Chemical Engineering - Industrial Organic Chemistry

Dispense Processi

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1 POLITECNICO DI MILANO Industrial Engineering Faculty Master Degree Course in Chemical Engineering Industrial Organic Chemistry: processes Prof: Enrico Tronconi By Giorgio Pastore Academic Year 201 6 – 201 7 2 Documentazione attinente ai seguenti riferimenti: • Appunti lezioni di “Industrial Organic Chemistry”, prof. E. Tronconi. • Slides lezioni di “Industrial Organic Chemistry”, prof. E. Tronconi. • Ullmann's Encyclopedia of Industrial Chemistry , VCH . • Chemical Process Technology , J.M oulijn, M.Makkee, A.Van Diepen. • Fischer -Tropsch refining, Arno de Clerk. • Petroleum refining, technology and economics , J.H.Gary, G.E.Handwerk . Il materiale che segue è strettamente da integrare con le lezioni del corso! 3 Contents Coal ................................ ................................ ................................ .............. 6 Coal gasification ................................ ................................ ................................ ................................ ............. 6 Gasification Reactions ................................ ................................ ................................ ............................... 6 Thermodynamics ................................ ................................ ................................ ................................ ....... 6 Gasification Technologies ................................ ................................ ................................ .......................... 7 Indirect liquefaction (Fischer -Tropsch syntesis) ................................ ................................ ............................ 8 Coal as raw material ................................ ................................ ................................ ................................ ...... 9 Natural gas ................................ ................................ ................................ . 10 Natural gas as raw material ................................ ................................ ................................ ......................... 10 Methanol ................................ ................................ ................................ ..... 12 Production ................................ ................................ ................................ ................................ ................... 13 Thermodynamics ................................ ................................ ................................ ................................ ..... 13 Kinetics ................................ ................................ ................................ ................................ .................... 13 Process Technology ................................ ................................ ................................ ................................ ..... 15 Production of Synthesis Gas ................................ ................................ ................................ .................... 15 Synthesis and processing of crude methanol ................................ ................................ .......................... 16 Reactor Design ................................ ................................ ................................ ................................ ......... 17 Formaldehyde ................................ ................................ ............................. 18 Physical Properties ................................ ................................ ................................ ................................ ...... 18 Chemical Properties ................................ ................................ ................................ ................................ .... 18 Producti on ................................ ................................ ................................ ................................ ................... 19 Silver Catalyst Processes ................................ ................................ ................................ .......................... 19 Formox Process ................................ ................................ ................................ ................................ ....... 22 Acetic Acid ................................ ................................ ................................ .. 24 Physical properties ................................ ................................ ................................ ................................ ...... 24 Producti on ................................ ................................ ................................ ................................ ................... 25 Acetaldehyde Process ................................ ................................ ................................ .............................. 25 Carbonylation of Methanol ................................ ................................ ................................ ..................... 26 Processes ................................ ................................ ................................ ................................ ................. 28 Fischer -Tropsch synthesis ................................ ................................ .......... 30 Mechanism ................................ ................................ ................................ ................................ .................. 31 Product Selectivity and Characteristics ................................ ................................ ................................ ....... 32 4 Refining of Crude Products ................................ ................................ ................................ .......................... 34 FT Operating Conditions ................................ ................................ ................................ .............................. 35 FT Reaction Engineering ................................ ................................ ................................ .............................. 35 Fischer –Tropsch Catalyst Formulation and Deactivation ................................ ................................ ............ 37 Analysis of the FTS reactor: ideal optimal reactor (Krishna method) ................................ ......................... 38 Level I: Catalyst design ................................ ................................ ................................ ............................ 38 Level II: Injection and dispersion strategies ................................ ................................ ............................ 40 Summary and existing reactors ................................ ................................ ................................ ............... 41 Petroleum ................................ ................................ ................................ ... 42 Crude Distillation ................................ ................................ ................................ ................................ ......... 45 Thermal an d catalytic cracking ................................ ................................ ................................ .................... 47 Thermal cracking ................................ ................................ ................................ ................................ ..... 48 Visbreaking ................................ ................................ ................................ ................................ .............. 48 Coking ................................ ................................ ................................ ................................ ...................... 49 Catalytic Cracking ................................ ................................ ................................ ................................ .... 49 Steam cracking ................................ ................................ ................................ ................................ ............ 52 Process ................................ ................................ ................................ ................................ ..................... 52 Catalytic Reforming ................................ ................................ ................................ ................................ ..... 53 Reactions ................................ ................................ ................................ ................................ ................. 54 Catalyst ................................ ................................ ................................ ................................ .................... 56 Process ................................ ................................ ................................ ................................ ..................... 56 Reactor design ................................ ................................ ................................ ................................ ......... 58 Hydrotreating ................................ ................................ ................................ ................................ .............. 58 Catalyst and operative conditions ................................ ................................ ................................ ........... 59 Separation of light hydrocarbons ................................ ................................ ................................ ................ 60 H2-CH 4 separation ................................ ................................ ................................ ................................ .... 61 C2H4-C2H6 separation ................................ ................................ ................................ ............................... 61 C3H6-C3H8 separation ................................ ................................ ................................ ............................... 61 C4 separation ................................ ................................ ................................ ................................ ........... 61 Separation of BTX fraction ................................ ................................ ................................ ........................... 62 Acetaldehyde ................................ ................................ .............................. 64 Physical Properties ................................ ................................ ................................ ................................ ...... 64 Chemical Properties ................................ ................................ ................................ ................................ .... 64 Production ................................ ................................ ................................ ................................ ................... 64 Production from Ethanol ................................ ................................ ................................ ......................... 65 Production from Acetylene ................................ ................................ ................................ ..................... 65 5 Production from C 1 ................................ ................................ ................................ ................................ .. 66 Production from Hydrocarbons ................................ ................................ ................................ ............... 66 Production from Ethylene ................................ ................................ ................................ ....................... 66 Ethylene Oxide ................................ ................................ ........................... 67 Physical Properties ................................ ................................ ................................ ................................ ...... 67 Chemical Properties ................................ ................................ ................................ ................................ .... 68 Production ................................ ................................ ................................ ................................ ................... 68 Catalysts ................................ ................................ ................................ ................................ ....................... 68 Operating conditions ................................ ................................ ................................ ................................ ... 69 Technolo gy ................................ ................................ ................................ ................................ .................. 69 Propylene Oxide ................................ ................................ ......................... 71 Chlorohydrin Process ................................ ................................ ................................ ................................ ... 71 Indirect oxidation process ................................ ................................ ................................ ........................... 72 Vi nyl Chloride (VCM) ................................ ................................ ................ 72 Balanced VCM Process ................................ ................................ ................................ ................................ 73 Maleic Anhydride (MA) ................................ ................................ .............. 77 Production ................................ ................................ ................................ ................................ ................... 77 Oxidation of Benzene ................................ ................................ ................................ .............................. 77 Oxidation of C4 Hydrocarbons ................................ ................................ ................................ ................ 78 Comparison ................................ ................................ ................................ ................................ .............. 80 Phthalic Anhydride ................................ ................................ ..................... 80 Styrene ................................ ................................ ................................ ........ 82 Production ................................ ................................ ................................ ................................ ................... 