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Liquid fuels are the power house of modern society. Presently, the bulk of the liquid fuel supply is from petroleum, and the untamed appetite for liquid fuel is pushing society toward the tolerance limit in terms of sustainable development. Meanwhile, vast amounts of biomass are produced annually through photosynthesis. In terms of empirical composition, biomass is similar to coal; hence it is envisaged that technologies hitherto developed for coal liquefaction could be adapted for liquefaction of biomass to hydrocarbon biofuels. However, the liquefaction process suffers from a low liquid fuel yield. In this chapter, we elaborate on this challenge and discuss emerging new opportunities to enhance the liquid fuel yield from biomass liquefaction processes. Thermodynamically, transforming a lower-energy-density feedstock into a higher-energy-density product is associated with a huge energy loss penalty. This is the heart of the problem of a low liquid fuel yield from biomass liquefaction. In this chapter, we explain how some of the energy loss penalty can be compensated for by renewable energy resources. We also discuss tandem processes for enhancing the economics of the biomass liquefaction process and highlight emerging new chemistry techniques for achieving this.

The history of energy use is closely linked to the history of human civilisation. Energy use per capita is a direct measure of living standards and prosperity.1  Advances in older civilisations and in present modern society are characterised by increasing energy consumption.2  Since the agrarian era, when wood was the main fuel for industrial economy and later when coal became the preferred fuel for powering steam engines, energy use grew exponentially to enable a higher standard of living and large-scale industrialisation. This was followed by petroleum becoming the dominant global fuel, after the invention of diesel and internal combustion engines. These engines run on liquid fuel that is obtained cheaply from petroleum.

Liquid-fuel-powered engines were poised to displace solid-fuel- (coal-) powered steam engines in the rapid industrialisation drive in the early 1900s. Moreover, petroleum is more portable than coal and is used as a feedstock for a plethora of chemicals and products. However, petroleum deposits are unevenly distributed across the globe,3  and perhaps the geographical scarcity of petroleum motivated early attempts at liquid fuel production from coal.4  Friedrich Bergius invented the direct coal liquefaction (DCL) process in 1913. Although the impact of DCL was not apparent during World War I, the future role of liquid fuels in military logistics was not missed. This spurred greater interests in coal liquefaction after World War I. Extensive investigations into coal liquefaction culminated in industrial-scale DCL development and the invention of indirect coal liquefaction (ICL) by Franz Fischer and Hans Tropsch in 1923. The two coal liquefaction processes were developed and operated at industrial scale to meet part of Germany's liquid fuel needs during World War II. South Africa was also able to meet her liquid fuel needs during the Apartheid era using coal liquefaction technology.

After World War II, the economic and political importance of petroleum grew rapidly.5,6  The long-term security of petroleum supplies moving from producing countries to consuming countries could not be guaranteed, with periods of smooth relations interspersed with interludes of disruption of petroleum supply.5,7–10  Hence, competence in coal liquefaction technologies was generally considered of strategic importance by petroleum-consuming countries. Instances of disruption of petroleum supply reawakened interest in coal liquefaction technologies, but whenever supply was restored, commercial interest in coal liquefaction technologies waned. However, memories of hardship suffered during the disruption interludes remained a motivation, for strategic reasons, for sustained research interest in continuous improvement and preserving competence in the technology.

The DCL process involves subjecting a slurry of dry coal powder and a hydrogen-donor solvent to high temperature and pressure hydrogenation in the presence of a catalyst. It may be considered a top-down approach to coal liquefaction, since it entails the depolymerisation of large molecules into smaller molecules. The type of coal, the nature of the solvent, the reaction conditions and the catalyst used are important variables that determine the liquid fuel yield. Low-grade soft coal – usually lignite – is the most widely used feedstock for DCL. The hydrogen-donor solvent is usually tetralin – a hydrocarbon that can transfer hydrogen to coal. The dehydrogenated donor solvent has low enthalpy of hydrogenation that allows for its regeneration at a high hydrogen pressure in the presence of a catalyst such as MoS2. This dehydrogenation and regeneration cycle of the donor solvent is crucial to the liquefaction process. Hence, research and developments on DCL are focussed on designing donor solvents and efficient catalyst for their regeneration, as well as on process optimisation to ensure a maximum liquid fuel yield.

The DCL liquid fuel production scheme is shown in Figure 1.1. The resulting DCL crude is made up of wide spectrum aromatic compounds and cannot be used directly in engines. It has to undergo an extensive refining process such as hydrocracking, hydrotreating, distillation, solvent extraction, blending etc. before it is suitable for use in engines.

Figure 1.1

Liquid fuel production scheme via DCL.

Figure 1.1

Liquid fuel production scheme via DCL.

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ICL involves a two-stage process: syngas production, followed by syngas-to-hydrocarbon conversion. The syngas-to-hydrocarbon process is called Fischer–Tropsch Synthesis (FTS). The syngas production stage is basically a coal gasification and cleaning process, while FTS is essentially high-temperature and high-pressure carbon monoxide hydrogenation over a catalyst to produce a mixture of hydrocarbons. In contrast to DCL, with this process, large molecules (liquid hydrocarbon molecules) are built or formed from small molecules (CO and H2); hence FTS may be considered a bottom-up approach to liquid fuel production. The syngas composition reaction conditions and catalyst are key variables that influence the product delivered by FTS. The ICL liquid fuel production scheme is shown in Figure 1.2. FTS crude – also known as syncrude – usually consists of a wide spectrum of linear hydrocarbon chains. The syncrude must be separated into hydrocarbon fractions, and further refining processes may be necessary before use, e.g. hydrocracking, reforming and alkylation.

Figure 1.2

Liquid fuel production scheme via ICL.

Figure 1.2

Liquid fuel production scheme via ICL.

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Although some studies have suggested that DCL is more energy efficient than ICL,11,12  the two processes essentially entail transforming coal (e.g. Turow brown coal, CH1.1O0.3)13  into liquid hydrocarbons (typically of H/C ratio ∼2.0) that are suitable for use in engines. A comparison of Figures 1.1 and 1.2 reveals that syngas production is common to the two processes, and it may be argued that DCL appears to be a more complicated process than ICL. DCL is a reductive pyrolysis/hydrogenolysis process that converts coal to liquids (–CH2–) directly.14  For a hypothetical coal feedstock – represented as CHxOy, – the DCL reactions can be represented by eqn (1.1) and (1.2):

Equation 1.1
Equation 1.2

Discussions on DCL are often silent about the source or method of hydrogen production for DCL, which is probably through gasification of a portion of the feedstock, as shown in Figure 1.1. The donor solvent intermediate in the hydrogenation process, its regeneration and recycling make it difficult to monitor the DCL fuel yield or carbon utilisation efficiency.

Gasification is a widely discussed feature of ICL, since it is the first step (degradation of the feedstock into syngas) before reassembling the syngas into fuel molecules via FTS. Two simple theoretical gasification scenarios can be considered: oxygen and steam gasification. With a coal feedstock, CHxOy, the corresponding reactions in the ICL process may be represented as indicated in eqn (1.3) and (1.4):

Equation 1.3
Equation 1.4

Eqn (1.3) shows that in oxygen gasification, hydrogen yield depends on the hydrogen content of the feedstock. Eqn (1.4) shows that both the hydrogen and oxygen content of the feedstock influences the hydrogen yield in steam gasification. ICL has the advantage of prerequisite cleaning that removes the sulphur content of the syngas before FTS, so that the resulting syncrude-derived fuel produces very low sulphur emissions. For a given coal sample, CH1.1O0.3, the theoretical hydrocarbon yields via oxygen gasification and steam gasification are 27.5% and 62.5%, respectively, provided that the energy required for the gasification is supplied from other means than burning a portion of the feedstock. If part of the feedstock is to be burned to supply the energy required for gasification and the other process steps in ICL (including crude refining), the final FTS fuel yield will drop drastically.15  The low fuel yield of the coal liquefaction process (DCL or ICL) is of little consideration during a time of crisis, such as a war or embargo. However, in noncrisis times, the carbon footprint is a significant consideration, especially in light of global warming and climate change.16 

Table 1.1 shows a comparison of the fuel properties of coal and biomass samples.13  Coal and biomass have s similar empirical formula, so biomass can effectively replace coal in either of the liquefaction processes.17  The resulting biomass liquefaction process is known as the biomass-to-liquid (BTL) process. Since biomass is a product of photosynthesis and has a low sulphur content, BTL biofuel products may be considered carbon neutral and more environment friendly than coal.18  However, replacing lignite with biomass as the feedstock (Figure 1.2) has little or no effect on the constraint (indicated in eqn (1.3) and (1.4)) on syncrude or fuel yield. A higher fuel yield is only achievable with an external energy and hydrogen input.

Table 1.1

Chemical composition and proximate analysis results of lignite, sunflower hulls and sunflower hulls pellets. Adapted from ref. 13 with permission from Elsevier, Copyright 2018.

PropertyCoalBiomass
Turow brown coalSunflower hullsSunflower hulls pellets
Composition 
67.50 50.26 44.31 
5.98 5.98 5.95 
24.59 42.23 48.86 
0.62 1.28 0.77 
1.31 0.25 0.11 
Empirical formula (CHxOyCH1.1O0.3 CH1.4O0.6 CH1.6O0.8 
PropertyCoalBiomass
Turow brown coalSunflower hullsSunflower hulls pellets
Composition 
67.50 50.26 44.31 
5.98 5.98 5.95 
24.59 42.23 48.86 
0.62 1.28 0.77 
1.31 0.25 0.11 
Empirical formula (CHxOyCH1.1O0.3 CH1.4O0.6 CH1.6O0.8 

In the following sections, we discuss the direct biomass liquefaction process and the indirect biomass liquefaction process (DBL and IBL, respectively). We elaborate on IBL: biomass gasification, the FTS mechanism, catalysts and operation modes. Along the way, we detail the limitations of the processes and the strategies used to enhance the BTL biofuel yield.

Analogous to coal liquefaction, biomass liquefaction can be achieved via direct or indirect methods. Gasification is crucial to both of these methods. Gasification is an old practice, and gasifier designs have evolved over the past five centuries. Excellent accounts of gasification technologies can be found in the literature,19–21  and this section focusses on key highlights relevant to biomass gasification.

Important factors for successful biomass gasification operations include:

  • – Composition of feedstock (biomass)

  • – Handling and feeding system

  • – Gasifying agent

  • – Choice of gasifier

  • – Gas clean-up system

  • – Ash or solid residue removal system

Generally, biomass may be represented by the empirical formula CH1.4O0.6 (see Table 1.1). The composition of biomass and coal are similar, so coal gasification technology could be adapted for biomass gasification. In rare cases in which the composition of the biomass feed is significantly different from that of coal, a special gasification system may be warranted. The main objective of biomass gasification is to transform the biomass feed into water gas (for the DBL application) or hydrogen-rich syngas (for the IBL application). It involves a sequence of overlapping processes: drying, pyrolysis, oxidation and heterogeneous and homogeneous reactions (see Figure 1.3,22 ). The nature of the gasifying agent influences the energy requirement and the H2/CO ratio of the product syngas. In the absence of a gasifying agent, i.e. pyrolysis, gas formation efficiency is very low. The merits and challenges of using common gasifying agents are highlighted in Table 1.2.23  Qualitatively, the product syngas composition is the same, but quantitatively it is different for the different gasifying agents. None of the gasifying agents satisfy all desirable objectives for biomass gasification for BTL. Among the gasifying agents, oxygen or steam or a mixture of the two are most suited and are commonly used for BTL applications. However, oxygen and steam gasification pose challenges in terms of the high cost related to the required air separation facility and high energy cost.

