Methanol Production – A Technical History
Methanol Production – A Technical History
A review of the last 100 years of the industrial history of methanol production and a look into the future of the industry
Global methanol production in 2016 was around 85 million metric tonnes (1), enough to fill an Olympic-sized swimming pool every twelve minutes. And if all the global production capacity were in full use, it would only take eight minutes. The vast majority of the produced methanol undergoes at least one further chemical transformation, more likely two or three before being turned into a final product. Methanol is one of the first building blocks in a wide variety of synthetic materials that make up many modern products and is also used as a fuel and a fuel additive. This paper looks at the last 100 years or so of the industrial history of methanol production.
Methanol has been produced and used for millennia, with the ancient Egyptians using it in the embalming process – it was part of the mixture of substances produced in the destructive distillation (pyrolysis) of wood. However, it was not until 1661 that Robert Boyle produced pure methanol through further distillation, and only in 1834 was the elemental composition determined by Jean-Baptiste Dumas and Eugene Peligot. At a similar time, commercial operations using destructive distillation were beginning to operate (2).
There are many parallels between the industrial production of methanol and ammonia and it was the early development of the high pressure catalytic process for the production of ammonia that triggered investigations into organic compounds: hydrocarbons, alcohols and so on. At high pressure and temperature, hydrogen and nitrogen will only form ammonia, however the story is very different when combining hydrogen and carbon oxides at high pressure and temperature, where the list of potential products is lengthy and almost all processes result in a mixture of products. Through variations in the process, the catalyst, the conditions, the equipment or the feedstock, a massive slate of industrial ingredients suddenly became available and a race to develop commercial processes ensued.
The First Drops
Early research into methanol production quickly focused on copper as a prime contender for the basis of a catalytic process to methanol, with Paul Sabatier and Jean-Baptiste Senderens (3) discovering in 1905 that copper effectively catalysed the decomposition of methanol and to a lesser extent its formation. A lot of the early testing looked at what catalysts could effectively destroy methanol, assuming they would be equally as effective under alternative conditions at forming methanol. Following the start of large scale ammonia production in Germany during 1913, the pace of research picked up and in 1921 Georges Patart patented the basis of a high pressure catalytic process that used a variety of materials including copper (along with nickel, silver or iron) for methanol synthesis (4). A small experimental plant was later built using this process in Patart’s native France, near Asnières (5).
The German Effort
The wood-based processes were always very limited in scale and it was 1923 before production could be considered ‘industrial’ with a catalytic process developed by Mathias Pier at Badische Anilin- & Sodafabrik (BASF), Germany (Figure 1).
The BASF process produced methanol from synthesis gas (syngas), which at the time was a mixture of hydrogen and carbon monoxide. The process works by the following reactions:
Methanol formation (Equations (i) and (ii)) is favoured by low temperatures and high pressures. All three equilibrium reactions occur simultaneously, although it is common to only consider two of the three to simplify any analysis, as it can be seen that Equations (ii) and (iii) combined are the same as Equation (i).
The BASF process operated at above 300 atm and 300–400°C, using a zinc chromite (Cr2O3-ZnO) catalyst developed by Alwin Mittasch (6), about a decade after his work on the first industrial ammonia synthesis catalyst. The high pressures benefitted conversion to methanol and to achieve sufficiently quick reaction rates, high temperatures also had to be used. Further increases in temperature would have drastic effects on the selectivity and equilibrium, so conditions were selected to be a compromise. Methanol production began on 26th September 1923 at the Leuna site (7).
The subsequent research into the catalyst was extensive, with the list of possible candidates covering large swathes of the periodic table, from antimony to zirconium, bismuth to uranium (itself a popular catalyst of the time) (5, 8). Given the extensive testing, it is perhaps unsurprising that in the list can be found many of the components that make up the modern catalysts used in methanol plants in the 21st century.
