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Chapter 1
Historical Profile
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1. Historical Profile
1.1 Introduction to Ethyl Benzene (EB):
Ethyl benzene is an organic chemical compound which is an aromatic hydrocarbon (HC). Its major use is in the petrochemical industry as an intermediate compound for the production of styrene, which in turn is used for making polystyrene, a commonly used plastic material. Although often present in small amounts in crude oil, ethyl benzene is produced in bulk quantities by combining the petrochemicals benzene and ethylene in an acid-catalyzed chemical reaction. Catalytic hydrogenation of the ethyl benzene then gives hydrogen gas and styrene, which is vinyl benzene. Ethyl benzene is also an ingredient in some paints.
Ethyl Benzene is used almost exclusively as intermediate in the production of styrene monomer. It is produced by liquid phase alkylation’s or vapor phase alkylation’s of benzene with ethylene. Commercial production started in the 1930s and has grown to over 23 million metric annually (MTA)
1.2 History of Ethyl Benzene:
The alkylation of HC with olefins in the presence of AlCl 3 catalyst was first practiced by M.Balsohn in 1879. However, Charles Friedel & James M. Crafts pioneered much of early research on alkylation & AlCl3 catalyst .Over a century later, the process that employ the classic Friedel-Crafts reaction chemistry remain a dominant source of EB. Ethyl benzene was first produced on a commercial scale in the 1930s by Dow Chemical in US and by BASF in the Federal Republic of Germany.
Until 1980s , almost all ethyl benzene was manufactured with an aluminium chloride catalyst using a Fridal Crafts reaction mechanism . A few EB production units employed a different Fridal-Crafts catalyst , boron trifluriede. Small amount of EB also recovered as a by product from mixed xylenes streams using a very intensive distillation process. In 1980s ,the first commercial facility using a zeolite based process and the absence of maintenance nad environmental problems associated with the Fridal-Creafts catalyst havre allowed zeolite catalyst and to completely displace the order catalyst in all modern production facilities.
The4first zeolite process was based on vapor –phase reactor at temp. of over 4000C. In this temperature rang, reaction such as isomerisation/cracking and hydrogen transfer produce a number of by products that contaminates the EB product.
Efforts were made to reduce by-product formation by changing reaction condition , but it as not until the advent of liquid phase processes operating at temperatures lower than 2700C that zeolite –catalysed processes were truly capable of producing high purity EB . The first high purity zeolite based on technology developed by UOP and ABB Lummus Global , started up in the 1990.
The Ethyl benzene- Styrene industry remained relatively insignificant until World War 2. The tremendous demand for synthetic SBR during world war prompted accelerated technology improvements and tremendous capacity expansion. This enormous wartime effort led to the construction of several large scale factories, turning styrene production quickly in to a giant industry. In 1965, 10% of EB production was from super fraction of mixed xylene streams produced by catalytic reforming of naphtha. In 1986, the world annual production capacity was 14*106 t.
EB was first produced on a commercial scale in the 1930s by Dow Chemical in US and by BASF in the Federal Republic of Germany.
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1.3 Natural Occurrences:
Ethyl benzene is a colorless, flammable liquid that smells like gasoline. It is naturally found in coal tar and petroleum and is also found in manufactured products such as inks, pesticides, and paints.
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Chapter 2 Applications
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2. Applications 2.1 Current Applications: All commercial EB production is captive consumed for the manufacture of styrene monomer Styrene is used in the production of poly styrene and a wide range of other plastics. Of the other minor applications, the most significant in the paint industry as a solvent. Even smaller volumes go towards the production of acetophenon, diethyl benzene & ethyl anthraquinon. 2.2 Product Specification The product specification on EB is set to provide a satisfactory feedstock to the associated styrene unit. Objectionable impurities in the EB can be grouped in to two categories: a) Haliedes b) Diethylbenzene
Purity 99.5 wt, % min
Benzene 0.1-0.3 %
Toluene 0.1-0.3%
O-xylene +Cumene 0.02% max
m,p- xylene 0.2% max
Allylbenzene 0.2% max
Diethylbenzene 20mg/kg max
Total Chlorides 1-3 mg/kg max
Total organic sulpher 4 mg/kg max
Reactivity density @ 15o c 0.869-0.872
APHA colour 15 max
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Manufacturing Capacities:
PRODUCER CAPACITY* BP Chemicals, Texas City, TX 1,100 Chevron, St. James, LA 1,800 Cos-Mar, Carville, LA 2,200 Dow, Freeport, TX 1,900 Huntsman, Odessa, TX 350 Lyondell Chemical, Channelview, TX 3,000 Nova, Bayport, TX 1,400 Sterling, Texas City, TX 2,000 Westlake, Lake Charles, LA 380 Total 14,130
*Millions of pounds per year of Ethyl benzene (EB).
In India: From Chemical Weekly Buyers Guide 2005- Vol.3 Pashim Petrochem Ltd.
Amol Chemicals & Polymers
Chemaroma Drug House
Chemico
Forum Enterprizes
Ganesh Trading Co.