Figure 1.3

Block diagrams of unit processes biomass gasification. Adapted from ref. 22 with permission from the Royal Society of Chemistry.

Figure 1.3

Block diagrams of unit processes biomass gasification. Adapted from ref. 22 with permission from the Royal Society of Chemistry.

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Table 1.2

Merits and demerits of common gasifying agents for biomass gasification. Adapted from ref. 23, https://iopscience.iop.org/article/10.1088/1757-899X/863/1/012028, under the terms of the CC BY 3.0 license https://creativecommons.org/licenses/by/3.0/.

Gasifying agentMeritsDemerits
Air Simple equipment, low cost High N2 content in syngas, low heating value 
O2 Good syngas quality High cost of air separation 
H2High H2 content, high heating value Indirect or external heat supply, high tar content 
CO2 High heating value, high CO content Indirect or external heat supply, low H2 content 
Pyrolytic Syngas with medium heating value Low system efficiency 
Gasifying agentMeritsDemerits
Air Simple equipment, low cost High N2 content in syngas, low heating value 
O2 Good syngas quality High cost of air separation 
H2High H2 content, high heating value Indirect or external heat supply, high tar content 
CO2 High heating value, high CO content Indirect or external heat supply, low H2 content 
Pyrolytic Syngas with medium heating value Low system efficiency 

Contrary to the simple representations of gasification reactions provided in Section 1.1, in practice, a cascade of reactions take place inside a gasifier (see Table 1.3 24 ): tandem homogeneous, heterogeneous, exothermic and endothermic processes take place inside a gasifier. The heterogeneous reactions – except carbon (char) hydrogenation – are critical to obtaining syngas. But the homogeneous gas reactions – except the water-gas-shift (WGS) reaction – are detrimental to syngas selectivity. Overall, syngas formation is endothermic, therefore energy is needed to run the gasification process. A high temperature is favourable with endothermic reactions as it produces syngas with a low H2/CO ratio and a low tar content. Tar poisons the FTS catalyst, so a high tar content is detrimental to syngas utility, especially in FTS applications. Ultimately, syngas selectivity depends on an intricate optimisation of the parallel competing reactions.

Table 1.3

Main chemical reactions involved in the biomass gasification process. Adapted from ref. 24 with permission from Elsevier, Copyright 2019.

S/NReactionNameΔH (kJ mol−1)
 Drying and pyrolysis   
Biomass → char + tar + H2O + light gases Pyrolysis and devolatilisation > 0 
 
 Heterogeneous reactions   
 Char oxidation   
C + ½O2 → CO Partial oxidation − 111 
C + O2 → CO2 Complete oxidation − 394 
 Char gasification   
C + H2O → H2 + CO Steam gasification + 173 
C + CO2⇌2CO Boudouard reaction + 131 
C + 2H2 ⇌CH4 Hydrogenation gasification − 75 
 
 Homogeneous reactions   
 Gas reactions   
CO + ½O2 → CO2 CO oxidation − 283 
H2 + ½O2 → H2H2 oxidation − 242 
CH4 + 2O2 → CO2 + 2H2CH4 oxidation − 283 
10 CO + H2O⇌CO2 + H2 WGS − 41 
11 CO + 3H2⇌CH4 + H2Methanation − 206 
 Tar conversion reactions   
12  Partial oxidation  
13  Steam reforming Highly 
14  Dry reforming Endothermic 
15  Thermal cracking  
S/NReactionNameΔH (kJ mol−1)
 Drying and pyrolysis   
Biomass → char + tar + H2O + light gases Pyrolysis and devolatilisation > 0 
 
 Heterogeneous reactions   
 Char oxidation   
C + ½O2 → CO Partial oxidation − 111 
C + O2 → CO2 Complete oxidation − 394 
 Char gasification   
C + H2O → H2 + CO Steam gasification + 173 
C + CO2⇌2CO Boudouard reaction + 131 
C + 2H2 ⇌CH4 Hydrogenation gasification − 75 
 
 Homogeneous reactions   
 Gas reactions   
CO + ½O2 → CO2 CO oxidation − 283 
H2 + ½O2 → H2H2 oxidation − 242 
CH4 + 2O2 → CO2 + 2H2CH4 oxidation − 283 
10 CO + H2O⇌CO2 + H2 WGS − 41 
11 CO + 3H2⇌CH4 + H2Methanation − 206 
 Tar conversion reactions   
12  Partial oxidation  
13  Steam reforming Highly 
14  Dry reforming Endothermic 
15  Thermal cracking  

Besides the gasifying agent, the selection of an appropriate gasifier is crucial to achieving the gasification objective. Based on the mode of contact between the biomass and gasifying agent, gasifiers are classified into the following types: fixed bed; fluidised bed; and entrained-flow bed. Each class is further subdivided according to the flow regime of gasifying agents (see Figure 1.4). Some gasifier types are shown in Figure 1.5.25 

Figure 1.4

Classification of gasifiers based on mode of contact between biomass and gasifying agent.

Figure 1.4

Classification of gasifiers based on mode of contact between biomass and gasifying agent.

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Figure 1.5

Illustration of some gasifier types based on contact mode between fuel (biomass) and gasifying agents. Adapted from ref. 25 with permission from John Wiley & Sons, Copyright © 2013 Wiley Periodicals, Inc.

Figure 1.5

Illustration of some gasifier types based on contact mode between fuel (biomass) and gasifying agents. Adapted from ref. 25 with permission from John Wiley & Sons, Copyright © 2013 Wiley Periodicals, Inc.

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The gasification process can be classified into an autothermal or allothermal process, according to the mode of energy supply. The influence of gasification mode and gasifying agent on syngas composition is summarised in Figure 1.6.24  In autothermal gasification, the energy is supplied internally through partial combustion reactions inside the gasifier. The advantages of autothermal gasification are that it is simple and easy process to operate under pressurised conditions, and direct heating of the reactants promotes efficient energy utilisation. If air is used as the gasifying agent, the resulting syngas will contain a high amount of nitrogen. Nitrogen-rich syngas will lower the kinetics of the FTS or necessitate the use of an expensive high-pressure system to compensate for nitrogen dilution of the syngas, so pure oxygen and steam are preferred for biomass gasification for BTL applications.

Figure 1.6

Influence of gasification agent and gasification mode on syngas gas composition. Adapted from ref. 24 with permission from Elsevier, Copyright 2019.

Figure 1.6

Influence of gasification agent and gasification mode on syngas gas composition. Adapted from ref. 24 with permission from Elsevier, Copyright 2019.

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Based on eqn (1.3) and (1.4), pure oxygen gasification is expected to yield syngas of H2/CO = 0.7; steam gasification should yield syngas of H2/CO = 1.1; oxy-steam gasification should yield 0.7 < H2/CO < 1.1. The H2/CO ratios indicated in Figure 1.6 are a reflection of the extent of the carbon loss reactions, such as the WGS reaction, complete oxidation and unconsumed carbon as char. It also explains the report by Dasappa and co-workers, who obtained hydrogen-rich syngas from oxy-steam gasification of biomass using an autothermal mode downdraft gasifier.26,27  The steam : biomass ratio (SBR) and air/oxygen equivalent ratio (ER) influence the H2/CO ratio. Syngas with H2/CO ∼ 2 that is suitable for FTS was obtained at ER ≈ 0.23 and SBR ≈ 1.5. A high ER and SBR favour the WGS reaction, which can lead to pure H2 after pressure swing adsorption (PSA) and CO2 scrubbing. This implies that more than half of the carbons in the biomass feedstock will not make it into the second stage (FTS) of the liquefaction process. The oxy-steam autothermal downdraft gasifier has the advantages of being simple and easy to operate and of being flexible to the H2/CO ratio; however, the need for an air separation facility adds to the energy demand and the cost of the process.

In allothermal gasification, heat production and gasification are separated. This is achieved by connecting a combustion reactor and a gasifier using heat exchangers or bed materials. Allothermal gasifiers are essentially a dual reactor system that produces two gas streams: the combustion exhaust and the syngas (see Figure 1.7 28 ). Separation of the combustion exhaust from the syngas product allows for the production of an N2-free syngas without having to install an air separation facility.

Figure 1.7

Schematics of mode of energy transfer in allothermal gasification. Adapted from ref. 28, https://doi.org/10.3390/app10010002, under the terms of the CC BY 4.0 license, https://creativecommons.org/licenses/by/4.0/.

Figure 1.7

Schematics of mode of energy transfer in allothermal gasification. Adapted from ref. 28, https://doi.org/10.3390/app10010002, under the terms of the CC BY 4.0 license, https://creativecommons.org/licenses/by/4.0/.

Close modal

The choice and mode of heat exchange between the two reactors dictate efficiency, operational flexibility and the economy of the process. Heat transfer from a combustor to a gasifier may be achieved via a sandwich or an envelope design, with the wall of the gasifier serving as the heat exchanger.29  This allows efficient heat transfer from the combustor to the gasifier but prevents the exchange of material between them. Using circulating bed material can permit the exchange of solid materials between the combustor and the gasifier but prevent mixing of the respective gas streams. The bed materials serve as the carrier of heat from the combustor to the gasifier. This requirement narrows the choice of reactor type to the fluidised bed. Unconsumed char particles from the gasifier can flow to the combustor, where they are combusted to produce heat for the gasifier. The dual fluidised bed (DFB) gasifier coupling through the heat exchanger enjoyed wider interest because of the possibility of recycling the biomass feed between the combustor and the gasifier, thus creating an in-built self-reinforcing and self-regulating system.

Since the bed material serves as the heat carrier, the choice of bed material is critical to achieving the gasification objective. The bed material is an additional process variable involved in controlling the composition of the syngas produced.29  The bed material is exposed to different chemical environments and attrition as it shuttles between the reactors. Resistance to attrition, sintering and fragmentation is a desirable mechanical property in a bed material, as are low cost, availability, recyclability and disposability. The chemical activity of potential bed materials vary widely (see Figure 1.8 29 ), and a chemical property could be advantageous or problematic in the gasification process. For example, SiO2-sand is commonly used as the bed material in fluidised bed systems because of its attrition resistance and it is inert to on the gas streams. However, at an operating temperature of 650–900 °C, SiO2-sand can react with the ash by-product from the gasifier and cause agglomeration.

Figure 1.8

Schematic of the activity spectrum for a range of oxide-based bed materials. Adapted from ref. 29 with permission from Jelena Marinkovic, Copyright © Jelena Marinkovic 2016.

Figure 1.8

Schematic of the activity spectrum for a range of oxide-based bed materials. Adapted from ref. 29 with permission from Jelena Marinkovic, Copyright © Jelena Marinkovic 2016.

