Catalyst design for maximizing C5+ yields during Fischer-Tropsch synthesis

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Catalyst design for maximizing C5þ yields during Fischer-Tropsch synthesis Janardhan L. Hodala a, Dong J. Moon a, Kakarla Raghava Reddy b,**, Ch Venkata Reddy c, T. Naveen Kumar d, Mohd Imran Ahamed e, Anjanapura V. Raghu f,* a

Clean Energy Catalysis Lab, Korea Institute of Science and Technology, Hworang-ro 14-gil, Seongbuk-Gu, Seoul, South Korea b The School of Chemical and Biomolecular Engineering, The University of Sydney, Sydney, NSW, 2006, Australia c School of Mechanical Engineering, Yeungnam University, Gyeongsan, 712-749, South Korea d Department of Chemistry, R.R. Institute of Technology, Bengaluru 560090, India e Department of Chemistry, Faculty of Science, Aligarh Muslim University, Aligarhe202002, Uttar Pradesh, India f Department of Basic Sciences, Center for Emerging Technology (CET), School of Engineering & Technology, JAIN Deemed-to-be University, 562112, Karnataka, India

highlights
Recent progress on the C5þ selectivity in Fischer-Tropsch synthesis is discussed.
Nature of support influences product distribution.
Support does not affect the intrinsic activity, but facilities secondary reactions.
Products beyond ASF distribution is achieved by situ upgrading/cracking.

article info

abstract

Article history:

Fischer-Tropsch (FT) process has great potential to accomplish energy security but also for

Received 29 July 2019

utilizing greenhouse gases to address the energy problem. Different kinds of feedstocks like

Received in revised form

coal, biomass (via gasification), CO2, methane (via reforming), and nonconventional energy

3 November 2019

sources are used to obtain the syn-gas (CO and H2). The formation of hydrocarbons in the FT

Accepted 3 December 2019

process follows ASF distribution over the majority of the catalysts. It can be overcome by the

Available online xxx

application of a suitable catalyst, controlling the active metal interaction with the support and interaction of formed hydrocarbon with the support. The ratio of syn-gas is important to

Keywords:

maintain the desired conversion and to have more selectivity towards C5þ products. Increase

Synthetic clean fuels

in the H2: CO ratios in the feed increases C5þ products and methane decreases. Whereas with

Fischer-Tropsch synthesis

the decrease in the ratios increases undesirable reactions and methane formation. In this

Functional catalysts

article, we have discussed the recent literature from the viewpoint of increasing the C5þ

C5þ selectivity

selectivity. Support has a profound influence on product distribution. With the application of

Syn-gas

suitable support and controlling the interaction of the active sites yields the good CO conversion with fewer lighters and higher C5þ hydrocarbons. © 2019 Hydrogen Energy Publications LLC. Published by Elsevier Ltd. All rights reserved.

* Corresponding author. ** Corresponding author. E-mail addresses: [email protected] (K.R. Reddy), [email protected] (A.V. Raghu). https://doi.org/10.1016/j.ijhydene.2019.12.021 0360-3199/© 2019 Hydrogen Energy Publications LLC. Published by Elsevier Ltd. All rights reserved. Please cite this article as: Hodala JL et al., Catalyst design for maximizing C5þ yields during Fischer-Tropsch synthesis, International Journal of Hydrogen Energy, https://doi.org/10.1016/j.ijhydene.2019.12.021

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Introduction

Chain growth mechanism

Fischer-Tropsch synthesis (FTS) was first reported by Franz Fischer and Hans Tropsch, had influenced profoundly for the energy demands [1]. Its attractiveness decreased with the supply of cheap and abundant supply of fuels from crude oil. It was necessary to synthesis fuels to meet the energy demands and security for South Africa due to the embargo laid [1,2]. To meet the energy demands and abundant supply of coal, coal to liquid was followed and found to be more attractive. Recently more renewables are catching up and can produce a lot of energy during a stipulated time e.g. wind energy. The energy produced by wind highly depends on the velocities of it, results in fluctuating supply, similarly tidal and solar. In addition to that, there will be surplus energy in the grid during non-peak hours due to underutilization in the cities. To harness and to utilize this renewable power, it has to store in various useable forms. Producing liquid hydrocarbons via hydrogenation of CO2 or CO can be one such energy carrier [3]. The synthesis of hydrocarbons is done directly by FTS. Due to the drastic effect on the environment from greenhouse gases, stringent emission standards are enforced by the leading countries. Even though the diesel engines emit less CO2 (~20%) compared to gasoline engines. It has a major pollution problem in terms of sulfur, nitrogen particulates, CO, and aromatics. Fuels from FTS are free of Sulfur, nitrogen hence sulfur and nitrogen emission are not a problem. Additionally, it is observed that it also reduces the CO, Hydrocarbons, Polyaromatic hydrocarbons and particulate emissions compared to oil-derived diesel [4]. In recent days, global warming leading to climate change is happening at an alarming rate. To mitigate this FTS becoming an attractive and promising solution, as the greenhouse gases are converted to fuels. Syngas required for FTS is supplied by gasification of coal, biomass, reforming of methane and CO2. Though it is not directly removing these gases, it is closing the carbon cycle artificially. This is not the case while using crude oil-based fuels. FTS catalytically combines CO and H2, at elevated pressures (20e40 bar) and temperatures (200e350  C) to give hydrocarbons. At present, the FT fuel supply makes approximately 2% of the world’s fuel supplies [5]. FT products range from methane to hydrocarbon carbon chain length reaching as long as C90. Products can be simply called as syn-crude, and are separated into gas, gasoline, kerosene, diesel, and wax. Most of the cases FTS products are observed to follow Andersone SchulzeFlory (ASF) probability distribution. It is also observed that the wide distribution can be controlled or tuned by the application of suitable catalysts. In general, FTS is carried over Co or Fe catalysts supported on SiO2, Al2O3, TiO2, ZrO2, Zeolites, clays, etc (see Table 1). CO conversion is correlated to the dispersion of metallic Co and liquid hydrocarbon selectivity depends on the diffusion within the catalyst pores and secondary reactions [6e9]. In this review, for the first time, we discuss recent studies on various parameters that affect C5þ selectivity.

