1. Introduction

The slurry bed reactors are gas-liquid-solid three-phase reactors, which are commonly used in many industrial processes [1, 2]. The research of slurry bed reactors can be traced back to the Second World War, where the most studied and the earliest application process is the low-temperature Fischer-Tropsch synthesis (FTS) to oil. The catalytic reactions in slurry bed reactors are highly complicated due to the presence of multi-phase mass transfer, intraparticle diffusion, adsorption, surface reaction and desorption of products [3, 4]. The difficulty in separating the catalyst and handling of the slurry limits the application of slurry reactors in continuous processes. However, the slurry bed reactors possess many advantages, such as the low investment cost and simple construction, high heat exchange efficiency, online catalyst addition and withdrawal [5]. Moreover, the catalyst consumption per unit product is only 30% of the fixed bed reactor, and the total amount of required catalyst for the slurry bed reactor can be greatly reduced, which make it very suitable for gas-to-liquid processes [6].

In this review, the difference in three types of slurry bed reactors is compared, including the mechanically stirred slurry reactor, bubble column slurry reactor and three-phase fluidized bed reactor. The second part shows the application of slurry bed reactors for C₁ chemistry, discussing the four typical exothermic reactions for syngas conversion to different valuable chemicals. The third part introduces the comparison between slurry bed, fixed bed, and fluidized bed reactors. The fourth part discusses the catalyst deactivation for syngas conversion in slurry bed reactors, which mainly includes the catalysts sintering, carbon deposition and effects of water on catalyst structure. While, the last part demonstrates the development trend of slurry bed reactors for carbon one chemistry.

2. Classification and comparison of slurry bed reactors

According to the way of catalyst particles suspending in the slurry bed, the slurry bed reactors can be divided into three categories, i.e., mechanically agitated slurry reactor, bubble column slurry reactor, and three-phase fluidized bed reactor. Figure 1 shows these three types of reactors [7, 8].

Figure 1.
Three different types of slurry bed reactors [7].

2.1 Mechanically agitated slurry reactor

The mechanically agitated slurry reactors are commonly used in many chemical processes, such as the hydrogenation of unsaturated oils and nitro compounds. It has the advantages of high mass and heat transfer efficiency, which are best suited for laboratory kinetic studies. And it is because the catalyst particles are small in the reactor, the diffusion resistance in the particles is very low, and the utilization rate of the catalyst is very high. The disadvantages are the high power required for stirring, the presence of significant liquid back-mixing, and the difficulty of catalyst separation in continuous operation. The commonly used agitator in the reactor is a turbine-type radial flow agitator, and there is an optimal ratio of paddle to vessel diameter. To keep the catalyst particles suspended evenly, the number of the baffles should be 4 in most cases and the size is 1/10 ~ 1/12 of the container diameter. Moreover, the use of baffles could provide a proper gas distribution and avoid the formation of vortex on the surface of the liquid. Generally, the gas inlet should be as close to the paddle as possible, using an annular sparger rather than a single nozzle. The width of the sparger should be about 0.8 times of the paddle diameter. Figure 2 shows the slurry bed reactor with multiple paddles, which is commonly used for the hydrogenation of unsaturated oils [7]. To reduce the back mixing in continuous operation, the horizontal inlets are provided at various locations in the reactor.

Figure 2.
Slurry bed reactor with multiple paddles [7]. 1. Paddle, 2. baffle, 3. cooling or heating coils, 4. annular pipes for supplying H₂.

2.2 Bubble column slurry reactor

In a bubble column slurry reactor, the dispersion of the feed gas takes place through a deep pool of inert liquid in which the catalyst particles are suspended, and the momentum is transferred to the inert liquid by the movement of the bubbles. The operation is usually implemented in the columns with a height-to-diameter ratio of 4 to 10. For the conversion of liquid reactants, the operation can be semi-batch or continuously. The advantages of this type of reactor are that there are no moving parts, no need to seal the stirrer, less floor space and low power consumption compared to stirred reactors. Its main disadvantage is that the liquid phase has considerable back-mixing, which needs to overcome a large pressure drop, and when the height-diameter ratio is higher than 10, the gas-liquid interface area will be decreased quite rapidly. The bubble column slurry reactors have been used in many industrial processes, such as FTS to hydrocarbons, and the production of caprolactam, etc. Figure 3(a) shows a continuous bubble column slurry reactor with solid-liquid separation [7]. In another type of bubble column slurry reactor, the catalyst and liquid are circulated by using external pumps, and the catalyst can be operated at higher loads, either counter-current or co-current. The schematic diagram is shown in Figure 3(b).