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Taking SiO2-sand as a reference inert bed material, the intrinsic catalytic activity of other bed materials influence the composition of the gas produced by the gasification process. For example, Fe-based and Ni-based bed materials are the oxidation and reduction catalysts, respectively. Fe-based bed material gives CO2-rich syngas, while Ni-based bed material gives CH4-rich syngas. Although limestone has poor attrition resistance, its CO2 sorption property is advantageous for obtaining H2-rich syngas. Limestone decomposes into quicklime and CO2 in the combustor but reabsorbs CO2 in the gasifier. This produces a net effect of transfer of CO2 from the gasifier to the combustor (see Figure 1.9,30 ). The CO2 is then expelled along with combustor exhaust. In addition, limestone promotes the WGS reaction by this in situ removal of CO2 from the gasifier. The WGS reaction is mildly exothermic and is favoured at a lower temperature. This explains the observed inverse relationship of the H2/CO ratio with the gasification temperature when using limestone bed material in a DFB gasifier (see Figure 1.9 30 ). The advantage of N2-free syngas production without the use of pure oxygen or the need to install an air purification system has made DFB gasification very attractive for hydrogen and biomethane production. Also, the feed recycling between the combustor and the gasifier allows for complete flexibility with the feed type utilised.

Figure 1.9

Illustration of sorption enhanced gasification effect of limestone in a DFB gasifier. Adapted from ref. 30, https://doi.org/10.1016/j.energy.2019.02.025, under the terms of the CC BY 4.0 license, https://creativecommons.org/licenses/by/4.0/.

Figure 1.9

Illustration of sorption enhanced gasification effect of limestone in a DFB gasifier. Adapted from ref. 30, https://doi.org/10.1016/j.energy.2019.02.025, under the terms of the CC BY 4.0 license, https://creativecommons.org/licenses/by/4.0/.

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Combinations of bubbling and circulating fluidised bed units can give four variants of combustor-gasifier dual-fluidised bed gasification systems: bubbling–bubbling, bubbling–circulating, circulating–bubbling and circulating–circulating. But the circulating bed (combustor)–bubbling bed (gasifier) appears to be superior in terms of feed/fuel utilisation and the heating value of the syngas output.31  This gasifier has become synonymous with dual-fluidised bed (DFB) biomass gasification. Several pilot and industrial biomass gasification plants that are based on the circulating bed (combustor)-bubbling bed (gasifier) DFB have been built and operated.

To enhance fuel flexibility, transitioning from the classical circulating bed (combustor)-bubbling bed (gasifier) to the circulating bed (combustor)-circulating bed (gasifier) is being investigated.32–34  The new features added to the classical DFB and the anticipated improvements in performance include (see Figure 1.10):24 

  • – Replacing the bubbling bed (gasifier) with a circulation bed allows counterflow contact between the bed material from the combustor and the syngas from the gasifier, which leads to longer contact time for tar conversion and WGS reactions.

  • – Gently sloped cyclones for loops to minimise the attrition of limestone for sorption enhanced gasification.

Figure 1.10

Classical (left) vs. advanced design (right) DFB gasifier for steam gasification of biomass. Reproduced from ref. 24 with permission from Elsevier, Copyright 2019.

Figure 1.10

Classical (left) vs. advanced design (right) DFB gasifier for steam gasification of biomass. Reproduced from ref. 24 with permission from Elsevier, Copyright 2019.

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Although the development of the DFB gasifier is a significant milestone in biomass gasification, the fundamental deficiency remains with using biomass as the feedstock for FTS biofuel: about 50% of the carbon in the feedstock may end up as CO2 during syngas production. The amount of carbon in the feedstock that will make it to the final hydrocarbon fuel depends on the efficiency of the FTS and refining processes.

The next section deals with the application of H2-rich syngas from DFB steam gasification and with strategies for improving carbon utilisation of biomass feedstock.

DBL is analogous to DCL, and the chemical basis of DBL is similar to that of DCL. The DBL reactions of a hypothetical biomass feedstock, CHxOy, may be represented by eqn (1.5) and (1.6) as DCL:

Equation 1.5
Equation 1.6

The technical feasibility of carrying out eqn (1.5) through CO2 sorption-enhanced DFB gasification was discussed in the previous section; however, there are technical challenges with carrying out eqn (1.6). Like DCL, DBL is a reductive pyrolysis process that involves subjecting biomass-solvent slurry to high temperature and H2 pressure. During the petroleum supply crisis of 1970s, many alternative means of obtaining liquid fuel were explored, including the DBL process.35  Based on the past success of DCL in Germany (1930–1940s) and South Africa (1960–1980s) and more recently the commercial DCL operation in Shenhua China (since 2009), it is anticipated that DBL will have a lower CO2 footprint than DCL. However, DBL has not enjoyed the same level of success as DCL, and although several DBL projects were initiated and operated on a pilot scale,36  none of the projects progressed to the commercial scale.

Although lignite and biomass have a similar empirical formula, they have a markedly different structure.37  Lignite has a fairly monolithic simple structure compared to biomass, which is a composite of cellulose, hemicellulose and lignin. While there are established hydrogen donor solvents (e.g. tetralin) that can solvate and facilitate the hydrogenolysis and hydrogenation of lignite, it is very challenging to find a solvent/compound that can simultaneously solvate the three components of biomass, also the solvent/compound need to possess reversible hydrogenation/dehydrogenation property. The poor success of direct adoption of the DCL process for DBL is traceable to the differences in the structure of the feedstock used. Pretreatment of the biomass feedstock into intermediates that are similar to lignite can improve the results achieved with DBL. This approach has been explored, albeit under a different process nomenclature – bio-oil formation and its upgrade.38  Consequently Figure 1.1 is revised to include process steps to make DBL feasible (Figure 1.11). Points of CO2 discharges in the scheme suggest that about three-quarters of the carbons in the feedstock may not be made into the final liquid fuel. An alternate approach to pretreatment of biomass is fractionation followed by conversion of each fraction (cellulose, hemicellulose and lignin) into niche liquid fuels.

Figure 1.11

Revised scheme for DBL.

Figure 1.11

Revised scheme for DBL.

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Since the objective of transitioning from coal to biomass liquefaction is to achieve neutral or negative CO2 emissions, considering the energy cost/CO2 emissions associated with biomass harvesting, transportation, processing and cultivation time, the carbon efficiency or biofuel yield of the current liquefaction processes is very low.39  The main CO2 discharge points are hydrogen and heat supply; hence sustainable complementary hydrogen and heat inputs for biomass liquefaction would improve the carbon efficiency.

Assuming the gasification and FTS reactions in eqn (1.3) are stoichiometric as written, only a fraction of the carbon in the feedstock can be converted into liquid fuel. This is because the syngas from oxygen gasification of coal or biomass is hydrogen deficient with respect to the hydrocarbon formation reaction. WGS and CO2 scrubbing are often used to reconstitute the syngas into H2/CO ∼ 2.1 that is suitable for FTS reactions or methanol synthesis. Similarly, the net reaction leading to DBL – eqn (1.7) (the sum of eqn (1.5) and (1.6)) – shows that about 50% of the carbon in the starting feedstock may end up as CO2.

Equation 1.7

The carbon loss steps from hydrogen production via gasification. However, alternate feedstocks for the hydrogen production step will help to save the feedstock and enhance the carbon efficiency and liquid fuel yield. Methane and water are important hydrogen-rich feedstock and use of the H2 feed obtained from methane or water to replace the gasification step in DBL or to augment the H2-deficiency of bio-syngas for FTS reactions is a viable strategy to improve the carbon efficiency of biomass liquefaction processes.

Methane is a hydrogen-rich molecule that has the lowest CO2 emission per unit mass among the fossil fuels. With the increased availability of low-cost natural gas (NG) due to new discoveries, the availability of technology for accessing finds from difficult terrains and the effort to discourage flaring of associated gases, gas-to-liquid (GTL) projects became popular during petroleum supply crisis. With methane as the main constituent, NG is rich in hydrogen, but GTL is not CO2 emission neutral. Biosyngas is H2 deficient, so the gas-biomass-to-liquid (GBTL) process – which is a hybrid of GTL and BTL – has the potential to improve the liquid fuel yield of BTL and also lower the CO2 footprint of GTL.40,41  Combustion of NG also produces more energy per unit mass than biomass.

GBTL can take different blends and forms; for example, hydrogen from steam reforming of methane can replace biomass gasification, or NG combustion can supply the energy required for fast pyrolysis and subsequent upgrade of the resulting bio-oil. A recent study showed that the economic and environmental performance of GBTL fuels depends on NG/biomass ratio blends.42,43  GBTL enhances the economic feasibility of the biomass liquefaction process, but a compromise is required to achieve environmental sustainability. A maximum of 28% NG/biomass blend is considered optimum for meeting the CO2 emission targets for renewable fuel emission standard.44 

Water is a carbon-free source of hydrogen. Splitting water into hydrogen and oxygen can be achieved using either thermal or electrical energy inputs. Electrical energy input for hydrogen production is the focus of this section, while thermal water splitting will be discussed in the next section. The goal of the power-biomass-to-liquid (PBTL) process is the integration of biomass liquefaction processes with water electrolysis using carbon-free electricity (hydro, wind, solar, nuclear, etc.).45  This is similar to other related power-to-fuel46  ideas, where the focus is co-electrolysis of CO2 and H2O to syngas and the production of gaseous fuels (methane,47–49  dimethylether50  or methanol51 ). The similarity is in the use of renewable electricity, but PBTL is different because its target product is liquid fuel.

The process of integrating power (electricity) and BTL depends on the type of biomass liquefaction process (DBL or IBL). The integration can take the form of replacing hydrogen production – via biomass gasification – with hydrogen supply from electrolysis (see Figure 1.12). Oxygen from the electrolysis process may be released into the atmosphere or sold as a by-product to improve the process economy.

Figure 1.12

Proposed scheme for integration of electricity with DBL.

Figure 1.12

Proposed scheme for integration of electricity with DBL.

Close modal

With IBL, a higher level of integration can be envisaged, as proposed in Figure 1.13. Oxygen from electrolysis then becomes a feed for oxy-steam gasification. The water by-product from FTS and the heat recovered from syngas cooling and FTS could be used to generate steam for the oxy-steam gasification process. The syngas produced by the oxy-gasification could be reconstituted, so that the hydrogen added is sufficient for stoichiometric conversion of CO + CO2 into hydrocarbon fuel. WGS and CO2 scrubbing are eliminated; as a result, carbon efficiency and liquid fuel yield are greatly enhanced.

Figure 1.13

Proposed scheme for integration of alkaline electrolysis with oxy-steam biomass gasification.

Figure 1.13

Proposed scheme for integration of alkaline electrolysis with oxy-steam biomass gasification.

Close modal

Water electrolysis technologies can be classified into alkaline, polymer membrane and solid oxide electrolysis. Polymer electrolyte membrane (PEM) electrolysis gives a higher purity H2 and O2; the noble metal electrodes and an expensive polymer membrane used in PEM make it more expensive than alkaline water electrolysis. But alkaline water electrolysis is a mature technology that is commonly used by industry. It has the advantage of low cost due to the use of non-noble metal electrodes that are stable in alkaline media. Purity of H2 and O2 obtained from alkaline electrolysis is suitable and recommended for the proposed PBTL shown in Figure 1.13.