Manish Chemicals
Mikit Chemicals
Om Chemi Pharma
Perfect Chemicals
Solvchem
Tanay Corporation
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Chapter 3 Economic Scenario
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3. Economic Scenario 3.1 Current & Projected Demand: Demand ( till 2004): 1999: 13,193 million pounds 2000: 13,444 million pounds 2004: 14,552 million pounds Growth Historical (1995 - 2000): 0.7 percent per year. Strength EB demand runs parallel to that of styrene, and styrene is a mature and stable commodity, used in many homopolymer, copolymer and terpolymer applications. These applications cover a wide scope in industrial, consumer and medical products. Weakness EB’s major shortcoming is that it is essentially a one market segment product - styrene. Moreover, as the styrene is produced with captive EB there is not much noticeable market activity. 3.2 Price & Price Variations: PRICE
• Historical (1995 - 2000): High, $0.25 per pound, bulk, f.o.b. Houston, TX, list ($ 0.5511 / Kg)
� By Friday 8 August, 2008: EB was valued at $1,554-1,570/tonne FOB NWE. ($ 1.554—1.570 / Kg) Price Variations: The US unit sales value of EB in $/Kg from 1960 – 1986: Year Sales ( $/Kg)
1960 0.13
1965 0.09
1970 0.09
1973 0.11
1974 0.37
1975 0.20
1978 0.24
1979 0.35
1980 0.51
1983 0.50
1986 0.48—0.51
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3.3 Application wise Consumption Pattern: Application Consumption ( in % )
Styrene Production 99 %
Paint industry as a solvent < 1%
Production of Acetophenon, Diethylbenzene, Ethyl anthraquinon
< 1%
3.4 Site Considerations: Before any site selection work begins the company should be organized for expansion planning in a way that depends on the size of the firm. A company may want to utilize a standing committee, a special project team or planning by one person. In any event the planning function must be clear cut responsibility of one individual. Site Selection Factor: 1. Markets 2. Work force 3. Unionization 4. Transportation 5. Energy 6. Business climate 7. Water and waste systems 8. Living conditions 9. Topography The basic aim of the site selector is to choose a location that maximizes income and minimizes cost compromises usually have to be made. No site is ever perfect, and it is the mission of the site selection team to weigh the alternatives and compromises on the best choice. Plant layout: Plant layout involves developing physical equipment for a processing facility. The development must effect a balance of equipment spacing and integration of specific systems related to facility as a whole. Some of the factors to be considered for designing the plant layout are: 1. Process 2. Economics 3. Client requirements 4. Operation 5. Erection and maintenance 6. Safety 7. Environment 8. Appearance 9. Expansion
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In-line plant layouts are made in various arrangements which often are referred to by letter designation. Various configurations are formed based on the main artery of the process unit i.e. the pipe rack, which contains long process and the utility lines that connect distant equipment and product piping entering and leaving the plant. Space for instrument and electrical feeders is allocated in the pipe rack such that they are connected to the related equipment. This area is kept free of piping and its related supports. Generally an I shaped plot is used for small process and an H-shape plot for larger units. In developing the plant layout for a chemical plant, it is essential that the firm decisions are made early as to equipments arrangement. This eliminates changes, which cost man-hours as the job progresses through engineering and design. The distillation sections are based on a grade-level process plant layout configuration. The steam generation and power facilities are housed in a building. The basic arrangement follows the equipment spacing charts and clearance tables. Based on all above factors we have selected Bombay High as our plant site due to good availability of raw materials as well as for better market.
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Chapter 4 Properties
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4. Properties 4.1 Physical Properties: Under ordinary conditions, EB is a clear liquid with a characteristic aromatic odor. EB is an irritant to the skin & eyes and is ordinary toxic by ingestion and skin adsorption. The properties are as follows: IUPAC Name Ethylbenzene
Other names Ethylbenzol, EB, Phenylethane
Identifiers
CAS Number [100-41-4]
RTECS Number DA0700000
SMILES c1ccccc1CC
Molecular Formula C8H10
Molar Mass 106.167 g/mol
Appearance Colourless liquid
Density (at 150 c) 0.87139 g/cc
Density ( at 200 c) 0.8669 g/cc
Density ( at 250 c) 0.86262 g/cc
M.P. -94.949 0 c
B.P. ( at 101.3 K Pa) 136.860 c
Refractive index (at 200 c) 1.49588
Refractive index ( at 250 c ) 1.49320
Critical pressure 3609 K Pa (36.09 bar)
Critical temperature 344.02 0 c
Flash point 150 c
Auto ignition temperature 4600 c
Flammability limit ( lower) 1.0%
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Flammability limit ( Upper) -----
Latent heat ( fusion) 86.3 j/kg
Latent heat ( vaporization) 335 j/kg
Heating value ( gross) 42999 j/kg
Heating Value ( net) 40928 j/kg
Kinematic viscosity (at 37.8 0 c ) 0.6428 * 10 -6 m2 / s
Kinematic viscosity (at 98.90 c ) 0.390 * 10 -6 m2 / s
Surface Tension 28.48 mN/m
Specific heat capacity ( ideal gas, 250 c )
1169 kg-1 K-1
Specific heat capacity ( liquid, 25 0 c)
1752 kg-1 K-1
Acentric factor 0.3011
Critical compressibility 0.264
LEL 1.2 %
UEL 6.8 %
4.2 Chemical Properties: The most important chemical reaction of EB is its dehydrogenation to Styrene. The reaction is carried out at:
1. High temprature ( 600-6700 c) 2. Usually over an iron oxide catalyst 3. Steam is used as dilutent
Commercially, selectivities to styrene range from 89-96%with per rpass conversion 65-70%
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. EB Styrene
� Another reaction of commercial importance is the oxidation of EB by air to hydroperoxides C 6H5 CH(OOH)CH3 .
