**Fischer-Tropsch Synthesis: Cd, In and Sn E**ff**ects on a 15%Co**/**Al2O3 Catalyst**

**Wenping Ma 1,\*, Gary Jacobs 1,**†**, Wilson D. Shafer 1,**‡**, Yaying Ji 1, Jennifer L. S. Klettlinger 2, Syed Khalid 3, Shelley D. Hopps <sup>1</sup> and Burtron H. Davis 1,§**


Received: 30 August 2019; Accepted: 11 October 2019; Published: 16 October 2019

**Abstract:** The effects of 1% of Cd, In and Sn additives on the physicochemical properties and Fischer-Tropsch synthesis (FTS) performance of a 15% Co/Al2O3 catalyst were investigated. The fresh and spent catalysts were characterized by BET, temperature programmed reduction (TPR), H2-chemisorption, NH3 temperature programmed desorption (TPD), X-ray absorption near edge spectroscopy (XANES), and X ray diffraction (XRD). The catalysts were tested in a 1 L continuously stirred tank reactor (CSTR) at 220 ◦C, 2.2 MPa, H2/CO = 2.1 and 20–55% CO conversion. Addition of 1% of Cd or In enhanced the reduction degree of 15%Co/Al2O3 by ~20%, while addition of 1% Sn slightly hindered it. All three additives adversely impacted Co dispersion by 22–32% by increasing apparent Co cluster size based on the H2-chemisorption measurements. However, the decreased Co active site density resulting from the additives did not result in a corresponding activity loss; instead, the additives decreased the activity of the Co catalysts to a much greater extent than expected, i.e., 82–93%. The additional detrimental effect on catalyst activity likely indicates that the Cd, In and Sn additives migrated to and covered active sites during reaction and/or provided an electronic effect. XANES results showed that oxides of the additives were present during the reaction, but that a fraction of metal was also likely present based on the TPR and reaction testing results. This is in contrast to typical promoters that become metallic at or below ~350 ◦C, such as noble metal promoters (e.g., Pt, Ru) and Group 11 promoters (e.g., Ag, Au) on Co catalysts in earlier studies. In the current work, all three additives remarkably increased CH4 and CO2 selectivities and decreased C5<sup>+</sup> selectivity, with the Sn and In additives having a greater effect. Interestingly, the Cd, In, or Sn additives were found to influence hydrogenation and isomerization activities. At a similar conversion level (i.e., In the range of 40–50%), the additives significantly increased 2-C4 olefin content from 3.8 to 10.6% and n-C4 paraffin from 50 to 61% accompanied by decreases in 1-C4 olefin content from 48 to 30%. The Sn contributed the greatest impact on the secondary reactions of 1-olefins, followed by the In and Cd. NH3-TPD results suggest enhanced acid sites on cobalt catalysts resulting from the additives, which likely explains the change in selectivities for the different catalysts.

**Keywords:** Fischer-Tropsch synthesis; Co; Al2O3; Pt; Cd; In; Sn; hydrocarbon selectivity; synergic effect; GTL; additives; reducibility; XANES

#### **1. Introduction**

Supported cobalt catalysts have received renewed attention in converting natural gas to liquid fuels (GTL) due to their high activity, high selectivity toward heavier hydrocarbons and excellent stability in long term operation. A large number of studies in past decades have focused on developing various supported Co catalysts aimed at high productivity of heavier hydrocarbons and good stability. Al2O3, SiO2, and TiO2 supports have been commonly used since the energy crisis of the 1970s. Because the reducibility and dispersion of cobalt, which remarkably affect Fischer-Tropsch synthesis (FTS) performance of cobalt catalysts, are closely related to the interaction between the support and cobalt, appropriate reduction additives such as noble metals (e.g., Pt, Ru, and Re) or Group 11 metals (e.g., Au, Ag), structural modifiers (e.g., Zr, Ce, K, Mn), or a high Co loading (e.g., >20%) are used to overcome the strong interaction issue, consequently leading to increased productivity [1–25]. Additives may also provide additional benefits by producing hydrocarbons with a desired product spectrum, thus significantly decreasing the catalyst cost. Generally, cobalt-based catalysts are reduced at 350 ◦C under hydrogen for several hours to activate cobalt metal sites. Under this standard reduction scheme, Co/Al2O3 catalysts were found to have a limited reducibility but a relatively higher Co dispersion due to the strong interaction between the support and cobalt oxides [3–8], whereas Co/SiO2 and Co/TiO2 catalysts usually exhibited higher reducibility but with a lower Co dispersion because of the weaker interactions between the support and cobalt oxides [3,4]. Therefore, additive effect studies have been an important topic ever since the discovery of FTS in the 1920s.

The impact of noble metal additives (e.g., Pt, Ru, Re and Pd) having an atomically equivalent loading as that of 0.5% by weight Pt, has been researched carefully in order to understand the catalyst structure-performance relationships [4,6,7,16,17]. It was found that all noble metal additives significantly improved the Co reduction degree in 25%Co/Al2O3 from 55 to 68–72% and increased Co dispersion from 5.5 to 9–10% [7], which was likely through a H2 dissociation and spillover mechanism [3,14]. The increased Co site density due to the reduction promoting capability of the noble metal additives led to a near doubling of the FTS activity in comparison with the unpromoted cobalt catalyst. However, Pd was an exception; Pd addition resulted in nearly unchanged FTS activity, and it accelerated the deactivation rate of the cobalt catalyst.

The noble metal additives were also found to alter hydrocarbon selectivity in different ways. At atomically equivalent loadings to 0.5% by weight Pt and about 50% conversion, the Ru and Re additives improved C5<sup>+</sup> selectivity and suppressed methane formation, while the Pd additive prohibitively worsened selectivities by increasing the formation of methane and light hydrocarbons at the expense of losing heavier hydrocarbons [7]; Pt also tended to worsen the selectivities, but this effect was only slight. Further studies by XANES/EXAFS indicated different structures of the noble metal additives. Both Pd-Pd and Pd-Co coordination were found in the spent Pd-Co catalysts, but for the other promoted catalysts (Pt, Re, and Ru-which performed significantly better than the Pd promoted one), only coordination from the additive to Co was detected by EXAFS [6,16]. Thus, the small Pd particles or Pd patches were deemed to be responsible in part for losses in catalyst performance.

The effect of Group 11 metals (e.g., Cu, Ag, and Au with a loading range of 0.5–5.0%) on 15%Co/Al2O3 catalysts have been also investigated with the aim of (1) finding a substitute for Pt and (2) possibly lowering light gas selectivity, thus potentially reducing the costs of the cobalt catalyst and process [5]. It was found that all levels of Group 11 metals improved the reducibility of cobalt oxides (50 to 70–90%), leading to improved catalyst activity and stability especially for Ag and low levels of Au. At a similar CO conversion level of ca. 50%, the addition of less than 1.5% or 2.8% Ag not only improved the CO rate, but also improved the selectivity towards heavier hydrocarbons relative to the unpromoted cobalt catalyst (81 to 83%). However, Cu (>0.5%) and higher loadings of Au significantly poisoned the cobalt catalyst and remarkably promoted CH4 formation (9 to 21%) in comparison to the unpromoted cobalt catalyst. This could be due to Cu metal being present on the Co surface and blocking Co sites and affecting the relative hydrogenation rates. Similar to the Pd additive effect addressed previously, EXAFS results revealed only Co-Co and Me-Me (Me = Cu, Ag and Au) structures in the cobalt catalysts after the standard H2 reduction. This study suggested that Ag is a promising potential additive as a substitute for the Pt that is used for commercial cobalt catalysts due to the better FTS performance of the Ag-Co catalyst and lower price compared with Au.

The effect of up to 5% Zr structural modifier on 25%Co/Al2O3 catalysts has been studied using wide and narrow pore Al2O3 supports (Puralox HP14/150 and Catalox 150) [8]. The cobalt catalyst prepared with the wide pore support was found to perform much better than the cobalt prepared with the narrow pore Al2O3. The addition of Zr made further improvement of performance of both cobalt catalysts by facilitating Co reduction or reducing cobalt cluster size, but the Zr additive slightly increased CH4 selectivity of the cobalt catalysts supported on a wide pore alumina.

In addition to the important additives reviewed above, other additives such as V, Mg and Ce [15,18,19], K [15], Cr, Ti, Mn, and Mo [9,20,21], Nb [22] and P [23] have been also studied. These additives served as structural or electronic additives to modify cobalt catalyst behavior. Most of them at low loadings were reported to enhance cobalt reduction and dispersion, and promoted activity and selectivity toward heavier hydrocarbons for cobalt catalysts.

According to this large number of additive effect studies, some of them such as Ag and Zr showed great benefits to the performance of the cobalt catalyst and are promising steps toward a potential replacement for Pt. Efforts continue to find potential substitutes for Pt additive and to better understand structure-performance relationships. In this work, we explore the effects of the Group 12–14 elements as potential additives and based on our success with Ag, have selected the Row 5 elements Cd, In, and Sn. The effect of these additives on physiochemical properties and FTS performance of 15%Co/Al2O3 catalysts were carefully studied using various characterization and testing techniques.

#### **2. Results and Discussion**

#### *2.1. BET and Porosity Measurements*

BET and porosity results of unpromoted and 1%Cd, 1%In, and 1%Sn catalysts are summarized in Table 1. BET of Catalox SBA 150 γ-Al2O3 is 149 m2/g [6–8,16,17,25]. After loading 15% Co, BET surface area decreased to 116 m2/g. A weight % loading of 15% is equivalent to ca. 20% by weight of Co3O4. If the support is the only contributor to the surface area, then the area of the 15%Co/Al2O3 catalysts should be 0.8 <sup>×</sup> 150 m2/<sup>g</sup> <sup>=</sup> 120 m2/g, which is close to the experimental values listed in Table 1. Thus, no significant decrease in surface area was observed due to pore blocking by Co oxide particles. The addition of 1%Cd, 1%In, and 1%Sn further decreased the BET of SBA 150 supported catalysts to 110, 111 and 115 m2/g, respectively. The pore volume (0.31 cm3/g) and the pore size (radius = 4.7 nm) remained nearly unchanged for the unpromoted and the three promoted 15% Co/Al2O3 catalysts, which were lower than those of the SBA-150 support (0.5 cm3/g and 5.4 nm from [8]). The results further suggest that the addition of Co and the Cd, In and Sn additives did not significantly block the pores of the alumina support.

**Table 1.** Results of BET surface area and H2 chemisorption pulse re-oxidation.


#### *2.2. TPR and Hydrogen Chemisorption*/*Pulse Reoxidation*

TPR profiles of the unpromoted and 1%Cd, 1%In, and 1%Sn cobalt catalysts are presented in Figure 1. Two peaks occurred at the temperature ranges of 220–380 ◦C and 400–700 ◦C, which represent the standard two-step reduction of cobalt: Co3O4 + H2 <sup>→</sup> 3CoO + H2O and 3CoO + 3H2 <sup>→</sup> Co0 + 3H2, where the second step consumes three times as much hydrogen as the first step. The broad high temperature peak for the reduction of CoO to Co<sup>0</sup> indicated strong interactions between CoO and the alumina support. It is interesting that the addition of 1%Cd, 1%In and 1% Sn did not significantly change the temperature for the first reduction step of Co3O4 to CoO, and the peak temperature for the unpromoted and three promoted cobalt catalysts remained at 320–330 ◦C (Figure 1). However, the addition of 1% Cd shifted the second reduction peak to lower temperature by 80 ◦C relative to the unpromoted 15%Co/Al2O3 catalyst (540 vs. 460 ◦C), and thus, as the atomic number of the additive is increased, the broad reduction peak occurs at higher temperatures, e.g., 535 ◦C for In and 570 ◦C for Sn, which indicated that only Cd and, to a lesser extent, In, facilitated reduction of CoO; on the other hand, Sn addition slightly hindered the reduction of CoO.

**Figure 1.** TPR profiles of, moving upward, 15%Co/Al2O3, 1%Cd-15%Co/Al2O3, 1%In-15%Co/Al2O3, and 1%Sn-15%Co/Al2O3. Catalysts prepared by SPI using 150 m2/g γ-Al2O3.

H2 chemisorption and oxygen titration results are also shown in Table 1. The cobalt reduction degree increased from 48.5% to 57–59% for the 1%Cd and 1% In promoted cobalt catalysts, but it slightly decreased to 46% for the 1%Sn promoted catalyst, as compared to that of the unpromoted catalyst. This could be due to the formation of cobalt-additive coordination, as this was shown for the noble metals by EXAFS and suggested for the Group 11 metal Cu based on chemisorption results [4–6,16,17]. The reduction results as determined by H2 chemisorption pulse re-oxidation are in agreement with the TPR results above. In terms of the TPR and H2 chemisorption/pulse oxidation results, it is likely that the Cd, In and Sn additives are highly dispersed on the catalyst surface, with a fraction strongly interacting with the support and resulting in little reduction (consistent with the XANES results in the next section which showed the presence of oxides), while a fraction was likely metallic and/or coordinated with Co to increase cobalt reduction. Interestingly, the hydrogen chemisorption/pulse reoxidation results show that all Cd, In, and Sn additives lowered cobalt site densities compared with the unpromoted catalyst, with H2 desorption amounts decreased from 57 to 37–53%. Assuming the desorption of hydrogen comes only from the surface of Co0, the "apparent" average cluster size is increased for the Cd, In and Sn promoted catalysts (11 to 14–16 nm). However, more likely, the decreased cobalt site density by Cd, In and Sn could arise by (1) the additives being located on the cobalt surface and covering some cobalt sites, or (2) blocking of pores. However, the BET results discussed above tend to rule out the latter explanation. That is, it is likely that the Co particles remained the same size but that Cd, In and Sn addition blocked surface sites.

#### *2.3. NH3 TPD Study of Cd, In and Sn Promoted Catalysts*

The NH3-TPD profiles of the unpromoted and the Cd, In and Sn promoted 15%/Al2O3 catalysts are shown in Figure 2. Two types of acidic sites (weak and mild to strong) are clearly identified with the corresponding NH3 desorption peaks present at ca. 90 ◦C and 200–260 ◦C, respectively. For the unpromoted cobalt catalyst, the first NH3 desorption peak is more intense than the second one, indicating the unpromoted cobalt catalyst surface was mainly covered by the weak acid sites (likely L sites). In the case of adding 1%Cd, the NH3 desorption amounts in both peaks increased greatly, but the first peak area exhibited a greater change, implying that addition of Cd resulted in a greater fraction of weak acid sites. However, for the In-promoted cobalt catalyst, the first NH3 desorption peak shows only a slight increase in the intensity, but the second one became more intense; moreover, the second peak slightly shifted to lower temperatures (i.e., 240 vs. 260 ◦C). Thus, addition of 1%In primarily promoted the strong acid sites. Interestingly, the Sn promoted cobalt catalyst displayed a similar NH3-TPD profile to that of the In promoted cobalt catalyst, suggesting a similar acid site density (e.g., B-sites) on the Sn and In promoted cobalt catalysts. However, the second peak moved to even lower temperatures relative to the Cd promoted and unpromoted cobalt catalysts. It can be observed that the higher the atomic number, the lower the peak temperature (i.e., 220 ◦C for the In-Co and 200 ◦C for the Sn-Co catalysts). The total amount of NH3 desorbed follows the order: Cd (119.3 μmol/g) > In (110.5 μmol/g) > Sn (104 μmol/g) > unpromoted (73.2 μmol/g).

#### *2.4. XRD Study of Unpromoted and Cd, In and Sn Promoted Cobalt Catalysts*

The XRD spectra of the four reduced cobalt catalysts are depicted in Figure 3. The unpromoted and the Cd, In and Sn promoted cobalt catalysts show similar XRD patterns in the 2θ range of 30◦–70◦. Five intense reflections are observed at 2θ of 36.9◦ and 61.9◦, 42.9◦, and 45.9◦, 67◦, representing the characteristic peaks of cobalt oxide, Co, and alumina, respectively. Interestingly, regardless of whether the unpromoted catalyst was used, or whether additives were used, the intensities for each peak for the four different cobalt catalysts are quite similar, suggesting that average cobalt cluster sizes are similar. Based on the Scherrer equation, cobalt cluster size for the unpromoted and Cd, In and Sn calculated at the 2θ of 42.9◦ are 12.7 nm, 11.4 nm, 12.5 nm and 11.1 nm, respectively, which suggests that addition of Cd, In and Sn on the cobalt catalyst only slightly decreased the cobalt cluster size, if at all. This result is not consistent with the H2-chemisorption results, which suggest an apparent increased Co cluster size with the addition of Cd, In and Sn. The discrepancy strongly suggests that the additives cover some surfaces of cobalt nanoparticles or block pores of the support. However, the latter reason can be excluded based on the BET results as discussed in Section 2.1.

**Figure 2.** NH3–TPD profile of unpromoted and Cd, In and Sn promoted 15%Co/Al2O3 catalysts.

**Figure 3.** XRD patterns unpromoted and Cd, In and Sn promoted 15%Co/Al2O3 catalysts.

#### *2.5. XANES Study of Cd, In and Sn Promoted Catalysts*

To explore the electronic structure of the Cd, In, and Sn additives by XANES, 1%Cd, 1%In and 1%Sn supported on high cobalt loading (25%) catalysts were used because the additives had a greater effect on the more highly loaded Co catalysts. After the catalysts were reduced ex-situ, the cobalt catalysts were transferred to a 1L-CSTR under inert gas where 3000 Polywax was previously charged and melted. In this way the in-situ state of the Cd, In and Sn additives in the catalyst following H2 activation were able to be reviewed by XANES. Figure 4 shows the XANES spectra of 1%Cd-25%Co/Al2O3 (left), 1%In-25%Co/Al2O3 (middle) and 1%Sn-25%Co/Al2O3 (right), respectively, along with the spectra of

reference Cd, In and Sn metal foils. The results show that metal oxides were observed for the three additives after activation. This result is different from the results of the noble metal additives Pt, Pd, and Ru [6,16,17] and Group 11 metals (Cu, Ag, and Pt) studied previously [4–6,17], In which only the metallic state rather than the oxidized state was found for the additives by XANES/EXAFS.

**Figure 4.** XANES spectra of 1%Cd-25%Co/Al2O3 (**left**), 1%In-25%Co/Al2O3 (**middle**) and 1%Sn-25%Co/Al2O3 (**right**) compared to the reference Cd, In and Sn metal foils.