The mechanism of FTS is revisited and disputed for a long time. Several studies carried out and being studied with recent advancements to understand the mechanism better. Two major mechanisms are carbide mechanism and C(H)O insertion mechanism. DFT calculation with microkinetic modeling the adsorption free energies of all species, the activation free energy barrier and the reaction free energy of all elementary reactions supports the Carbide mechanism. CO direct dissociation and H assisted dissociation and subsequent hydrogenation forms dominant monomers CHx (x ¼ 1e3) species. Chain growth occurs by coupling of monomer species. At higher hydrogen coverage termination occurs by the hydrogenation of RCH2CH2 species. CHO insertion and subsequent hydrogenation to alkane or alcohol are less favored compared to RCH2 coupling with CH2 for CeC chain growth [10]. During the reaction over Ru and Co surfaces chain grows rapidly than the initiation step. It is inconsistent with the carbene like mechanism, as CHx which inserts to the hydrocarbon chain, forms only after the initiation step. Mechanistic insights derived from theoretical calculation show that the growing hydrocarbon chain increases the C1 monomers by increasing CO activation due to the ability to disrupt dense chemisorbed CO. In the carbine mechanism, CO is activated over vacant sites (adsorbed CO can also get desorbed), adsorbed CO and H reacts to form hydroxy methylene (CHeOH), which form OH and CH by dissociation. This formed CHx is the monomer that inserts into the growing carbon chain. Dense CO adsorbed layers are disrupted by adsorbed CH and larger alkylidyne (CnH2n-1) homologs. This results in the formation of CO OH vicinal adsorbed species before desorption. By which the effective activation enthalpy and free energy barriers decreased respectively by 100 and 15 kJ/mol near growing carbon chain. This kind of mechanism favored only at high coverages of CO [11]. The mechanism for FT reaction is schematically represented in Fig. 1. CHx monomer is in-situ produced over catalyst surface, a coupling of these two monomers generates the ethylidene. Propagation of this polymerization reaction is seen in the formation of propylidene by CeC bond formation of ethylidene and CH. Chain termination can happen at any place, CH, ethylidene, propylidene, etc, resulting in the formation of respective alkanes or alkenes. Formed alkenes can readsorb to form alkylidene, growth of these will result in methyl branched hydrocarbons. Over Co catalysts, about 5% of branched hydrocarbons are formed whereas in Fe catalyzed FTS it is about 25% [12]. Navarro et al. monitored the reaction over Co model catalyst under reaction conditions in the scanning tunneling microscope. Cobalt terraces are covered by parallel arrays of stripes within 30 min of reaction. Accommodating large chain hydrocarbon depends on the particle size and nanoparticles of a below certain size would be too small for it. Molecular overlayer may no be developed over the smaller particles because the shorter molecules do not reach sufficiently high concentrations over the terrace, whereas the larger particles would display extended surface behavior. Hence, initiation

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Table 1 e Activity comparision of Co based catalysts towards C5þ selectivity. Catalyst system Co/SBA-15 Co/Al2O3 Co/Al-SBA-15 Co/Al2O3 Co/Al2O3 Co/SiC Ru/Co/Al2O3 Co-gAl2O3 Co-Re-gAl2O3 CoeB-gAl2O3 Co-aAl2O3 Co-Re-aAl2O3 CoeTiO2 Co-Re-TiO2 CoeBeTiO2 Co/Zr/SiO2 CoePt-MMS CoePteSiO2 Co/KIT-6 CoZrP/KIT-6

Co/Graphene-Silica CoeAl2O3/Silica Co/SBA-16 Al2O3-modified mesoporous Co3O4 CoePeAl

Metal Loading %

Co Particle size (nm)

CH4 Selec%

C5þ Selec %

Ref

30 15 e 20 20 15 Co-15 Ru-0.3 Co 12 Co 12 Re-0.5 Co 12 B-0.15 Co 12 Co 12 Re-0.5 Co 12 Co 12 Re-0.5 Co 12 B-0.15 Co-20 Zr-5 Co-15 Pt-0.1 Co-15 Pt-0.1 20 Co-20 Zr e 12 P e 2.4 Co-22 Graphene e 0.1 Co- 20 Al2O3 e 1 Co-15 Al - 5 Co - 20% P/Al2O3 e 0.005

12 9.3 10 7.9 7.9 15.6 3.5

28 40 50 22 99 36 88

25 28 12 2.6 17 14 20

58 65 82 96 79 79 76

[4] [6] [7] [8] [8] [9] [18]

11 12

15 30.3

10 10

79 80

[25] [25]

11

19.4

9

82

[25]

20 19

14 16

8.8 8.9

84 85

[25] [25]

17 14

14.7 23.1

8 7

85 87

[25] [25]

14

15.5

5

87

[25]

e

88

11

76

[26]

11

92

7

83

[27]

9

19

12.2

71

[27]

22 12

30 60

23 16

57 72

[29] [29]

16

50

4.2

92

[30]

8

54

10.7

80

[31]

12 e e

40 90 30

12 4 4

78 88 90

[33] [38] [42]

time and chain-growth probability, reaction rate, and selectivity shall depend on the particle size of the catalysts [5] (see Fig. 2). Cruz et al. studied the physisorption of n-alkanes over Fe3C surfaces. The energy calculated showed the sorption energy for given surfaces depends on the ratio of C/Fe atoms exposed in the top layer. Adsorption energy does not depend on the hydrocarbon chain length after a certain length, three carbon numbers for n-alkyl species, five for 1-alkenes, and four for nalkanes [13].