The reaction takes place in the liquid phase , with no catalyst required.
� With a suitable catalyst EB can be converted to xylenes. Commercial processes for isomerizing xylenes usually involve the catalytic isomerization or dealkylation of EB.
� EB may be dealkylated catalytically or thermally to benzene. 4.3 Environmental & Health Effects: Emergency Overview: Clear, colourless liquid with a characteristic, sweet, gasoline-like, aromatic odour. FLAMMABLE LIQUID AND VAPOUR. Liquid can accumulate static charge by flow, splashing and agitation. Vapour is heavier than air and may spread long distances. Distant ignition and flashback are possible. Liquid can float on water and may travel to distant locations and/or spread fire. Closed containers may rupture and explode in heat of fire. TOXIC. May be harmful if inhaled. Central nervous system depressant. Vapour may cause headache, nausea, dizziness, drowsiness, confusion, unconsciousness and possibly death. SKIN IRRITANT. May cause skin irritation. Aspiration hazard. Swallowing or vomiting of the liquid may result in aspiration (breathing) into the lungs. POSSIBLE CANCER HAZARD - may cause cancer, based on animal information. Potential Health effects: Effect of short Term Exposure:
Inhalation: Ethylbenzene readily forms high vapour concentrations and should be considered toxic by this route of exposure. Inhalation of the vapour or mists can irritate the nose and throat and produce symptoms of central nervous system depression such as mild unsteadiness, headache, nausea, dizziness and a feeling of drunkenness at approximately 100-200 ppm. Much higher concentrations can cause more severe symptoms including unconsciousness and death. Human volunteers exposed to 85 ppm for 8 hours reported no adverse health effects. Above 100 ppm, mild unsteadiness, sleepiness and headache were reported. In another report, approximately 100 ppm (cited as 400 mg/m3) produced a slight irritating effect on the respiratory tract, occasional headaches, sleepiness, slight drowsiness after 8 hours. More pronounced irritation, frequent headaches, sleepiness and a feeling of drunkenness were observed at 200 ppm (860 mg/m3). At 1150 ppm (5000 mg/m3), irritation of nose and throat was experienced. Exposure to 1000-2000 ppm (0.1-0.2%) for 6 minutes caused irritation of the nose and throat, fatigue and increasing unsteadiness, chest constriction and dizziness in 4-6 male volunteers. Exposure to 5000 ppm (0.5%) was considered intolerable Skin Contact: The liquid can cause moderate irritation, based on animal information. Ethylbenzene is absorbed through the skin to a small extent, but harmful effects are not expected to occur by this route of exposure.
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Eye Contact: The liquid can cause mild to moderate irritation, based on limited animal information. Volunteers reported that exposure to approximately 100 ppm was slightly irritating to the eyes. At 1000 ppm (0.1%) the vapour was very irritating to the eyes of 4-6 volunteers, producing smarting and burning, accompanied by profuse tearing. This irritation, gradually decreased until, after a minute or two, it was barely noticeable. At 2000 ppm (0.2%), the irritation was almost intolerable upon first exposure, but again became less irritating upon continued exposure, while 5000 ppm (0.5%) was considered intolerable. Ingestion: Ethylbenzene has relatively low toxicity following ingestion. As a central nervous system (CNS) depressant, it can cause nausea, vomiting, headache and dizziness. Very large amounts may cause unconsciousness and death. Ethylbenzene can cause severe lung damage or death if the liquid is accidentally breathed into the lungs (aspirated), based on physical properties. There are no reports of aspiration occurring in humans. Ingestion is not a typical route of occupational exposure. Effects of Long-Term (Chronic) Exposure: Nervous System: A number of human population studies involving painters and other occupational groups exposed to a wide range of solvents, including ethylbenzene, have led to some investigators to conclude that long-term exposure to solvents may cause permanent effects on the central nervous system (CNS). The signs and symptoms are ill-defined and include headaches, memory loss, fatigue and altered emotional reactivity. This syndrome is commonly known as Organic Solvent Syndrome. There are no specific studies that implicate ethylbenzene as a causal agent, although it is present in many of the paints and other solvent-containing products. The available studies tend to have a number of deficiencies including concurrent exposure to many different chemicals, and lack of exposure data. In a limited study, most workers exposed to up to 11.5 ppm (cited as 0.05 mg/L) complained of headaches, irritability and of tiring rapidly. Functional nervous system disturbances were found in some workers employed for over 7 years. Skin: Repeated or prolonged contact may cause dry, red, chapped skin (dermatitis). Skin Sensitization: No allergic skin reaction was observed among 25 volunteers exposed to 10% ethylbenzene in petrolatum. Hearing: Studies in rats have shown that simultaneous exposure to ethyl benzene and noise increases the potential for hearing damage above that for noise exposure alone. Guinea pigs do not appear to be sensitive to these effects.(54,55) The relevance of these observations to human exposures is not known. 4.4 Handling, Storage & Transportation:
� EB is an flammable liquid. � It is stored & transported in steel containers . � The US DOT identification number is UN 1175 . � Foam, Carbon dioxide , dry chemical , halon & water ( fog pattern) are used in fighting EB fires. � The use of NIOSH approved respirators is recommended at high concentration. � Skin contact should be avoided.
