fraction for Rb and Cs promoted catalysts than for the K-promoted one. This could potentially rule out the first explanation for the loss in conversion, as more carbide is present in the higher basicity promoted iron catalysts. The second route seems to best explain the decrease in the overall conversion to FTS products. This explanation can be further verified by the decarburization of the catalysts with H2. Fig. 15 shows very similar amounts of carbon being decomposed to methane by the addition of hydrogen. However, the spent catalysts (Fig. 16) display a very different picture – where the highest amounts of carbon are again found on the larger, more basic alkali-promoted iron catalysts. If results from the total carbon decomposition, as displayed in

3.0K 3.0Rb 3.0Cs 1.5Cs

Decarburization rate (a.u.)

mer for FT synthesis, although it suggests that more carbon is available from CO to participate in the FT synthesis. Discussions into the specifics of the FTS mechanism are difficult to prescribe, i.e. whether CO dissociatively adsorbs, or dissociates by means of hydrogen. In summary, these results reveal an initial positive role that these larger, more basic alkali play in promoting CO2 dissociation as demonstrated by the selectivity differences observed in Figs. 3– 10. However, the samples taken at later TOS, for the 3.0Cs and 3.0Rb, exhibit oil levels very similar to the K run, where little FTS oil is actually produced. Deactivation of the iron catalysts may be explained by the following four mechanisms, which have been described in the literature:

0

200

400

600

800

o

Temperature ( C) Fig. 16. Temperature programmed decarburization of CO2 hydrogenation reaction used various alkali loaded iron-based catalysts.

Fig. 16, are compared to results from the Mössbauer spectroscopy (Tables 2 and 3), the difference in carbon may be due to coke. However, issues arise when attempting to describe the distinctions between the different forms of carbon. Thus, the focus here is on the differences between freshly activated and used catalysts regarding the total amount of carbon versus the amount of carbon in the carbide phase. Carbon totals are provided in Figs. 15 and 16 by decarburization where the 3.0Rb/3.0Cs catalysts displayed a much larger amount of overall carbon material. However, the Mössbauer spectroscopy results, provided in Tables 2 and 3, display very similar carbide and oxide phase amounts for each catalyst, both used and fresh. The 3.0Cs/3.0Rb catalysts display much higher carbon content and the results of Mössbauer spectroscopy display very similar amounts of each C-containing phase. This highlights that the catalysts suffer from significantly more carbon laydown. These trends were further confirmed by the TEM images (Fig. 17), including representative images of the fresh and the spent catalyst from the 3.0-promoted catalyst series. Previous FTS-IR work showed how a basic promoter such as potassium, when working on an iron catalyst, promotes CO dissociation by donating electron density to the active sites. The effect of this promotion was proposed to be addition of electron density to the FTS metal, followed by increased back donation to the CO pi bond from the metal [55–57]. Thus, the overall CO bond was weak-

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Fig. 17. A TEM of a representative of the fresh catalysts (as all the catalysts before appeared the same), and the spent for 3.0 promoted catalysts.

ened. There are numerous publications on the addition of K and how it affects the FTS product distribution [58–60]. All of these investigations agree on an observed chain lengthening. A higher alpha was observed in the product distribution. One paper in particular provided results over the whole of group I. Observations were noted that Cs and Rb were actually worse promoters as they poisoned the catalyst during FTS [24]. Thus, degradation may be due to the formation of carbon deposits. Simply put, the rate of dissociation for CO/CO2 relative to H2 was higher when heavier alkali metals were employed. Mechanistic routes for both FTS and WGS/ RWGS, including the overall active sites, are still being debated. Confirmation of this trend may be present in the olefin-toparaffin ratio. If hydrogen is being suppressed, more olefin prod-

ucts would be expected. This is displayed in Table 4. The overall olefin content increased, yet the 1-olefin, though it is dominant, did not increase in the same manner. This could mean that, although the overall olefin did increase, indicating the lack of hydrogen on the surface, tuning of the particular olefin material is only affected so much by the addition of these group I alkali promoters. Further evidence of the inability of hydrogen to participate in the mechanism is displayed by the product selectivities for the different series of catalysts Figs. 6–10, where the samples at the beginning of the run (early TOS) display higher amounts of both the alcohols and olefins for the 3.0Cs/3.0Rb promoted catalysts. Previous studies suggested that potassium is the best promoter of the Group I alkali metals for FTS using an active iron catalyst.

Table 4 The totals for the olefin to paraffin ratio (ranging from C1 to C25) with a specific comparison of the 1-olefin to the total olefin with TOS. All numbers here have been put in terms of % mole of carbon – thus, a direct comparison can be made with these products to the carbon from CO. 3.0Rb

3.0Cs

3.0K

1.5 Cs

TOS (h)

O/P

Primary olefin to total olefin

O/P

Primary olefin to total olefin

O/P

Primary olefin to total olefin

O/P

Primary olefin to total olefin

18.75 42.25 64.75 88.00 112.00 136.25 162.00 186.00 214.75 232.25

1.92 1.53 1.69 1.54 1.45 1.47 1.41 1.38 1.33 1.27

79.79 76.80 77.18 77.19 77.91 76.92 78.20 78.66 77.98 79.39

1.85 2.17 1.82 1.56 1.45 1.38 1.10 1.29 1.08 1.04

78.26 79.36 78.36 79.67 79.02 77.62 73.01 77.72 79.88 79.45

0.81 0.96 1.03 1.00 1.14 0.96 1.01 1.06 1.09 1.27

84.91 84.42 84.27 83.52 85.36 79.47 79.56 79.21 80.29 78.96

0.33 0.37 0.40 0.42 0.46 0.48 0.51 0.54 0.53 0.57

77.27 77.94 77.53 77.53 75.81 76.51 75.69 74.61 75.67 75.46

W.D. Shafer et al. / Journal of Catalysis 369 (2019) 239–248 Table 5 A mole to mole comparison of the H2/CO ratio observed, calculated from the exhaust gas. TOS (h)