Pore filling by chain growth In low-temperature Fischer-Tropsch synthesis, it observed that the pores of the catalysts will be filled with the liquid hydrocarbons. The movement of molecules depends on the external diffusion barrier. The calculation is done by the € ßler et al. shows that for chain growth probability a value Ro below 0.8, the time required to completely fill the pores shall take many days. It may take even more depending on the

CO Conv %

process conditions such as temperature and H2 to CO ratio. For filling a degree below 0.8, the effective vapor pressures of the hydrocarbons are much lower than the saturation vapor pressure hence the pore filling is influenced by the multilayer adsorption/desorption and capillary condensation [14]. This complex process was studied in a magnetic suspension balance at industrially relevant conditions (1 and 2 MPa, 222  C, € hlmann et al. H2/CO ¼ 1) over cobalt particles of 2.8 mm. Po observed that the complete pore filling of the catalyst pores with waxes took about 2 weeks. The initial pore filling rate was high until the monolayer adsorption was achieved. Product distribution within the pores constantly changes until pore filling is complete. In other words, a steady-state is achieved only after the complete filling of the pores. Pore filling time is greatly governed by the chain growth probability (a). Most of the catalyst exhibits the a < 0.82, in this case, to completely fill € hlmann the pores it takes several months to years. Hence Po et al. suggest that data for a less than 0.75 should be carefully handled as the catalyst particles cannot be completely dry. As the monolayer coverage of the catalyst will be complete within 10 days [15].

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Fig. 1 e Schematic representation of alkylidene mechanism for FischereTropsch synthesis. Reproduced with permission from ref. [12].

The model developed by Olewski et al. showed that with an increase in the chain length of 1- olefins, henrys constant exponentially increases. Re-adsorption of the formed olefins increases with the increase in chain length of the olefin. Hence, a re-adsorption phenomenon controls the olefin to paraffin ratio and chain-growth parameter (a) [16]. Olefin and paraffin ratio varies during the reaction is depends on the diffusion, readsorption, and other rate based phenomena. During the 3000 h of time on stream studied by Muleja et al. at first changes in the temperature were carried out after the steady-state is observed. Followed by flushing the reactor with N2 after catalyst deactivation and then the reaction was resumed. The Yao plot (plot of Pnþ1/Onþ1 versus Pn/On, n is the carbon number from n ¼ 2e5) shows a linear relationship for n ¼ 2e5. This is linear relationship is independent of the change in temperature, deactivation of catalyst and also after the catalyst has been flushed with nitrogen. Hence, the olefin and paraffin distributions cannot be changed independently and are interdependent. From this, it is suggested that the product distribution is determined by the quasi-reaction equilibrium among species or vapor-liquid equilibrium (VLE) or both in combination. The Lu Plot (normalized mole fractions of On, Onþ1, and Pn for n ¼ 2e4, Fig. 3) shows the properties which are observed in the reactive distillation systems having interaction with the reactions, VLE and feed flow rate composition. In other words, reactive distillation with reversible reactions exhibits this type of typical data. In

addition to kinetics thermodynamic equilibrium and VLE can be considered when modeling the FT process [17]. Catalyst bed should be diluted with inerts with uniform distribution. Improper distribution affects the CO conversion, CH4 selectivity, C2eC4 olefin to paraffin ratio, and the chain growth parameter (a). With proper distribution and dilution of the catalyst bed, the pellet to pellet readsorption of a-olefin takes place, and the highest yield of C5þ products (75.95%) with a probability of chain growth at a ¼ 0.896 can be achieved [18].

Wax-free liquid hydrocarbons Duyckaerts et al. studied the hydrocracking of n-paraffins and a-olefin individual and collective reactivities over Pt/nano-HZSM-5 catalyst under standard H2 environment and in-situ hydroprocessing atmospheres. Under the H2 atmosphere, olefin hydrogenation takes place first to give paraffin followed by cracking of the formed paraffin along with the paraffin in the feed. Hence, the reactivities of the olefin and paraffins cannot be distinguished under the hydrogen environment. Additionally, these alkanes can be dehydrogenated leading to the equilibrium which determines the partial pressure of alkenes and also the hydrocracking intermediate reacting on acid sites. Under the syngas environment, CO poisons the Pt sites inhibiting the hydrogenation/de-hydrogenation sites. This leads to a considerable deviation in the reactivities and