82 6 Coal Coal gasification Gasification of coal to produce syngas (or coal gas or town gas) dates back to the end of the eighteenth century. During the mid -1800s, coal gas was widely used for heating and lighting in urban areas. The development of large -scale processes began in the late 1930s and the process was gradually improved. Following World War II, however, interest in coal gasification dwindled be cause of the increasing availability of inexpensive oil and natural gas. In 1973, when oil and gas prices increased sharply, interest in coal gasification was renewed and, especially in the last 10 –15 years, much effort has been put in improving this proc ess . The main reason for this interest is the dramatic increase in oil and gas prices, which seems to be a continuing trend, together with the wide availability of coal c ompared to oil and natural gas . Although coal gasification, like steam reforming of n atural gas, produces syngas, the incentive of the two processes is different: the goal of steam reforming is the production of carbon monoxide and hydrogen for chemical use, whereas coal gasification was primarily developed for the conversion of coal into a gas, which happens to predominantly contain carbon monoxide and hydrogen. Gasification Reactions In a broad sense, coal gasification is the conversion of the solid material coal into gas. The basic reactions are: The first two reactions with carbon ar e endothermic, whereas the latter three are exothermic. It is common practice to perform coal gasification in an “autothermal” way. The reaction is carried out with a mixture of O 2 (or air) and H 2O, so that combustion of part of the coal produces the heat needed for heating up to reaction temperature and for the endothermic steam gasification reaction. Besides these heterogeneous reactions, homogeneous reactions occur. Coal is a complex mixture of organic and mineral compounds. When it is heated to re action temperature (depending on the process, between 1000 and 2000 K), various thermal cracking reactions occur. Usually, the organic matter melts, while gases evolve (H 2, CH 4, aromatics, etc.). Upon further heating the organic mass is transformed into po rous graphite -type material, called “char”. This process is called pyrolysis . When reactive gases are present in the gas phase, several succeeding reactions occur. When less than stoichiometric amounts of oxygen are present, as in coal gasification, the at mosphere is overall reducing and the product contains a complex mixture of organic compounds. Thermodynamics It will be clear that the structure of coal and its products is very complex. For convenience, in thermodynamic calculations we assume that coal solely consists of carbon. Furthermore, under practical gasification conditions, it is usually assumed that only the following species are present in the gas phase: H 2O, CO, CO 2, H2, and CH 4. Of course, other hydrocarbons are also present, but their c oncentrations are relatively small and in thermodynamic calculations, these are all represented by CH 4. Figure 5.10 shows the equilibrium composition resulting from gasification of coal in an equimolar amount of steam as a function of temperature and pres sure. With increasing temperature, the CO and H 2 mol fractions increase due to the increasing importance of the endothermic H2O gasification reaction. Accordingly, H 2O decreases with temperature. Both CO 2 and CH 4 go through a maximum as a result of the exothermicity of their formation and the endothermicity of their conversion. As can be seen from Figure 5.10, a low pressure is favorable for CO and H 2 forming reactions due to the increase in number of molecules. The CO 2 and CH 4 maxima shift to higher tem perature with increasing pressure. 7 Gasification Technologies Coal gasifi cation processes differ widely. The three basically different reactor technologies that are used are fixed bed, fluidized be d and entrained flow gasifiers. In moving -bed gasifiers (also called fixed -bed gasifiers), the gasifying medium passes through a bed of granules . A countercurrent configuration exhibits excellent thermal efficiencies (outgoing ash heats the incoming gases and outgoing product gas heats the incoming so lid feedstock ). The long residence times of solid particles moving through the bed (typically 1 – 2h), together with the temperature profile of the countercurrent system, allow high carbon conversion efficiencies. On the other side, the bed temperature is non -uniform, in particular for the first particles (at lowest T). Fluidized beds for gasification of solids are characterized by linear velocities of gasifying medium sufficient to lift the solid particles. This requires smaller particle sizes than moving -bed proc esses, typically in the 0.1 – 5mm range . As a 8 result of fluidization, mixing of solids within the bed approaches that of a perfectly mixed reactor. Therefore, the solids composition and the temperature of the bed are nearly uniform, bed temp erature being chosen according to feedstock reactivity. Advantages : the technology is less complex than that of moving beds and does not involve moving parts, gas composition is steady due to uniform conditions in the bed and moderate gasification temperat ures can be used. Limitations : capacity flexibility is limited by entrainment at high gas velocities and by the gas velocity required to maintain fluidization . In entrained -flow gasifiers, solid particles are carried or entrained by the reacting gases. Thu s, solids and gases move in the same direction with approximately the same velocity. To achieve this, the particles must be smaller than in other systems (typically < 0.1 mm). This, together with high temperature, leads to acceptable carbon conversion during the short solids residence time in the gasifier (a few seconds). Advantages : the gasifier has no moving parts and a simple r geometry than a fluidized bed, the gasifier has the highest capacity per unit vo lume. Limitation s: the high gasification temperature causes thermal losses , necessitates higher quality construction materials and leads to increased oxygen consumption. Indirect liquefaction (Fischer -Tropsch syntesis) Processes for the conversion of syngas derived from coal or natural gas into liquid fuels such as gasoline and diesel (Figure 6.21) have been considered on and off for many years, usually as an alternative for oil -based fuels. These processes are known as coal -to-liquids (CTL) and gas -to-liquid s (GTL) processes. More recently, biomass -to-liquids (BTL) processes have been receiving more and more attention. The best known of such processes is the Fischer –Tropsch (FT) process, named after F. Fischer and H. Tropsch, the German coal researchers who discovered, in 1923, that syngas can be converted catalytically into a wide range of hydrocarbons and/or alcohols. Before and during World War II this process was extensively used on a commercial scale in Germany for the production of fuels from coal -deri ved gas. After the war, the plants were shut down because they became uncompetitive when large quantities of crude oil were discovered. A less known route is the conversion of syngas to gasoline via methanol. This methanol -to-gasoline (MTG) process was developed by Mobil (now ExxonMobil) during the 1970s and 1980s in response to the critical energy situation in the Western world, which triggered the search for non -oil -based processes for fuel production. Until recently, these synthetic fuels have never be en able to compete ec onomically with oil -based fuels . However, in recent years these processes have come into the picture again, for instance as means to convert natural gas from remote gas fields into liquid fuels. The transportation of the gas to possible consumer markets, either by pipeline or as liquefied natural gas (LNG) in special tankers, is costly and logistically difficult. An interestin g option then may be to convert this gas into more readily transportable liquids, such as the bulk chemicals ammonia and methanol, or liquid fuels. The latter have a much larger market and are, therefore, more attractive from the viewpoint of economy of sc ale. See “ Fischer –Tropsch synthesis” chapter. 