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Chapter 5 Manufacturing Processes
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5. Manufacturing Processes:
Currently, the primary source of ethyl benzene is the alkylation of benzene with ethylene. The only other source, the super fractionation of mixed C8 aromatic streams, supplies only a small portion of the ethyl benzene produced. Two distinct types of ethyl benzene alkylation processes are currently used commercially: liquid- phase alkylation and vapor-phase alkylation.
5.1 Liquid Phase Alkylation:
Liquid phase aluminum chloride processes have been the dominant source of ethylbenzene since the 1930s. Several companies have developed variations of this technology. Processes currently in use include those of Dow chemical, BASF, shell chemical, Monsanto, societe chimique des cahrbonnages, and union carbide/ badger. The Monsanto process is currently the most modern commercially licensed aluminum chloride alkylation technology. Alkylation of benzene with in the presence of an aluminum chloride catalyst complex is exothermic (∆H-114 kJ/mol); the reaction is very fast and produces almost stoichiometric yields of ethy lbenzene. In addition to AlCl3, a wide range of Lewis acid catalysts, including AlBr3, FeCl3, and BF3, have been used. Aluminum chloride processes generally use ethyl chloride or hydrogen chloride as a catalyst promoter. These halide promoters reduce the amount of AlCl3 required. The reaction mechanism has been studied in detail
Alkylation:
In the conventional AlCl3 process (see Fig 1), three phases are present in the reactor. Aromatic liquid, ethylene gas, and a liquid catalyst complex phase (a reddish brown material called red oil). A mixture of catalyst complex, dry benzene, and recycled polyalkyl benzenes is continuously fed to the reactor and agitated to disperse the catalyst complex phase in the aromatic phase. Ethylene and the catalyst promoter are injected into the reaction mixture through spargers, and essentially 100% of the ethylene is converted. Low ethylene: Benzene ratios are used to give optimum overall yield of ethylbenzene. Commercial plants typically operate at ethylene: because molar ratios of ca.0.3-0.35. As the ratio is increased, more side reactions, such as transalkylation and isomeric rearrangement, occur. Further alkylation of ethylbenzene leads to the reversible formation of lower molecular mass polyalkylbenzenes. The loss in net yield due to residue is minimized by recycling this material to the alkylation reactor. In addition, because the reaction occurs close to thermodynamic equilibrium, the traditional processes use a single reactor to alkylate benzene and transalkylate polyalkylbenzenes.
The reaction temperature is generally limited to 1300C; a higher temperature rapidly deactivates the catalyst and favors formation of non aromatics and polyalkyllbenzenes, which are preferential absorbed by the highly acidic catalyst complex, resulting in byproduct formation. Sufficient pressure is maintained to keep the reactants in the liquid phase. High –alloy materials of construction are also required for the piping and handling systems. The liquid reactor effluent is cooled and discharged into a settler, where the heavy catalyst phase is decanted from the organic liquid phase and recycled. The organic phase is washed with water and caustic to remove dissolved AlCl3 and promoter. The aqueous phase from these treatment steps in first neutralized and then recovered as a saturated aluminum chloride solution and wet aluminum hydroxide sludge.
Removal of dissolved catalyst from the catalyst from the organic stream has long been a problem for ethylbenzene producers. Recently CdF chime found that more complete recovery of AlCl3 could be achieved by first contacting the organic phase with ammonia instead of sodium hydroxide.
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Separation: Purification of the ethyl benzene product is usually accomplished in a series of three distillation columns. The unreacted benzene is recovered by the first columns as an overhead distillate. The second column separates the ethyl benzene product from the heavier polyalkylated components. The bottoms product of the second column is fed to a final column, where the recyclable polyalkylbenzenes are stripped from non recyclable high molecular mass residue compounds. The residue or flux oil, consisting primarily of polycyclic aromatics, is burned as fuel. Because the alkylation mixture can tolerate only minor amounts of water, the recycled benzene and fresh benzene must be dried thoroughly prior to entering the reactor. Water not only increases corrosion, but also decreases catalyst activity. Benzene dehydration is accomplished in a separate column.