Initially, iron was to be used for methanol production (as with ammonia production), but this along with nickel was phased out in successive patent applications until the requirement for the process to be ‘completely excluding iron from the reaction’ was included in the mid 1920s (9). During the early years there was a lot of effort looking at other combinations of carbon, hydrogen and oxygen. One major application was Fischer-Tropsch reactions: the creation of straight chain saturated hydrocarbons, for example for fuels. This is readily catalysed by iron at similar conditions to methanol synthesis. With early iron-containing methanol synthesis catalysts, it was found that the iron would react with the carbon monoxide to form iron carbonyl, which decomposes at high temperatures to iron metal. It was therefore easy to transform the catalyst into one much more efficient at making hydrocarbons than methanol; reactions that are even more exothermic and not equilibrium limited, hence at risk of thermal runaway. The catalyst is not the only source of iron in such processes, with the obvious choice for construction of the early reactor vessels being steel, which itself contains iron. Many of the early plants were therefore either lined or made of non-ferrous metals, such as copper, silver or aluminium (10).
The equilibrium limitations of the methanol formation reactions (Equations (i)–(iii)), especially under the early operating conditions, were such that conversion to methanol in a single pass through a reactor was very low. To overcome this, the gas had to be recycled over the catalyst a number of times. Each time, the gas is cooled to condense any product methanol and the consumed reactants are replaced with fresh synthesis gas. The gas is rarely pure hydrogen and carbon monoxide, and any non-reacting species, such as methane or nitrogen, introduced through the fresh gas supply accumulate in such a loop, so a small fraction of the gas must be purged, also losing some reactants. Figure 2 shows the basic components of a methanol synthesis loop, which are still used today.
The interchanger is a more modern concept, reducing energy consumption by using the hot gas exiting the converter to heat the inlet gas. Early patents (11) show a lot of the aspects of modern methanol production, including the recycle loop and the use of a guard bed of additional catalyst or absorbent to remove “traces of substances deleterious to the reaction”, early versions tending to be copper based. The loss of reactants through the purge was also considered in early processes, with Forrest Reed filing a patent in 1932 (12) for recycling the purged gas through an additional reactor in a loop with high concentrations of non-reacting components, complete with condensation and separation. This approach is now used to revamp and add capacity to modern methanol plants.
The general concept spread rapidly and plants could be found around the world by the end of the 1920s producing a total of around 42,000 metric tonnes per year of methanol in new, catalysed, high-pressure processes (13).
Early on it was recognised that the most effective catalysts used a combination of copper and another metal oxide, but the synthesis section and catalyst remained very similar for about 25 years. Eugeniusz Błasiak filed a patent in 1947 for a new catalyst containing copper, zinc and aluminium, manufactured by co-precipitation (14). The patent claimed a method for producing a “highly active catalyst for methanol synthesis” and further laboratory testing over the following decades proved this.
The biggest impediment to the use of copper catalyst was the rate of poisoning by sulfur compared to the zinc chromite catalysts typically used in those plants. The syngas generation process had moved on from coal and coke feeds to natural gas reforming, and it was accepted that sulfur in the feed would poison the reforming catalyst and reduce the activity. The reformers were therefore run at close to atmospheric pressure to prevent hydrocarbon cracking over the poisoned catalyst, which would cover the surface in a layer of carbon and remove all residual activity. Around this time, work was underway to create an alkalised reforming catalyst which was protected against carbon deposition and could therefore run at elevated pressure (initially 14 atm, but soon after up to 35 atm) (15). A second development at a similar time gave hydro-desulfurisation catalysts, which remove sulfur from the naphtha or natural gas feedstock and preserve the activity of the reforming catalyst. This gave a process for supplying high purity syngas at increased pressure. The cost of compressing syngas is much greater than the cost of compressing natural gas, so the opportunity to move compression duty upstream also provided an energy efficiency benefit to plant designs.
By the 1960s, methanol was being made almost solely from natural gas and naphtha using low pressure reforming and high pressure synthesis, with a broad range of process licensors all offering a very similar configuration. Substantial gains in process efficiency had been made since the very early plants, partly due to the larger scale of the later plants. One technology that the largest plants of the time could take advantage of was centrifugal compressors, offering much lower costs at high gas flow rates compared to the previous reciprocating machines (16). With these gains increasing with equipment size, the drive for bigger and bigger plants continued.