#### *2.6. E*ff*ect of Cd, In and Sn Additives on Fischer-Tropsch Synthesis*

The effects of 1%Cd, 1%In and 1%Sn additives and time on CO rate, CH4, C5<sup>+</sup> and CO2 selectivities are shown in Figure 5a–d and Table 2. The promoted catalysts were running at 20–30% CO conversion in the first ca. 50 h; afterwards, CO conversion was adjusted to the 40–50% level in order to better compare catalyst selectivities. The CO rate for the unpromoted catalyst was 0.022 mol/g-cat/h. The addition of 1%Cd, 1%In or 1%Sn additive drastically decreased the catalyst activity to 0.0015, 0.004 and 0.003 mol/g-cat/h, respectively, corresponding to very high rate loss percentages of 93%, 82% and 86% for the Cd, In and Sn additives, respectively. After adjusting CO conversion to 40–50% level after 50 h, the unpromoted cobalt catalyst and 1%Cd promoted catalyst displayed better stability, and the CO rate over the next time period of 100–150 h remained at ca. 0.017 mol/gcat/h for the unpromoted catalyst and ca. 0.0015 mol/gcat/h for the Cd promoted catalyst. However, the In and Sn promoted catalysts were slowly deactivating with time, with CO rate changes from 0.004 to 0.0028 mol/gcat/h and from 0.003 to 0.0025 mol/gcat/h for the In and Sn promoted catalysts, respectively. The significant decreases in cobalt catalyst activity caused by the additives are not consistent with the changes in the H2 chemisorption results as discussed in Table 1. In re-examining the H2 desorption results, the H:Co ratio is assumed to be 1:1, where "Co" refers to surface Co<sup>0</sup> atoms. The addition of 1%Cd, 1%In and 1%Sn only led to cobalt site density decreases of 8.2%, 20.4% and 36.1%, respectively. Thus, it was expected to have similar activity losses percentages for the promoted catalysts, since the change in FT activity has been generally consistent with changes in H2 chemisorption capacities of cobalt catalysts resulting from different loadings of Pt, Re, Ru, Ag, and Au, and even Zr additives [5–8,16,17]. The activity of the Cd, In and Sn promoted catalysts being 3 to 10 times lower than the expected results strongly suggests that other reasons may account for the lowering in catalyst activity. During an investigation of Group 11 additives (i.e., the coinage metals–Cu, Ag, and Au), while all the additives facilitated cobalt oxide reduction, only the Ag and Au additives increased the catalyst activity on a per g of catalyst basis. All the catalysts had higher metal site densities by H2-TPD relative to the unpromoted 15%Co/Al2O3 catalyst, but Cu decreased catalyst activity; however, the difference in the adverse effects of the Cu additive were much less pronounced than that of the Cd, In and Sn additives in this study. Thus, H2-TPD only reports metal site density, and in the case of Cu, Cu0 was likely on the surface of the cobalt particles so that, while it on the one hand promoted reduction of cobalt oxides, it decreased the cobalt surface site density by blocking sites on the surface. However, BET results do not suggest pore blocking by cobalt and the additives.


**Table 2.** Activity and selectivity of unpromoted and Cd, In, Sn and Pt promoted 15%Co/Al2O3 catalysts (a).

(a) Reaction conditions: 220 ◦C, 2.2 MPa, and H2/CO = 2.1.

Another possibility, however, is that the catalysts have the active site density as measured by chemisorption after activation, but that the site density is decreased due to reoxidation once FTS is started. A previous study of Co catalysts in an in-situ EXAFS/XANES flow cell by Huffman et al. [24], reported that Co catalysts promoted with K were much more susceptible to reoxidation (i.e., even at low conversion) compared to catalysts having no K, which only oxidized under high H2O partial pressure at high conversion. However, carefully examining the FTS activity data as shown in Figure 5a, the Cd, In and Sn promoted cobalt catalysts did not deactivate in the first 50 h; instead, the activities slowly increased with time for all cases. Therefore, rapid reoxidation of cobalt particles facilitated by Cd, In and Sn additives is only a possible explanation if it occurred prior to measurement of the first point.

According to the above discussion, blocking of pores by the additives can be excluded; thus, it was not a cause for the unexpected low activity of the promoted cobalt catalysts; the fast oxidation of cobalt particles is uncertain. The formation of M-Co (M = Cd, In and Sn) coordination or alloying can also be ruled out, because if this were the case, the H2 chemisorption/pulse re-oxidation results would show similar large differences between the unpromoted and the promoted cobalt catalysts-i.e., ~90%, but this was not observed. Furthermore, the XRD experiment did not show Co-M alloys peaks. Based on the XRD results, addition of Cd, In and Sn on the cobalt catalyst should have led to slightly increased cobalt site density by decreasing cobalt cluster size; thus, the most likely reason that explains the unexpected low activity for the Cd, In and Sn promoted catalysts is that the additives covered cobalt sites and poisoned the surfaces of the cobalt catalysts. Larger Sn atoms might contribute more significantly to cobalt site poisoning, resulting in the lowest catalyst activity. However, an electronic effect of the additives such as Cd, In and Sn on catalyst performance cannot be excluded. Additional study is needed to clarify the assumption.

From Table 2 and Figure 5b,c, CH4 selectivity for the unpromoted 15%Co/Al2O3 catalyst is 7.4% at 50% CO conversion, but it increased dramatically to about 10.5%, 13.8% and 13.5% for the 1%Cd, and 1%In or 1%Sn promoted catalysts, respectively. This caused corresponding drops in C5<sup>+</sup> selectivity to ca. 82.7 %, 75.3, and 76.6%. Thus, the In and Sn additives have a greater impact on increasing CH4 and suppressing heavier hydrocarbon formation relative to Cd. The greater increase in CH4 and light hydrocarbon selectivities for the In and Sn promoted cobalt catalysts than that of Cd is

likely associated with a greater density of strong acid sites on the two cobalt catalysts as determined by NH3-TPD (Figure 2). This conclusion is further evidenced by the In and Sn promoted cobalt catalysts having similar amounts of strong acid sites (located after 200 ◦C) and displaying essentially the same hydrocarbon selectivities at about 50% CO conversion, i.e., CH4 selectivity 13.5–13.8%, C2–C4 selectivity 10–11% and C5<sup>+</sup> selectivity 74–75%. It is likely that the strong acid sites on the surface of cobalt promoted H2 adsorption and promoted methane and light hydrocarbon formation.

The addition of 1%Cd, 1%In or 1%Sn additives also led to significantly increased CO2 selectivity (0.5 to 2–5%). The significant changes in catalyst selectivity also indicated that species other than metallic Co are present in the catalyst, since metallic cobalt does not possess intrinsic water-gas shift (WGS) activity. Note that higher WGS activity leads to higher methane selectivity, since WGS promotes the formation of hydrogen. The XANES results showed that the Cd, In and Sn additives were present in oxidized form after reduction. It is well known that a synergy between a partially reducible oxide and a metal results in WGS activity, and this might explain the higher CO2 selectivity during FTS.

The effects of 1%Cd, 1%In and 1%Sn additives and time on the contents of propylene, propane, 1-butene, 2-butene, total butene and butane are shown in Figure 6a–f, respectively. Table 3 also summarized mean values of these parameters at different time ranges. At 40–50% CO conversion level, C3 olefin content for the unpromoted 15%Co/Al2O3 catalyst was 60.7%. Doping 1%Cd, 1%In and 1%Sn to the catalyst decreased C3 olefin content to 58.6%, 53.9%, and 46%, respectively, but C3 paraffin content increased to 41.4%, 46.1% and 48.9% from 39.3% (Table 3, Figure 6a,b). Moving to C4 hydrocarbons, precisely the same trend is observed at 40–50% CO conversion. The Cd, In and Sn additives resulted in decreases in 1-C4 olefin selectivity to 45.4%, 40.1% and 25.6%, respectively, from 48.1% as compared to the unpromoted catalyst. This also led to a measurable increase in 2-C4 olefin content to 5.2%, 5.5%, 10.7% from 3.5% and C4 paraffin contents to 49.4%, 54.4%, 63.9% from 48.4%, respectively (Table 3 and Figure 6c–f). The results are interesting, as they clearly indicate that the Cd, In and Sn additives enhanced the hydrogenation and isomerization of 1-olefin reactions on the cobalt catalysts. The extent of these secondary reactions increased with increases in the atomic number of the additive. The Sn apparently is the most effective one among all three additives to greatly increase the hydrogenation and isomerization reactions. As discussed in terms of the TPR and XANES results, a fraction of the additives might be in a metallic state, while the remaining additive may be highly dispersed and strongly interacting with the support or cobalt surface, which are in oxide form and/or Co-M (M=Cd, In and Sn) coordinated states. Thus, it is postulated that addition of Cd, In and Sn to the cobalt catalyst likely creates new acid sites on catalyst surface, leading to a higher activity of secondary reactions of olefins relative to the unpromoted cobalt catalyst. This hypothesis is consistent with the NH3-TPD results as discussed in Section 3.3, which indicated a greater abundance of mild acid sites on the Sn-Co catalyst occurring at low temperature (200 ◦C) relative to the Cd and In promoted cobalt catalysts. Much higher 2-C4 olefin content for the Sn-Co catalyst suggests that mild acid sites are more active for the secondary reaction of olefins.


**Table 3.** Olefins and paraffins contents of unpromoted and Cd, In, Sn and Pt promoted 15%Co/Al2O3 catalysts (a).


**Table 3.** *Cont*.

Reaction conditions: 220 ◦C, 2.2 MPa, and H2/CO = 2.1. C4 olefin selectivity, % = 100 × rates of all C4 olefins/rates of all C4 hydrocarbons; 1-C4 olefin selectivity, % = 100 × rate of 1-C4 olefin/rates of all C4 hydrocarbons; 2-C4 olefin selectivity, % = 100 × rate of 2-C4 olefin/rates of all C4 hydrocarbons.

In our previous studies, Pd and Pt were found to increase CH4 (8–12%) and suppress heavier hydrocarbon formation (83–76%); and, Pd displayed much higher hydrogenation and isomerization activities (1-C4 olefin: 47–25%, 2-C4 olefin: 7–14.8%) [7,16]. Thus, the impact of Sn on the formation of olefins and paraffins resembles that of Pd.

**Figure 5.** Change in (**a**) CO rate, (**b**) CH4 selectivity, (**c**) C5<sup>+</sup> selectivity and (**d**) CO2 olefin selectivity over unpromoted and Cd, In and Sn promoted 15%Co/Al2O3 catalysts. Reaction conditions: 220 ◦C, 2.2 MPa, H2/CO = 2.1, XCO = 40–50%.

**Figure 6.** Change of (**a**) C3 olefin selectivity, (**b**) C3 paraffin selectivity, (**c**) 1-C4 olefin selectivity, (**d**) 2-C4 olefin selectivity, (**e**) total-C4 olefin selectivity and (**f**) C4 paraffin selectivity with time over npromoted and Cd, In and Sn promoted 15%Co/Al2O3 catalysts. Reaction conditions: 220 ◦C, 2.2 MPa, H2/CO = 2.1, XCO = 40–50%.

The Cd, In and Sn elements used as additives to modify FTS cobalt catalyst performance have been scarcely reported. However, some studies have employed them (e.g., as bimetallics such as Pt-In, Pt-Sn) for other associated reactions. Cho et al. [26] reported that Sn supported on a mesoporous zeolite (3Dom-I MFI) offered significant improvements for the isomerization of C5 and C6 sugars such as xylose and glucose. Srinivasan et al. [27] found that Sn at different loadings changed the activity and aromatics selectivity of Sn-Pt/Al2O3 for n-octane conversion. Passos et al. [28] studied In-Pt/Al2O3, and Sn-Pt/Al2O3 bimetallic catalysts for cyclohexane dehydrogenation, methylcyclopentane hydrogenolysis, and *n*-heptane conversion. It was found that after the catalyst was reduced by 1.5%H2/Ar at 500 ◦C for 30 min., 50–80% In and 25–50% Sn was in a zero-valent state in the bimetallic system. During

the methylcyclopentane hydrogenolysis and *n*-heptane conversion reactions at 500 ◦C, the In and Sn additives were reported to decrease the activity of the Pt/Al2O3 catalyst, which was explained by the In and Sn additives diluting Pt active sites on catalyst. Furthermore, addition of In or Sn led to decreases in the selectivity of hydrogenolysis, and an increase in the selectivity for dehydrogenation and aromatization products, but Sn was reported to greatly enhance the isomerization activity for the Pt/Al2O3 catalyst in the conversion of *n*-heptane. Coleto et al. [29] studied the transformation of 1-pentene, and 1-hexene over bimetallic Pt-Re/Al2O3, Pt-Sn/Al2O3 and Pt-Ge/Al2O3 catalysts. The Pt-Sn/Al2O3 catalyst was also reported to have high hydrogenation activity to n-pentane at low temperature (200 ◦C), while high isomerization activity was observed at the expense of hydrogenation at a high temperature of 500 ◦C, which is consistent with the study of Passos et al. [28]. The investigation showed by comparing with the result of the Al2O3 support that the double bond shift and skeletal isomerization of olefins are both acid-catalyzed reactions, while hydrogenation sites are metallic in nature. Thus, the hydrogenation sites and isomerization likely changed with temperature, with higher temperatures yielding more acid sites for the Pt-Sn catalyst. Mazzieri et al. [30] reported the same role of Sn in increasing isomerization activity for the reaction of naphtha reforming over a trimetallic Pt-Re-Sn catalyst supported on chlorided Al2O3. In this study, the improvements in secondary reactions of 1-olefins observed with the addition of Group 12, Group 13 and Group 14 elements (i.e., Cd, In and Sn) are consistent with the studies of Coleto et al. [26] and Passos et al. [28], which suggests that the additives not only boosted the hydrogenation rate of olefins, but they also produced new acid sites for isomerization, for example possibly as MOx, M-Co (M = Cd, In and Sn).

#### **3. Experimental**

#### *3.1. Catalyst Preparation*

The unpromoted 15%Co/Al2O3 catalysts and 1% Cd, 1% In and 1%Sn promoted cobalt catalysts were prepared by the slurry phase impregnation (SPI) method, as previously described in [3–8,16,17,25]. Catalox SBA 150 -Al2O3 was used as the catalyst support. The salts used for the Cd and In were nitrates, while SnCl2 served as the salt for Sn. The additives were incorporated into the catalyst by incipient wetness impregnation (IWI) following the addition of cobalt. The catalysts were calcined in air for 4 h at 350 ◦C.

#### *3.2. BET Surface Area and Porosity Measurements*

A Micromeritics 3-Flex system was used to measure BET surface area and porosity characteristics. Before testing, the temperature was slowly increased to 160 ◦C; then, a vacuum was pulled for at least 12 h until the sample pressure was approximately 50 mTorr. The BJH method was employed to determine the average pore diameter and pore volume. Additionally, pore size distribution was obtained as a function of pore diameter via the correlation dV/d(log D).

#### *3.3. Temperature Programmed Reduction*

Temperature programmed reduction (TPR) was carried out with a Zeton Altamira AMI-200 unit (Pittsburgh, PA, USA) with a flow rate of 30 cm3/min of 10%H2/Ar. The heating rate was 5 ◦C/min from 50 ◦C to 1100 ◦C, with a final 30 min hold.

#### *3.4. Hydrogen Chemisorption*/*Pulse Reoxidation*

Chemisorption with hydrogen was conducted in a Zeton Altamira AMI-200 unit using a thermal conductivity detector (TCD). The mass of the catalyst was ~0.220 g. Each catalyst was reduced at 350 ◦C for 10 h in 30 cm3/min of 33%H2 in He and the temperature was decreased to 100 ◦C. Argon was flowed to remove any physisorbed species and the temperature was increased back to 350 ◦C in argon to desorb the chemisorbed hydrogen. The temperature programmed desorption peaks were integrated and the # of moles of hydrogen evolved was calculated.

After TPD of hydrogen, pulses of pure O2 in He were sent to oxidize the catalyst. The extent of reduction was determined on the assumption that Co0 reoxidized to Co3O4. Uncorrected % dispersion assumes (erroneously) complete reduction while corrected % dispersion includes the extent of reduction:

%*Duc* <sup>=</sup> (# of Co0 atoms on surface <sup>×</sup> 100%)/(total # Co atoms)

%*Dc* <sup>=</sup> (# of Co<sup>0</sup> atoms on surface <sup>×</sup> 100%)/[(total # Co atoms)(fraction reduced)]

#### *3.5. NH3 Temperature Programmed Desorption*

A microreactor loaded with ca. 200 mg of powder catalyst was employed to analyze the acid sites on the surface of the catalyst by means of NH3-TPD. A total flow rate of 50 sccm was used for NH3 adsorption and desorption with effluent gases being analyzed with a quadrupole mass spectrometer (QMS 200, Pfeiffer Vacuum, Asslar, Germany). The catalysts were first reduced at 350 ◦C in a flow of 10% H2 in He for 2 h, followed by the gas mixture being replaced with pure He for purging and cooling prior to NH3 adsorption at 40 ◦C. When the catalyst was saturated with NH3 by flowing 2000 ppm NH3 in N2 (ca. 30 min), the He flow was switched back again for purging at 40 ◦C in order to remove weakly adsorbed NH3 (ca. 30 min). The temperature-programmed desorption was then carried out in He at 50 cc/min using a ramp rate of 10 ◦C/min up to 600 ◦C.

#### *3.6. X-ray Absorption Near Edge Spectroscopy (XANES)*

X-ray absorption near edge spectroscopy was carried out using transmission mode in the vicinity of the Cd, In, and Sn K-edges at the National Synchrotron Light Source (NSLS) at Brookhaven National Laboratory, Upton, New York, Beamline X18-b. The beamline utilized a Si (111) channel-cut monochromator. The catalysts were prepared and activated in the same way as if conducting a reaction test except that following treatment in hydrogen, the catalyst was cooled so that it became fixed in the solid startup wax. Samples were made into self-supporting disks. XANES spectra were analyzed using WinXAS software [31] by comparing the spectra qualitatively once normalized.

#### *3.7. X-ray Di*ff*raction (XRD)*

X-ray diffraction (XRD) on powder samples was performed for freshly reduced cobalt catalysts at room temperature using a Rigaku Diffractometer (DMAX-B, Tokyo, Japan) operating with Cu Kα radiation (1.54 Å), In order to identify cobalt structure and cluster size. All cobalt catalyst samples were reduced at 350 ◦C by 25%H2/He for 15 h followed by passivation using 1%O2/N2 gas mixture prior to conducting the XRD measurement.

#### *3.8. Catalytic Activity Testing*

Catalyst reaction tests were carried out using a continuously stirred tank reactor (CSTR, PPI, Fort Worth, TX, USA)) that makes use of a mag drive stirrer with turbine impeller, gas-inlet outlet lines with a stainless steel (SS) fritted filter (7 μm) placed outside the reactor. To withdraw wax, a stainless steel tube with a 2 micron fritted filter was placed below the liquid level of the reactor. Mass flow controllers controlled the H2 and CO flow rates. Reactant gases were thoroughly mixed prior to the reactor. CO was scrubbed of iron carbonyls using lead oxide-alumina. Reactants entered the CSTR below the impeller, which had a stirring speed of 750 rpm. Temperature was well controlled by a temperature controller.

The amount of catalyst used was 12–18 g in the size range of 45–90 μm. The catalyst was first reduced ex-situ in a tubular reactor at 350 ◦C at 1 atm for 15 h using a gas mixture of H2/He (60 NL/h) with a volume ratio of 1:3. Reduced catalyst was transferred by forcing the catalyst out with N2 to the CSTR, which held 315 g of melted Polywax 3000. The reactor was weighed prior to and following catalyst transfer. The transferred catalyst was further exposed to pure hydrogen (30 NL/h) for another 10 h at 230 ◦C to ensure reduction, prior to commencing FTS.

FTS reaction conditions were 220 ◦C, 2.2 MPa, H2/CO = 2.1. Space velocity was controlled to achieve a CO conversion of 50%. Reaction products were continuously removed from the reactor head space and sent to two collection vessels, a trap maintained at 100 ◦C and a trap held at 0 ◦C. Uncondensed vapor was decreased to atmospheric pressure. Gas flow was measured by a wet test meter and the gas was analyzed using online gas chromatography. Accumulated liquids in the CSTR were removed daily through a 2 micron sintered metal filter. CO conversion was determined on the basis of GC data using a micro-GC equipped with thermal conductivity detectors. Wax, oil and the water phase products were also collected and analyzed by three different gas chromatographs. To investigate the effect of Group 12–14 elements, Cd, In and Sn were selected from Row 5 based on our earlier success with Ag as an additive. The activity and product selectivities (e.g., CH4, C5+, CO2, 1-olefin, 2-olefin and paraffin) of unpromoted 15%Co/Al2O3 catalysts and 1% of Cd-, In, and Snsupported cobalt catalysts were studied at a reference CO conversion of about 40–50%.

