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Coal-to-Methanol-to-Olefins: Processes and Catalysts

Methanol-to-hydrocarbon conversion reactions were first discovered in the early 1970s using ZSM-5 (MFI) catalysts (Chang, 1977; Chang, 1994). Mobil commercialized a methanol-to-gasoline process in New Zealand in the 1980s and also developed methanol to olefins employing a ZSM-5 catalyst. With support of DOE funding, a 100-barrels per/day demonstration unit was operated in Germany (DOE, 1986). In thel980s, Edith Flanigen and her team (Wilson, 1982) at Union Carbide Corporation (UCC) discovered a new class of material, called silicoaluminumphosphate (SAPO). A particular structure with 3.8 A pore opening known as SAPO-34, a silicon-aluminum-phosphorous-based molecular sieve, showed excellent properties for conversion of methanol to light olefins, primarily ethylene and propylene (Kaiser, 1985, 1987). The structure of SAPO-34 and the small sizes of pore openings are the keys to the high selectivity to produce light olefins using a SAPO-34 catalyst.

The small pore size of SAPO-34 restricts the diffusion of heavy and branched hydrocarbons, which leads to high selectivity to the desired olefins. On the other hand, ZSM-5 molecular sieves produce much lower light olefin yields, primarily due to larger pore openings (about 5.5 A) in the MFI structure. Figure 9 shows a comparison of ZSM-5 and SAPO-34 structures (Vora, 2003). A further advantage of SAPO-34 is that the majority of the C4-C6 fraction is olefinic. This C4-C6 fraction can be converted to light olefins and, thus, increase the production of C, plus C3 olefins to near 90%. This reduces the net purge of C4-C6 fraction to about 5% of the carbon yield. Reaction product distributions for methanol processed over the ZSM-5 and SAPO-34 catalysts are compared in Figure 11 (Vora, 2006).

A number of technologies based on the use of ZSM-5 or SAPO-34 as a catalyst have been developed. These are the UOPFIYDRO MTO™ (“Methanol to Olefins”) Process, which employs a catalyst based on

Framework of SAPO-34 and ZSM-5 molecular sieves

Figure 10. Framework of SAPO-34 and ZSM-5 molecular sieves.

Once-through hydrocarbon yields for ZSM-5 and SAPO-34

Figure 11. Once-through hydrocarbon yields for ZSM-5 and SAPO-34.

SAPO-34 material (Vora, 1998; Chen, 2004), and the Luigi MTP™ (“Methanol to Propylene”) Process, based on a ZSM-5 type catalyst (Gronemann, 2005). Similar technologies also have been developed by Dalian and Sinopec in China (Ying, 2013). In China, one commercial unit of each of these technologies came into commercial operation during 2011, and by 2018 there were 7 or more units in operation. There are more than 10 units in design and construction and these are expected to be in operation by 2022. Both Chinese technologies use catalysts containing SAPO-34 in a fluidized bed reactor with continuous catalyst circulation and regeneration. In addition, there are two methanol to propylene units licensed by Lurgi in operation in China. The MTO Process mainly produces ethylene and propylene and some C4 olefins, while the MTP process mainly produces propylene with gasoline range C5-plus hydrocarbon byproduct. Honeywell’s UOP has announced the licensing of eleven MTO units in China. The first unit at Wison, Nanjing, successfully came onstream during the 2013. By 2019, five UOP licensed units were in operation (Funk, 2014; Senetar, 2019).

Because ZSM-5 catalysts allow larger molecular weight gasoline range materials to come out of the pores, there is less formation of coke on the catalyst relative to SAPO-34 which only allows n-butene and lower molecular weight hydrocarbons. Thus, for the ZSM-5 based catalyst system, it is feasible to design a fixed bed reactor system with cyclic regeneration, which is used in the Lurgi MTP design. On the other hand, SAPO-34-based catalyst systems employ a circulating fluidized bed reactor and regenerator similar to that used in the fluid catalytic cracking process (FCC) in petroleum refining.

Lurgi MTP process

The Lurgi MTP process uses a fixed-bed ZSM-5 catalyst manufactured by Sud-Chemie AG. It provides high propylene selectivity, low coking tendency, low propane yield and limited byproduct formation. An MTP plant with a methanol feed rate of 5000 MT/day produces 1410 MT/day propylene, 540 МТ/ day gasoline, and 109 MT/day LPG. Up to 60 MT/day of ethylene can optionally be recovered from the purge gas or used as fuel. Methanol, both fresh and recycled (as recovered from aqueous streams), is the feed to the MTP unit. Figure 12 (Lurgi, 2003; Wurzel, 2006) shows a schematic process flow diagram. The methanol is vaporized, superheated, and fed to a DME reactor. The DME reactor is a single- stage adiabatic reactor where most of the methanol is converted to dimethyl ether (DME) on an alumina catalyst. The reaction is exothermic and closely approaches thermodynamic equilibrium.