Most alkaline and PEM electrolysers usually operate at <100 °C; however, operating at a temperature of >100 °C can lower the electrical energy required for water splitting (see Figure 1.14,52 ). High-temperature operation of solid oxide electrolysis can permit co-electrolysis of H2O/CO2 mixtures into syngas.53,54  These features can be harnessed to modified the integration scheme, as shown in Figure 1.15. Here oxygen gasification is selected to obtain a H2-lean syngas using entrained-flow gasifiers. The hot syngas from the gasifier is fed into the solid oxide electrolyser for co-electrolysis of H2-rich syngas.

Figure 1.14

Thermodynamics for H2O electrolysis at atmospheric pressure. Reproduced from ref. 52 with permission from Elsevier, Copyright 2017.

Figure 1.14

Thermodynamics for H2O electrolysis at atmospheric pressure. Reproduced from ref. 52 with permission from Elsevier, Copyright 2017.

Close modal
Figure 1.15

Proposed scheme for integration of solid oxide electrolysis with oxygen biomass gasification.

Figure 1.15

Proposed scheme for integration of solid oxide electrolysis with oxygen biomass gasification.

Close modal

The proposed PBTL above are based on low renewable electricity.55  The cost of solar and wind electricity has decreased steadily over the past few decades, and a further decrease in price is expected in the future, as the technologies become more mature and competitive. However, sunshine and wind are intermittent in nature, and battery technologies for use in grid scale storage systems are still in their infancy; thus the proposed PBTL offers a means of storing excess renewable electricity as a portable liquid fuel.

A significant fraction of a given biomass feedstock is sacrificed to supply heat and hydrogen in biomass liquefaction processes. Solar thermal is carbon-free and can therefore be harnessed to improve the carbon efficiency of biomass liquefaction processes.56  Integration of the solar thermal process with the biomass-to-liquid process is termed the solar-thermal-biomass-to-liquid (STBTL) process. Solar thermal energy can be exploited as a source of heat energy or to generate hydrogen for the liquefaction process.57,58  Both processes require a high temperature: 600–1200 °C. Normal solar insolation on the surface of our planet is between 200 and 400 W m−2, which corresponds to a black body temperature of 18–35 °C. Thus normal solar insolation needs to be concentrated about 500–1000 times59  before it is useful in thermochemical processes.

Improvement of carbon efficiency has been demonstrated in solar-assisted biomass gasification.60  Different solar-powered gasifier configurations have been investigated, such as updraft and downdraft.61,62  Direct and indirect solar heat supply modes to the gasifier have also been reported. It has been shown that the use of solar thermal energy can reduce CO2 emissions from the gasification process; however, simulated concentrated solar radiators were used in these studies.63  Very few studies in this field have involved the use of real concentrated solar power for biomass gasification.64–68 Table 1.4 shows the performance parameters of some typical solar collector installations.69  The solar thermal installations occupy large areas, and integration with gasifiers would be very challenging. Moreover, most concentrated solar thermal installations are cited in arid areas where biomass is scarce. Nevertheless, electricity produced from concentrated solar thermal is renewable, and it can be transmitted to locations where there is a sustainable supply of biomass for PBTL.

Table 1.4

Performance parameter of typical solar collectors. Adapted from ref. 69 with permission from Elsevier, Copyright 2019.

Solar collectorConcentration ratioOperating temp (°C)
Parabolic trough 30–100 400 
Linear Fresnel 30 300 − 400 
Tower collector 500–5000 400–600 
Disc collector 1000–10 000 550–750 
Solar collectorConcentration ratioOperating temp (°C)
Parabolic trough 30–100 400 
Linear Fresnel 30 300 − 400 
Tower collector 500–5000 400–600 
Disc collector 1000–10 000 550–750 

We had identified the factors responsible for low carbon efficiency of biomass liquefaction processes and also highlighted strategies to reduce carbon loss in BTL. The cardinal focus of these strategies is to salvage biomass fractions that would have hitherto been sacrificed to supply heat or hydrogen for the process. Solar energy is the most abundant and sustainable renewable carbon-free resource that can be used to improve the liquid fuel yield from biomass liquefaction processes. Ab initio, biomass (and by extension, fossil fuels) are products of solar energy, so the strategies essentially consist of fossilisation of biomass with more solar energy, i.e. forcing more solar energy into biomass, thereby increasing its energy density. Next we discuss syngas cleaning, before taking a detailed look at the second part of the IBL process, i.e. FTS.

As indicated in Table 1.1, biomass generally has a low sulphur content (<0.25%) but a significant nitrogen content (∼1.28%). The resulting nitrogen and sulphur compounds (H2S, CS2, thiophenes, thiols, HCN, NH3, etc.) in the gasification product will not only poison the FTS catalyst but are also potential air pollutants. Hence removal of these gases from the syngas is an important step in the IBL process. Also, tar is a common undesired constituent in virtually all biomass-derived syngas irrespective of the gasifying agents but tar density will depend on the gasification technology employed.70  Tar comprises a wide spectrum of condensable organic compounds, which vary from small oxygenate molecules to large acyclic, cyclic, aromatic and polycyclic aromatic hydrocarbon molecules. The conditions inside the biomass gasifier promote the formation of hundreds or thousands of different tar molecules.70  Condensation of tar initiates fouling, which can result in malfunctioning or plugged equipment. Other important impurities in biomass-derived syngas are chlorine, ash and char particulates.

A number of strategies are being investigated for cleaning biomass-derived syngas. These strategies can be grouped according to the exit temperature of the syngas from the clean-up device: hot (T > 300 °C), warm (300 > T >100 °C) and cold (T < 100 °C) gas cleaning strategies. The cold strategy – also known as the wet method – involves the use of wet scrubbers or bubblers.71  It uses relatively mature techniques such as absorption or adsorption. Wet methods are generally effective but suffer from energy loss and wastewater generation. The hot cleaning strategy has the advantage of less energy loss because it does not include the process of cooling and reheating the gas stream. However, many of the techniques being used are still under development, and technical challenges are involved in handling gas at a high temperature.72–81  The warm cleaning strategy is a hybrid of the hot and cold methods.

Here, we focus on practical methods for bio-syngas cleaning for FTS applications. In general, the logical sequence of cleaning steps for producing syngas from biomass gasifiers is as follows: particulate removal; removal of organic (tar) impurities; removal of inorganic impurities; removal of volatile (alkali/heavy) metals. The most challenging of the impurities is tar, and it attracts more attention from researchers in terms of cleaning biomass-produced syngas. The oil-based gas washing process (OLGA), developed by the Energy Research Centre of the Netherlands (ECN),82  focussed on removing tar but also encompassed the removal of other categories of impurities. It is based on the control of dew point (the temperature at which tar starts to condense). OLGA as applied to biomass-derived syngas for FT applications is shown in Figure 1.16.83 

Figure 1.16

OLGA-based biomass-derived syngas cleaning protocol.83 

Figure 1.16

OLGA-based biomass-derived syngas cleaning protocol.83 

Close modal

Hot syngas from the gasifier is first passed over a solid filter, where ash and particulate impurities are removed. The syngas is then channelled into an oil scrubber or bubbler (the OLGA unit), where lipophilic tar molecules are removed. The exit syngas is then passed through water to remove NH3, HCl and other inorganic impurities. After the three scrubbing steps are completed, the syngas is considered suitable for FTS. The order of the scrubbing step is critical to ensuring a smooth gas cleaning operation. Biochar is a by-product of biomass gasification that is often employed as the solid filter, while vegetable oil is the most common oil scrubber or bubbler. Bio-oil, biodiesel and surfactant solutions have also been investigated for tar scrubbing.

The sulphur content in biomass is usually very low, and for most practical purposes, the cleaning step used to remove sulphur impurities in biomass-derived syngas is often unnecessary with FT applications. If the sulphur content in the biomass feedstock is high, the sulphur removal step becomes necessary. This can be accomplished using the established conventional acid gas wet scrubbing techniques, which removes acidic gas such as H2S, SO2, NOx and CO2. This step may be beneficial as part of the gas-conditioning step that reduces the CO2 load of the commonly CO2-rich and/or H2-deficient biomass-derived syngas. Where a low-cost H2 feed from water electrolysis is available, the rest of the gas-conditioning step involves augmenting the Ribblett ratio H2/(2CO + 3CO2) to unity.84 

In Section 1.2.2, biomass gasification that leads to syngas production and syngas cleaning was examined. The carbon loss associated with the process was noted, and ways to enhance carbon efficiency were discussed. In this section, the second stage of the IBL process, i.e. FTS, is examined. It involves the conversion of cleaned bio-syngas to hydrocarbon products.

Hydrocarbon formation is normally obtained at a temperature of 180–350 °C and 10–30 bar pressure, depending on the type of catalyst and reactor employed. More than one reaction is possible at FTS reaction conditions; Table 1.5 shows a list of likely reactions in an FTS reactor.85  Generally, product selectivity follows the order of alkanes > alkenes > alcohols, while the full product spectrum is a function of the catalyst type and reaction conditions.

Table 1.5

Fischer–Tropsch synthesis reactions. Adapted from ref. 85 with permission from Elsevier, Copyright 2010.

Main reactions  
1. Methane CO + 3H2 → CH4 + H2
2. Paraffins (2n + 1) H2 + nCO → CnH2n + 2 + nH2
3. Olefins 2nH2 + nCO → CnH2n + nH2
4. WGS CO + H2O → CO2 + H2 
 
Side reactions  
5. Alcohols 2nH2 + nCO → CnH2n+1O + nH2
6. Boudouard reaction 2CO → C + CO2 
 
Catalyst modifications  
7. Catalyst oxidation/reduction a. MxOy + yH2yH2O + x
 b. MxOy + yCO → yCO2 + x
8. Bulk carbide formation yC + xM → MxCy 
Main reactions  
1. Methane CO + 3H2 → CH4 + H2
2. Paraffins (2n + 1) H2 + nCO → CnH2n + 2 + nH2
3. Olefins 2nH2 + nCO → CnH2n + nH2
4. WGS CO + H2O → CO2 + H2 
 
Side reactions  
5. Alcohols 2nH2 + nCO → CnH2n+1O + nH2
6. Boudouard reaction 2CO → C + CO2 
 
Catalyst modifications  
7. Catalyst oxidation/reduction a. MxOy + yH2yH2O + x
 b. MxOy + yCO → yCO2 + x
8. Bulk carbide formation yC + xM → MxCy 

For purposes of simplicity, all the reactions in Table 1.5 are approximated as:

Equation 1.8

As pointed out in Section 1.1.2, biosyngas normally contains a significant amount of CO2, so taking the CO2 content into consideration, the stoichiometrically balanced syngas feed should have a Ribblett ratio H2/(2CO + 3CO2) = 1.86,87  Catalysts with low WGS activity require a stringent syngas composition of H2/CO ≈ 2. However, active WGS catalysts can form hydrocarbon with H2/CO < 2, especially at high-temperature mode, where WGS proceeds to equilibrium. For instance, the net reaction for hydrocarbon formation on active WGS catalyst using syngas with H2/CO = 1 can be written as per eqn (1.9):

Equation 1.9

Half the CO is ejected as CO2, which results in a maximum carbon efficiency of 50%, assuming stoichiometric conversion is achieved. For co-conversion of CO and CO2 to hydrocarbon, H2/(2CO + 3CO2) = 1 is required. The reverse WGS (rWGS) reaction of CO2 to CO may be a prerequisite for subsequent hydrogenation to hydrocarbon; however, alternate CO2 hydrogenation to hydrocarbon via the methanol intermediate has also been recognised.88–90 

The main FTS reaction conditions are temperature, catalyst type, H2/CO ratio and reactor type. There are two modes of operations in commercial FTS, which are based on temperature:

  • – High-temperature Fischer–Tropsch (HTFT) mode: the reactor temperature is between 300 and 350 °C. Gasoline and linear low-molecular-mass olefins are selectively produced with an Fe-based catalyst, and significant amounts of oxygenates are also produced.