#### **4. Conclusions**

TPR and hydrogen chemisorption/pulse reoxidation results showed that only a fraction of Cd, In and Sn was reduced. NH3-TPD results indicated that addition of Cd, In and Sn promoted mild to strong acid sites, which might be responsible for the enhancement of hydrogenation of 1-olefin during the FTS. Cd and In were found to promote CoO reduction to Co0, while Sn slightly hindered it, resulting in more unreduced cobalt than the other two additives. The XANES results showed oxidized states for the Cd, In and Sn additives after the catalysts were activated at 350 ◦C by H2. The TPR, XANES and reaction results suggest that M-M, MxOy and M-Co coordination (M refers Cd, In and Sn) may be present in the catalysts. XRD results showed only a slight decrease in cobalt cluster size with addition of 1% of Cd, In or Sn.

The FTS reaction was carried out on all research catalysts at 220 ◦C, 2.2 MPa, H2/CO = 2.1 and 25–50% CO conversion using a 1-L CSTR for about 200 h. Space velocity was adjusted if needed during testing. Addition of 1% the additives resulted in 3 to 10 fold activity losses relative to the unpromoted cobalt in comparison to the expected activity losses in terms of the decreases in cobalt sites as determined by H2 chemisorption capacities. The significant catalyst activity losses were explained based on the additives covering and poisoning cobalt and possible electronic effects resulting from the interaction of the additives.

Addition of Cd, In or Sn greatly modified the selectivity of the cobalt catalyst. All the additives remarkably promoted the formation of methane, light hydrocarbons and CO2, and suppressed heavier hydrocarbon formation. However, addition of Cd, In and Sn greatly improved secondary reactions of 1-olefins. The extent of the improvement increased with increasing atomic number (Cd < In < Sn). The selectivity changes are linked with acid sites on the cobalt catalysts, which were found to be promoted by the additives, while a fraction of reduced additives in the metallic phase might improve the rate of hydrogenation during FTS. It is concluded that mild to strong acid sites on the cobalt catalysts (i.e., In-Co and Sn-Co) enhanced H2 adsorption to a greater extent and promoted methane and light hydrocarbon selectivities, while mild acid sites on the cobalt catalysts (i.e., Sn-Co), enhanced the isomerization reaction of 1-olefins to a greater extent relative to other types of acid sites on the cobalt catalyst.

**Author Contributions:** Conceptualization, W.M., G.J., B.H.D., writing—original draft preparation, W.M., G.J., writing—review and editing, J.L.S.K., B.H.D., investigation, W.M., G.J., W.D.S., Y.J., S.D.H., resources, J.L.S.K., B.H.D, S.K., supervision, G.J., J.L.S.K., B.H.D., project administration, J.L.S.K., B.H.D., funding acquisition, G.J., B.H.D., formal analysis, W.M., G.J., W.D.S., Y.J., S.D.H., S.K., data curation, W.D.S., S.K., visualization, G.J., B.H.D., validation, W.M., W.D.S., Y.J., S.D.H., reaction testing, W.M., catalyst preparation, G.J., catalyst characterization, G.J., Y.J., S.D.H., product analysis, W.D.S., synchrotron beamline operation, S.K.

**Funding:** This research was funded by NASA (grant number NNX11A175A) and the Commonwealth of Kentucky. The APC was funded by UK-CAER.

**Acknowledgments:** This paper is dedicated to the late Professor Burtron H. Davis.

**Conflicts of Interest:** The authors declare no conflict of interest.

#### **References**


© 2019 by the authors. Licensee MDPI, Basel, Switzerland. This article is an open access article distributed under the terms and conditions of the Creative Commons Attribution (CC BY) license (http://creativecommons.org/licenses/by/4.0/).

#### *Article*

### **Selective CO Hydrogenation Over Bimetallic Co-Fe Catalysts for the Production of Light Para**ffi**n Hydrocarbons (C2–C4): E**ff**ect of Space Velocity, Reaction Pressure and Temperature**

**Seong Bin Jo 1,**†**, Tae Young Kim 2,**†**, Chul Ho Lee 2, Jin Hyeok Woo 2, Ho Jin Chae 1, Suk-Hwan Kang 3, Joon Woo Kim 4, Soo Chool Lee 1,\* and Jae Chang Kim 2,\***


Received: 6 August 2019; Accepted: 16 September 2019; Published: 19 September 2019

**Abstract:** Synthetic natural gas (SNG) using syngas from coal and biomass has attracted much attention as a potential substitute for fossil fuels because of environmental advantages. However, heating value of SNG is below the standard heating value for power generation (especially in South Korea and Japan). In this study, bimetallic Co-Fe catalyst was developed for the production of light paraffin hydrocarbons (C2–C4 as well as CH4) for usage as mixing gases to improve the heating value of SNG. The catalytic performance was monitored by varying space velocity, reaction pressure and temperature. The CO conversion increases with decrease in space velocities, and with an increase in reaction pressure and temperature. CH4 yield increases and C2<sup>+</sup> yield decreases with increasing reaction temperature at all reaction pressure and space velocities. In addition, improved CH4 yield at higher reaction pressure (20 bar) implies that higher reaction pressure is a favorable condition for secondary CO2 methanation reaction. The bimetallic Co-Fe catalyst showed the best results with 99.7% CO conversion, 36.1% C2–C4 yield and 0.90 paraffin ratio at H2/CO of 3.0, space velocity of 4000 mL/g/h, reaction pressure of 20 bar, and temperature of 350 ◦C.

**Keywords:** Synthetic natural gas (SNG); Cobalt; Iron; Fischer-Tropsch synthesis; C2–C4 hydrocarbons; paraffin ratio

#### **1. Introduction**

At present, the production of synthetic natural gas (SNG), mainly consisting of methane, has aroused extensive attention and been commercially produced from different starting materials, including coal and solid dry biomass (e.g., wood and straw) [1–5]. CH4 via synthesis gas (syngas, CO + 3H2) is an effective and environmentally friendly method, because it emits the smallest amount of CO2 per energy unit among all fossil fuels. However, the heating value of CH4 is typically below the standard heating value for power generation (especially in South Korea and Japan) [6–12]. For power generation, liquefied petroleum gas (LPG, C3–C4 hydrocarbons) must be added to SNG to enhance its heating value; however, the price of LPG is strongly correlated with that of oil. In principle, synthetic light hydrocarbons (C1–C4 ranges) via Fischer–Tropsch (FT) reaction could be added to SNG

as a substitute for LPG by using the same syngas source (H2/CO ratio = 3.0) for the SNG process. Furthermore, the gas products must maintain a high paraffin ratio, because olefins exhibit a low heating value, as well as being more susceptible to hydration with CH4 and liquefaction than paraffins of the same carbon chain length under pipeline conditions (-5 ◦C, 70 bar) [13]. Therefore, the FT product gas must have a high paraffin ratio in C2–C4 ranges, as well as a high light hydrocarbon yield (CH4 and C2–C4) if it is to be used to replace LPG for power generation.

Inui et al. reported a "high calorific methanation" process using Co-Mn-Ru/Al2O3 catalyst for the production of high-calorie gas comparable to natural gas with added C2–C4 hydrocarbons [6]. The Co-Mn-Ru/Al2O3 catalyst afforded high CO conversion (98.8%) and C2–C4 selectivity (19.1%). Lee et al. elucidated the role of each component in the Co-based catalysts, and proposed the 10Co-6Mn-2.5Ru/Al2O3 and 20Co-16Mn/Al2O3 as optimum catalysts for high heating value of SNG [7]. They also developed Fe-Zn and Fe-Cu catalysts, and the Fe-based catalysts were evaluated after caburization and reduction pretreatment [8–10]. In an earlier report, bimetallic Co-Fe catalysts supported on γ-Al2O3 were developed for the production of light hydrocarbons (C2–C4 ranges) at high CO conversion [11]. It was found that the reducibility of the iron phase was enhanced in the presence of cobalt, leading to enhanced catalytic activity. Of all catalysts, 5Co-15Fe/γ-Al2O3 exhibited the highest C2–C4 paraffin selectivity at high CO conversion. The high CO conversion and similar hydrocarbon distribution of 5Co-15Fe/γ-Al2O3 compared to 20Fe/γ-Al2O3 is due to improved iron reducibility. Moreover, the effects of the H2/CO gas ratio and the reaction temperature on the catalytic performance over 5Co-15Fe/γ-Al2O3 catalyst were investigated: the FT catalyst showed high paraffinic C2–C4 selectivity and CO conversion at H2/CO = 3.0, reaction temperature of 300 ◦C and pressure of 10 bar; but this led to substantial byproduct formation, such as C5<sup>+</sup> liquid and waxy hydrocarbons and CO2. Despite the high C2–C4 yield, a considerable amount of byproducts (C5<sup>+</sup> hydrocarbons and CO2) need to be condensed or separated to be used as mixing gases in SNG for practical processing. To overcome this problem, hybrid catalysts (FT + cracking) in a double-layered bed reactor system were introduced to minimize C5<sup>+</sup> and CO2 [12]. The layer of cracking catalysts (SAPO-34 zeolite and Ni catalysts) was loaded underneath the FT catalyst (5Co-15Fe/γ-Al2O3) layer in the double-layered bed reactor system. Compared with the FT catalyst in a single-layered bed reactor, cracking catalysts (SAPO-34 and Ni catalysts) convert C5<sup>+</sup> hydrocarbons into light hydrocarbons (CH4 and C2–C4) in the double-layered bed reactor system. In addition, the Ni catalyst improved the CO conversion and reduced the CO2 yield via methanation.

At present, few studies have made an effort to improve the heating value of SNG by producing paraffinic C2–C4 hydrocarbons and minimizing byproducts (C5<sup>+</sup> and CO2). Although catalytic performance, including CO conversion and hydrocarbon distribution, is strongly dependent on the operation conditions such as space velocity, reaction pressure and temperature in the practical process, these effects were not investigated in detail. Herein, catalytic performance over bimetallic Co-Fe catalyst (5Co-15Fe/γ-Al2O3) under different reaction conditions (SV, P and T) is evaluated to determine the optimum operating conditions for the production of high paraffinic C2–C4 yield, as well as reduction of byproduct (C5<sup>+</sup> and CO2). In addition, characterization of the catalysts was performed using inductively coupled plasma optical emission spectroscopy (ICP-OES), X-ray diffraction (XRD) techniques and Brunauer-Emmett-Teller (BET) analysis.

#### **2. Results**

Table 1 shows the metal content and textural properties, such as BET surface area, pore volume, and average pore size, of support material (γ-Al2O3) and FT catalyst (5Co-15Fe/γ-Al2O3). As listed in Table 1, the metal content in FT catalyst (5Co-15Fe/γ-Al2O3) is 5.2 wt.% Co and 14.1 wt.% Fe, which is almost consistent with the intended metal loading. The γ-Al2O3 shows a BET surface area of 156.9 m2/g, pore volume of 0.23 cm3/g and average pore size of 5.9 nm, and those values decreased after impregnation of γ-Al2O3 with cobalt and iron. In addition to γ-Al2O3 (JCPDS No. 10-0425), fresh 20Co/γ-Al2O3 and 20Fe/γ-Al2O3 catalysts showed Co3O4 phase (JCPDS No. 43-1003)

and Fe2O3 phase (JCPDS No. 52-1449), respectively (Figure S1) [11]. However, the XRD peaks of Fe phase are very broad compared to those of Co3O4 phase, due to the high dispersion of iron phase on γ-alumina [11,12,14,15]. In the case of the bimetallic Co-Fe catalysts, the XRD peaks of Co3O4 decreased and those of Fe2O3 increased slightly with increasing iron-to-cobalt ratio. However, fresh 5Co-15Fe/γ-Al2O3 showed XRD patterns of CoO (JCPDS No. 48-1719) and Fe2O3 phases, indicating the incorporation of Co into Fe2O3 after calcination, as shown in Figure 1 [11,12,14,15]. On the other hand, the reduced 5Co-15Fe/γ-Al2O3 showed XRD peaks of Co metal (JCPDS No. 01-1259) and Fe metal (JCPDS No. 87-0722), respectively. The crystallite size of CoO phase is 4.7 nm, while that of Fe2O3 could not be calculated because of the too-broad peaks of Fe2O3, as mentioned above. In reduced states of the catalyst, crystallite sizes of the Fe metal are ~20 nm. Based on the H2-TPR results, reduction temperatures of cobalt phase increased, and those of iron decreased with the increasing iron-to-cobalt ratio (Figure S2). Thus, it can be concluded that the incorporation of Co into the Fe2O3 phase results in a weaker interaction between iron and alumina, which enhance the reducibility of iron species and catalytic activity (Figure S3). Therefore, the 5Co-15Fe/γ-Al2O3 catalyst was chosen as the catalyst with optimum mass ratio for the following studies.


**Table 1.** Characterization of γ-Al2O3 support material and FT catalyst (5Co-15Fe/γ-Al2O3).

<sup>a</sup> Metal contents were determined by ICP-OES. <sup>b</sup> Average pore size were measured following the Barrett-Joyner-Halenda (BJH) method. <sup>c</sup> Crystallite sizes of metal phase were calculated using the Scherer equation.

**Figure 1.** XRD patterns of the FT catalyst (5Co-15Fe/γ-Al2O3) in fresh and reduced states; (-) Co3O4, (-) CoO, () Co metal, () Fe2O3, () Fe3O4, and () Fe metal.

Figure 2 shows the CO conversion of FT catalyst (5Co-15Fe/γ-Al2O3) as a function of time on stream after reduction at 500 ◦C for 1 h under a 10% H2/N2 gas mixture. Activity tests were conducted under conditions of H2/CO = 3.0 at different reaction parameters, such as space velocity, reaction pressure and temperature. Overall, CO conversion of the FT catalyst increased with the increase in reaction temperature. Under almost all operating conditions, the FT catalyst maintains its CO conversion and yield of hydrocarbons. However, the CO conversion of the FT catalyst decreased below 10 bar, with a space velocity of 8000 mL/g/h, at 300 and 350 ◦C (Figure 2c). CO conversion decreased from 64.9 to 55.9% at 300 ◦C, and from 97.9 to 81.4% at 350 ◦C as the reaction progressed. It is well known that the deactivation of the catalysts in CO hydrogenation is due to several factors, including sintering, re-oxidation of active materials, poisoning, coke formation on the surface of active materials, etc., but its causes and effects are not elucidated in this paper. The deactivation of the FT catalysts was compensated by the high reaction temperature (Figure 2c) and pressure (Figure 2f), leading to an increase in CO conversion [16].

**Figure 2.** CO conversion of FT catalyst at different reaction temperature under 10 bar (left) and 20 bar (right) at space velocity of (**a**,**d**) 4000, (**b**,**e**) 6000, and (**c**,**f**) 8000 mL/g/h as a function of time on stream.

The results of the catalytic behavior, such as the initial CO conversion and initial hydrocarbon yield of the FT catalyst, are shown in Figures 3 and 4, and summarized in Table 2. Figure 3 shows the initial CO conversion of the FT catalyst as a function of reaction temperature. As shown in Figure 3, the initial CO conversion of the catalyst increased dramatically to 300 ◦C, and then increased slightly above 350 ◦C at all pressures and space velocities. Furthermore, it was also found that the CO conversion increased with the increase in reaction pressure, and with the decrease in space velocity. At high temperature (≥300 ◦C), CO conversion appears to be largely independent of space velocity and reaction pressure. It was reported that CO conversion initially increased dramatically, and then decreased with increasing reaction temperatures (>400 ◦C) [17,18]. In addition, CO conversion increased with increase in reaction pressure and decrease in space velocity [18,19]. However, space velocity and reaction pressure have little effect on CO conversion, because high reaction temperature (≥300 ◦C) has a significant influence on CO conversion.

**Figure 3.** The initial CO conversion of the FT catalyst (5Co-15Fe/γ-Al2O3) under (**a**) 10 and (**b**) 20 bar at different space velocities as a function of reaction temperature.

**Table 2.** Summarization of catalytic performance over FT catalyst (5Co-15Fe/γ-Al2O3) at different space velocity, reaction pressure and temperature.


**Figure 4.** The initial yield of (**a**) CH4, (**b**) C2–C4, (**c**) C5<sup>+</sup> and (**d**) CO2 over the FT catalyst (5Co-15Fe/ γ-Al2O3) at different space velocity as a function of reaction temperature at (I) 10 bar and (II) 20 bar.

Figure 4 shows the initial product yields for (a) CH4, (b) C2–C4, (c) C5<sup>+</sup> and (d) CO2 under high CO conversion conditions (≥300 ◦C). The effects of reaction parameters such as space velocity, reaction pressure and temperature on product distribution are strongly dependent on the secondary reactions of primary products including hydrogenation, reinsertion, hydrogenolysis, and isomerization [19,20]. It is well known that CH4 increases and C5<sup>+</sup> hydrocarbons decreases with increasing space velocity under typical FTS conditions. In addition, an increase in reaction temperature shifts the hydrocarbon distribution towards light hydrocarbon products, whereas an increase in reaction pressure shifts the hydrocarbon distribution to heavier products in typical FTS reaction [19]. As shown in Figure 4a, CH4 yield increases with the increase in reaction temperature under all reaction pressures and space velocities. In contrast to typical FTS conditions, improved CH4 yield at high reaction pressure (20 bar) indicates that the higher reaction pressure is favorable for secondary CO2 methanation reaction, as discussed below in Figure 4d [21–24]. In addition, space velocity did not affect methane yield at relatively low temperature (300 and 350 ◦C), whereas CH4 yield increases with the increase in space velocity at 400 ◦C. As shown in Figure 4b, C2–C4 yield reached the highest values at 300 or 350 ◦C, and then decreased at higher reaction temperatures. Furthermore, reaction pressure enhanced C2–C4 yield, and space velocity is inversely proportional to C2–C4 yield. In the case of C5<sup>+</sup> hydrocarbon yield, although it is difficult to confirm the effects of reaction parameters, C5<sup>+</sup> yield decreased with the increase in reaction pressure and temperature (Figure 4c). As shown in Figure 4d, CO2 yield exhibited almost the same values (ca. 20%) at reaction temperatures between 350 and 400 ◦C. This is due to the fact that the increase in CO2 production with increasing CO conversion can mainly be attributed to increased water gas-shift reaction (WGS, CO + H2O ↔ CO2 + H2) at high water partial pressures [11,25]. At 20 bar, however, it is notable that CO2 yield decreases with the increase in reaction temperature, despite almost 100% CO conversion at all temperatures. These results show that higher reaction pressure and temperature are favorable conditions for secondary CO2 methanation reaction.

Figure 5 shows the yield of light hydrocarbons in C1–C4 range and ratio of C2–C4 to C1–C4 yield under different reaction parameters, such as space velocity, reaction temperature and pressure. At 10 bar, sum of CH4 and C2–C4 yield increased (45–53 to 63–66%), and ratio of C2–C4 to C1–C4 yield decreased (0.5–0.55 to 0.25–0.35) with the increase in reaction temperature between 300 to 400 ◦C, as shown in Figure 5a. The sum of CH4 and C2–C4 yield decreased from 53 to 45% at 300 ◦C, but these values did not change a lot at higher reaction temperature (350 and 400 ◦C) with an increase in space velocity. In addition, the ratio of C2–C4 to C1–C4 yield decreased from 0.35 to 0.25 at 400 ◦C, but its value remained constant (ca. 0.5–0.55) at lower temperature (300 and 350 ◦C). At 20 bar, the sum of CH4 and C2–C4 yield increased (58–63% to 73–75%), and the ratio of C2–C4 to C1–C4 yield decreased (0.49–0.55 to 0.30–0.35), with the increase in reaction temperature between 300 to 400 ◦C, as shown in Figure 5b. The sum of CH4 and C2–C4 yield remained constant at all reaction temperatures between 300 and 400 ◦C. In addition, the ratio of C2–C4 to C1–C4 yield decreased from 0.35 to 0.30 at 400 ◦C, but its value remained constant (ca. 0.5–0.55) at lower temperature (300 and 350 ◦C). Overall, the yield of light hydrocarbons in C1–C4 range increased with increasing in reaction temperature between 300 to 400 ◦C, because CH4 yield increases dramatically and C2–C4 yield decreases slightly, resulting in reduction of (C2–C4)/(C1–C4) ratio. With increasing in space velocity, on the other hand, the sum of CH4 and C2–C4 yield and (C2–C4)/(C1–C4) ratio decreased slightly. Furthermore, high reaction pressure enhanced the light hydrocarbon (C1–C4) yield and the (C2–C4)/(C1–C4) ratio. At both 10 and 20 bar, a space velocity of 4000 mL/g/h and 350 ◦C are considered to be appropriate conditions for high calorific methanation, since the 5Co-15Fe/γ-Al2O3 catalyst exhibit the highest light hydrocarbon (CH4 and C2–C4) yield and (C2–C4)/(C1–C4), respectively.