3.0Rb

3.0Cs

3.0K

1.5Cs

18.75 42.25 64.75 88 112 136.25 162 186 214.75 232.25

77.84 67.42 63.13 58.55 55.84 53.18 52.08 50.75 49.94 48.71

77.36 66.60 61.68 57.47 54.13 53.07 52.50 52.44 51.28 50.71

44.97 42.50 41.57 41.03 40.48 40.96 41.05 41.90 42.60 43.84

49.25 46.74 46.81 45.63 45.21 44.57 44.83 45.23 45.28 46.89

Only sodium trended with potassium; lithium was not effective, and the Rb/Cs may have provided too much CO scission, resulting in C laydown [24]. TEM images (Fig. 17) confirm that the differences might be due to the carbon laydown. Particle growth did seem to occur on all three catalysts, but the carbon deposits are more severe for the Cs and Rb promoted iron catalysts. The question remains as to whether these trends are due to size, basicity, or both. Thus, data from the deactivation once again indicates that these larger, more basic metals, perhaps through the inability for H2 to reach the surface (due to the high carbon surface coverage), coke themselves much more rapidly. The first observations in this deactivation trend are the overall CO2 conversion drops because the selectivity toward the C5+ selectivity decreases (Figs. 3–6). Rapid deactivation rates occur until the product distributions of the higher loading Cs and Rb promoted catalysts resemble that of the K-promoted iron catalyst (Fig. 5). As a confirmation of this effect, the lower Cs-promoted iron catalyst displayed somewhat higher CO2 conversion, but overall it mimicked the stability in CO2 conversion of the potassium-promoted iron catalyst (Figs. 5, 6, 9, 10). Eventually, toward the 200 h mark, all four catalysts end up displaying very similar conversions, selectivities, and rates (Figs. 2–14). Further evidence of the high carbon surface coverage is shown by comparing H2/CO ratios observed in the exhaust (Table 5). Initially, the H2/CO ratios for the 3.0Cs/3.0Rb promoted catalysts are much higher than for 3.0K and the 1.5Cs promoted iron catalysts. However, a clear decrease in the H2/CO ratio is observed for the 3.0Cs/3.0Rb promoted iron catalysts with TOS. The high ratio indicates a significant portion of the CO is being converted to FTS products. Evidence of the FTS conversion is further displayed through the rates of FTS (Fig. 7) and CO (Fig. 8) production. However, as time continues, the H2/CO ratios of the 3.0Rb/3.0Cs promoted iron catalysts begin declining to eventually match those of the 3.0K and 1.5Cs catalysts. The decrease in these ratios provide evidence for the decrease in FT selectivity in the 3.0Rb/3.0Cs promoted catalysts. Once steady state for CO2 conversion is met, all the catalysts display the same H2/CO ratios. The H2/CO trends suggest that the loss in CO2 activity for the 3.0Cs/3.0Rb promoted catalysts is directly tied to the decline in FTS activity. Further evidence for the decrease in FTS activity could be shown by the decreasing trends for the C+5 (Figs. 3–6) and increasing trends for the methane selectivities.

stirred tank reactor to prevent thermal issues. Lastly, initial carburization data indicate the iron catalysts had similar carbide/oxide fractions before reaction testing. By doing this, a direct comparison between the promoters, in how they affect the product selectivity and overall CO2 conversion, could be performed. The higher loadings of Rb and Cs acted very similarly, which led to an initial iron catalyst that was superior to the potassiumpromoted iron catalyst. Not only was the initial overall CO2 conversion much higher, but also the selectivity toward C+5 greatly increased. With the aim of developing a better catalyst in the future for CO2 utilization, these iron catalysts reveal a step in the right direction. However, the shortcomings with these catalysts indicate that carbon deposition is a problem that selectively causes the rate of FTS to decrease. Once the CO2 conversion reaches a steady state, all the catalysts display similar trends, where the main cause for CO2 conversion is RWGS. Thus, if the coking could be mitigated while still using the basicity strength of these alkali metal promoters, a much more effective and cheaper catalyst capable of converting CO2 to higher hydrocarbons may be implemented. Additional work is needed not only regarding catalyst development, but also in optimizing process conditions in order to ensure that the CAO dissociation rate is high enough to provide activity and selectivity, but not excessively high to cause carbon laydown that results in instability. Conflict of interest The authors declare no competing financial interest. Acknowledgments This work was supported by Asbury University, by the Commonwealth of Kentucky, Toposal, and Wichita State University. Dr. Gary Jacobs would like to thank UTSA, the State of Texas, and the STARS program for funding. The authors thank Chelsea Parsons for assisting with editing. References [1] [2] [3] [4] [5] [6] [7] [8] [9] [10] [11] [12] [13] [14] [15] [16] [17] [18]

4. Conclusions [19]

This work describes FTS catalysts that have the potential to effectively convert CO2 to higher hydrocarbon products. To make a direct comparison of promoters to FT synthesis, all of the catalysts were reduced and reactions were tested under the same conditions. Furthermore, all reactions were run in an isothermal

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