  • – Low-temperature Fischer–Tropsch (LTFT) mode: the reactor temperature is usually between 200 and 250 °C, and either Fe- or Co-based catalysts are used. This mode gives products with a high selectivity for paraffins and high-molecular-mass linear waxes.

The dominant design consideration in terms of a Fischer–Tropsch reactor is removing the large exothermic heat of reaction.91  The choice of reactor type for a commercial operation is largely dictated by the desired product selectivity. Prominent FTS reactor types are:

  • Circulating fluidised bed (CFB) reactor,

  • Fixed fluidised bed (FFB) reactor, also known as Sasol advanced synthol (SAS) reactor,

  • Multitubular fixed-bed (MTFB) reactor, and

  • Slurry phase reactor.

An accurate kinetic description is important for the evaluation of activity and comparison of catalysts. It is generally accepted that the specific activity of FT active metals follows the order of Ru > Ni > Co > Fe. Only iron and cobalt are used in commercially FTS operations. Although FTS is generally considered to be a polymerisation reaction in nature, polymerisation kinetic expressions92–94  are not directly applicable because monomer species are formed in situ on the catalyst surface in FTS, and more than one potential monomer is possible. Thus FTS kinetics focus on the initiation step and monomer formation. The rate law of such heterogeneous gas phase reactions are usually derived using the Langmuir–Hinshelwood or Eley–Rideal mechanisms, and the Langmuir–Hinshelwood mechanism has been shown to give a more accurate description of the monomer formation than the Eley–Rideal mechanism.95,96 

Although CO activation is the rate-limiting step in monomer formation, there is no agreement on the monomer species that undergoes repeated additions in order to give hydrocarbons. According to Krische et al.,97,98  organometallic intermediates that arise transiently in the course of catalytic hydrogenation of CO are intercepted and re-routed to C–C coupling. Potential intermediate candidates include HCO*, HCOH*, CH* and . These intermediates correspond to the proposed propagation mechanisms CO insertion, enol, alkylidene and alkenyl, respectively. The monomer formation and propagation mechanisms are illustrated in Figure 1.17.99 

Figure 1.17

Proposed FTS mechanisms. Reproduced from ref. 99 with permission from the Royal Society of Chemistry.

Figure 1.17

Proposed FTS mechanisms. Reproduced from ref. 99 with permission from the Royal Society of Chemistry.

Close modal

FTS products fall in a spectrum of hydrocarbons ranging from C1 to C50. A typical gas chromatogram of syncrude is characterised by a notable degree of order with respect to the carbon number of the molecules.100–102  The mole fraction of the components in each carbon number fraction declines exponentially with the carbon number, and this behaviour is typical of polymerisation reactions. This is what informed the description of FTS as a surface polymerisation reaction at the earlier stage of its development. The polymerisation ideal conceived as FTS proceeds via stepwise addition of a monomer carbon species and one termination constant for all products.103 

Instead of a tabular list or pictorial presentation of the products, a simple and unambiguous model known as Anderson–Schultz–Flory (ASF) distribution is used to provide a simpler comparison of the selectivity of catalysts at different reaction conditions. ASF is an empirical statistical distribution. It is described in eqn (1.10).

Equation 1.10

A semilogarithm plot of the product mole fractions as a function of the carbon number should give a straight line with a slope, α, the chain growth probability. The ASF distribution is based on the assumption that the chain growth probability is constant and that there is stepwise addition of monomer- leads to saturated hydrocarbons as terminal products. But product distribution of FTS catalysts usually deviate from the ASF distribution law (see Figure 1.18). The beauty of ASF is that, for most practical purposes, is its one model parameter, α, that provides a broad overview of the product spectrum, and it is sufficient for describing syncrude. ASF provides information that guides the selection of the refining technology needed to obtain the target fuel from the syncrude.

Figure 1.18

Theoretical ASF vs. Experimental FTS product distributions.

Figure 1.18

Theoretical ASF vs. Experimental FTS product distributions.

Close modal

Syncrude (short form of synthetic crude oil) is so-called because like conventional crude oil it is made up of mixtures of hydrocarbon molecules as the main constituent and oxygenates as the minor constituents. Syncrude that has not been refined or upgraded cannot be used as fuel to run engines. The practice in the petroleum refineries in dealing with crude oil is to sort out the constituents based on the boiling point range for different engine types. The next steps are specific processes such as reforming, cracking, extraction blending etc., which are applied in order to meet the requirements for engine fuel. However, refining syncrude demands a holistic approach, and there are very few dedicated syncrude refineries around the globe, consequently, there are very few experts in this area. Arno de Klerk has made a significant contribution to the literature on syncrude refining, based on his experience at Sasol (South Africa).104  Some of the highlights from his contribution are dealt with here.

Since FTS is a bottom-up process, the target fuel or chemical products should be in view right from syngas production so as to tailor the FTS mode and subsequent refining processes toward maximising the desired products.104  Furthermore, from a carbon efficiency and process cost point of view, it is desirable to obtain the target fuel using as few process steps as possible. The product distribution of some commercial Fischer–Tropsch technologies is shown in Figure 1.19.104  It shows that the HTFT mode delivers more straight-run naphtha (gasoline-range carbon number) fractions and lower distillate (diesel-range carbon number) fractions – and vice versa for the LTFT mode.104  Hence, having a target fuel in view from the start, allows for choice of appropriate FTS mode and catalyst selection to be made.

Figure 1.19

Carbon number distribution of some commercial Fischer–Tropsch technologies. Reproduced from ref. 104 with permission from Prof. Arno de Klerk.

Figure 1.19

Carbon number distribution of some commercial Fischer–Tropsch technologies. Reproduced from ref. 104 with permission from Prof. Arno de Klerk.

Close modal

Another important aspect of syncrude refining is to take cognisance of the differences between syncrude and petroleum. The blank adoption of established petroleum refining technologies for use with syncrude may prove unwise. Detailed attention should be given to the molecular differences between petroleum and syncrude when selecting refining technologies for syncrude. Although the first diesel engine ran on vegetable oil, development and improvement of the diesel engine (and, by extension, petrol and jet engines) were tailored to use petroleum-derived fuels. Generally, syncrude naphtha and distillate are poor in aromatics and isoparaffins; HTFT naphtha and distillate are rich in olefins, while those of LTFT are rich in n-paraffins. Because of the higher olefin content, syncrude fractions (especially HTFT) are more reactive than petroleum fractions.105  Hence mild catalyst activity and reaction conditions are more appropriate for refining such syncrude. HTFT naphtha and distillate have a higher oxygenates content than LTFTs; in fact, the water by-product of FTS contains some short-chain oxygenates. The aqueous (oxygenates) fraction of syncrudes is about 7% for HTFT and 3% for LTFT. These aqueous products are also FTS products that have to be refined separately from the syncrude. The naphtha fractions in syncrude do not meet the specifications for gasoline fuel. Upgrading the paraffins rich in LTFT naphtha to comply with the specifications for gasoline fuel requires hydroisomerisation of the paraffins and olefins to isoparaffins and then partial aromatisation. This can be accomplished using a Pt/MOR catalyst at 250–280 °C. For the olefin-rich naphthas, the upgrading may lead to greater aromatisation.105 

HTFT syncrude has a higher naphtha fraction than LTFT; thus it makes sense to build the HTFT syncrude refining process around naphtha and select the processes for converting naphtha to the specified gasoline fuel, as previously discussed. Other fractions of HTFT include distillate, gases and aqueous products. The gas fraction is rich in lower olefins, which can be separated into individual olefins and sold as valuable products (monomers), but this may require expensive cryogenic operations. The olefin mixtures may be oligomerised and hydrogenated into gasoline-range isoparaffins, while a fraction of the olefins may be aromatised into a BTX (benzene, toluene and xylene) mixture. Alkylation of BTX with olefins can give alkyl benzenes that are suitable for use as blending feed (gasoline, jet fuel or diesel), or they can be sold as a high-value chemical for the surfactant industry. The distillate fraction can be used directly or with mild hydroisomerisation as jet fuel or diesel fuel. However, the distillate fraction of HTFT syncrude is barely 7%, and the aqueous fraction may be fractionated into its major constituents (alcohol and carbonyls). The alcohol fraction can further be separated into individual alcohols, which can then be sold as valuable products. The carbonyls can be hydrogenated to deliver products that are suitable for use as gasoline fuel.106 

Table 1.6 presents a comparison of the fuel properties of Co-LTFT and Fe-HTFT distillates with the European Union's specifications for diesel fuel (EN590:2004).107  The most critical properties in syncrude distillate utility are density, cetane number and cloud point. As shown in Table 1.6, the Fe-HTFT distillate meets the specifications; however, the distillate fraction of Fe-HTFT syncrude is barely 7%. The Co-LTFT mode, where the distillate yield is higher, does not meet the requirements in terms of density and cloud point. EN590:2004 details the required properties of petroleum-derived diesel, which generally contain fairly equal proportions of n-paraffins, cyclic paraffins and aromatics. In this regard, the Fe-HTFT distillate is a close mimic in terms of paraffins, isoparaffins and aromatics. The Co-LTFT distillates comprise mainly n-paraffins, which accounts for the lower density and higher cloud point.

Table 1.6

Selected fuel properties: Diesel specifications in EN590:2004 vs. straight run distillates from Co-LTFT and Fe-HTFT. Adapted from ref. 107 with permission from American Chemical Society, Copyright 2009.

