Friday, July 19, 2019

Indirect Coal Liquefaction - The Production of Fuels From Coal

With indirect coal liquefaction, the coal is gasified to produce a raw gas, which after treatment and purification can be used as feed gas for a wide range of syntheses. A general schematic indicating major process units of a typical indirect coal liquefaction process chain is given in Figure 17. Gasification is the thermo-chemical conversion of carbonaceous feedstock in a reductive gas atmosphere by adding an agent, which can be oxygen, air, CO2 or steam. This produces a combustible gas normally containing larger amounts of CO, H2 and smaller amounts of CO2, steam, CH4 and trace component gases. High-purity oxygen (sometimes mixed with steam) is commonly used as the gasification agent and is typically provided by cryogenic air separation. The feed coal, which in some cases needs to be dried and crushed to the grain size required by the gasification process, is conveyed into the gasifier where it reacts with the gasification agent.

The raw product gas from the gasifier is subsequently cooled and cleaned to remove dust and/or tar prior to the removal of corrosive, catalyst poisoning gas components, such as sulphur compounds, and the composition of the cleaned syngas is adjusted to meet the requirements of the downstream synthesis. Besides the major process units, there are typically several auxiliary processes, such as balance of plant, steam, sulphur recovery, CO2 and or off-gas treatment, waste water treatment, as well as synthesis product upgrading and refining.

Figure 17 Schematic of an indirect liquefaction process (modified by author)

Suitable feedstock range

In principle, there is no restriction with respect to coal quality regarding its applicability to an indirect coal liquefaction process. That said, for specific coal gasification technologies, there are constraints as to acceptable ash and moisture content, and grain size. Since coal quality can vary significantly depending on coal rank, ash content and ash properties, not every coal is applicable to every gasifier design either for technological or economic reasons. Nevertheless, there is a wide range of commercial gasification technologies to cover the whole range of coal qualities.

As well as coal gasification, there are various technologies available or under development, which are designed to use biomass or other carbonaceous fuels as feedstock. There is also the option to co-gasify biomass with coal. This approach has been industrially tested, with the amount of co-fed biomass that can be accommodated strongly depending on the impact on the ash/slag behaviour, the availability of biomass and the level of pre-treatment needed since biomass has a lower energy density than coal.

Once the raw gas is provided, there are various mature and well commercialised gas treatment technologies to upgrade the syngas to the quality required by the synthesis process.

Description of underlying process principles

Coal gasification technologies

The main process of an indirect coal liquefaction route is the gasification process that converts the coal into a raw gas, which can be used as syngas after cleaning. The most important gasification reactions are given below. Besides the indicated heterogeneous and homogeneous gasification reactions, pyrolysis reactions also occur during heating of the coal particles, yielding mainly char, higher hydrocarbons (in particular aromatic compounds), water (from drying and decomposition of oxygen containing functional groups in the coal structure), carbon dioxide and methane. Because of the complex mechanism of pyrolysis reactions, they are not included in the list of reactions below. Most of the pyrolysis products will be reactants converted into syngas compounds during later gasification reactions. Hence, they are normally not found (or only to a very little extent) in the raw gas except for fixed bed gasification systems where higher hydrocarbons are commonly found in the raw gas because of the operating conditions created by the counter-current flow scheme inside the gasifier.

In situ combustion reactions partially taking place and covering the energy demand of endothermic gasification reactions are (Krzack and Smalfeld, 2008):

C + O2 → CO2                           -406.3 kJ/kmol (1)
2 C + O2 ↔ 2 CO                       -246.7 kJ/kmol (2)
2 CO + O2 ↔ 2 CO2                  -556.9 kJ/kmol (3)
2 H2 + O2 ↔ 2 H2O                   -483.7 kJ/kmol (4)
CH4 + 2 O2 ↔ CO2 + 2 H2O    -802.3 kJ/kmol (5)

Endothermic and exothermic gasification reactions:

C + CO2 ↔ 2 CO                       +159.6 kJ/kmol (6)
C + H2O ↔ CO + H2                 +118.5 kJ/kmol (7)
C + 2 H2O ↔ CO2 + 2 H2           +77.3 kJ/kmol (8)
C + 2 H2 ↔ CH4                           -87.7 kJ/kmol (9)
CO + H2O ↔ CO2 + H2              -41.1 kJ/kmol (10)
CO + 3 H2 ↔ CH4 + H2O         -206.2 kJ/kmol (11)
2 CO + 2 H2 ↔ CH4 + CO2      -247.3 kJ/kmol (12)

According to equations 1–12, the reactions comprise exothermic and endothermic reactions. A high share of the yielded syngas components can be attributed to heterogeneous gasification reactions, most of which are endothermic and therefore require provision of heat. For the gasification of solid fuels, the typical process principle is autothermic gasification where a fraction of the fuel is internally combusted to provide the heat required to carry out the endothermic reactions. For the conversion of gaseous or light liquid hydrocarbons, for example, naphtha or natural gas, the process schemes include an external heat supply and are called allothermic processes. In contrast, for autothermic processes, high-purity oxygen sometimes mixed with steam is used as the gasification agent. Although air was used in older plants, today’s gasification plants apply high-purity oxygen (≥98 vol%) to avoid dilution of the raw gas by the nitrogen contained in the air which would be disadvantageous for downstream gas processing and the synthesis unit.