9 Coal as raw material 10 Natural gas Natural gas as raw material 11 As a result of efforts to increase the value of natural gas in logistically favorable locations, the chemical liquefaction of natural gas (also the chemical reaction route) was developed on the basis of the Fischer – Tropsch process. This process creates h igh -quality liquid products and paraffin wax. 12 Methanol Methanol (CH 3OH, M 32.042 ), also termed methyl alcohol or carbinol, is one of the most important chemical raw materials. About 85% of the methanol produced is used in the chemical industry as a starting material or solvent for synthesis. The remainder is used in the fuel and energy sector; this use is increasing. Historical Aspects . Methanol was first obtained in 1661 by Sir Robert Boyle through the rectification of crude wood vinegar over milk of lime. From ca. 1830 – 1923, “wood alcohol”, obtained by the dry distillation of wood, remained the only important source of methanol. As early as 1913, A.Mittasch and coworkers at BASF successfully produced organic compounds containing oxygen, including methanol, from carbon monoxide and hydrogen in the presence of iron oxide catalysts during developmental work on the synthesis of ammonia. The decisive step in the large -scale industrial production of methanol was made by M. Pier and coworkers in the earl y 1920s with the development of a sulfur -resistant zinc oxide – chromium oxide catalyst. By the end of 1923 the process had been converted from the developmental to the production stage at the BASF Leuna Works. Processes based on the above work were perfor med at high pressure (25 – 35MPa) and 320 – 450° C. They dictated the industrial production of methanol for more than 40 years. In the 1960s, however, ICI developed a route for methanol synthesis in which sulfur -free synthesis was reacted on highly selective copper oxide catalysts. This and other related low -pressure processes are characterized by fairly mild reaction c onditions (5 – 10MPa, 200 – 300° C). Physical Properties . Methanol is a colorless, neutral, polar liquid that is miscible with water, alcohols, esters, and most other organic solvents; it is only slightly soluble in fat and oil. Because of its polarity, methanol dissolves many inorganic substances, particularly salts. The most important physical data for methanol follow: Liquid Density (101.3 kPa) 0.8100 g/cm3 (@ 0 °C) 0.78664 g/cm3 (@ 25 °C) 0.7637 g/cm3 (@ 50 °C) Melting Point −97.68 °C Boiling Point (101.1kPa) 64.70 °C Heat of vaporization (101.3 kPa) 1128.8 kJ/kg Standard Enthalpy of Formation (101.3 kPa) −200.94 kJ/mol (gas) −238.91 kJ/mol (liquid) Explosion limits in air 5.5 – 44 vol% Ignition temperature 470 °C Chemical Properties . Methanol is the simplest aliphatic alcohol. As a typical representative of this class of substances, its reactivity is determined by the functional hydroxyl group . Reactions of methanol take place via cleavage of the C −O or O −H bond and are characterized by substitution of the −H or −OH group. In contrast to higher aliphatic alcohols, however, β-elimination with the formation of a multiple bond cannot occur. Important industrial reactions of methanol include the following : 1) Dehydrogenation and oxidative dehydrogenation 2) Carbonylation 3) Esterification with organic or inorganic acids and acid derivatives 4) Etherification 5) Addition to unsaturated bonds 6) Replacement of hydroxyl groups 13 Main chemical use: formaldehyde, acetic acid, MTBE, methyl formate (HCOOCH 3), dymethil carbonate (CH 3)2CO 3, dymethil ether, solvents, gasoline additive, MTO process. Production Thermodynamics The formation of methanol from synthesis gas can be described by the following equilibrium reactions: ΔH300K =−90.77 kJ/mol ΔH°300K =−49.16 kJ/mol Both reactions are exothermic and accompanied by a decrease in volume. Methanol formation is thus favored by increasing pressure and decreasing temperature, the maximum conversion being determined by the equilibrium composition. In addition to the two methanol -forming reacti ons, the endothermic reverse water -gas shift reaction must also be taken into account: ΔH°300K =41.21 kJ/mol The conversion of carbon dioxide to methanol (2) is then the overall result of Equations (1) and (3), and the equilibrium constant K2 can be described as K2=K1·K3 (it provides an interesting route for CO 2 chemical utilization, provided the stoichiometric quantity of H 2). Syngas is mainly produced by CH 4 steam reforming, which ideally leads to H 2/CO ratio equal to 3. So this ratio has to be adjusted to meet the stoich iometric value of reaction (1), for example via reverse water gas shift reaction (3) (in general, minor amounts of CO 2 are always present in reformed gas products). Chemical equilibrium considerations are crucial for methanol synthesis ( ΔG°=0, @T=135°C for reaction (1)) since no active industrial ca talyst is available under 135°C (in general, hydrogenation reactions are o ften affected by relevant thermodynamic limitations; Methanol synthesis can be seen as a CO hydrogenation) : • Temperature: as low as possible, to favour K1; • Pression: as high as possible, to favour equilibrium yield. Kinetics The formation of methanol is a typical heterogeneously catalyzed reaction. The original catalysts (ZnO/Cr 2O3 based) were only active at high temperature (300 -400°C) . Therefore, the pressure had to be very high (250 −350 bar) to reach acceptable conversions ( classical methanol processes ). In the late 1960s, the ability to produce sulfur -free syngas allowed the use of m ore active catalysts ( Cu/ZnO/Cr 2O3 on Allumina ): this has led to a new generation of plants, the “low -pressure plants”. OH CH H CO 3 1 2 2 ⎯ →  + O H OH CH H CO 2 3 2 2 2 3 + ⎯ →  + O H CO H CO 2 3 2 2 + ⎯ →  + 14 The typic al temperature ranges for the classical and modern methanol processes are indicated in the figure. The development of catalysts that are active at lower temperature (200 -250°C) made it possible to operate at lower pressure (50 −100 bar), while maintaining t he same conversion as in the classical process. Temperature is critical. A low temperature is favorable from a thermodynamic point of view. Moreover, the high - activity catalysts are sensitive to “sintering”, which increases progressively with temperature. The temperature should not exceed 570 K becau se then unacceptable sintering of the catalyst will occur . Compared to ammonia synthesis, catalyst development for methanol synthesis was more difficult because, besides activity, selectivity is crucial. It is not surprising