**Figure 5.** C1–C4 yield and (C2–C4)/(C1–C4) ratio of the FT catalyst (5Co-15Fe/γ-Al2O3) as a function of space velocity at (**a**) 10 bar and (**b**) 20 bar

Figure 6 shows the paraffin ratio (P/(P+O)) with different reaction parameters, such as H2/CO ratio, space velocity, reaction pressure and temperature, where P and O represent the yields of paraffins and olefins in the C2–C4 range, respectively. As shown in Figure 6a, the paraffin ratio increased with increasing H2/CO ratio at a space velocity of 6000 mL/g/h, although this effect is not noticeable at H2/CO ratios above 2.0, due to the adjustment of H2/CO ratio by WGS conditions [11,25]. In addition, it was found that the paraffin ratio showed a positive correlation with CO conversion. The fact that the paraffin ratio increases with the CO conversion suggests that increasing the H2/CO ratio and temperatures leads to a second hydrogenation of the olefins into paraffin and an increase of the CO conversion. As shown in Figure 6b, space velocity, reaction pressure and temperature have little effect on paraffin ratio at H2/CO ratio of 3.0, but CO conversion is strongly dependent on reaction temperature. Furthermore, it is found that a higher reaction pressure (20 bar) improved the CO conversion above a reaction temperature of 300 ◦C, because higher reaction pressure enhanced hydrogen adsorption on support of the catalyst and improved second hydrogenation [19].

**Figure 6.** Paraffin ratio of the FT catalyst (5Co-15Fe/γ-Al2O3) as a function of CO conversion at different operating conditions: (**a**) effect of H2/CO ratio at SV of 6000 ml/g/h, and (**b**) effect of reaction pressure at H2/CO ratio of 3.0.

Table 3 shows the comparison of the 5Co-15Fe/γ-Al2O3 catalyst under optimum conditions with different catalysts published in other papers. As listed in Table 3, Co-Mn-Ru catalysts showed high CO conversion and CH4 yield, but relatively low C2<sup>+</sup> hydrocarbon yield, compared to Fe-based catalysts [7,8]. Lee et al. [7] reported that Mn promoter in cobalt catalysts acted as a Lewis acid, which increased the carbon chain growth and C2<sup>+</sup> hydrocarbon yield, but suppressed CO conversion. Ru promoter, one of the noble metals, provided the catalyst with hydrogen spillover ability, enhancing the reducibility of the cobalt site and CO conversion, but decreasing the C2<sup>+</sup> hydrocarbon yield. In addition, Ru promoters were substituted to save cost for the catalysts, and 20Co-16Mn/γ-Al2O3 catalyst reduced at 700 ◦C for 1 h showed similar reactivity to 10Co-6Mn-2.5Ru/γ-Al2O3 catalyst reduced at 400 ◦C for 1 h. The Fe-based bulk catalysts promoted by Cu and Zn showed high CO conversion and C2–C4 yield, but low paraffin ratio and high byproduct yield (C5<sup>+</sup> and CO2) [8,9]. On the other hand, the FT catalyst (5Co-15Fe/γ-Al2O3) affords high CO conversion (99.7%), high light paraffinic hydrocarbon yield (31.2% CH4 and 36.1% C2–C4), and low byproduct formation (C5<sup>+</sup> and CO2) under optimum conditions (SV: 4000 mL/g/h, T: 350 ◦C and P: 20 bar).


**Table 3.** Comparison of catalytic performance of FT catalyst (5Co-15Fe/γ-Al2O3) with different catalysts published in other papers.

<sup>a</sup> 10 wt.% Co, 6 wt.% Mn, and 2 wt.% Ru, reduced at for 400 ◦C for 1h. <sup>b</sup> 10 wt.% Co, 6 wt.% Mn, and 2.5 wt.% Ru, reduced at for 400 ◦C for 1h. <sup>c</sup> 20 wt.% Co, 16 wt.% Mn, reduced at for 700 ◦C for 1h. <sup>d</sup> Fe/Cu atomic ratio = 15, reduced at 500 ◦C for 1 h. <sup>e</sup> Fe/Zn atomic ratio = 5, reduced at 500 ◦C for 1 h. <sup>f</sup> Fe/Zn atomic ratio = 10, caburized at 500 ◦C for 1 h. <sup>g</sup> n/a: Not applicable.

#### **3. Materials and Methods**

#### *3.1. Catalyst Synthesis*

The bimetallic Co-Fe catalyst was synthesized by wet impregnation of γ-alumina (Sigma–Aldrich, St. Louis, MO, USA) with Co(NO3)2·6H2O (Sigma–Aldrich, St. Louis, MO, USA) and Fe(NO3)3·9H2O (Sigma–Aldrich, St. Louis, MO, USA) according to the same method as our previous papers [11,12]. During the impregnation procedure of the FT catalyst (5Co-15Fe/γ-Al2O3), γ-Al2O3 was added to an anhydrous ethanol solution containing cobalt and iron nitrates. The weight percentages of the cobalt and iron metal based on the catalyst were 5% and 10%, respectively. After stirring for 24 h, the solvent was vaporized in a rotary evaporator at 40–60 ◦C. The samples were dried at 120 ◦C for 12 h, and subsequently calcined at 400 ◦C for 8 h.

#### *3.2. Characterization*

The metal contents in FT catalyst were measured using inductively coupled plasma optical emission spectroscopy (ICP-OES; Perkin-Elmer, Waltham, MA, USA). Nitrogen adsorption-desorption isotherms at -196 ◦C were measured using a Micrometrics ASAP 2020 instrument (Norcross, GA, USA) to acquire the textural properties of the materials. Average pore size were measured following the Barrett-Joyner-Halenda (BJH) method. The crystal structure of the FT catalyst was analyzed via X-ray diffraction (XRD; PANalytical, Amsterdam, Netherlands) using a Cu Ka radiation source at the Korea Basic Science Institute in Daegu. Crystallite sizes of metal phase were calculated using the Scherer equation.

#### *3.3. Activity Tests*

Prior to the reaction, the catalysts (0.5 g) were placed in a fixed-bed stainless steel reactor (1/2 inch I.D.) and reduced with a 10 vol% H2/N2 gas mixture at 500 ◦C for 1 h. Then, the gas stream (H2, CO, N2) was fed to the reactor at a different total gas flow (33, 50 and 60 mL/min); N2 gas was used as an internal standard in the feed gas. The reactor was pressurized to 10 or 20 bar with the feed gas stream using a back-pressure regulator at constant pressure and heated to 200 ◦C. Then, the temperature was increased to 250, 300, 350, or 400 ◦C, and maintained during the FT reaction. All volumetric gas flows were measured at standard temperature and pressure (S.T.P). To prevent the condensation of water vapor and hydrocarbons, the inlet and outlet lines of the reactor were maintained at temperatures above 250 ◦C, and the liquid and wax products were collected in a cold trap (0 ◦C) before injection of the gas into the reactor and GC column. The outlet gases were analyzed using a gas chromatograph (Agilent 6890; Agilent, Santa Clara, CA, USA) equipped with both a thermal conductivity detector (TCD), and a flame ionization detector (FID). A packed column (Carboxen 1000; Bellefonte, PA, USA) was connected to the TCD to analyze the CO, H2, N2, and CO2 gases, and a capillary column (GS Gas Pro; Agilent, Santa Clara, CA, USA) was connected to the FID to analyze the hydrocarbon gases.

CO conversion, selectivity and yield for each product were calculated using Equations (1)–(4).

$$\text{CO conversion (carbon mole }\% \text{)} = \left(1 - \frac{\text{CO in the product gas (mol/min)}}{\text{CO in the feed gas (mol/min)}}\right) \times 100\tag{1}$$

$$\frac{\text{Selocity for hydrogenons with carbon number } n \text{ (carbon mole } \% \text{)}}{n \times \text{C } n \text{ hydrogen in the product gas } (mol/min)} \times 100 \tag{2}$$

$$\frac{(\text{total carbon} - \text{unmeasured CO}) \text{ in the product gas } (mol/min)}{(\text{total carbon} - \text{unmeasured CO}) \text{ in the product gas } (mol/min)} \times 100$$

$$\begin{array}{l} \text{Selocity for carbon dioxide (carbon mole }\% \text{)}\\ = \frac{\text{CO}\_2 \text{ in the product gas } (\text{mol/min})}{(\text{total carbon-unreacted CO}) \text{ in the product gas } (\text{mol/min})} \times 100 \end{array} \tag{3}$$

$$\text{Yield for hydrocarbons and carbon dioxide} = \frac{\text{CO conversion} \times \text{Selocity}}{100} \tag{4}$$

#### **4. Conclusions**

At present, few studies have made an effort to produce mixing gases consisting of paraffinic C2–C4 hydrocarbons into SNG for power generation. In this study, bimetallic Co-Fe catalysts supported on γ-alumina were developed, and the effects of operating parameters such as space velocity, reaction pressure and temperature on catalytic performance were elucidated for the production of light paraffin hydrocarbon yield (C2–C4 range) with high paraffin ratio, as well as reduction of byproduct formation (C5<sup>+</sup> and CO2). It was found that CO conversion increases with a decrease in space velocity, and with an increase in reaction pressure and temperature. CH4 yield increases and C2<sup>+</sup> yield decreases with increasing reaction temperature at all reaction pressures and space velocities. In addition, improved CH4 yield at higher reaction pressure (20 bar) implies that higher reaction pressure is a favorable condition for secondary CO2 methanation reaction. While paraffin ratio shows a positive correlation with the CO conversion according to increasing H2/CO ratio, reaction pressure and temperature have little effect on paraffin ratio at a H2/CO ratio of 3.0. Based on these results, the optimum conditions were determined to be H2/CO of 3.0, space velocity of 4000 mL/g/h, reaction pressure of 20 bar, and temperature of 300 ◦C, and the FT catalyst (5Co-15Fe/γ-Al2O3) affords a high light hydrocarbon yield (31.2 % CH4, and 36.1 % C2–C4) with high paraffin ratio (0.90). Based on these results, the bimetallic Co-Fe catalysts can be used for production of high paraffinic light hydrocarbons.

**Supplementary Materials:** The following are available online at http://www.mdpi.com/2073-4344/9/9/779/s1, Figure S1: XRD patterns of (I) fresh and (II) reduced monometallic and bimetallic Co-Fe catalysts; (-) Co3O4, (-) CoO, () Co metal, () Fe2O3, () Fe3O4, and () Fe metal; Figure S2. H2-TPR profiles of the monometallic and bimetallic catalysts supported on γ-alumina: (a) 20Co/γ-Al2O3, (b) 15Co-5Fe/γ-Al2O3, (c) 10Co-10Fe/γ-Al2O3, (d) 5Co-15Fe/γ-Al2O3, and (e) 20Fe/γ-Al2O3 (5 ◦C/min, pure hydrogen); Figure S3: CO conversion and hydrocarbon distribution of the monometallic and bimetallic catalysts supported on γ-alumina: 20Co/γ-Al2O3, 15Co-5Fe/γ-Al2O3, 10Co-10Fe/γ-Al2O3, 5Co-15Fe/γ-Al2O3, and 20Fe/γ-Al2O3 at H2/CO ratio = 3.0, 300 ◦C, and 10 bar.

**Author Contributions:** Conceptualization, S.B.J., T.Y.K., S.-H.K. and J.W.K.; Data curation, S.B.J., T.Y.K., C.H.L. and J.H.W.; Formal analysis, T.Y.K., C.H.L., J.H.W. and H.J.C.; Investigation, S.B.J. and T.Y.K.; Project administration, S.C.L. and J.C.K.; Supervision, S.C.L. and J.C.K.; Writing—original draft, S.B.J.; Writing—review & editing, S.B.J. and T.Y.K.

**Funding:** This research was funded by the Korea Institute of Energy Technology Evaluation and Planning (KETEP) and the Ministry of Trade, Industry & Energy (MOTIE) of the Republic of Korea (No.20173010050110 and the Basic Science Research Program through the National Research Foundation of Korea (NRF) funded by the Ministry of Science, ICT & Future Planning (No.2017R1A2B4008275).

**Acknowledgments:** This work was supported by the Korea Institute of Energy Technology Evaluation and Planning (KETEP) and the Ministry of Trade, Industry & Energy (MOTIE) of the Republic of Korea. (No.20173010050110). This research was also supported by the Basic Science Research Program through the National Research Foundation of Korea (NRF) funded by the Ministry of Science, ICT & Future Planning (No.2017R1A2B4008275).

**Conflicts of Interest:** The authors declare no conflict of interest.

#### **References**


© 2019 by the authors. Licensee MDPI, Basel, Switzerland. This article is an open access article distributed under the terms and conditions of the Creative Commons Attribution (CC BY) license (http://creativecommons.org/licenses/by/4.0/).

#### *Article*

### **E**ff**ect of Operating Temperature, Pressure and Potassium Loading on the Performance of Silica-Supported Cobalt Catalyst in CO2 Hydrogenation to Hydrocarbon Fuel**

#### **Rama Achtar Iloy and Kalala Jalama \***

Department of Chemical Engineering, Doornfontein Campus, University of Johannesburg, Doornfontein 2028, Johannesburg, South Africa; achtar2006@yahoo.fr

**\*** Correspondence: kjalama@uj.ac.za

Received: 31 August 2019; Accepted: 10 September 2019; Published: 26 September 2019

**Abstract:** Potassium (1–5 wt.%)-promoted and unpromoted Co/SiO2 catalysts were prepared by impregnation method and characterized by nitrogen physisorption, temperature-programmed reduction (TPR), CO2 temperature-programmed desorption (TPD), X-ray diffraction (XRD) and X-ray photoelectron spectroscopy (XPS) techniques. They were evaluated for CO2 hydrogenation in a fixed bed reactor from 180 to 300 ◦C within a pressure range of 1–20 bar. The yield for hydrocarbon products other than methane (C2+) was found to increase with an increase in the operating temperature and went through a maximum of approximately 270 ◦C. It did not show any significant dependency on the operating pressure and decreased at potassium loadings beyond 1 wt.%. Potassium was found to enhance the catalyst ability to adsorb CO2, but limited the reduction of cobalt species during the activation process. The improved CO2 adsorption resulted in a decrease in surface H/C ratio, the latter of which enhanced the formation of C2<sup>+</sup> hydrocarbons. The highest C2<sup>+</sup> yield was obtained on the catalyst promoted with 1 wt.% of potassium and operated at an optimal temperature of 270 ◦C and a pressure of 1 bar.

**Keywords:** CO2 hydrogenation; cobalt; potassium; pressure; temperature

#### **1. Introduction**

The promoting capabilities of alkali metals, namely potassium, have been investigated for a variety of catalysts and reactions, including steam reforming of bioethanol [1], water gas shift [2], N2O decomposition [3], Fischer-Tropsch synthesis (FTS) [4–6] and CO2 hydrogenation [7–11]. One of the earliest studies on the use of potassium as a promoter for the catalyst used in CO2 hydrogenation to hydrocarbons is that of Russell and Miller [12]. They investigated several copper-activated cobalt catalysts at atmospheric pressure from 448 to 573 K with H2/CO2 ratio varied from 2 to 3. All the catalysts mainly produced methane and liquid hydrocarbons were observed only after potassium addition to the catalyst in the form of either potassium carbonate or phosphate. Potassium was believed to selectively poison methane forming centres, and therefore, promote methylene radicals polymerization by the repression of the competitive hydrogenation reaction. Similarly, Owen et al. [13] studied the effect of potassium, along with that of lithium and sodium, on the performance of Co/SiO2 catalysts. The catalytic testing was carried out at 643 K, atmospheric pressure and using an H2/CO2 ratio of 3. They showed that with an alkali loading as low as 1 wt.%, the products distribution shifts towards longer chain hydrocarbons. Furthermore, C2 and C3 olefins, which did not form over the unpromoted catalyst, were detected in relatively significant amounts over the promoted catalysts. The authors attributed this behaviour to the ability of potassium to enhance the surface to molecule charge transfer, resulting in increased CO and reduced hydrogen binding strength. These findings were further corroborated by a more recent investigation by Shi et al. [8] on a CoCu/TiO2 system containing 1.5–3.5 wt.% K. Using CO2 temperature-programmed desorption, the authors were able to link an improved yield of liquid hydrocarbons (C5+) to the increased CO2 adsorption capacity of the catalyst, when loaded with potassium.

It appears that potassium has an enormous potential in the conversion of CO2 to liquid hydrocarbons. To derive most of the benefit from this promoter, the study of its effect on the reaction must be integrated with that of the effect of operating conditions. Most studies have reported the effect of potassium on cobalt-based catalysts under pre-selected operating conditions that were not optimized. Hence, the present study aims at systematically evaluating the promoting effect of potassium on a Co/SiO2 system used in CO2 hydrogenation under optimized temperature and pressure conditions.

#### **2. Results and Discussion**

#### *2.1. Surface Area and Porosity*

The information on the surface area and porosity of the catalysts investigated is presented in Table 1. The data show that cobalt incorporation into the silica support results in a significant drop in the surface area from 186.6 to 133.1 m2/g. This behaviour is generally explained by the growth of cobalt oxide particles within the pores of the support during catalyst calcination, leading to some level of pore obstruction. This agrees well with the pore volume data, which show a decrease from 1.5 to 1.0 cm3/g. The introduction of potassium, in amounts above 3% in the catalyst, further amplifies this phenomenon.


**Table 1.** Surface area and porosity data.

#### *2.2. X-ray Di*ff*raction*

Figure 1 shows the XRD patterns of unpromoted and promoted catalysts before and after reduction. All the unreduced catalysts showed diffraction peaks at 2θ values of approximately 18◦, 30◦, 36.6◦, 39◦, 44.5◦, 55.3◦, 60◦ and 65◦, attributed to Co3O4 [14].

After catalysts reduction, the diffraction peaks for Co3O4, which were present in the unreduced catalysts disappeared (Figure 1b). The only visible peaks are those of the lower oxide of cobalt (CoO) at 42.4◦ and metallic cobalt at 44.5◦.

The Scherrer equation was used to calculate the average crystallite sizes of cobalt species in the catalyst, using two theta values of 36.6◦, 42.4◦ and 44.5◦ for Co3O4, CoO and Co respectively. The data are reported in Table 2. Although there is no observable trend in the data with respect to Co3O4 and Co, it appears that the average crystallite size for CoO decreases with increasing potassium loading in the catalyst. This suggests that potassium controls the size of CoO in the catalyst.

**Figure 1.** XRD patterns for unpromoted and potassium-promoted 15%Co/SiO2 catalysts: (**a**) Before reduction, and (**b**) after reduction.


**Table 2.** The particle size of the calcined catalysts.

<sup>a</sup> Particle size in nm.