Autothermic solid fuels gasification processes can be distinguished by three principles for solid gas contacting, namely fixed or moving bed, fluidised bed or entrained-flow gasification. The major differences regarding grain size, residence time and other parameters are summarised in Table 5.

One of the most important parameters distinguishing the three process principles is the operating temperature. Whereas entrained-flow processes are operated above the ash melting temperature, fluidised and moving-bed processes require lower temperatures below the ash melting temperature (with some exceptions). Gas cooling is required because subsequent gas cleaning processes are operated at about room temperature or lower. The level of gas cooling needed is determined by the operating temperature, with very high gasifier exit temperatures above 1250°C from entrained flow systems and about 600°C for moving bed systems. The fluidised bed exit temperature ranges between 800°C and 1000°C, depending on the coal type used.

There are multiple concepts including direct cooling by injection of cold water into the hot gas (water quench), direct cooling by recirculation and injection of dedusted cooled raw gas into the hot gas (gas quench), direct cooling by secondary injection of carbonaceous fuel and taking advantage of subsequently occurring endothermic reactions or indirect cooling by radiant or convective cooling with recovery of heat for steam generation. Besides, there are hybrid systems combining some of the described options. A major issue for gasification processes operated above the ash melting temperature is the avoidance of deposition or fouling in heat exchangers caused by solidifying sticky ash particles. Another issue is the need to prevent corrosion caused by condensation of alkali vapour compounds like Na or K. For chemical applications, the widest applied option for gas cooling is water quenching because of lower equipment costs and advantageous conditions for the downstream gas treatment. In contrast, applications requiring the highest overall system efficiencies like IGCC power plants benefit from heat recovery cooling systems.

Besides differences in cooling of the raw gas exiting the gasification reactor, there are requirements to control coal grain size and moisture content, with the latter strongly dependent on coal rank and quality. There are different milling technologies applicable for different coal types. Whereas mills for hard coal often feature an integrated milling and drying approach, for lower rank coals typically there are separate processes.

The energy requirement for milling depends on the type of the applied mill, for example rod mill, hammer mill or roller mill, and coal hardness, typically with decreasing hardness for lower rank coals. The average energy consumption of widely applied roller mills is about 7 to 7.5 kWh/kg of coal.

For separate coal drying there are tubular dryers and fluidised bed dryers, which are particularly applicable for low rank coals like lignite. The thermal energy requirement is at least as high as the evaporation enthalpy of the water at the given process pressure, and normally is about 2500 to 3000 kJ/kg of evaporated moist water. As there are different binding forms of water in coal, the energy demand for drying can exceed the evaporation enthalpy with decreasing moisture content of the coal. In addition to integrated drying and milling being mainly applied to bituminous and higher-rank subbituminous coals, there are also fluidised bed dryers for low-rank subbituminous coals with high moisture content. Fluidised bed drying is used for fine-grained coal under atmospheric or elevated pressure. There is a high efficiency potential by fluidised bed drying as the drying heat demand can be satisfied by compression and utilisation of the latent heat of the vapour obtained from coal drying. If installed upstream of an entrained-flow gasification process, the dried coal will be milled to the required particle size after drying. As yet, there is not a commercial drying technology for lump coal. 

Whereas moving-bed and fluidised bed gasification are characterised by gravimetric dry feeding, coal injection into an entrained-flow gasification reactor can be realised by dry or slurry feeding of dust-grained coal (less than 0.2 mm). Gasification processes applying slurry feeding can achieve higher operating pressures than dry-fed gasifiers ranging between 6 and 6.5 MPa as the slurry consisting of approximately 60–65 wt% coal and 35–40 wt% water is simply pumped to the required pressure and injected into the reactor. Today’s dry feeding systems for entrained-flow gasification mostly rely on pneumatic or dense phase feeding using a transport gas that is often N2 limiting the maximum pressure to about 4 MPa for prevention of raw gas dilution with inert gas components. Dry-fed systems generally feature higher efficiencies because there is no need for evaporation of the slurry water consuming an additional fuel for heat provision in the autothermal processes. New developments aim for high-pressure dry-feeding systems by so-called solid feed pumps allowing for higher feeding pressures without a need for increasing carrier gas flow. Developments include the pumps by GE and by Rocketdyne-Aerojet (Gräbner, 2015). However, none of the new developments have achieved commercial application so far.