Fuel propertyEN590:2004Co-LTFTFe-HTFT
Density at 15 °C (kg m−3820–845 772 844 
Cetane number 51 min 80 53 
Cloud point (°C) − 10 < x < −34 − 11 
Viscosity at 40 °C (cSt) 2.0–4.5 2.11 3.62 
Lubricity HFRR wear scar (µm) ≤ 460  251 
Aromatic content (mass%)  < 0.1 34 
T95 boiling point (°C) ≤ 360 291 339 
Fatty acid methyl esters (vol%) ≤ 5   
Sulphur content (µg g−1≤ 10   
Fuel propertyEN590:2004Co-LTFTFe-HTFT
Density at 15 °C (kg m−3820–845 772 844 
Cetane number 51 min 80 53 
Cloud point (°C) − 10 < x < −34 − 11 
Viscosity at 40 °C (cSt) 2.0–4.5 2.11 3.62 
Lubricity HFRR wear scar (µm) ≤ 460  251 
Aromatic content (mass%)  < 0.1 34 
T95 boiling point (°C) ≤ 360 291 339 
Fatty acid methyl esters (vol%) ≤ 5   
Sulphur content (µg g−1≤ 10   

A portion of the paraffins can be isomerised into iso-/cyclic paraffins using the low-energy penalty hydroisomerisation step over Pt/ZSM-12 catalyst. This improves the cold flow property, but there is little impact on density. Multistep and high-energy processes are required to upgrade the density at a compromised cetane number. Thus the cetane number and density appear mutually exclusive in terms of diesel fuel. So with respect to low-energy penalty-refining of syncrude to on-specification diesel fuel from syncrude, the low-yield Fe-HTFT distillate meets all three specifications: the high-yield Co-LTFT distillate does not meet the requirements in terms of density. This conundrum of reconciling yield and density is what De Klerk described as the density-cetane-yield triangle.107 

In tropical countries, where the cold flow requirement for diesel fuel is not stringent, hydroisomerisation of the straight-run distillate becomes less important. Also, where a minimum density of diesel is a requirement, the straight-run LTFT distillate can be used directly in diesel engines, so that the focus is then only on maximising yield. However, when density and cloud point requirements are stringent, hydroisomerisation of the straight-run distillate becomes necessary to improve the cold flow property. Blending or co-refining is the least demanding way to improve density with little compromise in terms of the cetane number. Blending or co-refining involves adding or co-processing a denser feed with the less dense LTFT distillate. Since we have to be conscious of the CO2 emission footprint of the entire biomass liquefaction process, the complementary feed must be from renewable sources. Feedstock types that satisfy this criterion are bio-oil and biodiesel/vegetable oil.

Bio-oil is crude obtained via DBL, which is rich in aromatics and oxygenates. Aromatics are usually denser than paraffins, while oxygenates are beneficial for improving the lubricity of diesel fuel. Hence blending the LTFT distillate with refined bio-oil or co-refining (hydrodeoxygenation) bio-oil with syncrude can help to overcome the density-cetane-yield triangle problem.

Biodiesel/vegetable oils are renewable feedstock that is denser than the LTFT distillate. High viscosity of vegetable oils makes them less suitable for use directly in diesel engines, but they become suitable once transesterified into biodiesel. Most biodiesel products are denser than the LTFT distillate and have a high cetane number; so, blending the LTFT distillate with biodiesel will improve the density without compromising the cetane number. Hydrocracking is another means to lower viscosity of vegetable oil. Vegetable oils such as soybean, sunflower and camelina oils, which are rich in poly-unsaturated fatty acids, produce hydrocracked products that are rich in aromatics. Therefore co-hydrocracking over Pt/Al2O3/SAPO-11 of LTFT wax or vegetable oil, or co-hydroisomerisation of the LTFT distillate with vegetable oil will deliver a product with a sufficient aromatics content to boost the density of the resulting fuel.107 

Irrespective of the starting feedstock, all indirect liquefaction processes, i.e. feed-to-liquid (XTL), can be broken down into three distinct processes108  (see Figure 1.20). From process cost point of view, BTL, particularly where the energy density of the biomass feedstock is low, it is desirable to reduce the number of distinct categories of processes, so as to minimise the energy expenditure and CO2 emission of the overall BTL process. The most feasible way to achieve this objective is to merge the latter two processes into one, i.e. achieve direct conversion of syngas to the specified fuel or chemicals.

Figure 1.20

Overview of indirect liquefaction processes – XTL.

Figure 1.20

Overview of indirect liquefaction processes – XTL.

Close modal

Traditionally FTS is a polymerisation-type reaction, albeit the monomer formation takes place on the catalyst surface, along with chain propagation. Whether the active sites for the sequence of reactions are the same or different is an open question. The FTS-produced syncrude is a broad spectrum of mainly linear hydrocarbons and some minor oxygenates. However, only a very narrow range within the broad spectrum of products fits the requirements for usage as fuels or chemicals. The product distribution is defined by chain growth probability, α.100,109,110  Desired product range may be enhanced through catalyst formulation (use of structural and electronic promoters of an active FT metal), but selectivity to a useful product range (e.g. C2–C4 olefins, C5–C11 gasoline, C8–C16 jet fuel and C10–C20 diesel) is still limited.111  Hence refining (separation and even further conversion processes) becomes necessary to maximise the desired products obtained.

Although syncrude refining is said to constitute 22% of the overall cost of XTL processes, merging the syncrude refining process with the FTS or direct syngas-to-fuels or chemical process would be a huge boost to the process economy.112  To this end, much effort has been made to overcome the ASF product distribution limitation. A rational approach could be to modify FTS by introducing chain scissoring to curtail the unbridled chain growth or the departure from the polymerisation-type reaction mechanism. These two approaches are examined in the next sections.

Modified FTS approaches involve controlling product chain length by introducing a chain scissoring catalyst. Bifunctional catalysts that consist of FTS and C–C cleavage (hydrocracking) active sites are used to break the ASF product distribution. Although some measure of chain length control is obtainable with the use of promoters (for example, copper), a promoted iron-based FT catalyst usually displays high distillate selectivity and low wax selectivity.113,114  This is due to the increase in chain terminations as a result of high hydrogenation activity engendered by copper promotion. However, in the bifunctional catalyst, product selectivity is achieved not by manipulation of chain growth probability, α value, but by scissoring the chain to the desired length by coupling high α-value FT catalysts with zeolites.

The FT catalyst gives the ASF product spectrum defined by α-value, while the zeolite imposed a confining effect or C–C scissoring (hydrocracking) to redefine the product spectrum. Zeolites are framework solids with a varied but well defined pore size and structure. Depending on the acidity, pore size and structure of the zeolite, product selectivity can be tilted toward gasoline, jet fuel or diesel fuel chain length ranges.115  Another advantage of the zeolite is that it can catalyse other desirable refining reactions, such as hydroisomerisation, oligomerisation, aromatisation, dehydration and etherification. Thus by careful choice of the FT catalyst and zeolite, a one-step syngas process to produce the specified fuel can be achieved.

Spatial arrangement of the FT catalyst and the zeolite catalyst is critical to the performance of the hybrid catalyst. Four spatial arrangements of FT and zeolite catalysts have been reported:116 

  • – Mixing the FT catalyst with the zeolite

  • – Using two catalyst beds

  • – A core–shell catalyst design

  • – Using a zeolite-supported FT active metal

The proximity of the FTS and the acid sites of the zeolite is crucial to the efficiency of the tandem reactions in obtaining liquid fuel.117  Among the spatial arrangements, the core–shell catalyst design and zeolite-supported FT active metal offer the greatest chance of ensuring close proximity of the active sites. The zeolite-supported FT active metal is simply obtained by loading the FT active metal onto the zeolite using the impregnation or ion exchange method, while the core–shell catalyst consists of a conventional FT catalyst as the core, surrounded by zeolite as the shell.

A high gasoline range selectivity (C5–C11) with an isoparaffin/n-paraffin ratio has been reported for FT core–zeolite shell hybrid catalysts.118  However, core–shell structure bifunctional catalysts are difficult to produce using the hydrothermal method. Growing the zeolite shell around the FT catalyst pellets is a challenge because the strong alkaline condition that is usually employed may be detrimental to the FT core catalyst.119  Alumina and silica commonly used as the support for the FT active metal, including copper, which is a ubiquitous reduction promoter with iron-based FT catalysts – are prone to dissolution in strong alkaline media. For a silica-supported FT core, e.g. Co/SiO2, a prior silicalite-1 coating placed before growing the zeolite shell gave a catalyst with good performance. Adding the Fe50Al50 alloy to the zeolite growth recipe has been demonstrated as a facile one-pot method to FT core–zeolite shell hybrid catalyst.120  The strong alkaline media for growing the zeolite leaches the aluminium out of the alloy, and the zeolite that formed around the iron or iron oxide is left behind.

Recently, a synthetic method for achieving encapsulation of the metal oxide nanoparticles within the zeolite was demonstrated.121  Encapsulation is achieved by adding a metal nitrate solution containing the chelating ligand (N-[3-(trimethoxysilyl) propyl]ethylenediamine; TPE) to the zeolite recipe. The ligand prevents precipitation of the metal cations in the alkaline zeolite recipe. The trimethoxysilyl moiety of the ligand forms siloxane linkages with the nucleating zeolite precursors, thereby incorporating the metal oxide precursor into the crystallising frameworks. Oxidative treatment of the crystalline framework removes the ligand species and forms metal-oxide-encapsulated zeolite.

Conventional hydrocrackers are normally operated in the temperature range 350–450 °C.122  Incorporation of a (de) hydrogenation metal such as Pt enhances catalyst activity and stability. Hydrocracking of FT wax can take place at a lower temperature, depending on the strength of the zeolite acid sites. The formation of cracked products is preceded by an isomerisation step; hence hydrocracking and hydroisomerisation can take place on the same catalyst but at a lower temperature: 250–280 °C. Oligomerisation of lower olefins can also readily take place on zeolite with the LTFT operation mode. Hence hydrocracking, hydroisomerisation and oligomerisation can take place under both LTFT and HTFT conditions.

Typically, the HTFT product spectrum is characterised by high naphtha selectivity and high olefin and oxygenate selectivity. At HTFT temperatures, dehydration of oxygenates, isomerisation of naphtha, oligomerisation and aromatisation of lower olefins readily take place on the zeolite. However, strong acidic zeolite, such as H-ZSM-5 is prone to deactivation, due to coke forming when using this operation mode. Hence weaker acidic zeolite is more suited to achieving tandem FTS and refining reactions under the HTFT mode. With the LTFT mode, hydroisomerisation and oligomerisation are more favourable than hydrocracking over an H-ZSM-5 catalyst, as hydroisomerisation of olefins to iso-olefins discourages chain growth on the catalyst surface. Large linear hydrocarbon chains are also reactive to hydrocracking over H-ZSM-5 at LTFT temperatures. This may account for why most reports on syngas-to-fuel conversions that involve FT-zeolite bifunctional catalysts operated in the LTFT mode display selectivity toward the gasoline-range rather than the diesel-range hydrocarbons.

Direct syngas-to-jet fuel and diesel fuel processing over an FT-zeolite has been less successful, compared to the direct syngas-to-gasoline fuel process.123  The reactions required on zeolites to increase the distillate selectivity are oligomerisation and hydrocracking. Weaker-acid-strength and lower-acid-site-density zeolite is more suitable for oligomerisation, selectivity and hydroisomerisation, which discourage chain growth, may be favoured over hydrocracking. Use of hierarchical zeolites have been explored to enhance distillate selectivity.124,125  Hierarchical zeolites combine the advantages of microporous zeolites (e.g. acidity functions, molecular sieve features and high temperature stability) and mesopores for easier mass transfer of larger molecules. Syngas conversion over a FT-mesozeolite delivered promising distillate selectivity.126–128  The acid sites of mesoporous zeolite were treated with an alkali to minimise the unwanted acid-catalysed reactions. Reagents used for synthesis of mesozeolite are expensive, making this catalyst prohibitive in large-scale operations. Spatial separation of the FT and the hydroisomerisation active sites can minimise early shortening of the product chain; thus a LTFT mode that uses a two-bed catalyst arrangement is considered a more practical and plausible approach to direct syngas-to-diesel or syngas-to-jet fuel process.