#### *2.3. Temperature-Programmed Reduction (TPR)*

The effect of potassium addition on the reducibility of silica-supported cobalt catalysts was investigated using TPR analysis. TPR profiles of various potassium-promoted catalysts, along with that of the unpromoted sample are presented in Figure 2. For the unpromoted catalyst, an early and slow reduction process was observed from ca. 170 ◦C. It became significant from ca. 290 ◦C, where a fast reduction peak was observed to start and went through a maximum at 365 ◦C. Subsequent overlapping reduction peaks, with respective maxima at ca. 395, 425 and 466 ◦C, were also observed. These peaks can be attributed to a two-step reduction of Co3O4 species in the catalyst to CoO and Co0. The presence of more than two peaks observed for this reduction process could indicate that not all the cobalt species in the catalyst underwent reduction at the same time. For example, as N2 adsorption data suggest some level of pore obstruction in the catalyst, it is possible that some cobalt species only got reduced after the reduction of some of those that blocked some pores. Adding potassium to the catalyst reduced the reducibility of cobalt species as per the following observations: (i) The reduction temperatures for the catalysts shifted to higher values. For example, the start of the reduction process moved from 170 ◦C for the unpromoted catalyst to 210 and 255 ◦C for catalysts containing 1% and 3–5% K respectively; (ii) the area under the TPR profile below 500 ◦C decreased, indicating lower degree of catalyst reduction as the amount of potassium increased in the catalyst and (iii) the formation of cobalt species in strong interaction with the support, as shown by a broad reduction peak, with a maximum at ca. 512 ◦C, observed in the catalyst containing 5% K. The negative effect of potassium on the reduction of cobalt catalyst was also reported by Jacobs et al. [6] who found that (0.5–5%) K shifted

the reduction peak temperatures to higher values and lowered the extent of catalyst reduction. This suggests that potassium interacts with the cobalt species and possibly the silica support [15].

**Figure 2.** TPR profiles for (**a**) 15%Co/SiO2, (**b**) 15%Co-1%K/SiO2, (**c**) 15%Co-3%K/SiO2, and (**d**) 15%Co-5%K/SiO2.

#### *2.4. CO2 Temperature-Programmed Desorption (CO2-TPD)*

CO2-TPD profiles for 15%Co/SiO2 and 15%Co-1%K/SiO2 are presented in Figure 3. Both catalysts showed two desorption peaks, with the first one centred at 65 ◦C with near-identical areas. This low-temperature peak can be attributed to the desorption of physically adsorbed CO2. A second peak, observed for each catalyst, was attributed to the desorption of chemisorbed CO2 and was used as an indication of the strength and amounts of basic sites in the catalyst. As expected, the data show that the addition potassium to the catalyst increases the strength and amounts of basic sites in the catalyst. This is indicated by the large and extended CO2 desorption peak, which goes through its maximum at ca. 187 ◦C, compared to a corresponding small peak, with a maximum at ca. 134 ◦C, for the unpromoted catalyst. These data agree with earlier studies [8,16] that also reported an improvement in CO2 adsorption in cobalt-based catalyst upon potassium addition.

**Figure 3.** CO2-TPD profiles of (**a**) 15%Co-1%K/SiO2 and (**b**) 15%Co/SiO2.

### *2.5. X-ray Photoelectron Spectroscopy (XPS)*

XPS spectra, in the Co 2p region, for calcined and activated catalysts are shown in Figure 4.

**Figure 4.** XPS data for calcined catalysts (**a**) 10%Co/SiO2-calc., (**b**) 10%Co/1%K/SiO2-calc.; and reduced catalysts (**c**) 10%Co/SiO2-red., (**d**) 10%Co/1%K/SiO2-red.

The Co 2p1/<sup>2</sup> and 2p3/<sup>2</sup> peaks for the calcined and unpromoted catalyst were respectively observed at ca. 795.2 and 779.6 eV and are characteristic of Co3O4 [14,17], in agreement with XRD data, discussed in Section 2.2. A shift to lower binding energies can be observed for Co 2p1/<sup>2</sup> (to 793.5 eV) and 2p3/<sup>2</sup> (to 778 eV) following catalyst promotion with potassium. This suggests an electronic donation by potassium as also observed by other studies, where potassium was added to Co/Al2O3 [18] and Pd/Co3O4 and Co3O4 [19] catalysts.

Spectra of reduced catalysts (Figure 4c,d) display features of CoO with broader Co 2p1/<sup>2</sup> and 2p3/<sup>2</sup> peaks and increased intensities of the shake-up satellite features [17,20]. They look similar for both the unpromoted and the potassium-promoted catalysts. These findings indicate that the electronic properties of cobalt in the catalyst were modified by potassium during the calcination process, not during catalyst reduction, causing a different reduction behaviour for the promoted catalyst.

#### *2.6. Catalyst Testing*

#### 2.6.1. Effect of Temperature

In order to study the effect of temperature, CO2 hydrogenation was carried out over a 15%Co-3%K/SiO2 catalyst at atmospheric pressure from 180 to 300 ◦C. The temperature dependency of CO2 conversion and its corresponding Arrhenius plot are reported in Figure 5. As expected, the CO2 conversion continuously increased from 0.6 to 18.4% as the temperature was raised from 180 to 300 ◦C, in agreement with other earlier studies [21–23].

**Figure 5.** CO2 conversion during hydrogenation vs. reaction temperature (1 bar, SV = 0.92 NL/gcat/h, H2/CO2 = 3.1/1): (**a**) CO2 conversion vs. temperature; (**b**) Arrhenius plot (Ea = 78 kJ/mol).

From the Arrhenius plot, activation energy of 78 kJ/mol was obtained in a temperature range of 180 to 240 ◦C. A marked curvature in the Arrhenius plot can be observed at temperatures above 240 ◦C, suggesting that the catalyst surface underwent some changes, possibly including deactivation by carbon [23]. For comparison, activation energies reported by earlier studies involving cobalt catalysts for CO2 hydrogenation are summarised in Table 3.

The value of the activation energy obtained in this study is similar to most of those reported in earlier studies. Exceptions can be noticed for the data reported by Weatherbee and Bartholomew [23], who reported activation energies of 93 and 171 kJ/mol over 15%Co/SiO2 at 1 and 10 bar respectively.

The effect of the operating temperature on products selectivity and yields is summarized in Figure 6. The methane selectivity showed relatively little dependency on temperature from 180 to 225 ◦C, but continuously increased from 240 to 300 ◦C, while the selectivity to CO decreased almost linearly with increasing temperature (Figure 6a). Both C2-C4 and C5<sup>+</sup> selectivities increased as the temperature was raised and went through a maximum at 240 ◦C before decreasing. Figure 6c shows that up to 270 ◦C, the yields for C2+, CH4 and CO all increased with the rise in temperature, with the yield of CO remaining the highest of the three. The yield of methane and C2<sup>+</sup> hydrocarbons were similar up to 240 ◦C, above which the yield of methane quickly surpassed that of C2<sup>+</sup> in an exponential manner.


**Table 3.** The activation energy for CO2 hydrogenation over cobalt catalysts.

\* S1 and S3 are carbon supports obtained from saran copolymer. The difference between the two is in the burn-off percentage, i.e., 0 and 20% for S1 and S3 respectively [26].

**Figure 6.** Effect of temperature on products selectivity: (**a**) CH4 and CO; (**b**) C2<sup>+</sup> hydrocarbons; and products yields: (**c**) C2+, CH4 and CO.

The rise in CO yield flattened off around 285 ◦C and was surpassed by the fast-rising yield of methane around 290 ◦C. The C2<sup>+</sup> yield went through a maximum at 270 ◦C, indicating that, above this temperature, the reaction is turning into a preferential methanation process.

The mechanism of CO2 hydrogenation to hydrocarbons is still subject of some controversies. However, since CO formed during CO2 hydrogenation, it is most likely that hydrocarbons formed via a typical Fischer-Tropsch mechanism. Indeed, this is a plausible explanation, since some studies [27–29] have shown that, in the presence of CO, on cobalt-based catalysts, CO2 behaves like an inert gas and only reacts when CO is depleted. Also, the rapid increase in methane yields with the temperature at values above 240 ◦C is typical to FT reaction [22,23]. An operating temperature of 270 ◦C was selected as optimal for the rest of the study.

#### 2.6.2. Effect of Pressure

The effects of pressure on CO2 conversion, and products selectivities and yields are reported in Figure 7. An increase in operating pressure, from 1 to 15 bar, resulted in an increase in CO2 conversion. As can be seen from Figure 7a, the CO2 conversion measured at 1 bar was ca. 12.5%; it increased to ca. 21%, 22%, and 27% when the pressure was increased to 5, 10 and 15 bar, respectively. Further increase in the operating pressure to 20 bar resulted in a slight decrease of CO2 conversion to ca. 26%. The increase in CO2 conversion with the operating pressure from 1 to 15 bar was expected because of an increase in reactants partial pressures. However, the decrease in CO2 conversion observed when the operating pressure was increased from 15 to 20 bar was not expected; it is possible that some CO2 or reaction intermediate species irreversibly adsorbed on the catalyst surface, blocking some active sites.

An increase in operating pressure from 1 to 5 bar significantly decreased the selectivities to CO and C2<sup>+</sup> hydrocarbons from ca. 48% and 21% to 8% and 11%, respectively (Figure 7b). Further increase in pressure only resulted in slight decreases in CO and C2<sup>+</sup> hydrocarbons selectivities. An opposite behaviour was observed for CH4 selectivity, which increased from 30 to 81% when the operating pressure was increased from 1 to 5 bar. Further increase in operating pressure resulted in a relatively slight increase in CH4 selectivity. Similar trends can be observed for CO and CH4 yields as a function of the operating pressure (Figure 7c); however, the C2<sup>+</sup> yield was not significantly affected by changes in operating pressure. It remained between 2.1% and 2.7% over the range of pressure used. Under these conditions, operating at 1 bar is optimal.

**Figure 7.** Effect of the operating pressure on (**a**) CO2 conversion, (**b**) products selectivity and (**c**) products yields (catalyst—15%Co/3%K/SiO2, 270 ◦C, SV = 0.92 NL/gcat/h, H2/CO2 = 3.1/1).

2.6.3. Effect of Potassium Addition

Figure 8 shows the effect of potassium addition on CO2 conversion, and products selectivity and yields. It is observed that the presence of potassium at a loading of as low as 1% results in a significant drop in CO2 conversion from 39 to 16% when compared to the unpromoted catalyst (Figure 8a).

**Figure 8.** Effect of potassium loading on (**a**) CO2 conversion, (**b**) product selectivity and (**c**) product yields.

Adding more potassium further exacerbates this behaviour, but with an attenuated effect. The following can explain these observations: (i) Coverage of active sites by potassium, although considered to happen at a low extent because of the low potassium loading employed; (ii) increase in CO2 adsorption capacity: As discussed earlier in Section 2.4, CO2-TPD results have shown that the CO2 adsorption capacity for the catalyst improves upon potassium addition. As a consequence, the H/C molar ratio on the catalyst surface is decreased, leading to a drop in CO2 conversion. This is in agreement with numerous experimental [30–33] and theoretical [34,35] investigations that reported a drop in CO2 conversion with a decrease in H2/CO2 molar ratio. The decrease in CO2 conversion with potassium addition was also reported by Owen et al. [13] and Shi et al. [8] on Co/SiO2 and CoCu/TiO2 catalysts, respectively. Using H2 and CO2 temperature-programmed desorption analyses, Shi et al. [8] were able to show that potassium promotion decreases the H2 adsorption capacity of the catalyst, while that of CO2 is enhanced; (iii) the oxidation state of cobalt in the catalyst: XRD results have shown the presence of CoO and metallic cobalt phases in all the reduced and passivated catalysts. TPR analyses, on the other hand, confirmed that these catalysts have different reducibility properties. Promotion with potassium limits the reducibility of the catalysts, resulting in limited amounts of metallic cobalt sites for CO2 conversion. A similar relationship between catalyst reducibility and CO2 activity was reported

by Melaet et al. [36] who conducted CO2 hydrogenation on Co/SiO2 catalysts activated at 523 K and 723 K. They established, by means of XPS, that CoO and metallic cobalt formed upon activation at 523 K and 723 K respectively. The catalyst reduced at 723 K showed higher activity.

The selectivity towards methane decreased from 96 to 37.6% (Figure 8b) upon adding 1 % of potassium to the catalyst. Meanwhile, the selectivity of both CO and C2<sup>+</sup> hydrocarbons increased. Additional amounts of potassium resulted in a further decrease in methane selectivity and increase in CO selectivity. The selectivity of C2<sup>+</sup> hydrocarbons, on the other hand, decreased with further increase in potassium loading above 1%. The improvement of C2<sup>+</sup> hydrocarbons selectivity at 1% potassium can be attributed to a decreased surface H/C ratio as discussed earlier. This implies that carbon-containing species from CO2 dissociation can polymerize rather than being hydrogenated as is often the case in a hydrogen-rich environment.

The decreased C2<sup>+</sup> hydrocarbons selectivity at 3% and 5% potassium loading is also explained by an increased CO2 adsorption capacity of the catalyst, causing a decrease in the surface H/C ratio. Since the CO yield is shown to increase and undergo little variations with further increase in potassium loading, while both CH4 and the C2<sup>+</sup> yields decrease (Figure 8c), it is possible that there is not enough surface hydrogen to readily react with both CO2 and CO on the catalyst surface.

The highest C2<sup>+</sup> yield achieved in this study was ca. 5% and was measured on the 15%Co/1%K/SiO2 catalyst at 1 bar and 270 ◦C. This condition is compared to results reported in other studies that used cobalt-based catalysts for CO2 hydrogenation under various conditions, as summarized in Table 4. Of the catalysts that produced C5<sup>+</sup> products at low pressure (<2 bar), our catalyst had the lowest methane selectivity under the optimized operating temperature and pressure. This is particularly important, since it offers opportunities to limit the formation of the undesirable methane without the need for excessive operating pressures that will make the process more energy-intensive.

#### **3. Materials and Methods**

#### *3.1. Catalyst Synthesis*

The catalyst synthesis consisted essentially of two steps, namely support preparation and metal loading. Fumed silica with an average particle size range of 0.2–0.3 μm, supplied by Sigma-Aldrich South Africa, served as the catalyst supporting material. Given its small particle size, it was pre-treated with deionized water and agglomerated by drying overnight at 120 ◦C in the air before crushing and sieving to obtain a powder with particles within the size range of 212–500 μm. The powder so obtained was subsequently calcined at 400 ◦C for 6 h in the air to lock its properties before loading the metals. The addition of cobalt and potassium was done through co-impregnation with solutions of cobalt and potassium nitrates—both purchased from Sigma-Aldrich. After impregnation, the catalysts were dried overnight at 120 ◦C and calcined at 400 ◦C in the air for 6 h. All the prepared catalysts contained 15 wt.% cobalt with varying potassium loading (0–5 wt.%). The amount of silica used in catalyst preparation was reduced to account for the addition of potassium. This allowed for the cobalt loading to be kept constant for all the catalysts.



*Catalysts* **2019**, *9*, 807



Calculated from reported flow of CO2 over 24 h, H2/CO2 ratio and the mass of catalyst. Calculated from the reported milliliters of oil that formed during the reaction, assumingaverage chain length of 7 (density of 0.684). c Data read from graphs.

a

#### *3.2. Catalyst Characterization*

The surface area and the porosity of the catalysts were measured by nitrogen physisorption at −196 ◦C using an Accelerated Surface Area and Porosimetry System (ASAP 2460) from Micromeritics. Each analysis was preceded by degassing the sample at 150 ◦C for 4 h. The multipoint Brunauer-Emmett-Teller (BET) method was used to determine the surface area of the materials analysed.

The reducibility of the catalysts was studied by means of temperature-programmed reduction (TPR). An in-house built instrument, equipped with a thermal conductivity detector (TCD), was used for this purpose. In a typical analysis, 100 mg of catalyst was loaded in a stainless-steel reactor and heated to 300 ◦C for one hour under 70 NmL/min of helium to remove traces of moisture and other ambient contaminants. This step was referred to as degassing. After allowing the reactor to cool to room temperature, helium was switched with a gas mixture containing 5% H2 in argon at a flow of 65 NmL/min. In the final step, the temperature was raised from room temperature to 700 ◦C at a heating rate of 10 ◦C/min, while recording the signal of the TCD.

Temperature-programmed desorption (TPD) of CO2 was carried out using the same instrument as described for TPR analysis. Different to TPR analysis, the catalysts used in this analysis were first reduced at 335 ◦C for 17 h, using the same reactor and conditions as for the reduction of catalyst samples used in the CO2 hydrogenation testing as will be described in Section 2.3. The reduced catalysts were passivated using 5% O2 in helium for 2 h at ambient temperature before their transfer from the CO2 hydrogenation reactor to the TPD apparatus. Two hundred milligrams of catalyst sample was degassed in a similar manner as for TPR analysis. After degassing and cooling to room temperature, the temperature was raised to 335 ◦C at a heating rate of 10 ◦C/min and maintained at this value for 30 min under a flow of 5% H2 in argon. This step was necessary for the removal of the cobalt oxide layer formed during catalyst passivation. Thereafter, the reactor was cooled and maintained at 50 ◦C for at least 10 min before replacing H2 (5% in argon) with CO2 (10% in helium). CO2 adsorption was performed at 50 ◦C for 1 h before re-introducing helium, but this time to remove the physically adsorbed CO2 molecules. TPD was then performed under helium flow, after stabilization of the TCD signal, from 50 to 700 ◦C at a heating rate of 5 ◦C/min.

X-ray diffraction (XRD) analysis was performed to identify the oxidation state of cobalt species in the unreduced and reduced catalyst samples. The instrument used for this purpose was a Rigaku Ultima IV equipped with a copper target. The voltage and current at which the diffractometer was operated were 40 kV and 30 mA respectively. Spectra were acquired in the range of 2θ from 10◦ to 90◦ with a step size of 0.01◦ at the scanning speed of 1◦/min.

X-ray photoelectron spectroscopy (XPS) was used to determine the oxidation states of the elements present on the surface of the catalysts. This analysis was performed on a Specs Phoibos 150 spectrometer with a monochromatic X-ray source Al Kα at 1486.71 eV. A low-energy electron flood gun operated at 2.0–2.5 eV and 20 μA was used to stabilize the sample surface charge. The spectrometer was operated at constant pass energy of 40 eV. The shift in binding energy peaks position, due to the surface charging effect was corrected by setting the C 1s binding energy to 284.8 eV [14].

#### *3.3. Catalyst Testing*

Carbon dioxide hydrogenation was carried out in a system which consisted mainly of a stainless steel fixed-bed reactor (16 mm i.d. × 220 mm length) mounted in an electrical furnace, a mass flow controller (Aalborg), a back-pressure regulator and a product collection pot. The furnace temperature was controlled using a programmable temperature controller connected to a K-type thermocouple and the furnace heating element. Accurate reaction temperatures were measured by means of another K-type thermocouple in direct contact with the catalyst bed held in place by plugs of quartz wool. Any liquid product formed was collected in a cold pot mounted at the bottom of the reactor. There was no need for a hot trap since the products were mainly light hydrocarbons. The reactor outlet was connected to a three-way valve, which made it possible to either send the reaction products to vent or to an online gas chromatograph (GC) for analysis. The Dani Master GC used in this study was

equipped with a flame ionization detector (FID) connected to a capillary column (Supel-QTM PLOT) that separated hydrocarbons and oxygenates, and a thermal conductivity detector (TCD) connected to a packed column (60/80 Carboxen 1000) for the separation of H2, N2, CO and CO2.