An overview of commercial coal gasification technologies is given in Table 6. Whereas moving-bed gasification was the dominating technology in the past, the majority of newly installed plants today rely on entrained-flow gasification (Higman, 2014; GTC, 2015). Although there are a number of Western equipment providers with technologies developed mainly during the 1970s and 1980s, the more recent developments took place in China, resulting in a number of new processes. Current drivers are to reduce Capex, increase efficiency and optimise gas composition to specific applications. There is also focus on the development of new gasifiers to utilise lower feedstock quality.
As the most important component of a coal-to-liquids route is the gasification block, significant effort is put on reduction of capital expenditures in particular as it has the highest share of equipment costs ranging between 40% and 65%. Besides the development of more compact gasifiers, for example, by Aerojet Rocketdyne, reduction of specific capital costs shall be achieved by increasing the single-unit capacities (such as, GE, SIEMENS, ECUST, Air Liquide Global E&C Solutions), introducing dry-feed pumps (see above), replacing heat recovery systems for syngas cooling by less expensive quench systems or by increasing the gasification pressure to build smaller reactors with higher specific throughput (including CB&I, Air Liquide Global E&C Solutions).

Examples for increase of efficiency are illustrated by efforts to maximise carbon conversion or by developments for dry-feeding systems or by the introduction of heat recovery systems instead of quench cooling. Optimisation of the raw gas composition for a specific application is to some extent contradictory to efficiency increase by heat recovery steam generators as syngas composition adjustment is mainly addressed by introduction of quench systems reducing the need for downstream CO-shift conversion.

Adaptation to a lower quality feedstock, that is low-grade and low-rank coals, is addressed particularly by low or medium-temperature gasification, typically based on fluidised bed systems. Examples for new gasification concepts include the ICC-CAS fluidised bed gasifier, the E-Str concept of CB&I and INCI proposed by TU Bergakademie Freiberg. In addition, fixed bed dry bottom gasification technologies can cope with high ash content coal, albeit with adaptation of operating procedures and conditions (Gräbner, 2015).

Provision of high-purity oxygen

While some gasification processes operate with air (for example, to provide producer gas), modern autothermic gasification processes use high-purity oxygen as the gaseous reactant, especially for syngas or hydrogen provision. Typically, this is provided by cryogenic air separation, which comprises a low-temperature distillation of air according to the Linde principle. The air is first dried then subjected to stage-wise compression with intercooling to very high pressure (about 7 MPa) before it is expanded to take advantage of the Joule-Thomson effect leading to a very low temperature close to 80 K. Separation of oxygen and nitrogen takes place in two separate distillation columns, these being insulated (so-called cold box) and operated at different pressure levels, in the range 0.5–0.6 MPa for oxygen and 1–1.2 MPa for nitrogen. Thus high-purity nitrogen is recovered from the high-pressure column (because of its lower boiling temperature) while oxygen and lower-purity nitrogen are obtained from the low-pressure column. The purity required for the oxygen ranges between 98 and 99.5 vol%. If high-purity nitrogen is also a desired product, it can be obtained at up to 99.9 vol% purity. Today’s largest cryogenic air separation units are capable of providing up to 5000 t/d oxygen in a single train with new developments targeting 7000 t/d oxygen capacity. Auxiliary power consumption of state-of-the-art air separation produces high-purity oxygen is about 0.245 kWh/kg of oxygen. Future potential indicates 0.175 kWh/kg of oxygen for highly heat integrated, thermo- and fluid-dynamically optimised air separation units (Pardemann and Meyer, 2015).

Gas treatment and conditioning

Gas purification for liquid syntheses comprise mature technologies. The typical sequence starts with the preferentially separate removal of tars and dust followed by the removal of NH3 and HCl. Dependent on the concentration of sulphur species in the gas and the requirement for the final syngas H2 content, the next step is either removal of sulphuric acid compounds or gas conditioning by CO-shift conversion. If the sulphur removal unit is not capable of handling organic sulphur compounds, a COS and CS2 conversion reactor is required before scrubbing of H2S and subsequently of CO2.

For raw gas containing sufficient sulphur concentration and for syntheses or applications not requiring 100% H2 in the syngas, the typical CO-shift conversion process would be the sour shift process. For ammonia synthesis or H2 production, the CO-shift process would be a sweet CO-shift. 
Figure 18 Gas purification sequence for sour and sweet CO-shift (modified by author)

Figure 18 provides a schematic that indicates the typical sequence of gas purification steps. If the gas contains tar, as is the case for fixed bed coal gasification, tar oil compounds need to be removed, for example, benzene, toluene or xylene, either together with the dust or separately by removing the dust at elevated temperature and the tar compounds at lower temperatures being the optimal case. The most commonly applied process is a venturi-type wash cooler removing tar and dust and cooling down the raw gas for downstream gas conditioning steps. There is a requirement for extensive treatment of the tar-dust-water mixture due to the dissolution of organic compounds, especially phenol, in the scrubber water.