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Fig. 2 e Image (and detail) of the cobalt surface fully covered by a striped pattern after 40 min of reaction. Fourier analysis and histogram of the width of the stripes. a, 62 £ 62 nm2 STM topography image of the cobalt surface at 221 ± 10  C, 40 min after switching to a reactive 1:2:2 mixture of carbon monoxide, hydrogen, and argon at 4 bar total pressure. Vbias ¼ 0.3 V, Itun ¼ 100 pA, scan rate 3 £ 103 nm s¡1 b, Enlarged view of the 15 £ 12.5 nm2 region indicated by the dashed rectangle in a, where the internal structure attributed to the individual alkane molecules produced during the reaction and self-assembled on the Co(0001) surface can be seen. The blue lines serve as a guide to indicate the arrangement of the individual linear hydrocarbons. Inset: Fourier transform, which reflects the main periodic structures in the image. Green circles highlight the peaks corresponding to the periodicity of the striped pattern (1.8 ± 0.3 nm). Pink circles highlight the periodicity of the individual molecules within the stripes (0.46 ± 0.04 nm). c, Histogram of the period of the striped pattern in images a and b, expressed as n (number of carbon atoms of the alkane molecules). Reproduced with permission from Ref. [5]. (For interpretation of the references to color in this figure legend, the reader is referred to the Web version of this article.)

pathways for n-paraffins and a olefins. With the nonavailability of the hydrogenation sites, a olefins showed enhanced reactivity, additionally contributed to moderate secondary cracking by enhanced competitive adsorption on the acid sites. On the acid sites, a-olefin analogs readily undergo isomerization and cracking reactions. In the meanwhile, cracking reaction is governed by the partial pressure of the higher alkanes, likely via enhanced competitive adsorption on the acid sites. Hence, the presence of a-olefins in the hydrocarbon feedstock curbs secondary cracking and limits the undesired gaseous (C4) product formation. Due to the inhibition of the dehydrogenation sites, zeolites’ intrinsic olefin oligomerization activity predominates. Short-chain (e.g., C3, C6) a-olefins undergo oligomerization reactions, followed by subsequent cracking of their longer-chain oligomers. With this acid-catalyzed chain growth mechanism, undesired light product formation will be decreased. The product distribution of in situ hydrocracking under the syngas environment compared to hydrocracking under the hydrogenation environment varies to a greater extent. Product distribution under the syngas environment is controlled by the balance between hydrocracking and short-chain oligomerization. As a result, an increase in the C5þ yields can be observed [19]. The higher quantity of medium acid sites over zeolites with Pt loading of 0.2e0.1 wt% showed better performance towards jet fuel production. Compared to amorphous catalysts supports, zeolite supports shows more selectivity towards C9eC15 hydrocarbons (jet fuel) [20]. Co-precipitated Ru-, Pt- and Lapromoted CoeAl2O3/ZSM-5 (Si/Al ¼ 25) for in situ hydrocracking during FTS is studied by Ryu et al. CoeAl2O3ePt/ZSM-

5 showed high reducibility, weak acidic sites large pore size and pore volume. Over CoeAl2O3ePt/ZSM-5 hybrid catalyst highest catalytic activity, low C1 and olefin selectivity with high gasoline range hydrocarbons (C5eC9) selectivity were observed [21] (see Table 2). With an increase in the Si/Al ratios of zeolites, acid strength increases and decreases the available acid sites. Additionally, Si/Al ratios influence the zeolite and Co interaction and hence, the reducibility of Co. Therefore, the Si/Al ratio of the zeolite plays an important role in the activity and product selectivity. Lower ratios (Si/Al ¼ 15) showed higher CO

Fig. 3 e The Lu plot: A plot of the normalized mole fractions Onþ1 (corresponding to the vertical axis), On (corresponding to the horizontal axis) and Pn. Reproduced with permission from Ref. [17].

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Table 2 e Insitu upgrading of FT products over hybrid catalysts. Catalyst System CoePt/ZSM-5 CoeAl2O3/ZSM-5 CoeAl2O3eRu/ZSM-5 CoeAl2O3ePt/ZSM-5 CoeAl2O3eLa/ZSM-5 Co/ZSM-5 Co/SiO2 Co/AlSBA-16

Metal loadings %

Co Conversion

CH4 Selectivity %

C5eC9 Selectivity %

C10þ Selectivity %

Reference

Co-13 Pt-0.3 Co-13 Co-13 Ru-0.3 Co-13 Pt-0.3 Co-13 La-0.3 20 20 Co-15 Si/Al-10

42

9

14

68

[21]

33 31

22 10

17 21

35 55

[21] [21]

41

10

18

62

[21]

21

16

14

46

[21]

56 91 27

26 15 17

25 7 60

18 65 11

[22] [22] [33]

conversion and C5eC9 selectivity. With an increase in the ratio (>25) reducibility of Co decreased and so is the acid sites. Hence, the reduction in the CO conversion and C5eC9 selectivity is observed [22]. ZSM-5 (Si/Al ratio ¼ 40) was prepared by in situ hydrothermal conditions in presence of Co/SiO2 catalyst. During the hydrothermal synthesis from Co/SiO2, some of the Co species might have migrated over the zeolites. Over the catalyst with 25% zeolite, Co particle sizes were smaller with high reducibility and have optimum acidity. Hence, Co conversion showed high CO conversion and maximum selectivity to C5eC22 hydrocarbons with a minimum of C1eC4 [23] (see Fig. 4). The new hybrid catalyst CoeSiO2/MoePdePt-HZSM-5 (with a metal-metaleacid functionality) studied by Teiseh et al. and compared with that of a conventional CoeSiO2 catalyst. More of gasoline (C4eC12) Hydrocarbons, C8eC17 some extent diesel range (C15eC25) Hydrocarbons were produced over conventional catalysts. Under similar conditions, the CoeSiO2eMoePdePt/HZSM-5 catalyst exhibited increased selectivity to C8eC17 range hydrocarbons and increases further with an increase in the pressure. Interestingly product distribution produced straight-chain hydrocarbons rather than isomers.