Figure 3.
(a) Schematic diagram of continuous bubble column slurry reactor, (b) schematic diagram of a bubble column slurry reactor circulating with a pump [7].

2.3 Three-phase fluidized bed reactor

In a three-phase fluidized bed reactor, the catalyst particles are primarily fluidized by the liquid, while the gas stream flows co-currently in the form of intermittent bubbles. An important difference between the three-phase fluidized bed reactor and the bubble column slurry reactor is that the former transfers momentum through the movement of liquid, while the latter transfers momentum through the movement of the bubbles. Moreover, the relatively large catalyst particles can be used in three-phase fluidized bed reactors, which makes it suitable for continuous operation due to easy separation of catalyst particles from the liquid. The main industrial application of three-phase fluidized reactor is the hydrogenation of petroleum, including the hydrodesulfurization, hydrocracking to produce olefins, and the partial oxidation to produce aromatics. The three-phase fluidized bed reactors are generally operated in co-current flow, but in some cases, they can also be operated in counter-current flow (the liquid flows down while the gas flows up), such as when the density of catalyst particles (ion exchange resin) is lower than that of the liquid. In the industry, the particles are kept fluidized with internal circulation of the liquid, which can be accomplished by using a pump. This makes it possible to fluidize large catalyst particles. If the internal circulation is not used, the single-pass conversion rate will be too low due to the excessive liquid flow rate used.

2.4 Comparison of three types of slurry reactors

The catalyst particles in the mechanically stirred slurry reactor are suspended by the mechanical stirring of paddles, which shows high heat transfer and mass transfer efficiency. However, the power required for mechanical stirring is the largest, and the sealing problem of the stirrer needs to be considered. In addition, the wear of the catalyst during the mechanical stirring process is probably the most significant.
The catalyst particles in bubble column slurry reactor are suspended by the gas bubbling; however, the relatively small particles result in the difficulty in catalyst separation.
The catalyst suspension in three-phase fluidized bed reactor is the result of combination of the bubble movement and the co-current liquid flow. The main difference between a fluidized bed and a bubbling bed is that the suspension of the particles for the former is the result of liquid flow, thus, the relatively large catalyst particles can be used, which benefits the catalyst separation. Due to the absence of moving parts, the design of bubble columns and three-phase fluidized bed reactors is simpler than that of mechanically stirred reactors.

3. Slurry bed reactors for carbon one chemistry

3.1 Fischer-Tropsch synthesis

FTS has been practiced on a large scale by Sasol in South Africa since the mid-1950s in tubular fixed-bed reactors and circulating fluidized-bed reactors using iron catalysts. The commercial scale conventional bubbling fluidized-bed reactor and slurry bubble column reactor (SBCR) were placed on stream by Sasol in 1989 and 1993, respectively. At present, the multi-tubular fixed bed reactor, fluidized bed reactor, gas-liquid-solid slurry reactor, and microchannel reactor are used for FTS. The reactor type and the operating conditions are the governing factors in controlling the product distribution with chain growth probability, catalyst activity, and product selectivity during FTS [9, 10]. The multi-tubular fixed bed reactor and slurry reactor are used for the low-temperature process, while the circulating fluidized bed reactor, fixed fluidized bed rector, and microchannel reactor are used for the high-temperature process. The high-temperature process yields large amounts of olefins, a lower boiling range, and very good gasoline while producing the substantial amounts of oxygenates. The low temperature process yields much more paraffin and linear products and can be adjusted to very high wax selectivity. The primary diesel cut and wax cracking products can give excellent diesel fuels. The very linear primary gasoline fraction needs further treatment to attain a good octane number. Olefin and oxygenate levels for the low-temperature process are lower than those for the high-temperature process.