The next step after removal of the bulk content of solids from the raw gas is additional water scrubbing. Besides removal of the finest particles and droplets, water scrubbing aims for recovery of ammonia and halide species from the raw gas. Having passed the water scrubbing, the downstream gas conditioning differs dependent on the syngas quality required. The downstream gas purification does not only aim for removal of pollutants and catalyst poisoning components but also for adjustment of the gas composition, in particular the (H2-CO2)/(CO+CO2) ratio. The hydrogen content normally needs to be increased at the expense of the CO content by applying the homogenous water-gas-shift reaction (see equation 10) using steam as reactant and increasing the content of CO2 that needs to be removed after passing the CO-shift stage. The extent of hydrogen content adjustment depends on the H2/CO ratio of the raw gas provided by the gasifier and the syngas requirement. The H2/CO ratio differs with coal rank and gasification process, with lower ratios for entrained-flow gasifiers and higher hydrogen content for fluidised bed and moving-bed gasification. As CO-shift conversion is performed in fixed bed reactors filled with a catalyst bed, the sulphur content of gas entering the CO-shift unit is an important parameter. There are two process variants: the sulphur resistant catalyst is applied to the raw gas, known as sour CO-shift; and a sulphur sensitive catalyst is used for the sweet gas CO-shift. Whereas the sour shift catalyst requires a minimum H2S/steam ratio in the raw gas and a minimally achievable CO concentration in the exit gas in the range of 3 to 4 vol%, sweet CO-shift can yield as low as 0.2 to 0.4 vol% residual CO concentration if applied as a multi-stage process. The maximum H2S concentration should not exceed between 6-7 ppmv for the high-temperature stage and needs to be less than 0.1 ppm for the low-temperature stage. Most coals will yield a raw gas of sufficiently high sulphur content to apply a sour shift system. A minimum steam to dry gas ratio of 1.5 to 2 is required for both CO-shift types for optimal reaction conditions to approach thermodynamic equilibrium.

If the synthesis requires maximum hydrogen concentrations, for example, ammonia synthesis or hydrogen for refinery applications, sweet CO-shift conversion will always be required to reduce the CO content to a minimum, thereby maximising the H2 yield. Hence, the next step after water scrubbing includes the removal of sulphur components before the gas is saturated with steam and then completely passed through the CO-shift unit. This is a combination of multi-stage high-temperature conversion (280–490°C) and a one-stage low-temperature reactor (180–250°C). The significantly increased content of CO2 is reduced by low-temperature CO2 removal, see below. Residual concentrations of CO2 or CO are further reduced, either by conversion of both species into methane if increased inert gas content is acceptable, or by cryogenic scrubbing or application of special CO removing solvents.

All other syntheses for transportation fuels including synthetic natural gas feature lower requirements for the H2/CO ratio. One or two-stage reactor setups are applied to sour CO-shift conversion with the reactors operated at high temperature between 280°C and 490°C and bypassing a fraction of the raw gas around the CO-shift unit. The bypass ratio is determined by the exit CO and H2 contents of the converted gas, the raw gas H2/CO ratio, and the syngas requirement. Sulphur compounds as well as CO2 are removed following CO-shift conversion, see below. As CO2 often acts as a reactant, there are less stringent requirements for the CO2 capture rate, allowing for a slightly higher operating temperature of the CO2 removal section. Complete removal of CO2 is advantageous only for cobalt-based low-temperature Fischer-Tropsch synthesis where too high a CO2 content in the purified syngas leads to a need for larger equipment size.

As noted above, the final step to obtain a syngas meeting the synthesis requirements includes the removal of acid gas components including sulphur (organic and inorganic) and CO2. Currently, the widest applied acid gas removal process for coal gasification-based syngas is based on physical absorption applying methanol as the scrubbing agent. All such processes rely on the strongly temperature-dependent solubility of sulphuric acid gases and CO2 in methanol. Applying methanol as solvent results in residual sulphur concentrations of less than 0.1 ppmv, with the capability to simultaneously reduce the CO2 content down to 10 ppmv. Most of the sulphuric acid gases are removed at about 273 K whereas trace concentrations and the CO2 are removed at between 200 K and 240 K. Although this type of washing is energy intensive, due to the low-temperature cooling, it is the preferred technology solution because of its capability to also remove most other pollutants to minimum levels, for example, higher hydrocarbons such as tar compounds as well as carbonyls. Both types of compounds are collected from the raw gas before the sulphur removal stage. Moreover, it is possible to remove organic sulphur species without requirement for prior conversion to H2S.