Fig. 4 e Catalytic performance (CO conversion and product distribution) with respect to the acid site density on the ZSM5 modified Co/SiO2 catalysts. Reproduced with permission ref. [23].

Over Mo/Pt Pair, primary carbonium ion formed from olefins preferentially undergoes oligomerization and/or polymerization producing only straight-chain hydrocarbons [24]. The highest apparent site activity (STY) with high C5þ selectivity is observed over a Co-based catalyst. The C5þ selectivity changes with the Co conversion, significantly it depends on the water partial pressure. Higher Cn values are determined by the ratio of monomer production rate to the CeC coupling rate which depends on the shape, strain, and size of Co particles, which is controlled by the catalyst support material and its interaction with the Co particles. There exists a negative correlation between a C1 and the higher Cn values. This can be a result of the formation of methane being formed majorly by different mechanisms over sites other than where the chain growth takes place. But by a common carbon pool, both the mechanisms seem to be connected. Some of the features of the catalyst are also observed over Re and B promoted catalysts. With Re and B promotion C5þ selectivity increased and also the apparent STYs and decreased C1 values. Co particles over TiO2 smaller that the ~15 nm increases the C1 selectivity which has strong metal-support interaction. Even though Re exhibited SMSI, an increase in the higher Cn selectivity is observed. An increase in the C1 selectivity with a decrease in the Co particle size is attributed to the particle size induced C5þ selectivity decline. C5þ selectivity is found to be independent of bulk CoO or Co particle cluster size [25]. Major products of FTS over Co catalyst are linear alkanes and a-olefins based on this the theoretical molecular mass ratio of the total water to hydrocarbons is in the range of 1.125e1.286. A linear relationship exists between the weight percentage of C5þ hydrocarbons and the weight ratio of C5þ hydrocarbons to the total water. Lower the ratio indicates the formation of more gaseous products. If the linear equation is obtained and it is independent of the reaction condition and does not change, it indicates that the selectivities obtained are the feature of the catalyst. This mathematical method is proposed by Zhou et al. for shortening the evaluation cycle [26]. To shorten the chain length and to control the product distribution beyond ASF distribution in situ upgrading/ cracking can be applied. The key is mild hydrocracking to

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avoid the formation of light hydrocarbons. Bifunctional catalysts with FTS part and mild hydrocracking part are essential.

Activity over mesoporous silica-supported catalysts Heavy hydrocarbons deposited on the active cobalt active sites surfaces suppress the catalyst activity and selectivity by inhibiting the reactant access to the active sites. This nonpolar hydrocarbon layer formed on the cobalt active surfaces can change the product distribution. By enhancing the mass transport rate of heavier hydrocarbons formed inside the macropores of catalyst higher activity stability is observed by Koo et al. Compared to silica support Co/MMS (MMS ¼ MesoeMacroporous Silica) had larger wax trapping capacity due to large pore cavity, which results in and the suppression of the wax deposition on the active cobalt sites. Hence, CO conversion and C5þ selectivities enhanced. This wax transportation ability of the catalyst is a more important factor for the catalyst activity and stability than the oxidation states of the cobalt particles with their particle size. This catalyst can be regenerated by just an injection of liquid octane during the FTS reaction, which clears the hydrocarbons trapped in the pore cavity. This easy regeneration of the Co/ MMS catalyst is possible due to its Meso-Macropore Structure [27]. Co supported on a novel silica Nano spring (NS) had a major advantage over conventional silica gel in terms of unique helical structure and accessible surface. Even though it has less degree of reduction, a good FTS activity was observed with relatively higher WGS activity and CH4 selectivity [28]. Ordered mesoporous Co/KIT-6 catalyst prepared by zirconium phosphate (ZrP) modification. Over Co/KIT-6 catalyst the spatial confinement effects of cobalt nanoparticles make the catalyst to exhibit higher activity and stability with higher C5þ selectivity. Thermally stable hydrophobic ZrP nanoparticles make the catalyst resistant to coke formation by restricting the formation of coke precursors. ZrP makes the spatially confined cobalt nanoparticles resistant for sintering and agglomeration during the reaction [29]. Smaller Co particles (~10 nm) were obtained over Co/ graphene-Silica nanocomposites. The interactions between silica and cobalt oxide are weakened by the introduction of graphene. This facilitates the reduction of Co2þ to Co. The probability of formation of Co particles which are difficult to reduce was decreased by the graphene located at the cobalt and silica interface. Hence, the Co particles which are easy to reduce are formed over the prepared catalyst. Graphene introduction increased the hydrogen and CO adsorption ability as a result higher concentrations were observed over the prepared catalyst compared with the catalyst without the graphene. With 0.1e0.5 wt% of graphene introduction, quantity and the stability of the adsorbed CO enhanced. The graphene loading catalyst showed increased FTS activity, C5þ product selectivity and the fractions of C19þ hydrocarbons [30]. To the Co/SiO2 catalyst small amount of Al2O3 addition shows a promotional effect on the activity and C5þ selectivity. 1 wt% of alumina loading exhibited the highest promoting