The improved Monsanto process has distinct advantages compared to conventional AlCl3 processes. The most important of these is a significant reduction in the AlCl3 catalyst use, thus lessening the problem of waste catalyst disposal. Monsanto found that by an increase in temperature and by careful control of ethylene addition, the required AlCl3 concentration could be reduced to the solubility limit, thereby eliminating the separate catalyst complex phase.. There fore, alkylation occurs in a single homogeneous liquid phase instead of the two liquid phases is earlier processes.
Monsanto claims that a separate catalyst complex phase may actually prevent the attainment of maximum reactor yields. With a few exceptions, the flow scheme of the Monsanto process is nearly the same as that of more traditional processes. The process is also capable of operating with low- concentration ethylene feed. The process is also capable of operating with low concentration ethylene feed. Typically, the alkylation temperature is maintained at 160-1800C. This higher operating temperature enhances catalyst activity, with the additional benefit that the heat of reaction can be recovered as low- pressure steam. Whereas the traditional process accomplishes alkylation and transalkylation in a single reactor, the homogenous catalyst system must employ a separate transalkylation reactor. At lower catalyst concentrations, the recycle of substantial amounts of polyalkylbenzenes terminates the alkylation reaction. Therefore, only dry benzene, ethylene, and catalyst are fed to the alkylation reactor. The recycle polyethylbenzene stream is mixed with the alkylation reactor effluent prior to entering the transalkylation reactor. The transalkylation reactor is operated at much lower temperature than the primary alkylation reactor. After transalkylation, the reaction products are washed and neutralized to remove residual AlCl3. With the homogenous process, all of the catalyst remains in solution. The catalyst-free organic reaction mixture is then purified using the sequence described previously for the conventional AlCl 3 process. As with other AlCl3 process, the organic residue is used as fuel and the aluminum chloride waste streams are usually sold, or sent to treatment facilities.
5.2 Vapor Phase Alkylation:
Vapor-phase alkylation has been practiced since the early 1940s, but at that time processes were unable to compete with liquid-phase aluminum chloride based technology. The alkar process developed by UOP, based on boron trifluoride catalyst, had modest success in the 1960s, but fell from favor because of high maintenance costs resulting from the severe corrosion caused by small quantities of water. Nevertheless, some ethylbenzene units continue to use this process.
The Mobil –badger ethylbenzene process represents the latest and most successful vapor phase technology to be introduced. The process was developed in the 1970s around Mobil’s versatile ZSM-5 synthetic zeolite catalyst. Earlier attempts at using zeolite or molecular sieves for benzene alkylation had suffered from rapid catalyst deactivation because of coke formation and poor transalkylation capabilities. The Mobil catalyst combines superior resistance to coke formation with high catalytic activity for both alkylation and transalkylation by American Hoechst Corp. at their 408x10-3/t/a Bayport, Texas plant. Currently nine commercial plants have been licensed, representing ca. 3x10106 t/a of production capacity.
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ALKAR PROCESS:
This process produces a high purity ethylbenzene product and can use dilute ethylene feed stock. If the entry of water into the process is strictly prevented, the corrosion problems associated with aluminum chloride processors are avoided. However, even small amounts of water (<1mg/kg) hy6drolze the BF, catalyst. The alkyation reaction takes place at high pressure (2.5-3.5 Mpa; 25-35 bar) and low temperature (100-1500C). Dehydrated benzene, ethylene, and make up BF3, catalysts are fed to the reactor. Typically, ethylene; benzene molar ratios between 0.15 and 0.2 are used. The reactor inlet temperature is controlled by recycling a small portion of the reactor effluent.
Transalkylation takes place in a separate reactor. Dry benzene, BF3 catalyst, and recycled polyethlybenzene are fed to the transalkylation reactor. The effluent streams from the two reactors are combined and passed to a benzene recovery column, where benzene is separated for recycle to the reactors. Boron trifluoride and light hydrocarbons are taken over head as a vapor stream from which the BF3, is recovered for recycle. The bottom for the benzene recovery column is sent to a product column, where ethylbenzene of > 99.9% purity is taken overhead. A final column serves to recover polyethylebenzenes for recycle to the transalkylation reactor.
The alkar process can operate with ethylene feed containing as low as 8-10 mol% ethylene, enabling a variety of refinery and coke-oven gas streams to be used. However, purification of these streams is necessary to remove components that poison the BF3 catalyst, e.g., trace amounts of water sulfur compound, and oxygenates.
Mobil-Badger Process:
The fixed –bed ZSM-5 catalyst promotes the same overall alkylation chemistry as those used in the other processes; however, the reaction mechanism is different. Ethylene molecules are adsorbed onto the Bronsted acid sites within the catalyst, which activates the ethylene molecule and allows bonding with benzene molecules to occur. Hence, the range of higher alkylated aromatic byproducts formed by the Mobil – Badger process is some what different than that for the Friedel Crafts processes. These components do not affect the ethylbenzene product purity and are recycled to the reactor for transalkylation or dealkylation.