For common synthesis applications, especially fuel syntheses, this process is the last step before providing the purified and conditioned syngas to the reactor.

Fischer-Tropsch (FT) synthesis

This synthesis technique was developed and commercialised in Germany in the 1930s (de Klerk, 2011a,b) for processing a wide range of feedstocks (Table 7). It is characterised as a High-, Medium- and Low-temperature FT process, according to the synthesis temperature.
The product of FT synthesis (syncrude) comprises a broad range of hydrocarbons, the main products are alkenes, alkanes, aromatics and oxygenates. For simplification, hydrogenation of CO is shown in equation 13, which shows that water is a major by-product.

nCO + 2nH2 → -(CH2)- + H2O        -165 kJ/kmol (500 K)   (13)

Hydrogenation is a highly exothermic process and heat removal is critical for FT reactor design. As a result, all FT processes produce medium pressure steam in cooling coils (slurry reactors, fluidised bed reactor) or at the shell side of a multi-tubular reactor (fixed bed).

For all processes, the molar fraction xn of each carbon number n in the product depends on the chain growth probability 𝛼. This relationship is described by the Anderson-Schulz-Flory (ASF) distribution:

𝑥𝑛 = (1−) ∙ 𝛼 ^(𝑛−1)                     (14)

In practice, for all commercial FT processes, the methane selectivity is larger than predicted by the ASF relationship, Figure 19. In contrast, the concentration of C2-components is usually lower. Catalysts for LTFT processes can be characterised by two 𝛼 -values due to the large fraction of heavy hydrocarbons.
Figure 19 Anderson-Schulz-Flory carbon number distribution (de Klerk, 2000)

Cobalt catalysts have a lifetime of several years, which makes them more suitable for operation in fixed bed reactors. The water-gas-shift reaction is not catalysed and only low amounts of CO2 are produced. The desired H2/CO ratio in the syngas feed stream should be in the range of 2.06–2.10.

For FT processes using iron catalysts, the water-gas-shift activity needs to be considered. Thus:

CO2 + H2 ↔ CO + H2O           -40 kJ/kmol (500 K)     (15)

Iron catalysts have a high-water-gas-shift activity which should be taken into account when the syngas modulus is calculated. The so-called Ribblett ratio needs to be considered in this case: (H2)/(2 CO+3 CO2) ≈ 1 (Steynberg 2004). Iron catalysts are well-suited for feedstocks with a low hydrogen content such as coal. Usually, iron catalysts are replaced after several weeks and not rejuvenated. For HTFT a typical cycle time is 40–45 days, while for LTFT it is 70–100 days (de Klerk, 2000).

HTFT is undertaken in either circulating or stationary fluidised bed reactors at 330–360°C (de Klerk, 2011). All products are gaseous and can be separated after cooling outside the reactor. Only iron catalysts are used. Fluidised bed reactors have advantageous mass and heat transfer properties in comparison to fixed bed reactors. Gas-catalyst separation is achieved using cyclones. As shown in Table 7 a large fraction of the syncrude consists of short- and medium-length chain products, such as LPG and naphtha. This makes HTFT suitable for the production of gasoline and diesel products.

A flowsheet of the HTFT process is shown in Figure 20. Stepwise cooling of products is used to separate recycled components and heavy hydrocarbons. Liquid-liquid separation steps are required to remove liquid oil from aqueous products.
Figure 20 HTFT: Product separation and gas loop design (de Klerk, 2011a,b)

HTFT (also known as High Temperature Slurry Fischer-Tropsch) is conducted using an iron catalyst in the slurry phase. Operating conditions are around 270–290°C (Xu and others, 2015). The elevated reactor temperature can be used to generate high temperature steam. In addition, the developers state that the catalyst has a superior activity, low methane selectivity and a low oxygenate content in the syncrude. Slurry reactors have superior mass and heat transfer characteristics and can be used for online catalyst replacement. As can be seen from Table 7, MF-FT syncrude composition is similar to other Fe-LTFT processes.

For LTFT reactors both iron and cobalt have been applied in industrial multi-tubular fixed bed reactors and slurry reactors. The temperature regime of LTFT is usually in the range of 200–240°C (de Klerk, 2011). Liquid product-catalyst separation is complicated and catalyst entrainment to downstream process units can cause severe damages (de Klerk, 2011a,b).