7

effect and it influenced the C5þ product distribution. Paraffins in the range of C8eC25 increased and longer chain hydrocarbon fraction decreased. Alumina loading influences the Co particle size, narrow Co◦ distribution with 8 nm of mean particle size is observed. CO chemisorption on alumina promoted catalyst showed volcano relationship with the alumina loadings and maximum chemisorption is observed for 1% alumina. A relatively high concentration of CO on the surface decreases the ratios of H2 to CO on catalyst surface resulting in the higher C5þ selectivity [31]. In a few studies, it is observed that the pores of the support have a profound influence on product distribution. Over CoSBA-15 is observed to exhibit the shape-selective properties and product distribution is seen deviate from the ASF distribution. Pore diameter affects not only the Co reducibility but also product distribution. Catalyst supports having smaller pores limit the chain growth leading to higher selectivity towards lower hydrocarbon. With an increase in the pore size of the support, heavier hydrocarbons increased [4]. Increase in the bronsted sites over Al-SBA-15, Co interaction with the support and isomerization products increased. Whereas with the increase in the Lewis sites, an increase in C5þ selectivity is observed [32]. Ru promoted Co supported on SiO2 and SBA-15 showed enhanced activity but not the selectivity to C5þ hydrocarbons. RueCo supported over 3-D porous silica had no effect of water partial pressure rather C5þ selectivity correlated with the literature. Whereas, catalyst supported over SBA-15 deviated significantly at low conversions which are attributed to the CO diffusion limitations. Compared to 1D pores 3 D pores have a lesser diffusion barrier. 1 D pores are long enough at the molecular level to increase the H2/CO rations compared to that of bulk gas. And with a low partial pressure of water inside 1 dimensional (1-D porous network), C5þ selectivity is low due to the increase in H2/CO ratio because of CO diffusion limitation [33]. Zhao et al. reported that AlSBA-16 support plays an important role in controlling the activity and selectivity of the FTS. With Si/Al ratios of 10, high C5eC12 and olefin selectivity are obtained. With the decrease in the Si/Al ratio selectivity shifted towards lighter products, which is due to the increased acid sites which delay the olefins further transformations on the Co surfaces [33]. The addition of Fe to Co/SiO2 increases the WGS activity which increases the H2/CO ratio and increases the C5þ selectivity. Formed FeeCo bimetallic phases favor the formation of long-chain hydrocarbons. Whereas, with the metal rations reaching 1 (Fe/Co) interaction with the metals increases hence catalytic performance decreased [34]. Colloidal nanoparticles of Co (2.7 nm) was directly used in the reaction using water and binary mixtures as a solvent. By changing the water and organic co-solvents composition and nature, product distribution can be changed to a greater extent. With the selection of appropriate aqueous solvent mixture, based on solubility H2: CO ratios can be controlled. Hence, product distribution can be controlled from light to heavy hydrocarbon. Co-solvent chain length did not affect the product distribution whereas it slightly affected the activity. As the presence of water and hydrocarbons at the surface affects H2/CO coverage, hence, it directly influences their selectivity in FTS. The presence of water can decrease the

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chain length of products. With an increase in the chain length of alcohol, acetals become major products [35]. Co particle size depends on the interaction of Co with the support which can influence the product distribution, at the same time pore size of the support also affects C5þ selectivity. Controlling Co particle size, with proper pore width can increase the C5þ selectivity.

Activity over Co-based catalysts Different alumina supports viz., g, d and q Al2O3 were used to synthesize and control the Co particle size. The effect of Co particle size on the hydrocarbon selectivity was studied by Rane et al. Co particle size of 2e14 nm was used to study the relationship between methane, olefin and C5þ selectivities. With the increase in the Co particle size, the C5þ selectivity increases and the maximum is observed at 8e9 nm. And with further increase in the Co particle size C5þ selectivity decreases and no change in the C5þ selectivities is observed around 40e45 nm [36] (see Figs. 5e7). Small and medium pore g-Al2O3 was used as starting material for d, q, and a- Al2O3. These were prepared by heating at different temperatures. Cobalt catalysts 12 wt% Co was synthesized by using these alumina phases as supports. Cobalt metal particle prepared are approximately similar in sizes. Medium pore g-Al2O3 as starting material showed higher C5þ selectivity compared to small pore g-Al2O3 as a starting material. Higher C5þ selectivity was observed for d, and a-Al2O3 compared with g and q Al2O3. Concentrations of active surface intermediates were followed by steady-state isotopic transient kinetic analysis. From this, it is said that the C5þ selectivity correlates well with the concentration of active surface intermediates leading to products. Pore sizes and crystallite sizes of the precursor alumina is observed to be an important parameter for C5þ selectivity [37]. Highly ordered mesoporous Co3O4 was pillared with 5% of Al2O3 which was less than a monolayer on the catalyst

Fig. 5 e C5þ selectivity as a function of the alumina phases (g-Al2O3, d-Al2O3, q- Al2O3, and a-Al2O3) and Lewis acidity. 20 bar, 483 K, H2/CO ¼ 2.1, 45e50% CO conversion ((-) small pore alumina; (:) medium pore alumina). Reproduced with permission ref. [36].