The Mobil-Badger heterogeneous catalyst system offers several advantages when compared to the other commercially available processes. The most important are that it is noncorrosive and nonpolluting. The catalyst is essentially silica – alumina, which is environmentally inert. Because no aqueous waste streams are produced by the process, the equipment for waste treatment and for catalyst recovery is eliminated. In addition, carbon steel is the primary material of construction, high-alloy materials and brick linings are not required.
The reactor typically operates at 400-4500C and 2-3 Mpa (20-30 bar). At this temperature >99% of the net process heat input and exothermic heat of reaction can be recovered as steam. The reaction section includes two parallel multibed reactors, a fired heater, and heat recovery equipment. The high-activity catalyst allows transalkylation and alkylation to occur simultaneously in a single reactor. Because the catalyst slowly deactivates as a result of coke formation and requires periodic regeneration, two reactors are included to allow uninterrupted production: one is on stream while the other is regenerated. Regeneration takes ca. 36h and is necessary after 6-8 weeks of operation. The catalyst is less sensitive to water, sulfur, and other poisons than the Lewis acid catalysts.
The reactor effluent passes to the purification section as a hot vapor. This steam is used as the heat source for the first distillation column, which recovers the bulk of the unreacted benzene for recycle to the reactor. The remaining benzene is recovered from a second distillation column. The ethybenzene product is taken as the overhead product from the third column. The bottoms product from this column is sent to the last column, where the recyclable alkylbenzenes and polyalkylbenzenes are separated from heavy nonrecyclable residue. The low-viscosity residue stream, consisting mainly of diphenylmethane and diphenylethane, is burned as fuel. The Mobil-Badger process also
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can use dispute ethylene feedstocks. In semi commercial applications, the process has operated on streams containing as little as 15 mol% ethylene.
5.3 New Developments:
Dow Chemical and Snamprogetti are developing a process for making ethylbenzene/styrene from ethane and benzene. The process combines the dehydrogenation of ethane and ethylbenzene in one unit and integrates the processes for preparing ethylene, ethylbenzene and styrene. This process is claimed to have lower costs than the conventional route to styrene, largely stemming from the low cost of ethane in relation to ethylene. A pilot plant has been operating since 2002 and commercialisation could be possible by the end of the decade.
5.4 Comparison between Processes:
Although both the alkylation process i.e. liquid phase & vapor phase are of equal use commercially. Yet there are some differences:
1. In vapor phase alkylation, the reactors operates at higher temperatures(400—4500 c) which causes catalytic deactivation by fouling as a result catalyst required periodic regeneration.
2. In vapor phase alkylation process two reactors will be required so that processing and regeneration can proceed alternatively without interrupting production.
3. All the ethylene feedstock is reacted completely in the liquid benzene, thus eliminating off gas recovery equipment.
4. Ethyl benzene yield is 99.7% in liquid phase alkylation process while in vapor phse alkylation it is around 98%.
5. Zeolite as a catalyst can be used in any of the processes.
6. Up to 99.95wt% product purity in the with no xylene formation, in liquid phase alkylation process.
HAZARDS IDENTIFICATION
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Chapter 6 Selected Process
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6. Selected Process:
We have selected liquid phase alkylation process for ethyl benzene production, now we will discuss process under the followings:
Introduction to process:
Efforts were made to reduce by-product formation by changing reaction condition but it was not until the advent of liquid phase at temperature lower than 2700 c that zeolite –catalyzed processes were truly capable of producing. The first high purity zeolite based EB plant , based on technology developed by UOP & ABB Lummus Global started up in 1990.
Technology Supplier: UOP & ABB Lummus Global
Current Status: Currently 16 plants are using this technology.
Raw Materials: Ethylene & Benzene
Product quality achievable: 99.95wt% pure EB can be produced.
Catalyst: Zeolite
Material of construction: Carbon Steel
Process Effluents: Inert component of ethylene feed which will appear as benzene column vent.
6.1 Process Chyemistry:
EB made by the alkylation of benzene with ethylene in the presence of zeolite catalyst. Successive alkylation also occurs to minor extent, producing diethylbenzens, collectively termed polyethylene benzene (PEB)
Benzene + Ethylene → EB (Ethyl benzene)
EB + Ethylene →DEB (Di- ethyl benzene)
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6.2 Process Description
Benzene alkylated with the ethylene to yield a mixture of alkylated benzenes. This mixture is distilled to recover product EB, and higher ethylated benzenes (PEB).
The liquid phase alkylation reactor consists of multi[le beds of zeolite catalyst operating adiabatically. Process conditions are selected to keep the aromatic reaction mixture in the liquid phase. Excess benzene is used, and ethylene is injected before each bed. Multiple ethylene injection points improve selectively and enhance catalyst stability. In the alkylation reactor, ethylene reacts completely, leaving only the inert constituents of the feed, such as ethane. These inters pass through the reactor and are from the plant at a convenient point.
The alkylation effluents are fed to the benzene column, where benzene is taken as the over head product for recycle to the reactor. The benzene column bottoms feed the ethyl benzene column. Here EB is taken as the overhead product.
The reboiler of the distillation columns may used hot oil, high-pressure steam, or direct firing. Overhead vapors are condensed in waste heat boilers, generating valuable steam useful in a downstream SM or propylene oxide/styrene monomer plant. The EB unit has considerable flexibility to mmet a verify of local site conditions in an efficient manner. If no stream export is required, the net heat import can be reduced considerably.