that in the hydrogenation of carbon monoxide other products, such as higher alcohols a nd hydrocarbons, can be formed (t hermodynamics show that this is certainly possible ). Therefore, the catalyst needs to be ver y selective. The selectivity of modern copper -based catalysts is over 99%, which is remarkable considering the large number of possible by -products . However: • High catalyst stability issues with growing T (thermal sintering is crucial); • High sensibility to catalyst deactivation due to poison agents (S, Cl compounds). For c lassical processes: • Selectivity concerns due to undesired reactions (es. Methanation reaction above 400°C); • Thermal runaway concerns due to higher exothermicity of undesired reactions (es. Methanation); • Higher resistance to S, Cl compounds. Nowadays, most syngas for methanol production is produced by steam reforming of natural gas. The ideal syngas for Me thanol production has a H 2/CO ratio of about 2 mol/mol. A small amount of carbon dioxide (about 5%) increases the catalyst activity. A H2/CO ratio lower than 2 mol/mol leads to increased by -product formation (higher alcohols, etc.), a higher ratio results in a less efficient plant due to the excess hydrogen present in the syngas, which has to be purged. The composition of syngas depends on the feedstock used. When naphtha is the raw material, the stoichiometry is approximately right. When methane (natural gas) is used, however, hydrogen is in excess. In practice, eithe r the excess hydrogen is burned as fuel or the content of carbon oxides in the syngas is increased. This can be done by one of two methods: 1. When available, carbon dioxide addition to the process is a simple and effective way to balance the hydrogen and carbon oxides content of the syngas. Carbon dioxide addition can be implemented by injecting it either in 15 the reformer feed stream or in the ra w syngas. In both cases the stoichiometric ratio for methanol synthesis is achieved, although the compositions will be somewhat different. 2. Installing an oxygen -fired secondary reformer do wnstream of the steam reformer. The feed gas composition ( particular ly the proportions of CO 2 and H 2O) plays an important role in determining the activity and selectivity of catalysts in methanol production. In fact, i nvestigations have shown that various routes exist for the formation of methanol via car bon monoxide or carbon dioxide (this is one important reason according to which reverse WGS cannot be neglected in the presence of CO 2: the catalyst works also for reverse WGS reaction because it probably involves similar catalytic intermediate species from CO 2). Example of kinetic expression (BASF process, high P): • It accounts for the reversibility of Methanol production (i.e. it has thermodynamic consistency and considers equilibrium effects); • It accounts for the non -ideal gas behavior (fugacities in the rate expression); • It is derived from a simple LHHW model, provided the surface reaction as rate determining step; • It shows an inhibiting effect due to CO 2 competitive adsorption. Despite the fact this model has been used for several year s for his ability to efficiently fit experimental data, the logical statement “right model” = ”right mechanism” has not to be applied: the data -fitting ability of a model is not necessarily the result of a good kinetic mechanism, since it depends also on t he mathematical flexibility of the model itself (even though the model might be not correct). This is the case for the BASF model (for example, it considers molecular H 2 adsorption and surface reaction between the surface molecules of hydrogen and one of C O. This is not true since: 1) hydrogen is interested by dissociative adsorption, 2) such a surface reaction cannot be realistic of a physical situation). The actual mechanism is still unknown. It may imply the formation intermediate hydrogenated species from CO (ex. HCO, H 2CO, HCOOH …), which may be formed also by mean of a CO 2 pathway. In this sense, CO 2 would have, up to a certain concentration, a promoting effect on the formation of key surf ace intermediates, and thus on the reaction rate. see “optimal T profile” for adiabatic multi -stage reactors. Process Technology Methanol is currently produced on an industrial scale exclusively by catalytic conversion of synthesis gas. Processes are classified according to the pressure used: • High -pressure process 25 – 30MPa • Medium -pressure process 10 – 25MPa • Low -pressure process 5 – 10MPa The main advantages of the low -pressure process are lower investment and production costs, improved operational reliability, and greater flexibility in the choice of plant size. Industrial methanol production can be subdivided into three main steps: 1. Production of synthesis gas 2. Synthesis of methanol 3. Processing of crude methanol Production of Syn thesis Gas All carbonaceous materials such as coal, coke, natural gas, petroleum, and fractions obtained from petroleum can be used as starting materials for synthesis gas production. Natural gas is generally used in the large scale production of synthesis gas for methanol synthesis. The stoichiometry number. S, characterizes synthesis gases :        2 2 2 CO CO CO H S + − = 16 The stoichiometry number should be at least 2 for the synthe sis gas mixture. Values above 2 indicate an excess of hydrogen, whereas val ues below 2 mean a hydrogen deficiency relative to the stoichiometry of the methanol formation reaction. Natural Gas. Most methanol produced worldwide is derived from natural gas. Natural gas can be cracked by steam reforming and by partial oxidation . In steam reforming the feedstock is catalytically cracked in the absence of oxygen with the addition of water and possibly carbon dioxide . The reaction heat required is supplied externally. In partial oxidation, cracking takes place without a ca talyst. Reaction heat is generated by direct oxidation of part of the feedstock with oxygen. See Steam reforming and partial oxidation of natural gas… To reach the stoichiometry necessary for methanol synthesis, carbon dioxide, if available, is mixed with exit gas from the steam reformer. If carbon dioxide is not available, the conversion must be performed with an excess of hydrogen. Hydrogen accumulates in the synthesis recycle gas and must be removed. Other Raw Materials. Higher hydrocarbons (e.g., liquefied petroleum gas, refinery off -gases, and particularly naphtha) are also used ; they are processed mainly by steam reforming. Crude oil, heavy oil, tar, and asphalt products can also be converted into synthesis gas, but this is more difficult th an with natural gas (S content is considerably higher (0.7 – 1.5% H 2S and COS) and must be removed , s ynthesis gas contains excess carbon monoxide and must, therefore, be subjected to shift conversion). Synthesis and processing of crude methanol In one pass only about 50% of the synthesis gas is converted because thermodynamic equilibrium is reached; therefore, after methanol and water are condensed and removed, the remaining synthesis gas must be recycled to the reactor. In the scheme, t he make -up sy nthesis gas is firstly pre -treated by means of a two -stage condensation (separation of heavier condensable components in syngas) and brought to the desired pressure in a multistage