The British Intervention
In the 1960s, arguably the biggest change to the industry was introduced by Imperial Chemical Industries (ICI), UK. This began in 1963 when Phineas Davies and Frederick Snowdon filed a patent for a methanol production process operating at 30–120 atm (17). Using a copper, zinc and chromium catalyst, they had created a process capable of producing high quantities of methanol without the need for very high pressures. The lower pressures meant that fast reaction rates could be achieved at lower temperatures of 200–300°C, which reduced the formation of byproducts. This meant the catalyst was able to achieve a selectivity of greater than 99.5%, based on organic impurities in the liquid methanol.
At a similar time, ICI had developed its ‘high pressure’ steam reformer, capable of transforming naphtha or later, natural gas into syngas. The process was therefore not just a method of synthesising methanol, but a complete process from natural gas to methanol: the Low Pressure Methanol (LPM) process, which remains the leading route to methanol to this day.
The catalyst was soon revised with a patent application by John Thomas Gallagher and John Mitchell Kidd of ICI in August 1965 (18) to a catalyst containing the oxides of copper, zinc and another element from Groups II to IV of the periodic table, with aluminium being the preferred candidate. This was the catalyst that ICI installed in its own methanol plant constructed at the time and forms the basis of the KATALCOJMTM 51-series of catalysts sold around the world by Johnson Matthey today.
ICI constructed and commissioned the first LPM plant at its site in Billingham, UK, in 1966 (Figure 3) with a design capacity of 300 metric tonnes per day (MTPD) and an expected catalyst lifetime of six months. The synthesis section operated at only 50 atm (19). Two years later the catalyst was still operating and the plant could consistently produce 400 MTPD. This increased to 550 MTPD with the second catalyst charge and some further plant upgrades. The converter had 71 m3 of catalyst, with three cold shots of gas injected partway down the bed to cool the reacting gas. The plant operated until 1985.
At the lower pressure of the new process, the circulating gas volumes were greater and therefore centrifugal compressors were advantageous at lower plant capacities (16). Much more efficient plants were then available without needing to construct a large-scale facility.
ICI by this time had a long history of methanol production, stretching back to 1929 with its first high pressure plant operated under licence from IG Farben (then owners of BASF). Following a few years of successful operation of the Billingham plant, ICI licensed the technology and in 1970 a 130 MTPD plant was commissioned for Chang Chun Petrochemical Co Ltd in Taiwan (20). In spite of a challenging two weeks of commissioning, with “torrential rain, a typhoon and an earthquake”, this plant was to be the first of many and later that year a 1000 MTPD plant was commissioned for Monsanto at Texas City, Texas, USA. Only a single high pressure synthesis plant was built after 1966 (21).
The most distinguishing feature of most methanol plants (or licensors) is the type of converter used for methanol synthesis. Broadly the converters can be divided into two categories based on how they remove the heat of reaction to maximise conversion:
multiple adiabatic catalyst beds with external cooling of the gas
internal cooling within one or more catalyst beds.
Externally cooled converters come in a variety of configurations: quench converters inject cold, unreacted gas after each adiabatic bed to reduce the temperature, whereas series adiabatic converters use heat exchangers between the catalyst beds. Both externally and internally cooled types were used in the early low pressure plants, with the quench converters offered by ICI benefitting from the simple vessel design minimising cost. The early versions employed a single catalyst bed with gas injection points at multiple locations down the vessel. These designs were susceptible to large temperature distributions developing and propagating down the vessel. A subsequent improvement on the design therefore collected the gas, mixed it with the incoming quench gas and distributed it across the next bed. This prevented temperature variations propagating from bed to bed. Many reactors of this design operate around the world today as ARC reactors, a joint ICI and Casale SA, Switzerland, design from the early 1990s. Figure 4 shows the reaction pathway of a quench converter, with successive additions of cold gas taking it back away from the equilibrium line to maximise conversion.
Series adiabatic converters are more efficient users of catalyst as, without the need for quench gas that bypasses the early beds, all the gas passes over all the catalyst and the temperature control for each bed is truly independent. Additional heat exchangers in the loop contribute to higher capital costs and series adiabatic beds never really found favour in the industry.