6.3 Plant Capacity:
The Ethyl benzene plant capacity is 1000MTPY(109 kg per year) based on:
• Current demand and supply data for EB.
• As well as other capacities of other Ethyl benzene operating plants (or projected plants).
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6.4 Process Flow Diagram:
H-1 Heater
R-1 Reactor ( 270 0 c, 30-40 atmp)
C-1 Cooler
L-1 Light Column
T-1 Benzene Column
T-2 Ethyl Benzene Column
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Chapter 7 Mass & Energy Balances
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Mass & Energy Balances
7.1 Process information:
1. Stream (1) – Benzene is pure 2. Stream (2) – ethylene contains 7 mole % ethane as impurity. Ethane dose not react but moves in the process as inerts and vent in the light column. 3. Stream (3) – B: E ratio is adjusted to control reaction selectivity , 8:1. 4. Reactor (R-1) – the limiting reactant achivees 100% conversion. 5. Stream (8) –composition is; Inerts – 45% mole Benzene - 55% mole
6. Efficiency of T – 1 is such that 99.9% of benzene that is fed is in the overhead. 7. Stream(11) –
Ethyl benzene = 99.9% wt Benzene = 0.1% wt
8. All the EB fed in to the EB column (T—2) is in overhead.
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7.2 Mass balance Calculations:
The plant capacity for ethyl benzene plant = 1000 MTPA
For which we are suppose to produce the product as follows:-
The product ethyl benzene which we will produce will contain:-
1190 kmol/hr ethyl benzene
and 1.6 kmol/hr benzene
Since we are using liquid phase alkylation process for ethyl benzene production, for which yield and conversion are as follows:
Yield=99.7%
Conversion=100% (w.r.t ethylene)
Now consider PFD for ethyl benzene plant
Stream (11)
Ethyl benzene=1190 kmol/hr
Benzene = 1.6 kmol/hr
Now we are coming to the reaction part, here we have reactor:-
Reactor
R 1
Conversion = 100% B = ?, E= ?, Inerts = ? EB= 1190 kmol/hr,
Inerts=?, B= ?, DEB = ?
28
Now we know that:-
Yield = [{(moles of product produced)*(stochiomertric coefficient)} /(moles of reactant converted)]/100
99.7/100=1190/moles of reactant converted
Moles of reactant converted(ethylene)= 1193.5 kmol/hr
Now the reaction in the reactor are:-
C6H6 + C2H4 C6H5-C2H5
benzene ethylene Ethyl benzene
C6H5-C2H5 + C2H4 C6H4-(C2H5)2
Ethyl benzene ethylene Diethyl benzene
here,
For 1 mole of Ethyl benzene = 1 mole of ethylene needed
and
For 1 mole of diethyl benzene = 2 moles of ethylene needed.
Now after reactor 1190 kmol of ethyl benzene produced
So ethylene consume is = 1190 kmol
Thus rest ethylene is = 1193.5 - 1190
= 3.5 kmol of ethylene
Since conversion is 100%
It means that all ethylene had converted in to ethyl benzene and rest ethylene converted in to diethyl benzene.
Now
2 moles of ethylene will produce = 1 moles of diethyl benzene
So, 3.5 kmol will produce= 3.5/2
= 1.75 kmol/hr of diethyl benzene (DEB)
29
Now , to minimize the production of DEB we have choose high B : E ratio i.e. 8:1
Thus total benzene requirement is = 8 *1193.5
= 9548 kmol/hr
Since Ethylene is not pure it is having only 93mole % ethylene
Thus total ethylene stream (ethylene+ ethane) requirement is :-
=1193.5/0.93
= 1283 kmol/hr
Inert = 1283– 1193.5
= 89.5 k moles/ hr
After reactor will also have unreacted benzene which will be :-
=Benzene fed – benzene consumed
= 9548-1193.5
= 8354.5 kmol /hr (unreacted)
Now consider reactor-1 (R-1) again:-
Inputs:
Benzene = 9548 kmol
Ethylene = 1193.5 kmol
Inert = 89.5 kmol
Outputs:
Ethyl benzene = 1190 kmol
Benzene = 8354.5 kmol
Inert = 89.5 kmol
30
Diethyl benzene = 3.5 kmol
Stream:-(6)
Ethyl benzene = 1190 kmol
Benzene = 8354.5 kmol
Inert = 89.5 kmol
Diethyl benzene = 3.5 kmol
Reactor
R 1
Conversion = 100%
B = 9548 mol , E=1193.5 mol , Inerts =89.5 mol
EB= 1190 mol, B= 8354.5 mol, Inert =89.5 mol, DEB = 3.5 mol.
31
Stream 7
Thus Consider Lights Column:
Stream 6
Stream 8
Lights Column
L - 1
EB= 1190 mol, B= 8354.5 mol, Inert =89.5 mol, DEB = 3.5 mol.