Prior to testing, 500 mg of catalysts were reduced in flowing hydrogen (23 NmL/min) at 335 ◦C and atmospheric pressure for 17 h. Catalyst testing was done at temperatures ranging from 180 to 300 ◦C with an increment of 15 ◦C and at pressures within a range of 1–20 bar at a space velocity of 0.92 NL/gcat/h. The feed gas was premixed and contained 21.8% CO2, 68.6% H2 and 9.6% N2. After testing, all catalysts were passivated in 5% O2 in helium (23 NmL/min) at room temperature for 2 h. The nitrogen present in the feed gas was used as an internal standard for mass balance calculations. The CO2 conversion, the rate of CO2 conversion, the rate of products formation, selectivity and yield were calculated according to Equations (1)–(5), where F and X indicate the total molar gas flow rate and mole fraction respectively. The subscripts "in" and "out" refer to the gas streams entering or leaving the reactor.

$$\text{CO}\_2\text{ conversion } \left(\% \right) = \frac{\chi\_{\text{CO}\_{2,\text{in}}} - \frac{\chi\_{\text{N}\_{2,\text{in}}}}{\overline{\chi}\_{\text{N}\_{2,\text{out}}}} \times \chi\_{\text{CO}\_{2,\text{out}}}}{\chi\_{\text{CO}\_{2,\text{in}}}} \times 100,\tag{1}$$

$$\text{Rate of CO}\_2\text{ conversion} = \frac{\text{F}\_{\text{in}} \left[ \chi\_{\text{CO}\_{2,\text{in}}} - \frac{\chi\_{\text{V}\_{2,\text{in}}}}{\chi\_{\text{V}\_{2,\text{out}}}} \times \chi\_{\text{CO}\_{2,\text{out}}} \right]}{\text{Catalyst mass}},\tag{2}$$

$$\text{Rate of formation of product i} = \frac{\text{F}\_{\text{out}} \times \text{X}\_{\text{i,out}}}{\text{Catalyst mass}^{\prime}} \tag{3}$$

$$\text{Selectivity of product i } (\%) \, = \frac{\text{moles of carbon in product i per unit time}}{\text{Rate of CO}\_2 \text{ conversion} \times \text{Catalyst mass}} \times 100 \,\,,\tag{4}$$

$$\text{Yield of product i } (\%) = \frac{\text{Selocity of product i} \times \text{CO}\_2 \text{ conversion}}{100}. \tag{5}$$

After a change in operating conditions or in catalyst sample, the reactor was allowed to reach a steady state and maintained at the new conditions for at least two days. At least two data points were generated per day. To ensure reproducibility, each data point was an average of three independent measurements that were closer to each other within 5% error range.

#### **4. Conclusions**

The aim of this study was to investigate the effects of operating conditions (temperature, pressure) and potassium loading on the performance of silica-supported cobalt catalysts in CO2 hydrogenation. The highest yield in C2<sup>+</sup> hydrocarbons was measured at 1 bar and 270 ◦C. Potassium was found to negatively affect the reducibility of the catalyst, while enhancing its CO2 adsorption capacity. The improved CO2 adsorption capacity of the catalyst leads to a lower surface H/C ratio, which promotes chain growth reactions. The limited catalyst reducibility resulted in low catalyst activity and is explained by an electric donation of potassium to cobalt species during the calcination process of the catalyst. The optimal operating pressure and temperature determined in this study, combined with catalyst promotion with 1 wt.% of potassium, significantly lowered the undesirable methane selectivity when compared to other cobalt-based catalysts that also produced some C5<sup>+</sup> hydrocarbons at low pressures (<2 bar). This constitutes a significant further step in the development of efficient catalysts for CO2 hydrogenation to liquid fuels.

**Author Contributions:** Project conceptualization and methodology: R.A.I. and K.J.; Materials synthesis, experiments and data collection: R.A.I.; Data analysis and interpretation: R.A.I. and K.J.; Manuscript writing and editing: R.A.I. and K.J.; Project administration and supervision: K.J.

**Funding:** This project was funded by the National Research Foundation (Grant: UID 90757) and the University of Johannesburg Global Excellence Stature (GES) program.

**Conflicts of Interest:** The authors declare no conflicts of interest.

#### **References**


© 2019 by the authors. Licensee MDPI, Basel, Switzerland. This article is an open access article distributed under the terms and conditions of the Creative Commons Attribution (CC BY) license (http://creativecommons.org/licenses/by/4.0/).

*Article*

## **Kinetics of Fischer–Tropsch Synthesis in a 3-D Printed Stainless Steel Microreactor Using Di**ff**erent Mesoporous Silica Supported Co-Ru Catalysts**

#### **Nafeezuddin Mohammad 1, Sujoy Bepari 2, Shyam Aravamudhan <sup>1</sup> and Debasish Kuila 1,2,\***


Received: 14 September 2019; Accepted: 15 October 2019; Published: 21 October 2019

**Abstract:** Fischer–Tropsch (FT) synthesis was carried out in a 3D printed stainless steel (SS) microchannel microreactor using bimetallic Co-Ru catalysts on three different mesoporous silica supports. CoRu-MCM-41, CoRu-SBA-15, and CoRu-KIT-6 were synthesized using a one-pot hydrothermal method and characterized by Brunner–Emmett–Teller (BET), temperature programmed reduction (TPR), SEM-EDX, TEM, and X-ray photoelectron spectroscopy (XPS) techniques. The mesoporous catalysts show the long-range ordered structure as supported by BET and low-angle XRD studies. The TPR profiles of metal oxides with H2 varied significantly depending on the support. These catalysts were coated inside the microchannels using polyvinyl alcohol and kinetic performance was evaluated at three different temperatures, in the low-temperature FT regime (210–270 ◦C), at different Weight Hourly Space Velocity (WHSV) in the range of 3.15–25.2 kgcat.h/kmol using a syngas ratio of H2/CO = 2. The mesoporous supports have a significant effect on the FT kinetics and stability of the catalyst. The kinetic models (FT-3, FT-6), based on the Langmuir–Hinshelwood mechanism, were found to be statistically and physically relevant for FT synthesis using CoRu-MCM-41 and CoRu-KIT-6. The kinetic model equation (FT-2), derived using Eley–Rideal mechanism, is found to be relevant for CoRu-SBA-15 in the SS microchannel microreactor. CoRu-KIT-6 was found to be 2.5 times more active than Co-Ru-MCM-41 and slightly more active than CoRu-SBA-15, based on activation energy calculations. CoRu-KIT-6 was ~3 and ~1.5 times more stable than CoRu-SBA-15 and CoRu-MCM-41, respectively, based on CO conversion in the deactivation studies.

**Keywords:** Fischer-Tropsch synthesis; mesoporous silica based catalysts; kinetic studies; 3-D printed microchannel microreactor

#### **1. Introduction**

Although Fischer–Tropsch (FT) synthesis was discovered by Franz Fischer and Hans Tropsch in the 1920s in Germany [1], it has gained immense attention in last few years due to depletion of non-renewable energy sources. FT synthesis is an environmental friendly route for alternative fuels and can produce liquid fuels from carbon sources by coal-to-liquid (CTL), natural gas-to-liquid (GTL) and biomass-to-liquid (BTL) [2] processes. Three types of reactors have been utilized commercially for FT synthesis: Fixed bed, fluidized bed, and slurry bubble column bed by leading GTL companies like Shell, Sasol, Exxon Mobil, and Energy Int. [3]. There is a minimum scale limit of this FT process to be economical; for example, the Pearl GTL, a collaboration between Shell and Qatar petroleum, producing 140,000 bpd (barrels per day) is considered as a profitable economic scale for the FT GTL

process [4]. The limitation of the scale-up considerations to commercialize small scale plants to be more profitable has driven industries and researchers to pursue an interest in alternative technologies. Since FT synthesis is highly exothermic in nature, there is a need for much process intensification technologies. The microreactor platform, which contains microstructured units called microreactors, uses a large number of small, parallel channels with different channel designs. This technology provides an alternative platform for controlling highly exothermic reactions like FT synthesis with enhanced mass and heat transfer. It has gained much attention in process intensification of FT synthesis [5,6] as isothermal operating conditions are well maintained in a microreactor with good control over process parameters which favors quick screening of catalyst for different chemical reactions. The reaction zone for these microreactors are several parallel microchannels with small geometry. The specific surface area of the reaction zone is greatly enhanced by the design of microchannels resulting in an efficient FT synthesis. In addition to efficient heat and mass transfer with good heat dissipation, microreactors also have advantages such as high reaction throughput, easy scale-up, good portability, and lower cost over conventional reactors [6–10]. This has been demonstrated commercially and in R&D by Velocys and Micrometrics Corporations [11–14].

Iron, cobalt, and ruthenium catalysts have been extensively used for FT synthesis [15]. To increase the performance of catalysts, different supports have been used; some of the previous studies examined the role of Al2O3 [16–20], TiO2 [21–28], SiO2 [29–33], and CNTs [34–36] as supporting materials for the formation of higher alkanes. These supports tend to enhance the FT process by increasing the active number of catalytic sites and good metal dispersion with the high surface area. Therefore, the selection of support and study of its interaction with the incorporated metal ion plays an important role in catalysis. In our previous studies, sol-gel encapsulated catalysts were used in silicon microreactors for FT synthesis [37–39]. While Al2O3 and SiO2 sol-gel supports show similar behavior in formation of higher alkanes such as ethane, propane, and butane for the reactions at 1 atm, TiO2 has a profound effect on FT synthesis [40] and the stability of the catalysts are observed in reverse order from that observed with SiO2 and Al2O3. However, in all these studies, sol-gel coated catalysts in silicon microreactors tend to have challenges such as low surface area, clogging of microchannels and difficulty in reducing the metal oxides to expose active sites. In addition, the Si-microreactors are fragile and they break easily and require a large infrastructure for fabrication. Further, it's more difficult to increase pressure for FT studies using Si-microreactors. Thus, we have turned our attention to 3D printed stainless steel (SS) microreactors which are easy to fabricate by direct metal laser sintering *layer-by-layer* additive manufacturing technique. Recently, 3D printed microreactors have been used as flow devices in many chemical reactions such as fast difluoromethylation [41], a customizable Lab-On-Chip device for optimization of carvone semicarbazon [42], a micro fuel cell [43,44], and wide range of organic and inorganic reactions [45–47]. Further, these metal printed microreactors have been used for high pressure and temperature chemical reactions providing a new fast developing reactor technology in process development to industrial scale [47], which makes them suitable for reactions like FT synthesis due to its good mechanical and thermal properties. Although the specific surface area of stainless steel microreactor is less when compared to silicon microreactors used in our previous studies [40], the use of stainless steel material increases heat transfer, its chemical and mechanical resistances play a major role in process intensification of chemical processes. In order to increase specific surface area of the reaction zone in microreactors, we synthesized catalysts with surface area greater than 1000 m2/g using mesoporous MCM-41 support. The use of high surface area MCM-41 for FT catalysis stems from our previous studies, which can be prepared easily by one-pot hydrothermal procedure and are extremely stable, for steam reforming of methanol to produce hydrogen [48–51]. Bimetallic catalysts containing Co and one other metal—Fe, Ru, or Ni—were prepared to investigate the synergistic effect of bi-metallic species on the FT performance (manuscript submitted), The results show that CoRu-MCM-41 is more active than other bimetallic catalysts in producing longer-chain hydrocarbons at one atmosphere.

In this manuscript, we have focused on the kinetics of FT synthesis in a 3-D printed Stainless Steel(SS) microreactor using CoRu bimetallic catalysts supported by MCM-41, and two other mesoporous silica supports: SBA-15 and KIT-6. In order to translate new advancements in the laboratory as well as industry on both catalysis and microreactors for FT synthesis, chemical kinetics is a key issue in developing mathematical models for the reactors. However, to our knowledge, the kinetics of FT synthesis using mesoporous materials in a microreactor is relatively unknown in the literature. So, in order to understand more about the interaction between silica mesoporous materials and the metal, and especially kinetics, three different types of silica mesoporous materials (MCM-41, SBA-15, and KIT-6) containing cobalt and ruthenium metals were synthesized by one-pot hydrothermal method. To address the thermodynamic stability of the catalysts, CO-conversion using these three catalysts in the 3-D printed SS microchannel microreactor was also investigated.

#### **2. Results and Discussion**

#### *2.1. Catalyst Characterization*

#### 2.1.1. Textural Evaluation of Catalysts

The textural properties of the catalysts were evaluated using nitrogen Brunner–Emmett–Teller (BET) physisorption analysis. Table 1 shows the BET surface area, pore volume and the average pore diameter of all three catalysts. The surface areas of the catalysts are different depending upon the type of silica support. While the surface area of CoRu-MCM-41 was 1025 m2/g, that of CoRu-SBA-15 and CoRu-KIT-6 was around 691 m2/g and 690 m2/g, respectively. The general trend is consistent with that reported in the literature [48,52,53]. The pore diameter in the range of 3.2–5.3 nm was obtained from BJH desorption plot. The pore volume was in the range of 0.77–0.92 cm3/g. Figure 1a shows nitrogen adsorption–desorption isotherms for all the catalysts with pore size distribution of mesoporous materials. These isotherms represent the category of Type IV isotherms which is typical for mesoporous materials as mentioned in IUPAC classification [54]. These isotherms are classified into three different types of regions. The initial part of isotherm is a linear increment of nitrogen uptake at lower relative pressures (P/P0 = 0–0.2) called Type II isotherm. This is due to the adsorption of N2 on monolayer and multilayer within the pore walls of the catalyst. For relative pressure in the range of P/P0 = 0.2–0.4, there is an exponential increment in the isotherms which indicates the ordered mesoporous structure of the catalysts. Especially, the steepness of CoRu-MCM-41 is sharp when compared to the other two samples which indicate that MCM-41 support is more ordered in the nature of all the catalysts. Finally, the third region, in the relative pressure range of P/P0 = 0.4–0.95, has a long plateau for all the catalysts and it corresponds to the multilayer adsorption on the outer surface of the catalyst. The hysteresis loop for the samples is associated with condensation of N2 uptake in the interstitial voids of mesopores of the support [55]. Figure 1b shows pore size distribution obtained from BJH desorption plots. A sharp single peak for pore size with narrow distribution is observed for all the catalysts covering uniformly the pores with sizes in the range of 3.2 to 5.3 nm as shown in Table 1. The pore sizes of KIT-6 support appear to be larger and wider when compared to that of MCM-41 and SBA-15 supports. Furthermore, the pore distribution of MCM-41 is bi-modal, having major pores distributed in the range of 2 to 3 nm and minimal pore distribution between 3–4 nm.

**Table 1.** Brunner–Emmett–Teller (BET) surface area, pore size, pore volume and EDX metal loadings of synthesized catalysts.


<sup>a</sup> = Variation range ±2%, <sup>b</sup> = Variation range ±3%, <sup>c</sup> = Variation range ±5%.

**Figure 1.** (**a**) N2 adsorption–desorption isotherms of CoRu-S (S = MCM-41, SBA-15, KIT-6) and (**b**) pore size distribution of the catalyst.

#### 2.1.2. SEM-EDX Analysis

The metal loadings in the catalysts (wt%) and the surface morphology were obtained by SEM-EDX analysis. Figure S1 shows the SEM-EDX images of a typical MCM-41 catalyst showing uniform metal distribution with porous morphology. The actual and intended metal loadings are quite similar (Table 1) and suggest that the one-pot hydrothermal synthesis is one of the best routes to prepare mesoporous materials with uniform metal distribution. This uniformity plays a key role in the activity of FT catalysts; it not only decreases sintering but also increases the thermal stability of the catalysts for long-term studies.

#### 2.1.3. Transmission Electron Microscopic (TEM) Imaging

The size of the metal particles and the structure of the mesoporous support in all catalysts were obtained from TEM studies. The high magnification images in Figure 2 show the uniform ordered hexagonal pores present in the support. It is also worth noting that MCM-41 support has well defined hexagonal pores when compared to KIT-6 and SBA-15 and this is consistent with the BET surface area and the low angle XRD studies (discussed below). A uniform metal distribution with black dots, as shown in Figure S2, having almost circular in shape is clearly evident in the mesoporous silica matrix.

**Figure 2.** High magnification TEM images of CoRu-S Catalysts (S = (**a**) MCM-41, (**b**) SBA-15, (**c**) KIT-6.

2.1.4. Powder X-Ray Diffraction (XRD) Studies of Calcined Catalysts

In order to obtain information about the structural phases of the catalysts, XRD studies were carried out. Figure 3 shows the small angle XRD diffraction patterns for different mesoporous silica supported catalysts. The variations of peaks are probably due to the presence of metal nanoparticles present in the catalyst. For CoRu-MCM-41 catalyst, a sharp intense peak between 2-theta values 2–3◦ and two broad peaks between 2-theta values 4–5.5◦ corresponds to (100), (110), and (200) reflections of hexagonal mesoporous structure. This confirms that these catalysts are highly ordered mesoporous in nature with no deformation of hexagonal framework even after the addition of metals and this is consistent with the observed TEM images. For CoRu-SBA-15 catalyst, the peak between 2-theta value 1–2◦ indicates the mesoporous structure with 2D hexagonal symmetry with p6mm space group and long range ordered mesoporous structure [56]. For CoRu-KIT-6 catalyst, the peak at 2-theta value 0.94◦ corresponds to (211) plane and two low intensity peaks between 1.5–2◦ ascribes to (420) and (332) diffraction planes. These planes confirmed the characteristic three-dimensional nature of mesoporous KIT-6 reported in the literature [57].

**Figure 3.** Low angle XRD of three bimetallic mesoporous catalysts: (**a**) CoRu-MCM41; (**b**) CoRu-SBA15; (**c**) CoRu-KIT6.

The wide-angle XRD (WAXRD) analysis was carried out to determine the crystallinity of metal oxides in different mesoporous supports. Figure 4 shows the WAXRD patterns of these samples. The observed 2θ angles are compared with the JCPDS (Joint Committee on Powder Diffraction Standards) database. For all catalysts, the peaks at 18.90◦ (111), 31.09◦ (220), 36.74◦ (311), 38.36◦ (222), 44.72◦ (400), 59.25◦ (511), and 65.26◦ (440) correspond to the cubic structure of Co3O4 (JCPDS-42-1467) [58,59]. The orthorhombic structure of RuO2 (JCPDS-88-0323) is consistent with the observed peaks at 28.18◦ (110), 35.27◦ (101), and 54.56◦ (211) in all the catalysts.

#### 2.1.5. X-Ray Photoelectron Spectroscopy (XPS)

To determine the oxidation states of Co and Ru in MCM-41, KIT-6, and SBA-15, XPS studies were performed. Figure S3 shows the XPS spectra of Si 2p and O 1s containing a single spectrum which is centered at 104 eV and 532 eV, respectively, and confirms the presence of silicates in the sample. Figure 5a shows the Co 2p spectra for all the samples; the Co 2p3/2, and Co 2p1/<sup>2</sup> peaks are clearly observed to indicate the presence of cobalt in two oxidation states in the silica matrix [60–62]. The peaks centered at 780.5 eV and 796.2 eV are associated with Co 2p3/<sup>2</sup> and Co 2p1/2, respectively in the MCM-41 matrix. Whereas, in the case of KIT-6, the peaks for Co 2p3/<sup>2</sup> and Co 2p1/<sup>2</sup> are observed at 779.6 eV and 795.8 eV, respectively. For SBA-15, the similar peaks are noticed at 779.7 eV and 794.8 eV, respectively. It is clear from these data that the binding energy for cobalt in the MCM-41 matrix is distinctly higher when compared to that of cobalt in KIT-6 and SBA-15. This suggests that cobalt in

two oxidation states in MCM-41 is in a different environment from that of SBA-15 and KIT-6. This is also consistent with temperature programmed reduction (TPR) profile showing much higher reduction temperatures for CoRu-MCM-41 catalyst as discussed below. Similar XPS spectra for Co 2p were observed and analyzed by Bhoware et al., [63]. Figure 5b shows the XPS spectra for the ruthenium metal in the catalyst. The presence of Ru in the sample is confirmed by the Ru 3d spectra which is centered almost at 284.8 eV and it is associated with the Ru 3d3/<sup>2</sup> oxidation state [64]. However, in contrast to cobalt, there is no significant difference in the binding energy of the Ru metal ions in different mesoporous silica supports.