Syncrude from LTFT synthesis consists predominantly of long-chain hydrocarbons and the most desirable final products are Diesel, Kerosene and Wax. An illustrative Fischer-Tropsch case is shown in Figure 21. It shows that less intermediate cooling steps are required as the first gas-liquid separation steps takes place within the reactor.
Figure 21 LTFT: Product separation and gas loop design (de Klerk, 2000)

For this ‘open gas loop’ process, the tail gas is sent to a gas turbine to generate power. All commercial large-scale systems are operated by recycling at least a part of the unconverted syngas. An additional purge stream is needed to avoid build-up of inert components. The ‘closed gas loop’ design is chosen if the syncrude yield should be maximised. In this case the purge gas stream is kept at a minimum. Autothermal or steam reforming is applied to convert light hydrocarbons back to syngas and to increase the output of syncrude (Steynberg, 2004).

Product upgrading and refinery design depends strongly on the desired product yield, as does the chosen FT route. Product upgrading at FT facilities can involve production of intermediate products or fuel blending components, which are then shipped for final refining at conventional refineries. This minimum refining approach has been realised in the Oryx GTL-FT facility, where a single hydrocracker is used to produce intermediate products and LPG. On the other hand, the large Sasol Synfuels refinery complex produces motor-gasoline and diesel products, as well as numerous chemicals (de Klerk, 2011a,b).

All of the mentioned FT processes have been applied commercially by Sasol. As noted in the main report, other projects, primarily in China, are at the planning and construction phase, with a major unit recently starting operation.

Methanol synthesis

Methanol is an important chemical, especially as an intermediate in the production of future fuels such as DME, gasoline/petrol via the MTG process, and as a blend with gasoline/petrol. The first low pressure methanol process was introduced in 1966. Typical process conditions are in the range of 5-10 MPa and 200–300℃. Unconverted syngas is usually recycled with mass-related ratios of 3 to 7 (BTG, 2015). This is due to low conversion which is typically in the range of 4–14% for each pass. Currently, all commercially established methanol processes are low pressure using Cu/ZnO/Al composite catalysts (Ott and others, 2012). Syngas-based methanol formation can be described by the following reactions:

CO + H2 ↔ CH3OH                                 -90.77 kJ/kmol (300 K) (16)
CO2 + 3 H2 ↔ CH3OH + H2O                -49.16 kJ/kmol (300 K) (17)
CO2 + H2 ↔ CO + H2O                          +41.21 kJ/kmol (300 K) (18)

As can be seen from the equations above, the theoretical (H2-CO2)/(CO+CO2) ratio (syngas modulus) at the reactor inlet should equal 2. For most industrial processes, a slightly larger value up to 2.1 is chosen. By-products that can be formed during commercial processes include hydrocarbons through FT reactions, methane, alcohols, DME, esters and ketones. Syngas-based methanol formation is highly exothermic and a key issue in methanol reactor design is heat removal. Sintering of Cu particles on the catalyst surface reduces catalyst activity and lifetime and makes temperature control important (Ott and others, 2012).

Commercial methanol synthesis is conducted mainly in quasi-isothermal fixed bed reactors. Quench-type and multibed intercooled reactor configurations are utilised as well. In 2011, Lurgi was the major technology licensor, followed by Johnson Matthey/Davy Process Technologies and Haldor Topsøe. The original ICI process uses an adiabatic reactor with several fixed beds, which are directly quenched with cold syngas. This process is now owned by Johnson Matthey and offered in collaboration with Davy Process Technologies. Since modern methanol plants are mainly based on quasi-isothermal reactors, these technology providers now offer an axial or radial flow steam raising reactor (Bertau and others, 2014; Ott and others, 2012).
Figure 22 One-stage quasi-isothermal methanol synthesis (Wurzel, 2006)

Lurgi offers a quasi-isothermal one-stage process for medium-scale applications and large-scale, multi-stage concepts). The conventional one-stage reactor (shown in Figure 22) is suitable for methanol production capacities around 3000 t/d while large-scale methanol plants have capacities of 5000 t/d and more.

Medium-scale reactor inlet temperatures are 220–250°C and can reach temperatures up to 280°C within the reactor. By cooling the reactor outlet gas stream to about 40°C, liquid methanol and water can be retrieved before recycling of the unconverted syngas. A purge stream is needed to avoid accumulation of inert gases (Bertau and others, 2014; Wurzel, 2006; Chen, 2011).