3.1.1 Fixed bed reactor

The fixed bed reactor for FTS, such as the Arge reactor at Sasol with 12 m height and 0.5 m diameter, consists of a shell containing 2050 tubes packed with Fe-based catalyst that produced the 600–900 bbl/day/reactor [11]. Where, the heat generated in highly exothermic synthesis reaction of 220°C reaction temperature, and 25–45 bar pressure was used for the generation of steam on the shell side of the reactor. Similarly, some other reactors using Co-based catalyst were also introduced at Bintulu Malaysia, Las Raffan, Qatar, and the Pearl GtL facility which resulted the rate of 3000–140,000 bbl/day with 85–95% C₅⁺ selectivity [12]. Despite the robust nature and high productivity, the fixed-bed multi tubular reactors contain the disadvantages of design complexity, high cost, low catalyst utilization, high pressure drops, and insufficient heat removal due to poor heat conductivity [13].

3.1.2 Fluidized bed reactor

Fluidized-bed reactors being deployed for the high temperature FT process of entire gas phase, have been designed to comparatively improve the efficiency of fixed-bed reactors in terms of superior heat transfer and temperature control during highly exothermic FT reactions; avoiding intraparticle diffusion, pressure drop and the reaction rate and the better mixing of catalyst particles while giving high production capacity [14]. However, it finds some limitations of difficulty in scale-up for causing the agglomeration and blockage of the fluidization and containing a high risk of attrition and heavy product deposition on catalyst, while requiring the special equipment (cyclones) for catalyst separation. Since, the technical and economic problems of the industrial scale reactor, discouraged their application in 1956 by Hydrocol and later by SASOL in 1980s [15].

3.1.3 Microchannel reactor

The recent advancement in terms of a new type of reactor—the microchannel reactor, which consist of a large number of parallel channels with diameters <1 mm and with the catalyst on a thin layer inside the channel walls. It has been demonstrated in different studies that the large temperature gradients in the furnace-heated conventional fixed-bed reactor can be avoided in the microchannel reactor under the same operating conditions, achieving the best catalyst utilization and thus a high productivity for the large transfer surface area with high heat transfer coefficient between the bed and wall [9]. However, the main challenges of the difficulty in changing the catalyst and high cost of scaling up several microchannel reactors to be operated in parallel, still avoided its commercial FT plant.

3.1.4 Slurry reactor

To replace the fixed bed reactor for the low-temperature FT process, the Sasol slurry reactor was developed. In 1993, a single slurry reactor with 5 m in diameter and 22 m in height realized the industrial production for FTS. Compared with the other reactors, the slurry reactors for FTS are less expensive to construct, maintain and operate [16]. It was reported that the capital cost required for a large-scale slurry reactor was less than 40% of that needed for an equivalent multi-tubular fixed bed reactor: The catalyst usage of the slurry reactor is about a third of that of the fixed bed reactor with a promise of even better performance, due to the catalyst’s effectiveness and the higher average temperature used in the slurry reactor. Krishna and Sie compared the several reactor types for the FTS process and concluded that the slurry reactor was the best reactor type for large-scale plants [17].

The adopted catalysts for FTS in slurry bed reactors are commonly the iron-based or cobalt-based catalysts [18]. Parts of FTS using slurry bed reactors are listed in Table 1. The most common FTS catalysts have been usually regarded as Fe-based and Co-based catalysts, in which Fe-based catalyst have been synthesized by precipitation method [19]. While, the FTS process being carried out by cobalt-based catalysts were usually suspended in inert liquid [13], by impregnation method using Al₂O₃, SiO₂ or TiO₂ as support, with the Co loading of 10–30% (wt.). Sasol has developed cobalt-based catalyst SAC 2-100SB and applied in slurry bed reactor for FTS, providing this Co catalyst to two sets of natural gas-based synthetic oil production plants in Nigeria and Qatar. The Co catalyst show stable catalytic activity; however, it is only used for low-temperature FTS due to the high price of cobalt.