Fig. 6 e Olefin to paraffin (O/P) ratio for different medium pore alumina phases ((_) C3 ¼ /C3¡ ratio; (þ) C4 ¼ /C4¡ ratio). Co particle size equal to 8e10 nm except for _-Al2O3. 20 bar, 483 K, H2/CO ¼ 2.1, 12 wt.% Co, ~45% CO conversion. Reproduced with permission from Ref. [36].

surface. The pillared catalyst showed enhanced structural and catalytic stability. Even under the H2-rich reductive FTS reaction condition catalyst dis not exhibit significant structural damage. Interactions of Al2O3eCo3O4 partially form the catalytically inactive spinel-type CoAl2O4 phases in the matrices of the catalyst which gives the structural stability. And due to mesoporous nature, reactants and products have a better mass transfer. Hence, the fewer coke precursors deposit on the active site resulting in higher FTS activity [38]. Over Potassium promoted Fe and Cu catalyst (100Fe/5.1Si/ 2Cu/xK (x ¼ 1.25 or 3)), WGS reaction rate increased with increase in the potassium loadings. Under the reaction conditions, the low partial pressure of H2 is compensated by the WGS reaction. Even at low CO conversions of 30% H2 required for FTS is supplied by the WGS. With an increase in the potassium loadings, the secondary reaction of olefin is decreased resulting in increased olefin selectivity. With an increase in the potassium content, the product distribution shifted from oil to wax [39]. CO2 and CO hydrogenation over Co and Fe based catalyst under FTS reaction condition yield different product distributions. Over non-promoted Co/Al2O3, CO2 hydrogenation is faster compared to that of CO. However, CO2 hydrogenation yields methane as the major product whereas it is undesirable during CO hydrogenation which produces long-chain hydrocarbons. Over K-promoted 100Fe/10Zn/1Cu (at/at) catalyst Visconti et al. observed the CO2 hydrogenation produces middle distillate even though the CO2 hydrogenation rate is slower than the CO hydrogenation rate. Change in the rates of hydrogenation rates occurs due to a change in the surface H/C ratio occurring due to the change in the adsorption strength of the promoted catalyst. Under low partial pressure of CO chain growth sites act as hydrogenation sites, resulting inside produces. With Potassium promotion, the chain growth sites stability increases and decreases the secondary reaction. Over potassium modified catalyst CO hydrogenation leads to more wax products. It is anticipated that during CO2 Hydrogenation, CO2 is converted CO, which acts as intermediate and reacts to give products. Potassium promotion showed a marginal effect

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Fig. 7 e Olefin to paraffin (O/P) ratio and C5þ selectivity for medium pore q-Al2O3 as a function of the particle size of Co ((A) C5þ; (◊) C3 ¼ /C3¡ ratio; (,) C4 ¼ /C4¡ ratio) 20 bar, 483 K, H2/CO ¼ 2.1, 12 wt.% Co, ~45% CO conversion. Reproduced with permission ref. [36].

on the reactivities of CO and CO2 but has a pronounced effect on the product distribution [40]. The Pt addition (0.3%) to Co/ alumina catalyst showed higher reducibility of Co particles and weak acid sites increased. Due to which, CO conversion increased and selectivity increased for C5eC9 hydrocarbons [41]. The productivity of C5þ hydrocarbons increases over Co/ Al2O3 catalyst with Phosphorous modification. P-modified Al2O3 exhibits the homogeneous Co-particle size distribution and maintains it by avoiding the agglomeration even under FTS reaction conditions. During phosphorous modification Al2O3 surface partially transforms into aluminum phosphates. Tridymite aluminum phosphate (AlPO4) phases are formed on the phosphorous modified Al2O3 which decreased the surface hydrophobicity and increased Co dispersion. Highly dispersed Co nanoparticles exhibited strong interaction with the support, as a result, Co agglomeration decreased. Due to an increase in the hydrophilicity of the catalyst, the formation of heavy hydrocarbon deposits decreased and resistance to coke formation increased. On the hydrophilic surfaces of phosphorous modified catalyst heavy hydrocarbon and water formed macro-emulsion, this makes it easier to remove during the FTS reaction [42]. Cobalt supported alumina catalysts preparation condition and calcination conditions have shown to affect the CO conversions. Catalysts calcined under different calcination atmospheres resulted in different cobalt micro-structures. The strong interaction is observed for air calcined catalyst, which resulted in FCC metallic phases after reduction. Whereas weak interaction is observed for the catalyst calcined under a hydrogen atmosphere and resulted in the formation of HCP structure. The hydrogen calcined catalyst showed better conversion and had marginally affected the product distribution [43]. Methane selectivity increases with increasing temperature, pressure, GHSV, and H2/CO ratio, whereas the C5þ selectivity decreased [8].

Modification of hydrotalcite with kaolin increased the inter voids of hydrotalcite inducing bimodal pores increasing porosity. The highest dispersion was observed over hydrotalcite and decreased with the modification. But due to the effect of bimodal pores, Co-active sites were more active due to the easy diffusion of heavy hydrocarbons. Overall product distribution and catalyst performance depend on the reducibility and particle size of cobalt, catalyst porosity, and product diffusion barrier [44]. Silicon carbide having the highest thermal conductivity, which is essential for the exothermic FTS, used as support shows less interaction with the active meal. Due to this cobalt sintered, as result pore volumes decreased, the fouling of catalyst by waxes resulted in deactivation [9]. With respect to hydrogen FTS, the reaction is first order. With a decrease in the hydrogen, CO conversion slightly varied but affected the olefin/paraffin ratio. Under varying nitrogen ratios in the feed over Co-based Al2O3 supported catalyst, increased olefin/paraffin ratio is observed with an increase in nitrogen. It is due to a decrease in the partial pressure of hydrogen that decreased the chain growth probability [6].

Activity over Fe based catalysts Morphology controlled FeeMn catalyst synthesis is carried out by changing the composition (see Fig. 8). Fe2O3 without Mn addition shows spindle shape and with an increase in the Mn spindles changes to cubes and then to nanoparticles (Figs. 8 and 9). Mn addition decreased CH4 formation and increased carbon chain growth. An increase in Mn content increases the extent of chain growth. Low Mn catalyst showed higher selectivity to C2eC4 hydrocarbons whereas the higher Mn catalyst showed better selectivity to C5þ selectivity. Despite the better selectivity exhibited by the Mn modification, the stability of the catalyst is decreased. Optimizing the preactivation condition can increase the stability of the catalyst

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Fig. 8 e Schematic representation of mesoporous Fe-based spindle formation during synthesis. Reproduced with permission from Ref. [47].