Syngas conversion aimed at breaking ASF product distribution is achieved by departing from the polymerisation mechanism: the quest for a non-ASF product distribution is approached from monomer formation step. The polymerisation mechanism is avoided by using non-FTS catalysts to effect different pathways of CO activation and monomer formation. For example, FTS and methanol synthesis present the two extremes of syngas reactions. FTS gives an ASF product distribution, while methanol synthesis gives simple oxygenates as the major product. In principle, between these two extremes there is a continuum of possibilities of tuning the product selectivity through catalyst design and varying the reaction conditions. In practice, the catalyst design for such selective syngas conversion has been less successful until recently. The successful catalysts have features that resemble some earlier known syngas conversions, such as methanol synthesis, higher alcohol synthesis and isosynthesis. Since methanol synthesis is a well developed and mature process, and, because methanol can be readily converted into hydrocarbons and chemicals, it is also an attractive and plausible route for non-ASF product selectivity via one-pot tandem syngas conversion to fuels and chemicals. Some of the tandem syngas conversion methods are examined in the following section.

Dimethyl ether (DME) is a gas at room temperature, which can be liquefied at a low pressure; hence it is considered an environmentally friendly refrigerant and aerosol propellant alternative to Freon and R-134.129,130  DME is an intermediate to a plethora of chemicals, such as methylacetate, aromatics, acetic acid, light olefins and aromatics.130  DME has a high cetane number; hence it has attracted attention in recent years for use as a diesel fuel, and pilot tests of DME-powered buses have been carried out in China. It is clean burning and sulphur free, and it gives lower NOx and particulate emissions.131,132  DME is the smallest of the ethers and is usually obtained by dehydration of methanol. Conventional DME involves a two-stage process: methanol synthesis and methanol dehydration. Commercial synthesis of methanol from syngas is usually carried out using a CuO–ZnO–Al2O3 catalyst, while methanol dehydration can be achieved using a variety of acid catalysts (H2SO4, H3PO4, γ-alumina, zeolites, polyoxometallate etc.).133,134 

The reactions involved in syngas conversion to DME are provided in Table 1.7.135  Maximum equilibrium production of methanol is attained at an initial H2/CO = 2 and DME of H2/CO = 1. Biomass-derived syngas usually has an H2/(CO + CO2) ∼ 1; thus biomass-derived syngas is more suitable for DME synthesis than methanol or FTS.

Table 1.7

Reactions for syngas conversion to DME. Adapted from ref. 135 with permission from Springer Nature, Copyright 2000.

Methanol synthesis 2CO + 2H2 ↔ 2CH3OH ΔH° = −90.4 kJ mol−1 
 CO2 + 3H2 ↔ CH3OH + H2ΔH° = −49.4 kJ mol−1 
DME formation 2CH3OH ↔ CH3OCH3 + H2ΔH° = −23.04 kJ mol−1 
WGS CO + H2O ↔ CO2 + H2 ΔH° = −41.0 kJ mol−1 
Overall 3CO + 3H2 ↔ CH3OCH3 + CO2 ΔH° = −258.312 kJ mol−1 
Methanol synthesis 2CO + 2H2 ↔ 2CH3OH ΔH° = −90.4 kJ mol−1 
 CO2 + 3H2 ↔ CH3OH + H2ΔH° = −49.4 kJ mol−1 
DME formation 2CH3OH ↔ CH3OCH3 + H2ΔH° = −23.04 kJ mol−1 
WGS CO + H2O ↔ CO2 + H2 ΔH° = −41.0 kJ mol−1 
Overall 3CO + 3H2 ↔ CH3OCH3 + CO2 ΔH° = −258.312 kJ mol−1 

Direct or one-pot syngas-to-DME processing requires integration of the two catalysts. As indicated in the previous section, there are four possible spatial arrangements for the two catalysts and the proximity of the methanol synthesis and dehydrate sites that are crucial to ensuring efficiency of the tandem reactions to DME. However, the catalyst used in methanol synthesis presents an interesting situation. The established methanol synthesis catalyst is CuO–ZnO–Al2O3 with a Cu/Zn/Al molar ratio of about 4/5/1. CuO is the active site precursor; ZnO is an electronic promoter, while Al2O3 is considered a textural promoter. However, Al2O3 is also a methanol dehydration catalyst; so increasing the proportion of Al2O3 in the catalyst formulation should give a bifunctional DME catalyst. This was the strategy adopted by Shizuoka University, albeit using the sol–gel method to develop a high-DME-selective catalyst (Cu–Zn(36–4 wt%)/Al2O3).137 

There have been reports regarding stability issues when using γ-Al2O3 as a dehydration catalyst because of its hydrophilicity. This issue only occurs when using syngas of H2/CO > 1, when the overall reaction is 2CO + 3H2 ↔ CH3OCH3 + H2O, and the by-product is water. However, silica-supported polyoxometalates or H-ZSM-5 with stronger acid sites than γ-Al2O3 are good candidates for methanol dehydration to DME.136  Integration of these acid catalysts with a methanol synthesis catalyst is more demanding, and fine-tuning the CuO–ZnO–Al2O3 composition is still preferable, based on cost and simplicity of process. Most commercial DME synthesis technologies use a syngas of H2/CO = 1, a fixed bed or slurry reactor at 220–290 °C and pressure of 1.4–6 MPa.137 

Light olefins (C2–C4) are important bulk chemicals that are traditionally produced by cracking petroleum naphtha.138  In recent times, conversion of syngas to light olefins has attracted attention as a viable alternative.139–143  Direct syngas-to-olefins require tandem catalytic processes. On the basis of intermediates, three syngas-to-olefin processes have been proposed: the methanol-to-olefin (MTO) route, non-FT routes and the FT route. MTO is a well researched process that has proved successful in commercial operations.144–146  The intermediates involved in the non-FT routes are methanol/DME, carbonylation and ketene.115  The non-FT and FT routes are briefly examined.

The current commercial syngas-to-olefin process is a two-stage process: methanol synthesis and MTO. There are existing commercial technologies for each of these processes, and integration of the two can lower the energy cost of the overall process.147–149  However, technical issues must be addressed in order to achieve integration. In principle, in an integrated process, the MTO process should drive the preceding process, according to eqn (1.11):

Equation 1.11

The current commercial MTO process operates at ≥400 °C and uses SAPO-34 and SSZ-13 zeolites; the current methanol synthesis process operates at 250–280 °C.150,151  A combination of low-temperature methanol synthesis catalysts (CuO–ZnO–Al2O3, CuO–ZnO–ZrO2, CuO–ZnO–Cr2O3, Pd/ZnO) and the MTO catalyst SAPO-34 gave methane and paraffins, as the catalyst is too hydrogenating at the MTO temperature of 400 °C. However, excluding the hydrogenating metal (Cu) from the catalyst formulation led to selective direct syngas to olefins. Using such catalyst recipes that are without hydrogenation metal were reported about the same time by Wang et al. (ZnZrOx/SAPO-34)151  and Bao et al. (ZrCrOx/SAPO-34).152  Bao et al. reportedly obtained 80% C2=–C4= and 94% C2–C4 selectivity at a CO conversion of 17%. Wang et al. reported 74% C2–C4 olefin selectivity with 11% CO conversion at 673 K. In the two catalysts, the CO activation site is attributed to partial reduction of the mixed metal oxides. It is interesting to note that a similar CO activation site was invoked in old and familiar syngas-to-i-C4 hydrocarbon conversion (isosynthesis) over sulphated or doped ZrO2 catalysts.153,154 

In most isosynthesis studies, zinc is not investigated, and the reported isobutene selectivity is very low. But Wu et al. recently attributed Zn doping to isobutene selectivity in a Zn1Zr300Oz catalyst.155  This shows the critical role of Zn to olefin selectivity in the catalysts. In a follow-up report, Wang and co-workers replaced SAPO-34 with SSZ-13, with the resulting Zn–ZrO2/SSZ-13 catalyst showing 87% and 77% C2–C4 olefin selectivity at 10% and 29% CO conversion, respectively, at 400 °C.156  Olefin formation on a Zn–ZrO2/zeolite catalyst has been attributed to the methanol-DME intermediate. CO and H2 activation to methanol or DME takes place on the Zn–ZrO2, while the C–C bond formation takes place on the zeolite acid site via the CO carbonylation and hydrocarbon-pool mechanism. Here again, the proximity of the active sites and the strength of the acid sites is crucial to olefin selectivity.115 

In another report, Bao and co-workers expanded the metal oxide-zeolite catalyst design, by replacing ZnCrOx with MnO, and investigated MnO2/SAPO-34 for syngas-to-olefin conversion.157  An optimal performance of 10.1% CO conversion with a C2 − C4 olefin selectivity of 78.9% was reported at an oxide/zeolite ratio of 2/1. Recently, Su et al. replaced SAPO-34 with similar-pore-sized but less acidic, low-Si AlPO-18. The resulting bifunctional catalyst, ZnCrOx/AlPO-18, displayed C2–C4 olefin selectivity of 45.0% (86.7% conversion, CO2 free) and a high olefin/paraffin ratio of 29.9 at CO conversion of 25.2% at 4.0 MPa and 390 °C.158  In a demonstration of the effect of zeolite pore topography on C–C coupling and olefin selectivity, Bao and co-workers reported 73% ethylene selectivity at 26% CO conversion, using ZnCrOx-mordenite (MOR).157  This is more than double the maximum ethylene selectivity that is possible based on the ASF product distribution law.

Using highly sensitive synchrotron-based vacuum ultraviolet photoionisation mass spectrometry (SVUV-PIMS), Bao and co-workers detected ketene (CH2CO) species in the gas phase after the interaction of syngas with ZnCrOx.157  They explained that the ketene was converted to C2H4 and C3H6 over the SAPO-34. The hydrocarbon pool mechanism was invoked for ketene-to-olefins transformation, but no explanation was provided for the syngas-to-ketene mechanism. With respect to the MnO2/SAPO-34 catalyst, it was stated that two CO molecules dissociate over adjacent oxygen vacancy sites that are generated by partial reduction of MnOx forming a surface C and carbonate species.157  The carbonate is discharged as CO2, thereby conserving H2 consumption for the removal of surface oxygen for syngas H2/CO ratio <2. Overall, separation of the activation or monomer formation from CO and C–C coupling sites gives greater control of olefin selectivity.

This route is known as Fischer–Tropsch-to-olefins (FTO). Among the FTS active metals (Ru, Co, Fe and Ni), Fe-based FT catalysts have the least hydrogenation activity and display selectivity to α olefins.159,160  In commercial FTS, Co-catalysts (Co or Fe based) contain noble metal reduction promoters, such as Pt, Ru, Pd, Ir, Cu, etc., and these metals possess high hydrogenating activity, promote reduction and enhance the hydrogenating activity of the FTS active metal via a hydrogen spillover effect. Other common components of commercial FTS catalysts are oxides of metals such as Zn, Mn, Zr, Ti, Cr, Sn, Mg, Ca, K and Na. These oxides are textural promoters that increase the dispersion of active metals and sometimes exact an electronic effect through strong metal-support interaction (SMSI).161–163  An alkali metal (especially K) is usually added to modulate hydrogenation activity of active FTS metal. Overall, commercial FTS catalysts are designed to ensure high hydrogenation activity to maximise C5+ selectivity and deliver a high α-value.