Internally cooled reactors began with Lurgi GmbH, Germany, shortly after the first LPM plant from ICI. The Lurgi reactor was one that had already been used for many years in Fischer-Tropsch synthesis and consisted of catalyst-filled vertical tubes surrounded by a shell of boiling water, with the reaction heat transferred into the shell to generate steam to be used elsewhere in the process. A steam drum local to the converter provides a constant supply of water at boiling temperature through natural circulation. This design achieved a more even temperature distribution and lower peak temperature. Whilst the converter was more complicated than the ICI design, and therefore more expensive, the steam it generated at about 250°C could be used elsewhere for an efficiency benefit or even exported. The design also required a lower catalyst volume. Figure 5 shows the reaction pathway in such a converter, following more closely the temperature for maximum reaction rate compared to quench converters. Many variations exist on this theme today, some with the catalyst and boiling water reversed, such as in the Variobar of Linde AG, Germany, which uses helical tubes in an axial catalyst bed to achieve pseudo-cross flow.
Other internally cooled converters use process gas on the cooling side, including ICI’s subsequent tube cooled converter, in which cold gas rises inside empty vertical tubes, absorbing heat from the surrounding catalyst bed before turning over at the top of the converter and flowing back down through the catalyst bed. The large amount of heat generated by the synthesis reactions requires a high flow rate on the cooling side, which for gas-based cooling is typically only available within the synthesis loop, with different designs utilising gas from different parts of the loop.
Most modern converters use internal cooling, either with circulating gas or by raising steam, which broadly allows the temperature in the catalyst bed to track the point of maximum reaction rate, a balance of the kinetic limitations of low temperature and the thermodynamic (equilibrium) limitations of high temperature.
The basic formula was now set and so the plants could grow in size and scale. By the early 1970s the plants had gone from the 150 MTPD of the early low pressure plants to 1500 MTPD. The second plant ICI built at Billingham in 1972 had a design capacity of 1100 MTPD and used 110 m3 of catalyst (22) operating at 100 atm. This second plant operated through to 2001 and struck a better balance of operating pressure and equilibrium, with the vast majority of plants since having been designed for 80–100 atm. This heralded the start of the first golden age of methanol expansion in the early part of the 1970s as people recognised the benefits of the new LPM process. Figure 6 shows the approximate capacity added each year using LPM technology, with a notable peak in the 1970s and further peaks in the 1980s and around 2010 that will be explored in the second half of this history.
By the early 1980s all new plants were being constructed using low-pressure technology and almost all of the high-pressure plants had been converted to low pressure (23). Interestingly the pyrolysis of wood had not completely ceased as the use of ‘synthetic’ methanol had not yet been accepted as an alcohol denaturant in some countries. British Law to this day (24) is based on the use of ‘wood naphtha’ to denature pure ethanol, a process whereby it is made unsuitable for human consumption and therefore exempt from beverage sales taxes. Wood naphtha is the mixture of substances derived from pyrolysis, primarily methyl alcohol (methanol).
The 1980s saw the impact of the second oil crisis that followed the Iranian Revolution in 1979 and the Iran-Iraq war that started soon after. The increased oil price meant that oil producing nations had significantly increased revenues and this allowed them to increase petrochemical production, including methanol. Thus began the second golden age of methanol expansion. But the oil crisis also prompted countries to start looking at how they could become less reliant on imported oil and to start looking at production of synthetic fuels.
The expansion of methanol is driven by demand for derivatives and a recurring theme throughout the history is its potential use as an intermediate in the production of synthetic automobile fuel. Whilst interest has peaked on a number of occasions, typically when a nation struggles with domestic supply, there have been few plants actually constructed. One example is the two methanol plants in Motunui, New Zealand, which were constructed for synthetic fuel production in 1985, using the Mobil licensed methanol to gasoline (MTG) process (25). Both plants now solely produce methanol and the MTG equipment has been removed. Whilst the production of a direct petrol replacement has never found lasting favour, many plants today are being constructed to feed methanol to olefins (MTO) processes to produce olefins from coal instead of from naphtha or ethane, and an increasing amount of methanol is blended into gasoline supplies around the world to meet legislative requirements.