**Figure 4.** Wide angle XRD patterns of three bimetallic catalysts: (**a**) CoRu-MCM41; (**b**) CoRu-SBA15; (**c**) CoRu-KIT6.

**Figure 5.** XPS spectra of metals incorporated MCM-41, KIT-6, and SBA-15: (**a**) Co 2p spectra and (**b**) Ru 3d spectra.

2.1.6. H2 Temperature Programmed Reduction (H2 TPR)

Temperature programmed reduction (TPR) is an ideal technique to analyze the reduction behavior of metal oxides in mesoporous silica. It helps to investigate the interaction between metal and the support by providing information on physiochemical properties of the material. All the calcined catalysts are treated with 10%H2 to record TPR profiles for Co and Ru metal oxides shown in Figure 6. All the samples contain well defined peaks for ruthenium at low reduction temperatures and cobalt at much higher reduction temperatures. The TPR profiles of all the catalysts show that the ruthenium

oxide is reduced with H2 at a relatively lower temperature between 100 ◦C and 250 ◦C with main hydrogen consumption peaks for Ru3<sup>+</sup> to Ru<sup>0</sup> which is also reported by Panpranot et al. [65]. However, the reduction behavior of Co3O4 (Co3O4 <sup>→</sup> CoO <sup>→</sup> Co0) [66] in three samples to Co0 is remarkably different depending on the type of support. For MCM-41, as reported by Lim et al., the small peak centered around 310 ◦C ascribes the reduction of cobalt to CoO, while the second main peak corresponds to the reduction of CoO to metal ions Co2<sup>+</sup> into the silica network [67]. The last hydrogen uptake has a peak centered almost 780 ◦C which suggests that the cobalt and the MCM-41 support have strong interaction which is also confirmed by the binding energy obtained from XPS in Figure 5a [48,68]. This could be due to the formation of a spinel structure as cobalt silicates [69] and consistent with the XPS and XRD data. Unlike MCM-41, the reduction temperatures of Co inSBA-15 and KIT-6 were quite low around 365 ◦C and 375 ◦C, respectively, confirming weaker metal interactions with SBA-15 and KIT-6 supports [53]. However, no separate three peaks were observed for the reduction of cobalt, this may be due to the absence of silicates in SBA-15 and KIT-6 samples. The shift in the reduction peaks to the lower temperatures can also be due to the incorporation of Ru metal in the support [70]. Although the 5% weight of Ru is maintained in the catalyst sample, there might be a slight difference in the actual loadings of the Ru metal as shown in EDX Table 2. Qin et al., have studied the effect of the Ru metal on Co-SBA-15 catalyst at different loading and found a remarkable effect on the activity of catalyst during the FT synthesis [70]. Thus, the overall interactions of metal–metal and metal–support have a strong influence on the reducibility and reactivity of the catalysts for FT synthesis. Since the operating temperature zone of the FT synthesis is less than 350 ◦C, the activity of the catalyst is more dependent on the ease of reducibility of the metal oxides to pure metals (active sites) in the support.

**Figure 6.** H2-TPR(temperature programmed reduction) profiles of CoRu-S (S = MCM-41, KIT-6 and SBA-15) Catalysts.

The amount of hydrogen consumed by CoRu-MCM-41, CoRu-SBA-15, and CoRu-KIT-6 catalysts was calculated and quantified to be 0.108, 0.05, and 0.032 mmol H2/gm, respectively in the temperature range of 25 to 1000 ◦C. The amount of hydrogen consumed by CoRu-MCM-41 was found almost two times more than that of CoRu-SBA-15 and 3 times more than by CoRu-KIT-6. Higher hydrogen uptake by CoRu-MCM-41 is most likely due to the reduction of Co-silicates ~750 ◦C.

Although MCM-41 exhibits higher hydrogen consumption than other catalysts, the amount of hydrogen adsorbed by CoRu-KIT-6 in the temperature range of 25–311 ◦C is higher when compared to that by CoRu-MCM-41 and CoRu-SBA-15 in the same range of temperature.

#### *2.2. FT Reaction Mechanism*

In order to have a better understanding of the effect of metal and support interaction on catalysts, kinetic studies were carried out in the SS microchannel microreactors. The main difficulty to describe the FT kinetics is the complexity of its mechanism and a larger number of possible chemical species involved. Kinetic models of FT synthesis using Co based catalysts are less abundant than Fe based catalysts in literature [71]. Most of the existing models are mainly based on power-law models where Langmuir–Hinselwood (LH) type equations have been used by different researchers [72–76]. Although the simple power-law expression is widely recognized in the field of catalysis, it was recognized to have limited application in FT synthesis due to the narrow range of reaction conditions [77,78]. However, LH type equations are widely used for prediction of rates over a wide range of reaction conditions. As an example, Yates and Satterfield [72] worked on Co-catalyst and fitted the rate data obtained at 220–240 ◦C. They found that the rate data were best fitted with simple LH expression. Rautavuoma and van dar Baan [79] reported the rate of reaction at 1 atm pressure and 250 ◦C. They observed that reaction proceeds through CO dissociation and formation of -CH2- surface intermediate.

#### 2.2.1. Reaction Mechanism

In the microchannels of the microreactor, the flow of reactants is basically laminar. The complexity of the microchannel microreactor increases due to the parabolic type velocity profile; so, an average velocity profile is approximated during the development of the model for the microreactor system [80]. The outlet concentrations of the limiting reactant (CO), which was related to the rate of reaction, were calculated by an in-line GCMS. The following differential equation was used for a reactor model defined as Equation (1):

$$\frac{W\_{\rm cat}}{F\_{\rm in,CO}} = \int\_{X\_{\rm CO,in}}^{X\_{\rm CO,out}} \frac{dX\_{\rm CO}}{-r\_{\rm CO}} \tag{1}$$

*Wcat* = Wt. of the catalyst *Fin*,*CO* = Molar feed rate of CO *XCO* = Conversion of CO −*rCO* = Disappearance rate of CO

Equation (2), below, is used to calculate the disappearance rate of CO

$$1 - r\_{\text{CO}} = \frac{X\_{\text{CO}} F\_{\text{in,CO}}}{W\_{\text{cat}}} \tag{2}$$

The following boundary conditions (BC) were used:

*W* = 0; *Fi* =Fi(inlet) *W* = W*cat*; *Fi* = Fi(exit)

The partial pressure of the compound was calculated using the following equations:

$$p\_i = \frac{m\_i}{\sum\_{i=1}^{N\_c} m\_i} P\_T \tag{3}$$

where *pi* is the partial pressure of the component, *PT* is the total pressure of the reactor at the inlet (1 atm) and *Nc* is the total number of components. *mi* is the number of moles of component *i*.

#### 2.2.2. Mechanism and Kinetics

In order to determine the most suitable kinetic model for a particular catalyst, all possible combinations of FT reactions were considered and rate equations were developed based on CO conversion. A number of Langmuir–Hinshelwood and Eley–Rideal models have been developed for kinetics of CO hydrogenation to hydrocarbons in different types of reactors over the past few years [80–82]. In this study, it was assumed that the FT reactions occur only at active sites and proposed six possible mechanisms for FT reactions to develop kinetic models in the microchannel microreactor as shown in Table 2. In order to derive an appropriate model that describes a suitable FT equation, we considered six cases with different elementary reaction steps for each case as shown in Table 2.


**Table 2.** Elementary reaction steps for Fischer–Tropsch synthesis.

In FT-1, CO is adsorbed on active site (\*) of catalyst to form a CO\* intermediate. Then, CO\* reacts with H2 to give the products C (hydrocarbons) and D (H2O). Similarly, H2 can be adsorbed on the catalyst site (\*) to form H2\* intermediate and this intermediate subsequently reacted with CO to yield products in the FT-2 mechanism. There are two steps of adsorption in the FT-3 mechanism. In 1st and 2nd steps, CO and H2 both are adsorbed on catalyst active site (\*) to form two intermediates (CO\* and H2\*). In the last step (surface reaction), these two intermediates react with each other to give products C and D. The FT-4 model consisted of three different steps. In the 1st step (adsorption), CO is adsorbed on catalyst site (\*) to form CO\*. In 2nd step (surface reaction), the intermediate (CO\*) reacts with H2 to form another intermediate (COH2\*). In the last step (desorption), the final intermediate gave products (C and D). Like FT-4, FT-5 also consists of three different steps—adsorption, surface reaction, and desorption. In the 1st step (adsorption), H2 is adsorbed on catalyst site (\*) to form the H2\* intermediate. This intermediate reacts with CO to form another intermediate (COH2\*) in the 2nd step (surface reaction). In 3rd step, the products (C and D) are formed from the intermediate (COH2\*). In contrast to other models, FT-6 consists of four different steps. The 1st and 2nd steps are like that of the FT-3 mechanism. The 3rd step is the surface reaction where two intermediates (CO\* and H2\*) react with each other to give another intermediate COH2\* and released one active site (\*). In last step (desorption), the intermediate (COH2\*) yields products (C and D).

Using the six models described above, six rate equations can be deduced for FT reactions by considering surface reaction and rate-limiting desorption as shown in Table 3. (See Appendix A for the rate equation derived using the FT-3 kinetic model).


**Table 3.** Proposed kinetic equations for Fischer–Tropsch synthesis.

All the models presented in Table 3 were verified against experimental data to obtain the best suitable mechanism with the best fit. The models FT-1, FT-2, FT-4, and FT-5 are based on Eley–Rideal-type mechanism. In this case, one reactant gets adsorbed and another reactant reacts directly from the gas phase to form intermediates that yield products. Other models FT-3 and FT-6 are based on the Langmuir–Hinshelwood mechanism, which means all reactants are adsorbed on the catalyst surface before the products are formed. The kinetic parameter (*k*) and equilibrium constants (*K1, K2, K3*) at each temperature were evaluated by non-linear regression analysis based on Levenberg–Marquart algorithm in POLYMATH software by minimizing the sum of squared residuals of reaction rates [83,84]. The objective function is defined as:

$$F = \sum\_{i=1}^{N} \left( r\_{cal\_i} - r\_{\exp\_i} \right)^2 \tag{4}$$

where, *N* is the number of total observations, *rcali* and *r*exp*<sup>i</sup>* are calculated from the model equation and experimental rates at the different reaction temperatures.

The rate constants and the equilibrium constants can be related to The Arrhenius equation and van't Hoff laws as shown below:

$$k\_i(T) = A \exp\left(-\frac{E\_{\rm ui}}{RT}\right) \tag{5}$$

$$K\_i(T) = K \exp\left(-\frac{\Delta H\_i}{RT}\right) \tag{6}$$

where *ki* and *Ki* are reaction and equilibrium constants, respectively. *Eai* and Δ*Hi* are the apparent activation energy and standard enthalpy change of *i* species.

#### *2.3. E*ff*ect of Space Velocity on CO Conversion*

The influence of the weight hourly space velocity (WHSV) on the CO conversion for three different catalysts at 1 atm and H2/CO molar ratio 2 with error bar is shown in Figure 7a–c. The reactions were carried out at three different temperatures (210 ◦C, 240 ◦C, and 270 ◦C). CO conversion increases with the increase of space velocity and temperature. While CO conversion increases quickly with the increase of space velocity at the beginning, it remains almost constant at higher space velocity as the reaction reaches the equilibrium state. The variation of CO conversion was within 10% as observed during these reactions.

(**c**)

**Figure 7.** Effect of space velocity on CO conversion: (**a**) CoRu-MCM-41; (**b**) CoRu-SBA-15; (**c**) CoRu-KIT-6.

#### *2.4. FT Kinetic Model*

The kinetic models derived from Langmuir–Hinshelwood and Eley–Rideal mechanisms consider elementary reactions consuming CO and H2 to produce hydrocarbons and water. Recently, the mechanistic aspects of FT synthesis were well investigated by computational catalysis studies using DFT-based quantum chemical models [85–87]. However, the use of Langmuir–Hinshelwood and Eley–Rideal models facilitates understanding of the FT mechanism more easily. In this work, all the kinetic models derived based on these mechanisms are investigated and fitted with experimental data to check the feasibility of the proposed mechanism. The objective function (F) in Equation (4) was utilized to measure the goodness of the model to select the best-fitted mechanism for FT synthesis. Table 4 shows the experimental data obtained for all catalysts at 210 ◦C, 240 ◦C, and 270 ◦C.


**Table 4.** Kinetic experimental data for all the catalysts.

The kinetic parameters obtained for all the mechanisms for CoRu-MCM-41 are shown in Table 5. It can be inferred from data that the value of the rate constant (*k*) increases with increasing temperature with only one of the 6 mechanisms which is FT-3. Therefore, for CoRu-MCM-41, the model FT-3 is best fitted with the kinetic data.


**Table 5.** Kinetic parameters obtained from proposed mechanisms for CoRu-MCM-41

In order to determine the activation energy and the frequency factor from the Arrhenius equation, the logarithm of the rate constant was plotted against the inverse of reaction temperature as shown in Figure 8. The activation energy was determined to be 32.21 kJ/mol and the frequency factor was 4099

kmol/KgCat.hr. (atm)2 for CoRu-MCM-41 catalyst. Table 6 shows that the rate constant increases with the increase of reaction temperature. However, the two adsorption equilibrium constants (*K1* and *K2*) decrease with the increase of reaction temperature. Since adsorption is an exothermic process, the adsorption equilibrium constant decreases with rise in temperature.

**Figure 8.** Arrhenius plot-based on Langmuir–Hinselwood (LH) model for FT synthesis over CoRu-MCM-41 catalyst.

Table 4 shows the experimental data of the kinetic runs for CoRu-SBA-15 at 210 ◦C, 240◦C, and 270 ◦C. When the data are fit against all the kinetic models, FT-2 was the best-fitted model obtained for CoRu-SBA-15. The kinetic parameters for all of the proposed mechanisms are shown in Table 6.


**Table 6.** Kinetic parameters obtained from proposed mechanisms for CoRu-SBA-15.

The activation energy and frequency factor of this catalyst were evaluated from the Arrhenius equation by plotting the logarithm of the rate constant to the inverse of reaction temperature as shown in Figure 9. The activation energy was determined to be 13.39 kJ/mol and the frequency factor was 254 kmol/KgCat.hr. atm. Table 6 shows that for FT-3 mechanism, the reaction rate constant increases with reaction temperature and adsorption equilibrium constant (*K1*) decreases with the increase of reaction temperature.

**Figure 9.** Arrhenius plot-based on LH model for FT synthesis over CoRu-SBA-15 catalyst.

Table 4 shows the experimental data for CoRu-KIT-6 catalyst. All the models were fitted with this experimental kinetic data and FT-6 was found to be the best fitted model for CoRu-KIT-6. The kinetic parameters for all the proposed mechanisms are shown in Table 7 and the activation energy from Figure 10 was determined to be 12.59 kJ/mol and the frequency factor was 39 kmol/KgCat.hr. atm.


**Table 7.** Kinetic parameters obtained from all the proposed mechanisms for CoRu-KIT-6.

**Figure 10.** Arrhenius plot-based on LH model for FT synthesis over CoRu-KIT-6 catalyst.

Based on our experimental and kinetic model data, it can be concluded that only one of the six mechanisms for each catalyst is statistically relevant for fitting the model. The rate constant (*k*), for some of the other five mechanisms, does not show an increasing trend with the increase in temperature or remains constant, while the equilibrium constants, *K1* and *K2*, did not show the decreasing trend with the increase in temperature. Thus, FT-3, FT-2, and FT-6 mechanisms were considered as kinetically relevant model equations for CoRu-MCM-41, CoRu-SBA-15, and CoRu-KIT-6, respectively.

The results from our studies in a microreactor are similar to those reported in literature. Mansouri et al., developed a similar mechanism to estimate kinetic parameters for FT synthesis using cobalt-based catalyst with silica support and found that the experimental data were best fitted with surface reaction mechanism proposed based on Langmuir-Hineshelwood model and the optimal activation for the proposed kinetic model was found to be 31.57 kJ/mol [88]. Very recently, Sonal et al., detailed mechanistic approach for FT synthesis based on Langmuir–Hinshelwood–Hougen–Watson (LHHW) and Eley–Rideal using Fe–Co based catalyst. They claimed that a mechanism based on the adsorption and desorption have a satisfactory fit to the experimental data with the activation energies for the formation of methane, paraffin and olefin to be around 70 kJ/mol, 113 kJ/mol, and 91 kJ/mol, respectively [89]. In order to improve the efficiency of FT synthesis significantly, detailed kinetic rate expressions were derived which is very similar to our work reported in literature for both fixed bed reactor as well as microreactor using iron or cobalt-based catalyst [80,81,90,91]. The elementary steps in the above studies were used to develop kinetic mechanisms considering FT reactions with and without water gas shift (WHS) reactions occurring on the surface of the catalysts forming intermediates with active sites. A similar approach was considered in this present study where CO\*, COH2\* are assumed to form as intermediates, where \* is an active site of the catalyst. From the activation energy calculations, shown in Figures 8–10, the FT activation energy is observed in the order, CoRu-MCM-41 > CoRu-SBA-15 > CoRu-KIT-6. Almost 20 kJ/mol less activation energy was obtained for SBA-15 supported catalyst than that of MCM-41 catalyst. The activation energy of KIT-6 supported catalyst is a bit less than that of SBA-15 supported catalyst. This reflects that activation energy depends on metal–support interactions in different mesoporous catalysts. The variation of activity in different mesoporous catalysts might arise due to experimental uncertainties and operating conditions [92]. In addition, the FT activation energy is sensitive to the reactor system. Sun et al., [93] reported that the activation energy in a microchannel microreactor is smaller than that observed in a fixed bed reactor (FBR).

Figure 11 shows the variation of the model predicted rate with the experimental rate for all catalysts. The best fitted models i.e., FT-3, FT-2, and FT-6 for CoRu-MCM-41, CoRu-SBA-15, CoRu-KIT-6, respectively, were chosen to plot the graph for predicted and experimental rates. The correlation

coefficient (R2) for all cases was above or equal to 0.95 with an error band of <sup>±</sup>15% to <sup>±</sup>30%. This indicates that the error between experimental and predicted values lies within the statistical permissible limits at all reaction temperatures for all the catalysts and consistent with the mechanistic models proposed in the literature. Moazami et al., conducted kinetic studies for FT synthesis in a fixed bed reactor with cobalt-based catalyst over silica support and found that 60% of the results were predicted with a relative error of less than 15%, while the rest of the proposed kinetic models has error less than 32% with confidence interval of 0.99 [94]. They also proposed a pseudo-homogenous one-dimensional model to evaluate the kinetic performance of the catalyst and achieved less than 8% error with the predicted data for kinetic experiments [95]. More recently, Marchese et al., performed kinetic studies with Co-Pt/γ-Al2O3 catalyst in a lab-scale tubular reactor and reported an error band around ± 25% with a confidence level of 0.95 stating it lies in the suitable acceptable limits with many mechanistic models proposed in the literature [89,96–98].

(**c**)

**Figure 11.** Experimental rate vs predicted rate for all temperature of three different catalysts: (**a**) CoRu-MCM-41; (**b**) CoRu-SBA-15; (**c**) CoRu-KIT-6.

#### *2.5. Deactivation Studies*

In order to further understand how the interaction of Co and Ru metals with different mesoporous silica supports affects the stability of the FT catalysts, deactivation studies were performed. Figure 12 shows the deactivation rates of the catalysts tested continuously for 60 h at 240 ◦C, 1 atm, and H2:CO ratio of 2:1. All the catalysts maintained fairly consistent CO conversion with very little fluctuation during the first 10 h. More specifically, the catalysts maintained 65%–79% CO conversion in the first 10 h of the reaction with CoRu-KIT-6 exhibiting the highest conversion and CoRu-MCM-41, the lowest. The activity of all the catalysts dropped by 20% after 24 h and then started to decline further. At the end of 60 h, the activity of the CoRu-MCM-41 dropped by 70% whereas, in the case of CoRu-KIT-6 and CoRu-SBA-15, the CO conversion decreased by 64% and 84%, respectively. Our results suggest that in terms of stability, the support has a significant impact on FT performance. More significantly, MCM-41 and KIT-6 supports are more stable when compared to the FT stability studies with SBA-15.