compressor . The unreacted recycle is also added and compressed with the fee d. The syngas flow stream is then pre -heated (recovery of effluent sensible heat) and fed to the reactor. Depending on the reactor configuration, single or multiple feed streams may be present (multiple streams are representative of multi -stage adiabatic r eactor with intermediate quench cooling). All modern processes are low -pressure processes with plant capacities in the order of 10 6 t/y; t he plants differ mainly in reactor design, and, interrelated with this, in the way the heat of reaction is removed. The effluent mixture is then further cooled after passing through a series of heat exchanger s (effluent sensible heat recovery) and sent to separation (flash) stages in order to recover most of the incondensable , not -reacted gases and by products . In the scheme, a two -stage flash separation is present; the first one (high P) , allows recovering most of the not -reacted syngas, the second one (low P), allows recovering and separating low -boiling by products (mainly methane and C 2-C4 alkanes). The syng as is recycled into the process; low -boiling alkanes are usually recycled directly to the upstream steam reforming section together with natural gas feed. Since an internal recycle loop is present, a purge flow is necessary; t he quantity of purge gas from the loop is governed by the concentration and absolute amount of inert substances and the stoichiometry number. Stoichiometry number s different from the ideal value of 2 leads to higher recycle flow rates due to higher amounts of excess unreacted spec ies; this has a significant impact on the production cost, with higher compressing costs (and higher costs due to oversizing of the equipment involved in the recycle loop). 17 It can be simply vented (and burned for example for reformer heating ) or further recycled again to the upstream steam reforming section. Crude methanol is then treated in the distillation train section. Two columns are used: the first one is devoted to the separation of residual light products, mainly CH 3OCH 3 and CH 3COOH, removed as distillates (the scheme shows in particular a partial condenser with total reflux configuration in the head of the column, with a further equilibrium stage at low T for the head gases), the second one is devoted to the separation of heavier bypr oducts, removed as bottom stream ( the heaviest product is water, removed together with mainly C2-C4 higher alcohols). Pure methanol is removed as distillate. This configuration has two main advantages: • the reflux ratio can be used a s variable to efficientl y control the purity of the distillate; • since methanol is removed as distillate, heavier impurities are present in much less extent (respect to an indirect distillation sequence). If needed, higher alcohols may also be recovered and removed as further prod ucts; this is usually done by mean of a side stream in the second column (in order not to add an additional column). The process overall conversion is about 90% (conversion per pass of ~50 -70%), with a selectivity of about 97 -98%. As previously said, met hanol is not the thermodynamic favorite product from a syngas mixture: high process selectivity is because of the high catalyst efficiency in kinetically favoring the desired reactions to Methanol. Final remark : since the methanol market demand is very high, it is convenient to have very high capacity processes, ~10 6 t/y, instead of a great number of smaller plants, in order to have benefits from a scale economy. In fact, as a rule of thumb: Reactor Design Current industrial processes for producing methanol differ primarily in reactor design. Many different reactors are available; they may be either adiabatic (e.g., ICI) or quasi -isothermal (e.g., Lurgi). Adiabatic Reactors. The ICI process uses an adiabatic reacto r with a single catalyst bed . The reaction is quenched by adding cold gas at several points. Thus, the temperature profile along the axis of the reactor has a saw tooth shape. In the Kellogg process, synthesis gas flows through several reactor beds that ar e arranged axially in series . In contrast to the ICI quench reactor, the heat of reaction is removed by i ntermediate coolers. The Haldor -Topsoe reactor operates on a similar principle, but synthesis gas flows rad ially through the catalyst beds . Quasi -Isothermal Reactors. The Lurgi process employs a tubular reactor with cooling by boiling water . The catalyst is located in tubes over which water flows. The temperature of the cooling medium is adjusted by steam pressure control . In both cases, design efforts are for the development of a reactor temperature profile as close as possible to optimal one and the heat of reaction is used for the production of high -pressure steam (except for quench cooling reactors). In addition, pressur e losses concerns are always present, due to the presence of a process gaseous recycle loop (radial flow reactors allows lower ∆P). Remarks • For multi -bed reactors, the flow is downstream in order to avoid any possible entrainment and fluidization phenomen a. These phenomena are almost negligible for normal catalyst particle size but might become relevant whenever, during normal operating conditions, catalyst particles are damaged, with the subsequent creation of smaller particles and powders. • Adiabatic bed reactors have a much simpler and controllable configuration then external cooled reactors, since no direct cooling of the catalyst bed is present. On the other side, additional complications and costs are present due to external heat -exchanger (or furnace passes) or quench feed injection (oversizing of the reactor, with the same initial feed flow); • Catalyst deactivation with time is inevitable; in order to minimize the impact of shut down phases, some reactors might present particular configuration that all ows a continuous replacement of “spent” catalyst (quasi - mobile bed configuration, with catalyst withdrawal on the bottom side). 6.0 0 1 0 1        = P P c c 18 Formaldehyde Formaldehyde was first synthesized in 1859, when Butlerov hydrolyzed methylene acetate and noted the characteristic odor of the resulting solution. In 1867, Hofmann conclusively identified formaldehyde, which he prepared by passing methanol vapor and air over a heated platinum spiral. This method, but with other catalysts, still constitutes the principal method of manufacture. The preparation of pure formaldehyde was described later by Kekul è in 1882. Industrial production of formaldehyde became possible in 1882, when Tollens discovered a method of regulating the methanol vapor: air ratio and affecting the yield of the reaction. In 1886, Loew replaced the platinum spiral catalyst by a more efficient copper gauze. The German firm, Mercklin und L osekann, started to manufacture and market formaldehyde on a commercial scale in 1889. Another German firm, Hugo Bl ank, patented the first use of a silver catalyst in 1910. Industrial development continued from 1900 to 1905, when plant sizes, flow rates, yields, and efficiency were increased. In 1905, BASF started to manufacture formaldehyde by a continuous process emp loying a crystalline silver catalyst. The methanol required for the production of formaldehyde was initially obtained from the timber industry