Autothermal Reforming and Alternative Reforming
For a typical natural gas to methanol plant using steam reforming technology, roughly a half of the capital cost is in the steam reformer and it also accounts for a large part of the footprint. Available technology limited the maximum economic size of a single reformer and a new technology was therefore required to allow plant capacities to expand beyond about 2500 MTPD (26). This limit was first identified in the early 1970s when methanol was being considered as a way to move energy around in the face of global imbalance. To produce sufficient quantities of methanol to achieve this, production capacity would need to increase rapidly with plants of up to 5000 MTPD, which would have required 2000 tube steam reformers. The largest constructed at that time had only 600 (27).
The gap was ultimately filled with autothermal reforming; the controlled introduction of oxygen into (partially) reformed gas to combust some of the hydrogen, providing the heat for further reforming reactions across another bed of catalyst. As the heat is produced and retained within the process, a lot of the equipment associated with reformers is not needed, although a supply of oxygen is required, typically from an air separation unit. The technology is deployed in various configurations:
parallel reforming – a steam reformer and autothermal reformer (ATR) are used in parallel
combined reforming – the steam reformer is partially bypassed and the bypass and reformed gas are combined and fed to the ATR to complete the reforming process.
A further development by ICI in the 1980s was to completely remove the traditional steam reformer in the Leading Concept Methanol (LCM) process. Rather than burning fuel gas to provide the heat for the reforming reactions, the hot, autothermally reformed gas was used to heat the catalyst tubes in a gas heated reformer (GHR). The feed gas first passes through the catalyst in the GHR, then the ATR and finally the heating side of the GHR to provide the heat for the initial reaction.
It is possible to take these concepts even further and some plants have only an ATR. Autothermal Reforming is susceptible to soot formation if significant quantities of higher hydrocarbons are present and so a simple adiabatic pre-reformer is required to de-rich the natural gas. This arrangement produces a gas very rich in carbon oxides and is therefore most effective where a source of additional hydrogen is present to balance the stoichiometry of the gas.
Typically, combined reforming gives a plant with a reasonably sized steam reformer, a low level of methane in the syngas and a stoichiometrically balanced syngas for methanol formation.
The speed of catalyst development had greatly increased since the mid 1970s when testing equipment began to be automated, greatly increasing the amount of test work that could be conducted. This led to a number of step changes in the performance of methanol synthesis catalysts, although the base recipe of copper with a combination of zinc and aluminium or chromium oxides remained very similar. One such step change was in the early 1990s, with a new generation of catalysts being introduced, just as capacities were ramping up and plant operators were looking to uprate their original low pressure plants (28). ICI introduced a new, more active catalyst using a four-component system, adding magnesium to the existing copper, zinc and aluminium (Figure 7).
Modern catalysts are expected to last at least three years and typically between four and six years is achieved, although six to eight years is not uncommon. The catalysts are highly selective towards methanol synthesis and the effects of some of the early catalyst candidates (iron and nickel) are better appreciated, especially their role in the formation of paraffinic hydrocarbons, and these are now seen as catalyst poisons. Despite the selectivity of modern catalysts being in excess of 99.5%, there is still a need to remove various impurities from the condensed product methanol to achieve either chemical or fuel sales grades. Generally, this is achieved at low pressure with one, two or three distillation columns in series. Dissolved gases are removed first, along with low boiling point byproducts and then the difficult methanol-ethanol separation must be conducted, along with water removal. The water can be reused in the steam system, the light ends as fuel and the ethanol (actually a mixture of many heavier organic compounds) can be added back into the process before the reformer, to be reformed and reused.
In 2004, the long destined capacity of 5000 MTPD was achieved when the Atlas plant was commissioned in Trinidad, only for it to be overtaken the following year by M5000, also in Trinidad, producing up to 5400 MTPD. This latter plant achieved its capacity with only a steam reformer containing less than 1000 tubes, showing the simultaneous improvements in reforming catalyst and technology. Figure 8 shows the twin synthesis converters on M5000.
China – The Coal Story
A lot of the growth in the methanol industry through the early 21st century (the third golden age of methanol expansion) came from China and its booming economy. China’s petrochemical industry had been heavily dependent on imported crude oil, although China had plentiful supplies of cheap coal. China began to embrace new technologies for converting their coal into other chemicals and one key building block in that process was methanol. Rapidly increasing demand for a wide range of methanol derivatives, particularly olefins via the MTO process, has required a continuous supply of new methanol plants using coal gasification to provide the syngas for methanol synthesis.