**Figure 12.** Deactivation studies of the catalysts at T = 240 ◦C, P = 1 atm and H2:CO = 2:1.

Many deactivation mechanisms have been proposed for FT studies that include catalysts poisoning, sintering, oxidation, the effect of water, carbon deposition and surface reconstruction [99]. The syngas used in our studies is a mixture of ultrahigh pure 5.0 CO and H2 gases; therefore, there is very little or no chance of catalyst deactivation due to poisoning by the gas feed at the inlet to the reactor. It was also observed in our previous studies that the support (SiO2 vs TiO2) can enhance the stability of the catalyst to resist deactivation [38,40]. Iglesia et al., noticed that silica supported materials are less stable when compared to the other supports like Al2O3 [100] in their FT studies with Co-based catalyst. They also reported that CoRu-TiO2 and CoRu-SiO2 were found to have almost the same activation energy upon the addition of Ru to the Co catalyst; however, there were strong differences in the deactivation rates of catalysts depending upon the support [101]. Based on our CO- conversion studies, the ability of catalysts to withstand or retard the FT deactivation rate was in the order of CoRu-KIT-6 > CoRu-MCM-41 > CoRu-SBA-15.

#### **3. Materials and Methods**

#### *3.1. Materials*

The reagents used for catalysis synthesis were of analytical grade with no further purification. Tetramethyl orthosilicate, 99% (TMOS) and ammonium hydroxide, Tetraethyl orthosilicate reagent grade, 98% (TEOS), Pluronic acid (P123), Hydrochloric acid (HCl), Cetyltrimethylammoniumbromide (CTAB), Co(NO3)2.6H2O, RuCl3·xH2O were purchased from Sigma Aldrich. Ethanol (anhydrous), Butanol and acetone, ACS grade, were obtained from Fischer Scientific, Branchburg, New Jersey, USA.

#### *3.2. Fabrication of Microreactor*

The microchannel microreactor and the respective cover channel were fabricated using 3D printing technology. Typically, the microreactor and its cover channel are designed using AutoCAD software which is schematically shown in Figure 13. The design is based on the split and recombination principle which has 11 microchannels of 500 μm × 500 μm × 2.4 cm as reaction zone in between them. This stainless-steel 3D printed microreactor is assembled in a custom-built heating block with an inlet and outlet system which facilitates the flow of syngas through the channels.

**Figure 13.** (**a**) and (**b**): AutoCAD design of the microreactor and cover channel (**c**) 3D printed reactor, (**d**) SEM image of the microchannels coated with catalyst prior to FT studies.

#### *3.3. Catalyst Synthesis and Loading*

Three types of Co-Ru based nanocatalysts supported by different mesoporous silica supports i.e., MCM-41, SBA-15 and KIT-6 supports are used in this study. A constant metal loading of 10%Co and ~5% Ru in weight was maintained in all preparations and this metal loading was also determined using the amount of the precursor. Three catalysts using different mesoporous support—10%Co5%Ru-MCM-41, 10%Co5%Ru-SBA-15 and 10%Co5%Ru-KIT-6—were synthesized using the one-pot hydrothermal procedure ( as shown below) [48]. The catalysts were labeled as CoRu-MCM-41, CoRu-SBA-15, and CoRu-KIT-6 in this manuscript.

For the synthesis of CoRu-MCM-41, TMOS, CTAB, DI-water, and ethanol were used in a molar ratio of 1:0.13:130:20 as described elsewhere [48]. In short, CTAB was dissolved in DI-water at 30 ◦C to produce a clear solution. The metal precursors were dissolved in ethanol in a separate beaker. The precursor, TMOS, which is a limiting agent for this synthesis, was added dropwise to the mixture of the two solutions prepared previously. Ammonium hydroxide was added dropwise to precipitate metal hydroxides in the solution, till the final pH was ~10. The precipitate was stirred for 3 h, followed

by 18 h of aging at 65 ◦C. The precipitate was then washed with DI water till the filtrate reached a pH of 7, finally washed with ethanol and filtered. The filtered material is air-dried for a day and then oven-dried at 110 ◦C for 24 h. The dried catalyst is calcined at 550 ◦C for 16 h with a ramp rate of 2 ◦C/min to remove the CTAB template.

For the synthesis of CoRu-SBA-15: TEOS, CTAB, water, ethanol, pluronic acid, and hydrochloric acid were mixed in molar ratios of 1:0.081:41:7.5:0.0168:5.981. In a typical synthesis procedure, P123 was dissolved in 2M HCl at 35 ◦C till a clear solution was obtained. Another solution was prepared by dissolving CTAB in DI water at 35 ◦C until a homogenous mixture was produced. These two solutions were mixed and stirred for 35 min. Ethanol containing metal precursors were added dropwise into the solution and stirred for 30 min. Afterwards, TEOS which was limiting reagent in this procedure was also added dropwise and stirred for 20 h at 35 ◦C. The aqueous solution was aged for 48 h at 98 ◦C followed by air drying for 24 h. The material was then oven-dried at 110 ◦C for 24 h. Finally, the dried material is then calcined in stepwise fashion with heating rate of 1 ◦C/min at 350 ◦C, 450 ◦C, and 550 ◦C for 8 h each, respectively to remove CTAB and pluronic acid.

In the case of CoRu-KIT-6, TEOS, P123, HCl, DI water, and butanol were mixed in a molar ratio of 1:0.017:1.83:195:1.31 [102]. For a typical procedure, P123 was added to HCl at 35 ◦C till a clear solution was obtained. A separate solution was prepared with butanol containing metal precursors and poured to the previous solution and stirred until a homogeneous solution was obtained. To this mixture, TEOS, which was the limiting reagent, was added dropwise and stirred at 500 rpm for 24 h. The final solution was aged for 24 h at 100 ◦C, followed by air drying for 24 h under the fume hood. The material is oven-dried at 110 ◦C for 24 h and then calcined at 550 ◦C for 4 h, to remove P123, the structure directing agent, SDA, with heating and cooling rates of 1 ◦C/min.

The catalyst is loaded into the microchannels of the microreactor using a PVA suspension containing the catalyst, DI water, binder PVA (polyvinyl alcohol 98%–99% hydrolyzed MW: 31000) and acetic acid of weight ratio 1:5:0.25:0.05. This suspension with well-dispersed catalyst was dip-coated and dried in air and then calcined in presence of air at 400 ◦C for 2 h with heating and cooling rates of 5 ◦C/min. Figure 13d shows the SEM image of the catalyst coated microreactor prior to the in-situ reduction.

#### *3.4. Catalyst Characterization*

Specific surface area, pore size, pore volume and TPR studies of the catalyst were carried out using Micromeritics, 3-Flex instrument. The Brunner–Emmett–Teller (BET) method was used to calculate the surface area of the catalyst where an equation was obtained from adsorption isotherm in the relative pressure range of 0.07–0.03. The surface area was calculated from adsorption isotherm in the relative pressure range of P/P0 = 0.07–0.3 using the Brunner–Emmett–Teller (BET) equation. The total volume per gram of catalyst was determined from the amount of N2 adsorbed at P/P0 = 1. The N2 desorption from the catalyst surface provides information about the pore size distribution using BJH (Barret–Joyner–Halenda) plots [103]. The H2 temperature programmed reduction (TPR) analysis was also done with the same instrument which has a TCD detector to monitor the reduction signals of the catalyst. Around 50 mg of the catalyst was loaded into the quartz sample tube in which a stream of 10% H2/Ar at flowrate 110 mL/min was passed through and the temperature is increased to 1000 ◦C with 10 ◦C/min ramp rate. The small and wide-angle powder x-ray diffraction (XRD) were recorded using D8 Discover X-ray and Rigaku SmartLab X-ray diffractometers, respectively, with Cu K-alpha radiation (wavelength = 0.15418 nm) radiation generated at 40 mA and 40 kV. The step size and time per step used in these measurements are 0.05◦ and 3 secs/step, respectively. The crystal sizes of the metal oxides were determined using the Scherrer equation. In the Scherrer equation below, τ stands for the crystal size, λ is the wavelength of the Cu Kα radiation, β is the full width half maximum and θ is the Bragg diffraction angle.

$$
\pi = \frac{0.9\lambda}{\beta \* \text{Cost}\theta} \tag{7}
$$

The morphology and the size of the catalysts were analyzed using transmission electron (TEM Carl Zeiss Libra 120) at 120 KeV and scanning electron microscopy (Zeiss Auriga FIB/FESEM). The sample for TEM was prepared by dispersing a small quantity of catalyst in 3 mL of ethanol followed by vortex dispersion and sonication for a few minutes. Then the suspension was drop coated on a carbon-coated copper grid of 300 μm mesh size, followed by drying in an oven at 100 ◦C for 12 h.

The elemental composition and oxidation states of the metals were analyzed using Energy Dispersive X-ray spectrometry (Zeiss Auriga FIB/FESEM obtained from Carl Zeiss, Oberkochen, Germany) and oxidation states by X-ray photoelectron spectroscopy (XPS-Escalab Xi+-Thermo Scientific obtained from Thermo Scientific, West Sussex, UK), respectively.

#### *3.5. Fischer–Tropsch Synthesis in Microreactor and Kinetic Data Collection*

An in-house LabVIEW automated experimental setup was built to carry out the FT experiments for precise control over the operating conditions. The experimental setup is shown in Figure 14. The flowrates of the syngas mixture which is a mixture of hydrogen and carbon monoxide were controlled by precalibrated mass flow controllers obtained from cole parmer with flow rates ranging from 0–1 sccm. Nitrogen was used as a carrier gas into the system and was controlled by Aalborg mass flow controller with a maximum flow rate of 10 sccm. The upstream and downstream pressures were continuously monitored by pressure gauges obtained from Cole–Parmer and the data are fed to Aalborg solenoid valve from which the reaction pressure is controlled and kept constant throughout the reaction. All these controllers are operated by LabVIEW 2018 program. The product stream is directly fed to the GC-MS (Agilent Technologies 7890B GC and 5977 MSD). Prior to the start of the kinetic experiments, the microreactors were reduced ex-situ in Carbolite Gero tubular furnace with 10% H2Ar. To compensate the losses while transferring the microreactor to the heating block the microreactor containing the catalyst was reduced again in-situ for 6 h at 350 ◦C before the start of FT reaction. The kinetic studies were performed by varying the weight hourly space velocity (WHSV = Wcat/FCO,in, where WCat = weight of the catalyst and FCO,in = molar flow rate of CO in feed) in the range of ~25.2–3.15 kgcat.h/kmol. The reactions were performed with syngas having a feed molar ratio (H2/CO) of 2:1 at 210 ◦C, 240 ◦C, and 270 ◦C while the reaction pressure was maintained at 1 atm. Based on our previous FT studies using this setup and preliminary runs, all reactions reached a steady state after an hour at each setpoint of WHSV. Deactivation studies were also performed for all three catalysts at 240 ◦C using syngas feed molar ration of 2:1. CO conversion was calculated based on following equation:

$$X\_{\rm CO}\% = \frac{F\_{\rm CO,in} - F\_{\rm CO,out}}{F\_{\rm CO,in}} \times 100\tag{8}$$

**Figure 14.** Experimental set-up for FT synthesis in a 3-D printed stainless steel microreactor.

#### **4. Conclusions**

Three different types of mesoporous silica supported Co-Ru based catalysts were synthesized using the one-pot hydrothermal method and performance for FT was evaluated. These catalysts resulted in high surface area with hexagonal ordered mesoporous structure as supported by BET, low angle XRD and TEM studies. The interaction between the metal and different types of support has a significant effect on the kinetic and stability studies of FT synthesis. Six mechanistic models were developed based on the Langmuir–Hinshelwood and Eley-Radiel mechanisms. The best fitted model for all catalysts was obtained on the basis of non-linear regression by comparing with the objective function which is equation-7 in this paper. The proposed model FT-3 was best fitted with the kinetic data for CoRu-MCM-41 catalyst. Whereas, FT-2 and FT-6 were well fitted with the kinetic data for CoRu-SBA-15 and CoRu-KIT-6, respectively. CoRu-KIT-6 was found to be more active than other catalysts with a low activation energy of 12.59 kJ/mol, whereas for CoRu-SBA-15 and CoRu-MCM-41 the activation energies are 13.39 and 32.21 kJ/mol, respectively. An average error of 5.65%, 1.76%, and 3.70% was obtained for catalysts CoRu-MCM-41, CoRu-SBA-15, CoRu-KIT-6, respectively considering the best fitted model explained above for FT synthesis. The predicted data provided by kinetic models were satisfactory with the experimental data. These results highlight the potential of the mechanistic FT models as well as reaction mechanisms to further improve the performance of FT synthesis. In addition, this information can help to design more active and selective catalysts for the optimized FT process. Furthermore, all catalysts exhibited significant resistance to the deactivation rate following the order CoRu-KIT-6> CoRu-MCM-41>CoRu-SBA-15. This study suggests that even if the support is of same type, the structure of the support plays a vital role in catalyst performance for FT synthesis.

**Supplementary Materials:** The following are available online at http://www.mdpi.com/2073-4344/9/10/872/s1, Figure S1: title, Table S1: title.

**Author Contributions:** Methodology, analysis and experimental investigation by N.M. and S.B.; project administration and supervision by S.A., and D.K.; all the authors contributed to the discussion of the experimental results as well as writing and editing of the manuscript.

**Funding:** This project received funding from NSF CREST (#260326) and UNC-ROI (#110092). This work is performed at North Carolina A&T State University and Joint School of Nanoscience and Nanoengineering, a member of the Southeastern Nanotechnology Infrastructure Corridor (SENIC) and National Nanotechnology Coordinated Infrastructure (NNCI), which is supported by the National Science Foundation (Grant ECCS-1542174).

**Acknowledgments:** The authors gratefully acknowledge Dr.Kyle Nowlin and Mr. Klinton Davis for TEM support, and Dr. Xin Li for XRD support.

**Conflicts of Interest:** The authors declare no conflict of interest

#### **Appendix A**

Derivation of Kinetic model: FT-3 kinetic model Dual-site adsorption of CO and H2, the surface reaction is rate controlling step (RCS)

$$\text{CO} \, + \ast \, \frac{k\_1}{k\_{-1}} \, \text{CO} \, \ast \quad K\_1 = \frac{k\_1}{k\_{-1}} (\text{Adsorptions 1}) \tag{A1}$$

$$H\_2 + \*\underset{k\_{-2}}{\overset{k\_2}{\rightleftharpoons}} H\_2 \* \quad K\_2 = \frac{k\_2}{k\_{-2}} (\text{AdSorption 2}) \tag{A2}$$

$$\text{CO} \ast + \text{H}\_2 \ast \xrightarrow{k} \text{C} + \text{D} + \text{2} \ast \text{(Surface reaction)} \text{(RCS)}\tag{A3}$$

From step 1 (Adsorption 1 (rapid reaction)),

The rate of formation of CO\* is,

$$r\_{\rm CO\*} = k\_1 p\_{\rm CO} \mathcal{C}\_\* - k\_{-1} \mathcal{C}\_{\rm CO\*} \tag{A4}$$

From step 2 (Adsorption 2 (rapid reaction [89])), The rate of formation of H2\* is,

$$r\_{H\_2\*} = k\_2 p\_{H\_2} \mathbb{C}\_\* - k\_{-2} \mathbb{C}\_{H\_2\*} \tag{A5}$$

From step 3 (Surface reaction 3),

$$-r\_{\text{CO}} = k \mathcal{C}\_{\text{CO}\*} \mathcal{C}\_{H\_{2}\*} \tag{A6}$$

According to pseudo steady-state hypothesis (PSSH), the rate of formation of the intermediate is zero.

So,

$$
\sigma\_{\rm CO\*} = 0\tag{A7}
$$

$$r\_{\mathbb{H}\_2\*} = 0\tag{A8}$$

So, putting the value of *rCO*<sup>∗</sup> from Equation (A4) in Equation (A7), we have,

$$\begin{aligned} &k\_1 p\_{\rm CO} \mathbf{C\_\*} - k\_{-1} \mathbf{C\_{CO\*}} = 0 \\ &\Rightarrow k\_1 p\_{\rm CO} \mathbf{C\_\*} = k\_{-1} \mathbf{C\_{CO\*}} \\ &\Rightarrow \mathbf{C\_{CO\*}} = \frac{k\_1 p\_{\rm CO\*} \mathbf{C\_\*}}{k\_{-1}} \\ &\Rightarrow \mathbf{C\_{CO\*}} = K\_1 p\_{\rm CO} \mathbf{C\_\*} \end{aligned} \tag{A9}$$

So, putting the value of *rH*2<sup>∗</sup> from Equation (A5) in Equation (A8), we have,

$$\begin{aligned} &k\_2 p \mu\_2 \mathcal{C}\_\* - k\_{-2} \mathcal{C}\_{H\_2\*} = 0\\ &\Rightarrow k\_2 p\_{H\_2} \mathcal{C}\_\* = k\_{-2} \mathcal{C}\_{H\_2\*}\\ &\Rightarrow \mathcal{C}\_{H\_2\*} = \frac{k\_2 p\_{H\_2} \mathcal{C}\_\*}{k\_{-2}}\\ &\Rightarrow \mathcal{C}\_{H\_2\*} = K\_2 p\_{H\_2} \mathcal{C}\_\* \end{aligned} \tag{A10}$$

Taking the values of *CCO*<sup>∗</sup> and *CH*2<sup>∗</sup> from Equations(A9) and (A10) and putting in Equation (A6, we have,

$$-r\_{\text{CO}} = k \mathcal{K}\_1 \mathcal{K}\_2 p\_{\text{CO}} p\_{H\_2} \mathcal{C}\_\*^2 \tag{A11}$$

Making the catalyst active site balance. Considering, the total site is,

$$\begin{array}{l} \mathsf{C\_{T}} = 1 \\ \Rightarrow (\mathsf{No.of\ vacant\ sites}) + (\mathsf{No.of\ occuapied\ sites}) = 1 \\ \Rightarrow (\mathsf{C\_{\*}}) + (\mathsf{C\_{CO\*}} + \mathsf{C\_{H\_{2}\*}}) = 1 \\ \Rightarrow (\mathsf{C\_{\*}}) + (\mathsf{K\_{1}p\_{\complement}}\mathsf{C\_{\*}} + \mathsf{K\_{2}p\_{\varplement}}\mathsf{C\_{\*}}) = 1 \\ \Rightarrow \mathsf{C\_{\*}}(1 + \mathsf{K\_{1}p\_{\complement}} + \mathsf{K\_{2}p\_{\varplement}}) = 1 \\ \Rightarrow \mathsf{C\_{\*}} = \frac{1}{(1 + \mathsf{K\_{1}p\_{\complement} + \mathsf{K\_{2}p\_{\varspacebox{R\_{2}}}})}} \end{array} \tag{A12}$$

Putting the value of *C*<sup>∗</sup> from Equation (A12) in Equation (A11), we get,

$$\Rightarrow -r\_{CO} = \frac{kK\_1K\_2p\_{CO}p\_{H\_2}}{\left(1 + K\_1p\_{CO} + K\_2p\_{H\_2}\right)^2}(\text{FT} - 3)$$

[Taking values of *CCO*<sup>∗</sup> and *CH*2<sup>∗</sup> from Equations (A9) and (A10)]

#### **References**


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