by carbonizing wood. The development of the high -pressure synthesis of methanol by BASF in 1925 allowed the prod uction of formaldehyde on a true industrial scale. Physical Properties Monomeric Formaldehyde. Formaldehyde (CH 2O, Mr 30.03) is a colorless gas at ambient temperature that has a pungent, suffocating odor and an irritant action on the eyes and skin. Formaldehyde liquefies at -19.2°C, the density of the liquid being 0.8153 g/cm 3 at -20°C and 0.9172 g/cm 3 at −80°C. It solidifies at −118°C to give a white paste. The liquid and gas polymerize readily at low and ordinary temperatures up to 80°C. Liquid Density (101.3 kPa) 0.8153 g/cm3 (@ -20 °C) 0.9172 g/cm3 (@ -80 °C) Melting Point -118 °C Boiling Point (101.1kPa) -19.2 °C Heat of vaporization (101.3 kPa) 23.32 kJ/mol Standard Enthalpy of Formation (101.3 kPa) -115.9 kJ/mol Heat of combustion (25°C) -561.5 kJ/mol Heat of solution (23°C) In water: -62 kJ/mol In methanol: -62.8 kJ/mol Explosion limits in air 7 – 72 vol% Ignition temperature 430 °C At a low temperature, liquid formaldehyde is miscible in all proportions with nonpolar solvents such as toluene, ether, chloroform, or ethyl acetate. However, solubility decreases with increasing temperature and at room temperature, polymerization and vola tilization occur, leaving only a small amount of dissolved gas. Polar solvents, such as alcohols, amines or acids, either catalyze the polymerization of formaldehyde or react with it to form methylol compounds or methylene derivatives. Aqueous Solutions. At room temperature, pure aqueous solutions contain formaldehyde in the form of methylene glycol HOCH 2OH (hydrate formaldehyde) and its oligomers (H-(OCH 2-)n-OH , para -formaldehyde) and of cyclic trimer ( (CH 2O) 3, trioxane) . Monomeric, physically dissolved formaldehyde is only present in low concentrations of up to 0.1 wt %. Dissolution of formaldehyde in water is exothermic (heat of solution -62 kJ/mol) . Concentrated aqueous solutions containing more than 30 wt% formaldehyde become clo udy on storage at room temperature, because larger poly(oxymethylene) glycols ( n≥8) are formed , which then precipitate out. Chemical Properties Formaldehyde is one of the most reactive organic compounds known and, thus, differs greatly from its higher homologues and aliphatic ketones. Reduction and Oxidation. Formaldehyde is readily reduced to methanol with hydrogen over a nickel catalyst. It is easily oxidized by nitric acid, potassium permanganate, potassium dichromate, or oxygen to give formic acid or CO 2 and water. 19 In the presence of strong alkalis, formaldehyde undergoes a Cannizzaro reaction with formation of methanol and formic acid ( ). Production Formaldehyde is produced industrially from methanol by the following three processes: 1) Partial oxidation and dehydrogenation with air in the presence of silver crystals, steam and excess methanol (BASF process: 680 -720°C, methanol conversion = 97 -98 %). 2) Partial oxidation and dehydrogenation with air in the presence of crystalline silver or silver gauze, steam and excess methanol (600 -650°C, primary conversion of methanol = 77 -87 %). The conversion is completed by distilling the product and recycling the unreacted methanol. 3) Oxidation only with excess air in the presence of a modified Fe-Mo -V oxide catalyst (250 -400°C, methanol conversion = 98-99%). Methanol used for formaldehyde production , according to processes 1-3, must be usually AA grade . Processes that employ oxidation of methane do not compete with methanol conversion processes because of the lower yields of the former processes ( : at CH 4 Tactivation =600 °C , CH 2O is decomposed ). Silver Catalyst Processes The silver catalyst processes for converting methanol to formaldehyde are generally carried out at atmospheric pressure and at 600 -720°C. The reaction temperature depends on the excess of methanol in the methanol – air mixture. The composition of the mixture must lie outside the explosive limits. The following main reactions occur during the conversion of methanol to formaldehyde: ΔH°R = +84 kJ/mol ΔH°R = -243 kJ/mol ΔH° R = -159 kJ/mol The extent to which each of these three reactions occurs, depends on the process data (reaction 3 is the overall reaction from 1 and 2) . Byproducts are also formed in the following secondary reactions: ΔH°R = +12.5 kJ/mol ΔH° R = -159 kJ/mol ΔH°R = -519 kJ/mol Other important byproducts are methyl formate, methane, and formic acid. The endothermic dehydrogenation reaction (1) , despite it leads to the direct formation of formaldehyde, is not directly carried out in industrial processes. It is always “associated” to reaction 2 for two main reasons: • Endothermic reaction: for high production plants, relevant endotermicity is not usually a desired feature since it requires energy addition in order not to h ave activation problems. This is particularly true for catalytic processes. • Equilibrium effects : reaction 1 is highly temperature -dependent and its equilibrium constant has value greater than 1 only at the T above ~480°C. Thus, at low T it is negatively affect by thermodynamic equilibrium and by kinetic effects. For these reasons, oxygen (air) is provided in order to make reaction 2 possible. H2 combustion is highly exothermic and irreversible: thus, it allows to push CH 2O equilibrium yield (since H 2 is continuously removed by combustion) and to produce an overall exothermic effect. OH CH HCOOH O CH 3 2 2 + ⎯ →  O H O CH O CH 2 2 2 4 + ⎯ →  + 2 2 1 3 H O CH OH CH + ⎯ →  O H O H 2 2 2 2 5.0 ⎯ →  + O H O CH O OH CH 2 2 3 2 3 5.0 + ⎯ →  + 2 4 2 H CO O CH + ⎯ →  O H CO O OH CH 2 2 5 2 3 2 5.1 + ⎯ →  + O H CO O O CH 2 2 6 2 2 + ⎯ →  + 20 The presence of air in the process leads to two major concerns related to safety and selectivity. From the safety point of view, for a process involving O 2 as reactant, the choice of pure oxygen or air is usually related to the potential formation of explosive mixtures inside the process. Industrial processes use air due to its economicity and its heat carrier effect (large amount of inert N 2), despite much higher compression, separation and purge costs respect to pure O 2. According the overall reaction (3), in stoichiometric conditions Methanol feed is ~29.5% , which fal ls inside the explosion limits (at room T and atm. P, LFL=6.7%, UFL=36%) . Methanol must be fed in concentration such that to fall out this range, thus exceeding the UFL or under the LFL. From the selectivity point of view, oxygen must not react with the organic components in the system, mainly CH 2O and CH 3OH, according to undesired further oxidation reactions. The catalyst plays a pivotal role in controlling the rate of oxidation of Methanol selectively to Formaldehyde but not further to the carb oxylic acid, HCOOH, or the final oxidation products, CO/CO 2. Complete Conversion of Methanol (BASF Process) . A mixture of methanol and water is fed into the evaporating column. Fresh process air and, if necessary, recycled off -gas from the last stage of the absorption column enter the column separately . A gaseous mixture of methanol in air is thus formed in which the inert gas content (nitrogen, water, and CO 2) exceeds the explosive limit (Methanol concentration lower than LFL) . The heat required to evaporate the methanol and water is provided by a heat