To take advantage of the economies of scale, and in some cases to fit in with the economic size of a downstream MTO plant, the demand for higher and higher capacity synthesis loops has grown. With the methanol plants typically near to the coal in remote locations, the main process equipment must be transported to the sites by rail, where bridges in particular limit the maximum diameter and the infrastructure can limit the maximum weight. Whilst vessels can be made taller and taller, for catalyst beds this will soon result in very high pressure drops. For synthesis loops above about 3000 MTPD the catalyst requirement is too great to use a single vessel and multiple converters in a single loop are required. Initially and at modest capacities, two identical parallel converters were sufficient. As capacities continued to increase, so did the complexity, with multiple converters of different types used within single loops to reduce the capital cost of the loop equipment, as shown in Figure 9 with the Johnson Matthey Combi Loop. Other loops were designed using the Johnson Matthey Series Loop where product is recovered between converters to reset the equilibrium and increase production. The largest plants in operation by 2010 would typically have two or more converters to make up to 5500 MTPD of methanol. To minimise pressure drop and therefore compression duty in large synthesis loops, larger water cooled reactors are now available in radial flow configurations.
The second aspect of the growth in China is the coal to methanol story, which uses gasification technologies to convert coal and steam at very high temperature to syngas. Modern purification systems now allow the syngas to be substantially cleaned of sulfur and other impurities and a very pure gas is fed to the synthesis loop, unlike the systems from the 1920s and 1930s. Typically, coal-fed plants give a much more carbon monoxide-rich syngas compared to steam reforming of natural gas, the more exothermic route to methanol and so the ability to remove heat is even more important.
Energy and Environmental Efficiency
Since the introduction of the low pressure process, the focus turned to energy efficiency, especially during increasing energy prices in the 1970s and 1980s. Table I shows the progression of efficiency over these years by ICI through successive improvements to the integration of the whole plant.
|Flow sheet||Year||Consumption, GJ MT–1|
|LP – 50 atm||1966||36|
|LP – 100 atm||1972||36|
|Tube cooled converter||1983||29.3|
With the ever increasing focus on environmental performance, there are a number of designs and new plants in recent years which aim to set new standards for efficiency or emissions. One particular plant is Carbon Recycling International’s (CRI) George Olah Plant in Iceland, fully commissioned in 2012. Using electricity from the fully renewable Icelandic grid, it electrolyses water to provide hydrogen, which is combined with carbon dioxide recovered from a local geothermal power station (30).
Other new plants are considering the emissions benefits of avoiding a steam reformer and using the GHR technology to set new standards for low emission natural gas-based plants. The plans for Northwest Innovation Works (NWIW), USA, use the technology and will be among the largest plants in the world (31).
With the imminent start-up of the 7000 MTPD plant of Kaveh in Iran (32), the scale of plants continues to grow.
Methanol demand has grown steadily for many years fuelled by economic growth in major countries around the world, a trend which is likely to continue. Many of the current plant licensors and designers have flow sheets capable of scaling up to 10,000 MTPD, but after a number of purported projects, it remains to be seen if the economy of scale is ready to be stretched that far or if the security of multiple trains once again wins out.
At least for now, the production of methanol via the LPM process remains dominant, despite research interest into other themodynamically attractive routes. Recent examples based on the partial oxidation of methane to methanol include the work of Zhijun Zuo et al. (33) and Patrick Tomkins et al. (34). Whilst work such as this could open up a new, low temperature route to methanol, no such new routes have so far left the laboratory.
KATALCOTM is a trademark of the Johnson Matthey group of companies.
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Daniel Sheldon is a Senior Process Engineer at Johnson Matthey, Chilton, UK. He obtained his MEng (Hons) in chemical engineering from the University of Manchester, UK. He joined Johnson Matthey on the graduate training scheme in 2011 and has spent time in catalyst manufacturing and technology development for the ammonia and methanol industries. Currently he provides technical support to Key Methanol Customers. He is a Chartered Member of the Institute of Chemical Engineers (IChemE).