Company	Location	Capacity/(kt⋅a⁻¹)	Catalyst	Reactor
Institute of Coal Chemistry, Chinese Academy of Sciences/Synfuels China	Taiyuan, Shanxi	1.0–1.5	LTFT Fe	Slurry
			HTFT Fe
	Ordos, Inner Mongolia	160	HTFT Fe	Slurry
	Luan, Shanxi	1600	HTFT Fe	Slurry
	Luan, Shanxi	7.5	LTFT Co	Fixed bed
YangKuang group	Thengzhou, Shandong	5	LTFT Fe	Slurry
	Thengzhou, Shandong	5	HTFT Fe	Fluidized bed
SINOPEC group	Zhenhai, Zhejiang	3	LTFT Co	Fixed bed
Shenhua group	Shanghai	0.08	LTFT Fe	Slurry
	Ordos, Inner Mongolia	180	LTFT Fe	Slurry
Rentech	US	0.5	LTFT Fe	Slurry
Syntroleum	Tulsa, US	3.5	LTFT Co	Slurry
Exxon-Mobil	Baton Rouge, US	10	LTFT Co	Slurry
Conoco-Phillips	Ponca, US	20	LTFT Co	Slurry
BP	Nikiski, US	15	LTFT Co	Fixed bed
Statoil/PetroSA	Mossel Bay, South Africa	50	LTFT Co	Slurry

Table 1.

Installations for FTS with slurry reactors [18].

LTFT: low temperature Fischer-Tropsch; HTFT: high temperature Fischer-Tropsch.

The pure iron catalysts are easily worn and deactivated in FTS, which needs to enhance the catalytic activity, selectivity, and stability by introduction of additives (electronic additives and structural additives). Fe/Cu/K catalyst is one of the most successful industrial catalysts for low-temperature FTS (200–250°C) [20]. The precipitated iron-based catalyst prepared by Sasol in South Africa exhibits high activity, selectivity, and stability, and it has been industrialized. Subsequently, Synfuels China Technology Co., Ltd. proposed a high-temperature FTS in a slurry bed reactor (260–290°C), which greatly improve the steam pressure generated by the FTS reaction and significantly enhance the overall energy utilization efficiency of the FTS process. And based on the development of industrial iron-based catalyst for high-temperature FTS reaction, a complete set for high-temperature FTS process and product processing technologies were developed.

3.2 Methanol synthesis

The syngas conversion to methanol is a strongly exothermic reaction (CO + 2H₂ = CH₃OH, ΔH298KΘ = −90.37 kJ/mol). Inspired by the FTS in a slurry bed reactor, a liquid-phase methanol synthesis was first proposed in 1975. Compared with the gas-solid phase fixed-bed reactor, the gas-liquid-solid three-phase methanol synthesis selected long-chain hydrocarbons with high heat capacity and thermal conductivity liquid inert medium to remove the reaction heat, which makes the catalyst bed operated at a uniform temperature and easy to control. The slurry methanol synthesis can use the feed gas with high CO concentration to improve the single-pass conversion of CO, and the outlet methanol concentration is as high as 15–20%; moreover, there will be no local overheating and excessive temperature rise in catalyst bed [21].

In early studies, the catalysts applied in slurry methanol synthesis were mainly commercial fixed-bed CuZnAl catalysts [22]. In recent years, the slurry methanol catalysts have been intensively studied [23]. It is found that that the precipitation sequence of CuZnAl catalyst affected the catalytic performance. The catalyst prepared by precipitation of Al first and then co-precipitation of Cu and Al exhibit the highest activity and stability; the precipitation and aging temperature can not only affect the phase composition of CuZnAl catalyst precursor, but also affect its crystallinity. In addition, the introduction of microwave radiation heating can make the copper-based catalyst particles size smaller and more uniform, and the strengthened synergistic effect between copper and zinc is beneficial to improve the catalytic activity and stability.

3.3 DME synthesis

In 1991, a 10 t/d pilot-scale plant for synthesizing DME from syngas in a slurry bed was developed by Air Products and Chemicals, Inc., and Cu-based methanol synthesis catalysts and methanol dehydration catalysts γ-Al₂O₃ were adopted. Taiyuan University of Technology, Tsinghua University and Shanxi Coal Chemical Institute of Chinese Academy of Sciences have also done leading research work on slurry bed for DME synthesis. The catalyst for the one-step synthesis of dimethyl ether in the synthesis gas slurry bed is similar to the one-step catalyst in the fixed bed, which includes a methanol synthesis catalyst and a methanol dehydration catalyst. The methanol synthesis catalyst generally adopts CuZnAl catalyst or modified CuZnAl catalyst, and the catalyst for methanol dehydration generally adopts γ-Al₂O₃, zeolite, etc. Table 2 summarizes the progress of DME synthesis in a slurry-bed reactor [24]. It can be seen that JFE (formally NKK) Corporation produced DME while Air Products Chemicals produced DME and methanol.