Contrarily, FTO catalysts should have a low α-value and low hydrogenating activity.164  Hence, the first step in tuning a FTS catalyst to a FTO catalyst is to remove the reduction promoter, and it has already been shown that CO activation can take place on ZrO2, ZnO, CrOx and MnOx. The second step is to enhance the status of the metal oxide from a mere textural promoter to a co-partner with the FTS active metal. Third, the alkali is retained to modulate the hydrogenation activity. These three steps can explain the documented higher olefin selectivity displayed by bulk and supported FeMnK catalysts.165,166  Similar olefin selectivity was reported on FeK promoted with the reducible metal oxides: V, Nb, Pb, Sn and Bi.167–170  Combined (Na, K) and S promotion of Fe-based FT catalysts has also been reported as displaying high selectivity to lower olefins.171–174 

Recently, an FT-zeolite core–shell catalyst concept was implemented as the FTO catalyst. Compared to a simple mixture catalyst (Fe-FT/SAPO-34) and a bare Fe–FT catalyst, the new Fe–FT core–zeolite shell Fe@SAPO-34 catalyst displayed a higher lower olefin selectivity of 52.6% and an O/P ratio of 6.4.175  The nonzeolite core–shell catalyst design has also been explored for the FTO route to olefins.176  The Fe3O4@MnO2 core–shell is reported to show an exceptionally high lower olefin yield and a high lower olefin selectivity at high CO conversion.176  The catalyst exhibited 79.60% total alkene selectivity and 64.95% for C4+ alkene selectivity at 75% CO conversion. The performance of the catalyst was attributed to electron transfer from the MnO2 shell to the Fe3O4 core, which promotes dissociative adsorption of CO molecules on Fe3O4. Additionally, reciprocal H atom spillover from Fe3O4 onto MnO2 enhances C–C coupling, weakens the hydrogenation activity of the catalyst and improves olefin selectivity. Co-based FTS catalysts are generally known for their high wax and paraffin selectivity and have also been tamed as FTO catalysis. Supported CoMnOx displayed a high lower olefins selectivity, and the selectivity to lower olefins increased when promoted with alkali (Na, K) ion.177  The lower olefin selectivity of Co-based catalysts has been attributed to the Co2C phase.178,179  The widely agreed FT activity of the Co-based catalyst is in the Co0 phase. However, in CoMnOx mixed oxide, in addition to the Co0 phase, the presence of MnOx promotes Co2C formation during catalyst activation.180,181 

BTX are bulk aromatic chemicals used as the primary feedstock in the production of a wide variety of intermediates, fine chemicals and materials such as phenolic resins, Kevlar, polystyrene, Bakelite etc. Current industrial production of BTX involves catalytic reforming of naphtha and, more recently, using shale gas via cracking process. With a projected growing demand for BTX, quest to secure the supply of BTX via direct syngas-to-aromatic conversion is attracting increasing research attention across the globe.138,182  Aromatics are not primary FTS products, while little aromatics are often obtained in the Fe-HTFT condition.183  They are formed from aromatisation of lower olefins. Aromatisation of methanol/DME has also been investigated extensively. Thus lower olefins and methanol can be used as intermediates to aromatics in a direct syngas-to-aromatics via a tandem one-pot conversion process.115  The tandem catalytic processes for syngas-to-aromatics via olefin and methanol/DME intermediates are briefly explained in the next sections.

Catalyst designs for direct syngas-to-olefin processes eliminate the presence of strong acid sites because they promote olefin-to-aromatics conversion, which decreases olefin selectivity. But in direct syngas-to-aromatics processes, a deliberate effort is made to incorporate strong acid sites into a FTO catalyst to engender a direct syngas-to-aromatics process.

H-ZSM-5 is a prominent aromatisation zeolite catalyst. H-ZSM-5 and its cation- (Zn-, Ga-, Mo-, La- etc.) exchanged forms have been extensively studied for aromatisation of lower olefins and other substrates, such as methanol, methane and lower alkanes.184–187  Any of the FTO catalysts indicated in Section 1.5.2.2 can be integrated with H-ZSM-5 or its cation exchanged form and qualifies as a potential combination for tandem syngas-to-aromatics.188–191  As discussed, the spatial arrangement of the active sites is critical to the performance of the catalyst combinations in a tandem syngas conversion process. The single-bed catalyst mixture is the simplest arrangement, and core–shell design can also be adopted. However, since aromatisation catalysts are easily prone to deactivation due to coke formation, a two-catalyst bed arrangement is recommended because it allows for flexible adjustment of reaction conditions and independent regeneration and recycling of the catalyst beds.

The methanol/DME intermediate is another plausible route for direct formation of aromatics from syngas. Methanol synthesis is an established technology, and direct syngas-to-DME was discussed in Section 1.5.1. Methanol and DME aromatisations over H-ZSM-5 and related cation-exchanged forms are well documented in the literature.192,193  Therefore, integration of Cu-based methanol/DME synthesis catalysts and H-ZSM-5 should make the direct formation of aromatics from syngas possible. However, this integration faces a temperature mismatch. Aromatisation over H-ZSM-5 usually takes place at a temperature of ≥400 °C, while methanol/DME synthesis over a Cu–Zn–Al oxide catalyst is typically carried out at 250–280 °C. The use of the two-catalyst-beds arrangement, which allows for operating the catalyst beds at different temperatures, may help to circumvent the temperature mismatch. But in principle, operating two catalysts at different temperatures is easy to achieve at the laboratory scale but proves very challenging with commercial-scale operations. Thus an isothermal catalyst bed reactor is preferable for large-scale operations.

As discussed in Section 1.5.1, one of the syngas-to-olefins explained involved a methanol/DME intermediate. The reaction is a high-temperature methanol formation process; hence it is more suited than the Cu–Zn–Al-oxide catalyst for tandem conversion with H-ZSM-5. Therefore, replacing the smaller-pore-size and weaker-acidic SAPO-34 with H-ZSM-5 tunes the product selectivity from olefins to aromatics. Thus in a one-pot syngas conversion process, aromatics selectivity of 80% at a CO conversion of 20% was reportedly achieved using a ZnO–ZrO2/H-ZSM-5 catalyst.139  The catalyst was integrated by simple dispersion of Zn-doped ZrO2 nanoparticles dispersed on zeolite H-ZSM-5, and no catalyst deactivation was observed in 1000 h of operation. As previously suggested, methanol and DME are major intermediates with Zn-doped ZrO2 and are subsequently transformed on H-ZSM-5 into aromatics via olefins.194–196  The easy propensity for coke formation has been a long-standing challenge with aromatisation over H-ZSM-5. The high selectivity and stability obtained with ZnO–ZrO2/H-ZSM-5 are attributed to the self-promotion mechanism of CO in the selective formation of aromatics. It was explained that CO facilitates the removal of hydrogen species formed on H-ZSM-5 in the process of dehydrogenative aromatisation of olefins. The overall reaction is an elegant tandem process and a new frontier in syngas conversion chemistry.

Attention is increasingly being given to the use of biomass for the production of fuel and chemicals. For purposes of the security of liquid fuel supply and because of environmental concerns, it is becoming increasingly imperative to develop technologies for transforming biomass into liquid fuel and chemicals. Due to its similarities with coal, established coal liquefaction processes are being adopted and modified into biomass liquefaction processes. These can be classified into two broad categories: direct and IBL processes. Generally, biomass is a composite of three large molecules: cellulose, hemicellulose and lignin. The liquefaction process basically involves transforming the biomass composite into small hydrocarbon molecules that can be used in engines (petrol, jet and diesel). The direct liquefaction process is a top-down approach to transforming biomass into liquid fuel that involves biomass pyrolysis and refining/upgrading processes. It is inflexible in terms of obtainable products. Contrarily, IBL can be considered a bottom-up approach to fuel production from biomass. It involves biomass gasification (syngas production) and FTS (syncrude production), followed by refining (syncrude upgrade to fuels) FTS leads to a wide but predictable product spectrum, it makes IBL more versatile and flexible. Both approaches have challenges in terms of low carbon utilisation.

To address this challenge, renewable energy inputs (electricity, hydrogen and solar thermal) are being advocated and encouraged. To this end, some complementary variant biomass liquefaction processes have been proposed, including GBTL, PBTL and solar-biomass-to-liquid (SBTL). These inputs complement biomass deficiencies and enhance carbon efficiency, thereby reducing CO2 emissions from the liquefaction process. Another approach to addressing the challenge of low-carbon-efficiency biomass liquefaction processes is the integration of FTS with syncrude refining processes, i.e. the development of direct syngas conversions to a target fuel or chemical.

This task goes to the heart of FTS and requires a thorough understanding of the chemistry involved in the process. Conventional FTS is a polymerisation-type reaction, with the statistical product distribution generally described by ASF distribution law. All the reaction steps (monomer formation/initial chain propagation and chain termination) take place on the same active site. In order to control product formation, efforts have been made to truncate the chain growth step or to delineate the reaction steps onto separate active sites. This often involves replacing the conventional one-catalyst FTS with two- or more-catalyst FTS. The integration of the multicatalyst system for syngas conversion is not without some challenges. The tasks constitute emerging frontiers in syngas conversion chemistry.

ASF

Anderson–Schultz–Flory

BTL

Biomass-to-liquid

BTX

Benzene, toluene and xylene

CFB

Circulating fluidised bed

DBL

Direct biomass liquefaction

DCL

Direct coal liquefaction

DFB

Dual fluidised bed

DME

Dimethyl ether (DME)

ECN

Energy Research Centre of the Netherlands

ER

Equivalent ratio

FFB

Fixed fluidised bed

FTO

Fischer–Tropsch to olefins

FTS

Fischer–Tropsch synthesis

GBTL

Gas-biomass-to-liquid

GTL

Gas-to-liquid

HTFT

High-temperature Fischer–Tropsch

IBL

Indirect biomass liquefaction

ICL

Indirect coal liquefaction

LTFT

Low-temperature Fischer–Tropsch

MTFB

Multitubular fixed bed

MTO

Methanol-to-olefin

NG

Natural gas

OLGA

Dutch acronym for the oil-based gas washing process

PBTL

Power-biomass-to-liquid

PEM

Polymer electrolyte membrane

PSA

Pressure swing adsorption

RFS

Renewable fuel standard

rWGS

Reverse water-gas-shift

SAS

Sasol advanced synthol

SBR

Steam-to-biomass ratio

SMSI

Strong metal-support interaction

STBTL

Solar-thermal-biomass-to-liquid

SVUV-PIMS

Synchrotron-based vacuum ultraviolet photoionisation mass spectrometry

WGS

Water-gas-shift

XTL

Feed-to-liquid

S. Maity is thankful to the director of his Institute for allowing the book chapter to be published. O. O. James thanks KWASU management for the conducive work environment that enabled this contribution.

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