The product of the DME reactor is sent to three MTP reactors in parallel: Two of the reactors are in operation, while a third one is in regeneration or on stand-by. For the purposes of reaction control, each MTP reactor features six zeolite-based catalyst beds, over which the methanol/DME mixture is converted to a mixture of olefins, typically from ethylene to octenes, but such that the carbon distribution peaks at propylene. The operating temperature is about 450 °C and the operating pressure is 0.15 MPa (about 20 psia). Side products from the reaction include naphthenes, paraffins, aromatics, and light ends. The oxygen chemically bound in the methanol results in process water.

Lurgi MTP process flow scheme

Figure 12. Lurgi MTP process flow scheme.

The regeneration of the MTP reactors is performed in situ by the controlled combustion of coke with an air/nitrogen mixture at temperatures similar to the normal reactor operating temperature. The MTP reactor effluent is cooled in a heat recovery system and, finally, through a quench section, in which the hydrocarbons are separated from the bulk of the water. The water is condensed and sent to the methanol and DME recovery column, from which they are recycled to the DME reactor. The water with traces of oxygenates is routed to battery limits.

The hydrocarbon vapor from the quench section is compressed to about 2.5 MPa (365 psia) by a multistage centrifugal compressor with intercoolers and partial condensers. The liquid and vapor hydrocarbons are sent to the purification section. The hydrocarbon streams are first dried by using molecular sieves before the hydrocarbon liquid is fed to a debutanizer column and the vapor is processed through a DME recovery system. The C4+ bottom product is fed to a dehexanizer where aromatics and C7+ are separated from the C6-stream. The majority of the C6-fraction is sent back to the MTP reactors while the C7+ fraction is the gasoline byproduct.

The compressed hydrocarbon vapors, including light olefins and DME, and the overhead C4+/DME from the debutanizer are fed to a DME removal system, in which C3-hydrocarbons are separated from C4+ hydrocarbons and oxygenates. The methanol and DME stream are routed to the methanol recovery column for recycle to the DME reactor. The C4 hydrocarbon fraction is recycled to the MTP reactor for further propylene production, except for a small purge that is added to the LPG byproduct stream.

The C,-fraction is fed to the deethanizer, in which a C^-stream is recovered as top product; one part of this stream is recycled to the MTP reactor while the rest can optionally be sent to a two-column ethylene purification unit or to fuel gas. The C3 bottom product from the deethanizer contains about 97% propylene and 3% propane, but no methylacetylene or propadiene; it is routed through a guard bed of activated alumina and fed to the C3 splitter for the recovery of polymer-grade propylene.

l/OP/HYDRO MTOprocess

As mentioned earlier, Edith Flanigen and her associates at UCC discovered SAPO-34 material and showed it to be a good catalyst for methanol to olefins reaction during the early 1980s. This group at UCC, then known as Catalyst, Adsorbent and Process Systems, was merged with UOP LLC, Des Plaines,

Illinois. As a result, further development for the MTO technology took place at UOP. In 1992 UOP formed a partnership with Norsk Hydro of Norway for further joint development of the technology. A fluidized bed reactor-regenerator demo for a one ton per day methanol feed was built and operated at Norsk Hydro facility in Norway for several years. At the end of 1995, a joint team of UOP and Norsk Hydro first presented their data at the World Natural Gas Symposium held in South Africa (Vora, 1997). This demonstration was not sufficient to convince potential licensees, as questions were raised regarding the quality of ethylene and propylene coming from methanol, an oxygenate feed which may result in some unknown impurities that could be detrimental to polymerization catalysts. In order to demonstrate polymer-grade propylene and ethylene from the MTO process, a partnership with Total of France was formed and a large fully integrated MTO demonstration imit was built with high purity ethylene and propylene recovery, including polymerization reactor for polyolefins. With successful demonstration by 2010, UOP licensed its first imit in China. The Total partnership was expanded to include Total’s Olefins Cracking Process (OCP), a process for cracking higher C4-C6 olefins to propylene and ethylene. Integr ation of this process with MTO allowed UOP to increase ethylene-propylene yield to near 90% and the integrated process is called UOP Advanced MTO Process.

The UOP HYDRO MTO Process can use “crude” methanol, “fuel-grade” methanol. Grade AA methanol, or even DME as feed. The choice of feedstock generally depends on project-specific situations. Figure 13 illustrates a simple flow diagram for the UOP HYDRO MTO Process (Vora, 1998). The MTO process utilizes a circulating fluidized bed reactor that offers a number of advantages over both fixed bed reactors and other types of fluidized bed reactors. The circulating fluidized bed reactor provides better mass transfer than bubbling bed fluidized bed reactors as well as better temperature control than riser and fixed bed reactors, especially given the highly exothermic nature of the methanol-to-olefins reactions. This type of reactor has been widely used in the Fluid Catalytic Cracking (FCC) process units in petroleum refineries.