	JFE	Air Products	Tsinghua University
H₂/CO ratio	1	0.7	1
Catalysts		CuZnAl+γ-Al₂O₃	LP201 + TH16
Operating pressure, MPa	5	5–10	4.35–4.6
Temperature, °C	260	250–280	255–265
CO conversion, %	40	22	54–63
Selectivity, %	90	40–90	89–95
Reactor type	slurry bubble column	slurry bubble column	slurry airlift reactor
Reactor diameter, m	15	15.2	21.6
Reactor height, m	0.55	0.475	0.6
Design scale, t/d	5	10	10

Table 2.

Comparison of industrialization Technology of Dimethyl Ether Production in slurry bed [24].

3.4 Synthetic natural gas synthesis

Slurry bed reactor is suitable for the strongly exothermic methanation reactions (3H₂ + CO = CH₄ + H₂O, ΔH298KΘ= − 206 kJ/mol) [25]. The Chemical Systems Inc. (United States) developed a liquid-phase methanation process [26], as shown in Figure 4. The syngas produced in a coal gasifier was introduced into the catalytic liquid phase methanation reactor (LPM) along with a circulating process liquid (mineral oil), which absorbs the heat of reaction. The product gas was separated in a liquid phase separator and in a product gas separator. The process liquid was pumped through a filter to remove any catalyst fines and recirculated back to the LPM reactor. After a methanation reaction time for more than 300 h, the results showed low conversion and high catalyst loss, and thus the LPM-project was terminated in November 1981.

Figure 4.
Liquid phase methanation concept, adapted from [26].

Taiyuan University of Technology and SEDIN Engineering Co., Ltd. have jointly developed a slurry-bed process for syngas methanation [27], as shown in Figure 5. The generated methanation gas together with the catalyst and the liquid-phase components enters through a gas-liquid separator, where the gas-phase product was condensed to produce synthetic natural gas, the liquid-phase product is mixed with fresh catalyst in the storage tank and added into the slurry bed methanation reactor to preheat the fresh catalyst. Due to the excellent heat transfer performance of the slurry bed, the adaptability of the feed gas for methanation of slurry bed is stronger, and the content of CO in feed gas can be adjusted within a wide range of 2–30%. The results show that the CO conversion for methanation reached 96% at the reaction temperature of 280°C in a slurry bed reactor.

Figure 5.
Flow diagram of slurry-bed methanation for coal to synthetic natural gas. 1. Methanation reactor, 2. heat exchanger, 3. gas-liquid separator I, 4. gas-liquid separator II, 5. fresh catalyst storage tank, 6. heat exchanger II, 7. circulation pump, p1-p9 pipeline [27].

For industrial fixed-bed methanation technology, the commercial catalysts are Ni-based catalyst, which is due to its cheap and excellent catalytic activity at high temperatures. However, the traditional industrial Ni-based catalysts exhibit poor catalytic performance in slurry bed reactors [28]. By optimizing the preparation method and conditions [29], adding additives, a slurry bed methanation catalyst with a high CO or CO₂ conversion of more than 99.5% can be obtained.

4. Comparison of fixed bed, circulating fluidized bed and slurry bed reactors for FTS

A schematic diagram for a common fixed bed reactor is shown in Figure 6(a). A gas sparger at the reactor inlet is used to remove the initial kinetic energy of the gas stream. A certain gas space is left between the catalyst bed and the gas sparger for fluid buffering and uniform mixing. The packing (e.g., ceramic balls) spreads on the upper part of the bed can further make the stream enter the catalyst bed in a more uniform state. The small and uniform particle size of catalyst makes the resistance of each part of catalyst bed the same to enhance the reaction efficiency. After reacting in the catalyst channels, the unreacted syngas and the formed products leave the catalyst bed and reactor and release a large amount of heat. There is 1 ~ 2 thermocouple inserted in the reactor to monitor the temperature parameters of each section of the catalyst bed in time.

Figure 6.
Reactor types used for FTS. (a) Fixed bed reactor. (b) Circulating fluidized bed reactor. (c) Fixed fluidized bed reactor (d) slurry bubble column reactor [2].