[45]. Fe2O3@MnO2 spindles core-shell catalysts exhibited enhanced catalytic performance, in terms of C5þ hydrocarbons selectivity. Added Mn promoter can increase the CO dissociation, which enhances the concentration of active intermediates required for chain growth and suppress the methane formation. Mn promoter can be used as an oxygen carrier, while bonding, CO partially reduces manganese oxide with its oxygen, thus favoring CO dissociation and an increase in catalytic activity [46].

In general, catalytically active materials are coated on to the mesoporous support to achieve high dispersion, which results in the enhanced catalytic performance. Better dispersion leads to strong metal-support interaction. In the case of iron catalysts, a higher interaction leads to decreased activation of iron oxides resulting in lower catalytic activity. Novel catalysts prepared by the assembly of numerous iron carbide nanoparticles giving rise to mesoporous spindles. Catalytic activity over this catalyst is increased due to decreased

Fig. 9 e SEM and TEM images of (a) Fe, (b and c) Fe/CTAB, and (d) Fe/2CTAB. Reproduced with permission from Ref. [47]. Please cite this article as: Hodala JL et al., Catalyst design for maximizing C5þ yields during Fischer-Tropsch synthesis, International Journal of Hydrogen Energy, https://doi.org/10.1016/j.ijhydene.2019.12.021

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Table 3 e Activity over Fe based catalysts. Catalyst system Fe2O3 FeMn10 (Fe/Mn-0.1) Fe Fe/2CTAB Fe Precipitated Fe Nano spheres Fe Nano polyhedrons

CO Conversion %

CH4 Selectivity %

C5þ selectivity %

Reference

94.3 97.6 94 94 37 71 43

17.4 9.5 17 14 21 8 11

50.3 64 59 65 40 64 54

[45] [45] [47] [47] [49] [49] [49]

interaction with the support. With an increase in the pore size, Co conversion is increased, C5þ selectivity is increased and CH4 selectivity is decreased. Pore sizes of the catalyst support materials have a profound effect on the activity and C5þ selectivity in FTS [47]. The Iron and carbon synergistically formed Carbonencapsulated iron-core nanoparticles showed interesting catalyst performance under the wide range of studied conditions. Under the FTS reaction conditions, bio-syngas (19% H2, 20% CO, 12% CO2, 2% CH4 and 47% N2) were used as feed. The catalyst showed excellent activity and stability towards liquid yields. Liquid yields were 65% while the CO conversions were 90%. The main products were olefins (49.11 mol%), isoparaffins (11.82 mol%) and aromatics (7.83 mol%) and alcohols (3%). Due to the unique structure and synergetic effects deviation in the standard product distribution is observed. C/Fe catalyst is observed to be resistant to sintering due to the carbon shell and Fe core prevents the collapse of the shell [48]. Fe3O4 nanocatalyst with nanospheres (FNS) and nanopolyhedrons (FNP) was prepared by solvothermal method and applied in FischereTropsch synthesis (FTS) without calcination. Reduction of Iron, dispersion, and formation of surface cFe5C2 was better over FNS compared to FNP. Hence a number of active sites were observed over the FNS resulting in better activity and product distribution shifted towards C5þ hydrocarbons. Carbon-encapsulated iron nanoparticles with iron core/cementite/carbon shell showed Olefins as the dominant products in liquid hydrocarbons [49] (see Table 3). Catalytic conversion of the synthesized catalyst followed ASF distribution [50]. For the FTS process, CO and H2 are prepared by the reforming of various feedstocks. One of the other potential sources of hydrogen is the bio electrodes, here microorganisms consume the organic matter and yield hydrogen as the product [51e64]. With bioelectrode as a hydrogen source, the FTS process becomes cleaner and energy-efficient without the application of high-temperature reforming.

Conclusions The fuels produced by FTS can be sustainably produced and is clean in terms of vehicular emissions. FTS is gaining attention and renewed interests from time to time for various reasons. Selectivity of products is limited to the ASF probability distribution, but some time with the application of suitable catalyst deviation is observed. But still deviation is not sufficient to produce the hydrocarbons in the narrow range.

Presently catalyst developed can maximize gasoline, middle distillates and wax fractions based on the catalyst properties. With no evidence to chain termination step, it is difficult to restrict the hydrocarbons chain growth in a narrow range. Insitu hydrocracking catalysts are applied to crack longer chain to shorter chain but at the same time lighter hydrocarbons are of lesser use. Support has a limited role in the intrinsic activity of the catalyst but has a considerable effect on the product distribution, due to the diffusion barrier, readsorption of olefins, cracking. With the application of suitable support and controlling the interaction of the active sites yields the good CO conversion with less lighters and higher C5þ hydrocarbons.

Acknowledgments This work was supported by Board of Research in Nuclear Sciences and Department of Atomic Energy (BRNS-DAE) [Grant No. 2013/34/4/BRNS/0483], Bhabha Atomic Research Centre, India.

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Please cite this article as: Hodala JL et al., Catalyst design for maximizing C5þ yields during Fischer-Tropsch synthesis, International Journal of Hydrogen Energy, https://doi.org/10.1016/j.ijhydene.2019.12.021