Constant catalyst activity and product composition can be maintained via continuous regeneration of a portion of used catalyst by coke burning with air. UOP’s MTO catalyst has demonstrated the required selectivity, long term stability, and attrition resistance necessary for attractive economics with low operating costs.

The overall selectivity of the UOP, HYDRO MTO process is about 75-80% to ethylene and propylene on a carbon basis, and about 15% C4 plus hydrocarbons. The balance is Cj-C3 paraffins plus coke on catalyst. The C4 plus material is mostly linear butenes and some pentenes. These olefins make an ideal feed to the OCP unit to further increase the yields of ethylene and propylene. Propylene to ethylene ratios

UOP HYDRO MTO process flow scheme

Figure 13. UOP HYDRO MTO process flow scheme.

Olefin selectivity vs. operatuig severity of the UOP MTO process with and without olefin cracking process (OCP)

Figure 14. Olefin selectivity vs. operatuig severity of the UOP MTO process with and without olefin cracking process (OCP)


in the product can be adjusted within the range of 0.80 to 1.33 (Figure 14) in order to reflect the relative market demand and values for ethylene and propylene. The reactor temperature is the key variable for controlling propylene to ethylene ratios, with higher temperatures leading to a higher ethylene yield. The temperature requirements have to be balanced with higher coke formation at higher temperatures.

The reactor pressure is normally dictated by mechanical considerations. Lower methanol partial pressure leads to higher selectivity to light olefins, especially ethylene. Therefore, a slight yield advantage occurs when using a crude methanol feed compared to high purity methanol. The reactor effluent is cooled and quenched to separate water from the product gas stream. The reactor provides very high conversion, so there is no need for a large recycle stream.

A small amount of unconverted oxygenates are recovered in the oxygenate recovery section, after which, the effluent is further processed in the fractionation and purification section. Conventional treating methods have been shown to be effective for removing by-products to the specification levels required for producing polymer-grade ethylene and propylene products.

As shown in Figure 14, the total ethylene plus propylene yield can be further enhanced by incorporating a cracking process to convert C4 plus material to propylene and ethylene. Overall carbon selectivity for the integrated flow scheme approaches 90% ethylene plus propylene.

Integration ofCTL/GTL with CTO/GTO

GTL and CTL processes offer large product market opportunities for natural gas and coal utilization but are challenged by high capital costs and the relatively low transportation-fuel product values. Since synthesis gas production is a common step in the manufacture of GTL and methanol, there are possibilities for integrated complexes. Figure 15 illustrates such a complex, using coal or natural gas as feedstocks, and producing both olefins and liquid fuels. Both options—coal or natural gas liquid fuels (CTL, GTL) and coal or natural gas to polymers (CTP, GTP) facilities—incorporate sizeable front-end synthesis gas imits for the processing of natural gas. Over 60% of the capital cost is related to the production of synthesis gas. These units are the major contributors to the relatively high investments required for these

Integrated GTO/GTL complex for production of liquid fuels and olefins from coal or natural gas

Figure 15. Integrated GTO/GTL complex for production of liquid fuels and olefins from coal or natural gas.

complexes. It follows that the integration of these facilities to co-produce fuels and chemicals could offer substantial synergistic savings.

A rough rule of thumb is that the quantity of synthesis gas required to produce 25,000 barrels per day (BPD) liquid hydrocarbons could also be used to produce in excess of one million MTA light olefins. A typical liquid fuels plant is likely to have a capacity of 100,000 BPD or more. Thus, the addition of one million MTA light olefins represents roughly 25% additional synthesis gas production capacity. The incremental production will require substantially less capital cost than a stand-alone smaller unit. The integration of methanol and liquid fuels facilities combined seen in today’s oil refining sector, where there is increased focus on opportunities for petrochemicals production as some regional fuel demands change with the conversion of methanol to olefins, can provide cost saving synergies together with the production of high value-added olefins and polymer products. This would follow the current trend.


Though there are no direct routes for the conversion of coal or methane to liquid hydrocarbons or petrochemical products, these raw materials can be converted to liquid fuels or several high value- added petrochemical intermediates via synthesis gas. Examples are: Acetylene, ethylene, propylene, and methanol. These are the primary raw materials for a vast number of petrochemicals and polymer industry products, like PVC, polyethylene, polypropylene, acetic acid, acrylonitrile, formaldehyde, ammonia and many more. The vast resources of coal in areas of high demand, namely the USA, China and India, could provide long-term raw material security.


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