exchanger, which is linked to the first absorption stage of the absorption column . After passing through a demister, the gaseous mixture is superheated with steam and fed to the reactor, where it flows through a 25 -30mm thick bed of silver crystals. The crystals have a defined range of particle sizes and rest on a perforated tray, which is covered with a fine corrugated gauze, thus permitti ng optimum reaction at the surface (adiabatic reactor) . Inlet gas distribution is fundamental, in order to avoid non homogenous catalyst utilization and possible zones of local catalyst superheating; the flow is downward. The bed is positioned immediately above a water boiler (cooler), which produces superheated steam and simultaneously cools the hot reaction gases to a temperature of 150°C . This configuration allows very rapid cooling of the effluent gas, which is important to prevent any possible degradation of CH 2O to CO, favoured at high T. The gas from the cooler passes to the first stage of a four -stage packed absorption column ; Formald ehyde is eluted countercurrent to water or to the circulating formaldehyde solutions whose concentrations increase from stage to stage. The final product contains 40 – 55wt% formaldehyde, as desired, with an average of 1.3 wt% methanol and 0.01 wt% formic acid. The yield of the formaldehyde process is ~ 90%. Some of the off -gas is removed at the end of the fourth stage of the column and is recycled due to its extremely low formaldehyde content. The residual off -gas (4.8 vol% CO 2, 0.3 vol% CO, and 18.0 vol% H 2 as well as nitrogen, water, methanol, and formaldehyde ) is fed to a steam g enerator, where it is combusted. 21 (In a possible alternative configuration, the formaldehyde solution from the third or fourth stage of the absorption tower is recycled to t he evaporator; a certain amount of steam is used in the evaporation cycle. The resulting vapor is combined with the feed stream to the reactor to obtain an o ptimal methanol :water ratio .) The average life time of a catalyst bed depends on impurities such as inorganic materials in the air and methanol feed and on long exposure to excessively high reaction temperatures . Sometimes, H2O is added to feed: it preserves the catalyst, avoids sintering of Ag crystals and slows down the deposition of C; the catalyst has a lifetime of 2 ÷ 4 months, and is regenerated by electrolysis. (Since formaldehyde solutions corrode carbon steel, all parts of the manufacturing equipment that are exposed to formaldehyde solutions must be made of a corrosion -resistant alloy .) • Methanol concentration lower than LFL allows pushing to almost complete conversion (outlet Methanol very low); • No need for Methanol recycle; • Excess air acts as thermal diluent, thus allowing adiabatic conditions inside the reactor; • Excess air and high amount of off -gases (due to complete conversion) leads to additional costs related to the off -gas recycle and off -gas venting and burning. Incomplete Conversion and Distillative Recovery of Methanol . Formaldehyde can be produced by partial oxidation and distillative recovery of methanol. A feed mixture of pure methanol vapor and freshly blown -in air is generated in an evaporator. The resulting vapor is combined with steam and then fed into the reactor. The reaction mixture contains excess methanol and steam . The vapor passes through a shallow catalyst bed of silver crystals or through layers of silver gauze. Conversion is incomplete and the reaction takes place at 590 -650° C, undesirable secondary reacti ons being suppressed by this comparatively low temperature. Immediately after leaving the catalyst bed, the reaction gases are cooled indirectly with water, thereby generating steam. The remaining heat of reaction is then removed from the gas in a cooler and is fed to the bottom of a formaldehyde absorption column. Since Methanol and Formaldehyde are very soluble in water, process water is used to preferentially remove these components. In the bottom section of the column, the bulk of the methanol, water and formaldehyde separate out (water -cooled section , since most of Formaldehyde is here removed and due to its exothermic mixing enthalpy) . At the top of the column, all the condensable portions of the remaining formaldehyde and methanol ar e washed out of the tail gas by countercurrent contact with process water. A 42wt% formaldehyde solution from the bottom of the absorption column is fed to a distillation column equipped with a steam -based heat exchanger and a reflux condenser. Methanol is recovered at the top of the column and is recycled to the bottom of the evaporator. A product containing up to 55 wt% formaldehyde and less than 1wt% methanol is taken from the bottom of the distillation column and cooled. The formaldehyde solution is th en usually fed into an anion -exchange unit to reduce its formic acid content to the specified level of less than 50 mg/kg. The off -gas from the absorption column has a similar composition to that described for the BASF . The off -gas is combusted to generate steam, thus avoiding environmental problems caused by residual formaldehyde. Alternatively, the tail gas from the top of the absorber can be recycled to the reactor. This inert gas, with additional steam, can reduce the excess methanol needed in t he reactor feed, consequently providing a more concentrated product with less expenditure on distillation. The yield of the process is 91 -92 mol %. • Methanol concentration higher than UFL doesn’t allow complete conversion (relevant amount of outlet Methano l); • Need for Methanol recycle; • Excess Methanol acts as thermal diluent; water can be added to raise the amount of inert in the feed thus reducing that of Methanol (thus favoring also the final distillation step and Methanol recycle costs); • Since defect air is used, almost no oxygen is present in the off -gas, with an almost complete dilution in N 2. In order to sustain a combustion process, it is necessary to add additional O 2 and a fuel. As result, off -gas processing and burning has much higher costs. 22 Formox Process In the Formox process, a Fe-Mo metal oxide ( possibly also vanadium oxide) is used as a catalyst for the conversion of methanol to formaldehyde. The Formox process is a two -step oxidation reaction which involves an oxidized (K OX) and a reduced (K red) catalyst : Leading to the overal l reaction: Thus, the catalyst must have a selective oxidation activity and must be able to continuously undergo a redox cycle, i.e. must be easily reoxidized. In the temperature range 270 -400° C, conversion at atmospheric pressure is virtually complete. Conversion is temperature dependent , since at T>400 -470° C formaldehyde oxidation to CO/ CO 2 takes place (M ethanol oxidation is inhibited by wate r vapor ). The methanol feed is passed to a steam -heated evaporator. Freshly blown -in air (and possibly recycled off -gas from the absorption tower ) is mix ed and, if necessary, preheated before being fed into the evaporator. The gaseous feed passes through catalyst -filled tubes in a heat - exchanging reactor. A typical reactor for this process has a shell with a diameter of ca. 2.5 m that contains tubes only 1 -1.5m in length. A high - boiling heat -transfer oil circulates outside the tubes and removes the heat of re