Fluidized bed reactors possess high heat transfer efficiency. Generally, the fluidized bed reactors are classified to circulating fluidized bed reactor and fixed fluidized bed reactor, as shown in Figure 6(b and c). The feed gas is used as the power source to drive the catalyst, which is suitable for the strong exothermic reactions by using the intermediate heat exchange device to remove the reaction heat in time. However, the device structure of circulating fluidized bed reactor is complex, the investment and maintenance cost are high; Moreover, the reactor is difficult to operate and enlarge. In view of the limitations and defects of circulating fluidized bed reactor, a fixed fluidized bed reactor was designed by Sasol Corporation in 1995 and named as Sasol Advanced Synthnol reactor. Since the diameter of the fixed fluidized bed reactor can be much larger than that of circulating fluidized bed reactor, the space for installing the cooling plate is increased by more than 50%, which benefits to improve the CO conversion. However, the entrainment and attrition of the catalyst during the fluidization process is still very serious, which makes the fluidized bed methanation process is not practical for the expensive catalysts.

Since the heat transfer in the slurry bed reactor is more efficient than that in the fixed bed reactor, the temperature control is relatively easy. Therefore, the temperature of the catalyst active site in the slurry bed reactor could be uniform, which could avoid the formation of hot spots in the catalyst bed. In a continuous operation of the slurry bed reactor, the biggest difficulty and obstacle is the separation of fine catalyst, while in a fixed bed there is no such problem. However, the replacement of the catalysts in a slurry bed reactor is easier than in a fixed bed reaction. If the catalyst needs to be removed and withdrew frequently in a process, the fixed bed reactor may not be appropriate, it is because that the replacement of the catalyst usually requires to stop and remove the reaction.

The key parameters and catalytic methanation performance of different reactors [27] are compared and summarized in Table 3. It can be seen that the heat exchange rate of fluidized bed and slurry bed reactor is significantly higher than that of fixed bed reactor, indicating that the fluidized bed or slurry bed reactor is very suitable for the strong exothermic reaction. However, the cost of fluidized bed reactor is high, the catalyst entrainment and wear are serious, and the catalyst cannot be recovered, which limits the application of large-scale production.

Characteristic	Fixed bed	Circulating fluidized bed	Fixed fluidized bed	Slurry bed
Heat exchange rate or heat dissipation	slow	medium ~ high	high	high
Heat conduction in the system	bad	good	good	good
Reactor diameter limits	~8 cm	none	none	none
Pressure drop at high gas velocity	low	medium	high	medium ~ high
Gas phase residence time distribution	narrow	narrow	wide	narrow ~ wide
Axial mixing of the gas phase	low	low	high	low ~ medium
Axial mixing of catalysts	none	low	high	low ~ medium
Catalyst weight percentage/%	55 ~ 70	1 ~ 10	30 ~ 60	<30
Particle size of solid phase/mm	1.5	0.01 ~ 0.5	0.003 ~ 1	0.1 ~ 1
Catalyst regeneration and replacement	intermittent	continuous	continuous	continuous
Catalyst loss	none	2% ~ 4%	wearred, cannot be recycled	low

Table 3.

Key parameters and performance comparison of different reactors [27].

5. The deactivation of slurry bed catalyst

Dealing with the catalyst deactivation in a slurry bed reactor is a challenging phenomenon [30]. The detailed mechanistic insights in realizing the origin of the deactivation are much more complex, for the simultaneous involvement of various factors. Where, the most common factors could be relied on the S and/or N₂-containing compounds in the feed gas; the oxidation of the active metal (Co) components; strong metal-support interaction resulting into the hard to reduce species of silicates, aluminates; sintering, attrition, and carbon deposition on the nano-crystallites particles [31]. Since, these aforementioned factors of catalyst deactivation are mainly dependent on the type of reactor to be employed, the nature of the support, and the partial pressure of H₂O that can be further explained as follows.

5.1 Sulfur and nitrogen containing poisons

In carbon one chemical reactions, the poisons that may cause severe damage to the catalyst must be removed from the feed gas prior to the reaction, as these are independent of the catalyst nature to be used and operating conditions of the reactions. Since, the sensitive nature of cobalt-based catalysts to the sulfur content (<10 ppb) needs more attention of the additional gas cleaning steps for being expensive than Fe and also for its higher activity and a longer life time. In here, the different adsorption phenomena of the sulfur containing compounds (organic, inorganic) is more important for physically blocking the sites after being strongly adsorbed on the catalytic active sites, where one adsorbed S atom could deactivate the more than two Co atoms in a Co/Al₂O₃ catalyst. In a CO-rich syngas conversion to methanol reaction, the iron carbonyl, nickel carbonyl and carbonyl sulfide are severe catalyst poisons, which must be removed from the feed gas to avoid the catalyst deactivation [32]. Since, the removal of H₂S as the probe molecule and the nitrogen compounds such as NH₃, HCN (<50 ppb) are the crucial steps in gas purification, as they could lead to the severe damaging effect on the catalyst reducibility, activity, and selectivity.