The MegaMethanol™ technology consists of an integrated two stage reactor concept, as shown in Figure 23. The first reactor is a conventional steam raising reactor. The methanol-containing outlet stream enters a second reactor which is cooled by the syngas feed steam for the first reactor. As a result of the countercurrent flow of the cold syngas, reactor temperature is reduced and a high equilibrium driving force can be realised. The GigaMethanol concept has been developed for production capacities of up to 10,000 t/d but it can so far only be applied to high pressure (up to 10 MPa) autothermal reforming. In comparison to one-stage synthesis higher per-pass conversion rates can be achieved (Wurzel, 2006; Air Liquide Global E&C Solutions, 2015).
Figure 23 Illustration of the two-stage Lurgi MegaMethanol™ concept (Pardemann, 2013)

Adiabatic methanol reactors are commercially available from Haldor Topsøe and comprise three reactors with inter-stage cooling (Aasberg-Petersen, 2013). Liquid phase methanol synthesis (LPMEOH™) is a process developed by Air Products and Chemicals Inc, specifically for syngas with a low H2/CO ratio and for IGCC operations. As shown in Figure 23, the process comprises a slurry bubble column reactor, in which the catalyst is suspended in a mineral oil. Syngas enters the bottom of the reactor and conversion occurs at the surface of the suspended catalysts. Due to the hydrodynamic conditions, a homogeneous temperature distribution can be achieved within the reactor vessel. Steam is generated in tubes immersed within the slurry.

These conditions result in higher conversion for each pass than for conventional methanol synthesis routes. Another distinct feature is the possibility to operate the synthesis with gas with a low hydrogen content. As can be seen from Figure 24, additional steam can be added to the synthesis gas to generate more hydrogen through the homogeneous water-gas-shift reaction. Operation with a syngas modulus of 0.34 has been demonstrated successfully (Air Products, 2003). However, this process has not been commercialised due to the engineering complexity.

Figure 24 PMEOH process (Pardemann, 2013; Air Products, 1998)

Product upgrade of raw methanol can be realised through either two-column or three-column distillation, which offer low investment costs and energy savings respectively. The methanol purity standards A (99 wt% methanol; 0.1 wt% water; 0.05 wt% alcohols) and AA (99.85 wt% methanol; 0.1 wt% water; 10 ppm alcohols) can be achieved with either product upgrading route (Pardemann, 2013).

Dimethyl ether synthesis

There are two production processes, direct and indirect DME synthesis, for which the latter represents the large-scale technology application.

For indirect DME synthesis, methanol derived from syngas-based synthesis is fed into the DME reactor. Dehydration of the methanol takes place in a fixed bed Al2O3 or zeolite catalyst, typically at 1–1.2 MPa. Equation 19 shows the overall chemical equation.

2 CH3OH → CH3OCH3 + H2O                   ∆RH = -23 kJ/mol    (19)

Achieving a conversion of up to 80% in a single pass, the unreacted methanol is separated in two stages via distillation and recirculated. The feed stream is preheated to 250°C and the product temperature rises to 350–400 C due to an exothermal reaction and adiabatic conditions. The DME can be upgraded to various purity levels. The overall conversion rate of the methanol can reach 99%. A simplified flowchart of the indirect DME synthesis is shown in Figure 25.

Figure 25 Schematic of indirect DME synthesis (modified by author)

Table 10 lists the various production processes for indirect DME synthesis.

DME is primarily used for domestic fuel as a substitute for liquefied petroleum gas with the focus on China.

There is one established production plant that provides DME for use as a transportation fuel, which has an annual capacity of 80,000 t. This is located in Niigata, Japan, is based on Mitsubishi Gas Chemicals technology, and has been operational since 2008 (Ishiwada, 2011).

Direct DME synthesis differs from the indirect process in that there is only one reactor for the conversion of syngas to DME. Therefore, the bi-functional catalyst consists of a mixture of both methanol and dimethyl ether catalyst. Additionally, the homogeneous water-gas-shift reaction (compare to equation 10) supports the process. Equation 20 shows the overall chemical equation.

3 CO + 3 H2 → CH3OCH3 + CO2                   ∆RH = -246 kJ/mol      (20)

The molar ratio of H2 and CO equals 1:1 instead of 2:1 for methanol synthesis. In practice this value ranges from 0.7 to 1.0 depending on temperature and pressure within the reactor. In advance, the syngas thereby has to be treated the same way as for the syngas-based methanol synthesis to prevent catalyst contamination. For a once-through process, the conversion of syngas to DME reaches at most only 50% (based on CO conversion). Thus, there is also a recirculation loop but the separation of syngas, CO2, DME and partly entrained slurry is more complex. Initially the gas separation is performed by an absorber using parts of the produced DME as solvent in order to let the unconverted syngas pass and to capture the CO2. There then follows a separation of the remaining fractions via distillation. Residual methanol can also be mixed with the initial feed stream to the process. The integration of heat follows the same principle as the indirect DME synthesis. A simplified flowchart of the direct DME synthesis is shown in Figure 26.
Figure 26 Schematic of direct DME synthesis (modified by author)

Table 12 provides a comparison of the available direct DME processes, which indicates that this technology is at the pilot stage.