5.2 Sintering

Besides the poisoning, the catalyst could also be deactivated due to the catalyst sintering at high temperature. Sintering is a common phenomenon of deactivation in metal catalysis, that is usually based on the minimization of surface energy of the crystallites. Sintering is usually accelerated by the high temperature and the partial pressure of water vapor; however, it can be controlled by interactions with the metal support. There are two empirical rules for the effect of temperature on highly dispersed metal crystallites: [1] When the temperature reaches 0.3T_m (called Hüttig temperature, T_m is the melting point), the migration of particles on the catalyst surface will occur [2]. When the temperature reaches 0.5T_m (Tammann temperature), the particle migration within the lattice phase will occur. For the methanol synthesis catalyst CuZnAl, the melting point of metallic Cu is 1358 K, and its Hüttig temperature is 407 K, and its Tammann temperature is 679 K; while the operation temperature of CuZnAl catalyst is generally at 480–553 K. Though the operation temperature is much lower than the Tammann temperature, it is higher than the Hüttig temperature, which inevitably resulted in the catalyst sintering and thus deactivated the catalyst. Lewnard et al. [33] investigated the stability test on CuZnAl catalyst, and obtained that the activation energy for catalyst deactivation is 91.3 kJ/mol. These activation energy data all demonstrate that catalyst sintering is an important reason for the deactivation of methanol synthesis catalyst in a gas-liquid-solid three-phase slurry bed.

5.3 Carbon deposition

Generally, the slurry bed reactor needs to be operated at the temperatures lower than 350°C due to the limitation of boiling point of mineral oil. Even at this low temperature, there are still some carbon deposited on the catalyst surface. For example, the result of CO methanation in a slurry-bed reactor over the Ni-Al₂O₃ catalyst [34] shows that, the catalytic activity of Ni-Al₂O₃ catalysts decreased slowly after a reaction time of 450 h. The thermogravimetric analysis and microscopic morphology results show the carbon deposition on catalyst surface, which was attributed to the amorphous carbon. The carbon deposition occupies the surface of the catalyst and covers the active sites, resulting in a decrease in Ni metal surface area and thus reduce the methanation activity. The regeneration of the spent catalyst shows that the catalyst carbon deposit can be removed by calcination in the air, and the catalyst structure and catalytic performance can be recovered.

5.4 Effects of water on catalyst structure

The reoxidation of Co metal active sites usually occurs when the oxygen atom of CO is eliminated mainly as H₂O (either from surface oxygen or OH species) during the FTS, where the influence of water contents in terms of its higher partial pressure than that of H₂ and CO become more crucial at high CO conversion. Since, the effect of water contents on the catalyst deactivation is mainly adopted from the possible reoxidation of the surface depending on the operating conditions, presence of different promoters (Pt, Mn, Zn, Mg, etc.), nature of the support (Al₂O₃, SiO₂, TiO₂) to be used, dispersion of metal components, and on the size of the pores of the support. For the methanol synthesis from syngas in a slurry-bed reactor over Cu-based catalysts, the addition of water in the reaction shows that too much H₂O accelerates the growth of grain and agglomeration of Cu-based catalyst, and a certain degree of carbon deposition, which leads to a fast catalyst deactivation, as shown in Figure 7 [35].

Figure 7.
Effect of H₂O on CO conversion for methanol synthesis [35].

For the syngas to DME in a slurry bed reactor, the methanol dehydrated and formed DME in the reaction, if the generated water cannot be removed in time, the pores of the catalyst will be blocked and the reaction performance will be affected [36]. If the partial pressure of CO₂ in the feed gas is high, more ZnCO₃ will be appeared on the surface of CuZnAl catalyst, which is due to that the solubility of ZnCO₃ in water is several times higher than that of ZnO. Therefore, the leaching of Zn in the presence of water is one of the important reasons for the rapid deactivation of CuZnAl catalyst [37]. In addition, in the methanation reaction, the presence of water reacts with the γ-Al₂O₃ support of Ni-Al₂O₃ catalyst and thus formed AlOOH, which is one of the reasons for the deactivation of Ni-Al₂O₃ methanation catalyst.