Gasoline synthesis

The methanol-to-gasoline (MTG) process is based on the methanol acting as an intermediate product, which is converted to gasoline in one or more stages. The successive reaction steps are described in equation 21 (Spivey, 1992; Keil, 1999).

The dehydration of methanol results in DME and water (see section above). The DME is then converted into light olefins and subsequently into higher paraffins, olefins, aromatics and naphthenes, together with additional water. These steps are highly exothermic so that the product distribution strongly depends on the prevailing reaction temperature, as shown in Figure 27 (Joseph and others, 1985). A maximum yield of the higher aromatics that comprise gasoline can be obtained at 400°C and high pressure (Stöcker, 1999).
Figure 27 Temperature dependence of MTG product distribution (Pardemann, 2013)

The methanol for gasoline synthesis is produced from syngas either based on coal gasification or natural gas reforming. From 1986 until 1997 a small commercial plant in Plymouth, New Zealand provided 14,500 bbl/d of gasoline based on methanol production from natural gas reforming. The raw methanol was evaporated and fed into an adiabatic DME reactor at 2.6 MPa and more than 300°C. Without any treatment, the products were transferred into five fixed bed adiabatic reactors in parallel for conversion to gasoline. Inlet temperatures of 320–340°C were adjusted by controlling the mixture of unconverted and recirculated DME.

Generating high pressure steam, the product hydrocarbons, gases and water were cooled and separated in a flash drum. The gases were recirculated, the water was post-treated and the hydrocarbons treated further. Within two distillation columns, the light ends were removed and the remaining gasoline fractions were handled depending on their chain length. Light gasoline was sent to an alkylation reactor and split into propane, butane and alkylate. Heavy gasoline was purified from durene. A specific mixture of heavy and light gasoline was sold as the final product. Figure 28 shows a flowchart of the ExxonMobil MTG process as applied at the New Zealand plant.
Figure 28 Schematic of ExxonMobil MTG synthesis (modified by author)

In China, the Jincheng Anthracite Mining Group has taken out a licence for the Mobil process. It first established a 2,500 bbl/d unit in 2009 and recently began operation of a commercial 25,000 bbl/d plant in Shanxi Province (Helton and Hindman, 2014). Lurgi developed a downstream module for gasoline synthesis as part of their process chain of methanol-based large-scale applications (Wurzel, 2006). Methanol is converted into olefins which then pass oligomerisation, ultimately to provide kerosene, diesel, gasoline and LPG. The overall process, called MegaSyn®/MtSynfuels®, is in operation at the Mossel Bay plant (South Africa).

Efficiency and environmental performance

After the synthesis process, the gasification process has the biggest impact on overall process efficiency. The major parameter for assessment of energetic efficiency is cold gas efficiency – the chemically bound energy of the produced raw gas (not considering sensitive heat) related to the chemical. Other criteria for gasification technology assessment include raw gas (syngas) yield, carbon conversion, specific oxygen consumption and steam to oxygen ratio. All parameters are dependent on the type of coal and the applied gasification process, as shown in Table 13.
As for all coal liquefaction routes, water consumption is another crucial issue. The minimal water consumption is reported as five litres of water per litre of Fischer-Tropsch product. Major uses in indirect liquefaction plants are process water (for example, water required for provision of gasification steam or water scrubber make-up water), steam for CO conversion, boiler feed water to recover exothermically released heat of reaction and make-up water to compensate cooling water losses. Significant amounts of waste water are also produced, for example, from black water treatment after raw gas quenching, condensates or strip water from water scrubbing or CO-shift conversion, waste water from acid gas removal and recovery, and from synthesis product purification. The waste water is often organically loaded with hydrocarbons, has elevated ion or salt concentration and can contain adsorbed gases. Specific water treatment processes exist dependent on the gasification process and synthesis route.
The specific product yield of different syntheses is summarised in Table 14.
Overall environmental process chain performance parameters are provided in Table 15.

The major solid waste product from gasification plants is either ash or slag, depending on the gasification technology employed. An important issue is the immobilisation of hazardous components. From this point of view, slagging gasifiers produce the best residues. In other cases, thermal upgrading steps or landfilling might be required (Gräbner, 2015).

Used catalysts from the synthesis steps are another solid waste source. Metal recovery from catalysts is needed if the materials pose a risk to the environment. In the case of metals such as platinum or cobalt, recovery is also mandatory from an economic point of view (Maitlis and de Klerk, 2013).

The 10 largest coal producers and exporters in the Indonesia:
  1. Bumi Resouces
  2. Adaro Energy
  3. Indo Tambangraya Megah
  4. Berau Coal
  5. Bukit Asam
  6. Baramulti Sukses Sarana
  7. Harum Energy
  8. Mitrabara Adiperdana 
  9. Samindo Resources
  10. United Tractors

Source: IEA Clean Coal Centre
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