5.5 Attrition of catalyst particles

One of the serious concerns of catalyst deactivation is the catalyst attrition by mixed particle fragmentation or surface abrasion, for breaking the particles into various fragments. It is a common phenomenon in fluidized or slurry bed reactors, where the interparticle collisions and bed-to-wall impacts cause the high mechanical stresses to catalyst bed particles, thus resulting the loss of valuable material, generation of fine particles, and the degradation of catalyst efficiency [38]. Usually, it is a time-dependent process that may systematically change with time, being more severe in the start of reaction and tends to a constant value of the nonsteady-state attrition rate with time. However, different strategies can be applied for increasing the attrition resistance and catalytic performance and decreasing the fraction of fines by spray-dried Fe-based FTS catalysts in a stirred-tank slurry reactor.

6. Current developments in slurry bed reactor for C₁ chemistry

Since, in a relative comparison it can be demonstrated that, the slurry bed reactors show the advantages of high efficiency of mass transfer and heat transfer, uniform reaction temperature, online replacement catalyst and low investment costs as compared to that of fixed-bed/fluidized-bed reactors [3, 6]. However, the industrialization process of the slurry bed reactor is relatively slow, it is because the enlargement of slurry bed reactor can only be achieved by increasing the diameter of the reactor, and the flow behavior in a large-scale industrial slurry bed is extremely complex under the working conditions of high temperature and high pressure. The blockage of the reactor distributor, uneven gas distribution led to the unstable operation state, which puts forward higher requirements for the rapid heat transfer of the slurry reactor [4, 39].

The process intensification is one of the methods to improve performance or solve problems in slurry bed reactors. The development of new high-efficiency internal components, the optimization of performance based on the existing internal components, and the coupling arrangement of multiple internal components can enhance the mixing and mass/heat transfer in slurry bed reactors [40]. Now, the research in slurry bed reactor mainly focuses on the fluid mechanics such as gas holdup, bubble size distribution, internal mass transfer, and heat transfer characteristics. Some or all the above problems can be solved by introduction of suitable internals into the reactor; however, the critical internals are difficult to obtain in published reports. Therefore, the structural optimization and function enhancement technology of internal components has become a challenging hot issue in the process of slurry bed design and scale-up [41]. Moreover, though the research of particle size on the effect of gas-liquid flow behavior in slurry bed reactors has been carried out earlier, however, these results in different reports are often inconsistent, which is due to the complexity of the gas-liquid-solid three-phase system, and in most cases, the influence mechanism of catalyst particles is based on analysis and prediction, without direct and sufficient proofs. Therefore, the research on the effect of particles is always the focus of attention and research.

7. Conclusions

Since, the comparative analysis of three different types slurry bed reactors, has been well discussed in the first part of this review while revealing the application of slurry bed reactors for C₁ chemistry, along with the four typical exothermic reactions for syngas conversion to different valuable chemicals. Similarly, the comparison between slurry bed, fixed bed, and fluidized bed reactors in second part also emphasized the catalyst deactivation for syngas conversion in slurry bed reactors, including the catalysts sintering, carbon deposition, effects of water on catalyst structure and the catalyst attrition by mixed particle fragmentation. While, the current development trend of slurry bed reactors for C₁ chemistry in the last part demonstrates the optimization of the reactor structure, higher active catalysts, and improvement of process conditions that could promote the development of slurry bed reactor. Since, the various techno economic aspects of slurry bed reactor make it very suitable for gas-to-liquid processes, especially for the highly exothermic reactions in C₁ chemistry for different valuable chemicals.

Acknowledgments

This work was financially supported by the Natural Science Foundation of Shanxi Province (202103021224073, 201801D121056). The authors would like to thank Liping Wang, Baozhen Li, Lina Wang, Zhiyuan Gong, Qian Wang, and Pengfei Ding from Taiyuan University of Technology for the help in finding and collecting the information.

Conflict of interest

The authors declare no conflict of interest.

Kwan Han Yoon

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