Inerts 89.5 mol
EB= 1190 mol, B= 8354.5 mol, DEB = 3.5 mol
32
Material balance of Benzene Column:
Benzene Column
Benzene 8330 mol
EB 1186.5 mol
DEB 3.5 mol
Benzene 4.15 mol
EB 1186.5 mol
DEB 3.5 mol
Benzene 8325.85 mol
EB 0.6 mol
33
Now consider Ethyl benzene (EB) column:
Stream no :(10)
Stream 10 will contain:
EB= 1186.5 mol DEB= 3.5 mol Benzene= 4.15 mol
Now as we already know that stream (11) will be:
Stream (11):
EB = 1186.5 mol
Benzene= 0.15 mol
Stream (12):
DEB = 3.5 mol
Benzene = 4 mol
34
Thus finally EB column is:
Stream 11
Stream 10
Stream 12
Thus finally :
EB Column EB= 1186.5 mol DEB= 3.5 mol Benzene= 4.15 mol
DEB = 3.5 mol Benzene = 4 mol
EB = 1186.5 mol Benzene= 0.15 mol
Ethyl Benzene Plant Input Benzene = 9548 mol Ethylene = 1193.5 mol
Inerts = 89.5 mol
Output EB = 1186.5 mol DEB = 3.5 mol Benzene = 4.15 mol Inerts = 89.5 mol
35
Mass balance at a Glance:
For Reactor
Input :
Stream:
Component Kg/hr (10-3) Kmol//hr
Ethylene (E) 33420 1193.5
Benzene (B) 742560 9520
Inert (I) 2685 89.5
Total 778665
Output:
Stream:
Component Kg/hr (10-3) Kmol//hr
Ethyl Benzene (EB) 125768 1186.5
Di-ethyl Benzene (DEB) 472 3.5
Benzene (B) 649740 8330
Inert (I) 2685 89.5
Total 778665
36
For Light Column:
Input:
Stream:
Component Kg/hr (10-3) Kmol//hr
Ethyl Benzene (EB) 125768 1186.5
Di-ethyl Benzene (DEB) 472 3.5
Benzene (B) 649740 8330
Inert (I) 2685 89.5
Total 778665
Output:
Stream:
Component Kg/hr (10-3) Kmol//hr
Inert (I) 2685 89.5
Total 2685
&
Stream:
Component Kg/hr (10-3) Kmol//hr
Ethyl Benzene (EB) 125768 1186.5
Di-ethyl Benzene (DEB) 472 3.5
Benzene (B) 649740 8330
Total 775980
37
For Benzene Column:
Input:
Stream:
Component Kg/hr (10-3) Kmol//hr
Ethyl Benzene (EB) 125768 1186.5
Di-ethyl Benzene (DEB) 472 3.5
Benzene (B) 649740 8330
Total 775980
Output:
Stream:
Component Kg/hr (10-3) Kmol//hr
Benzene (B) 649416.3 8325.85
Total 649416.3
&
Stream:
Component Kg/hr (10-3) Kmol//hr
Ethyl Benzene (EB) 125768 1186.5
Di-ethyl Benzene (DEB) 472 3.5
Benzene (B) 323.7 4.15
Total 126563.7
38
For Ethyl Benzene Column:
Input:
Stream:
Component Kg/hr (10-3) Kmol//hr
Ethyl Benzene (EB) 125768 1186.5
Di-ethyl Benzene (DEB) 472 3.5
Benzene (B) 323.7 4.15
Total 126563.7
Output:
Stream :
Component Kg/hr (10-3) Kmol//hr
Ethyl Benzene (EB) 125768 .0 1186.5
Benzene (B) 11.7 0.15
Total 125779.7
&
Stream:
Component Kg/hr (10-3) mol//hr
Di-ethyl Benzene (DEB) 472 3.5
Benzene (B) 312 4.00
Total 784
39
7.3 Energy Balance Calculations:
Energy balance across Heater:
Now we perform similar calculations for all the I/P components:
Thus
s.no. Component Mass in
(mol/hr)
Qin
(MJ/hr)
Mass out
(mol/hr)
Qout
(MJ/hr)
1 Benzene (B)
1193.5 0 1193.5 668.758
2 Ethylene (E)
9520 0 9520 19.719
3 Ethane (I) 89.5 0 89.5 1.750
4 Total 0 690.228
Energy balance across Heater:
Q = Q out – Q in
= (Q B + Q E + Q I )out – (Q B + Q E + Q I )in
where QB ,
QB =[ nB ( ∫ c p dT + λ)
Heater
543K
I/P O/P
Stream 3
40
Energy Balance Calculations across Reactor:
Qr = Qout - Qin
= 1305.06 – 1441.28
= - 136.22 MJ/hr
It means in the reactor we are having exothermic reaction.
So we need dowtherm as an utility to make reactor at operating condition, which would be:
Qr = mw * cp * ∆T
mw = 2173 Kg/ hr
Energy balance across Cooler:
Q c = Qout – Qin
= -357.697 MJ
Thus water needed for cooler:
Qc = mw * cp * ∆T
mw = 5677.73 Kg/ hr
Energy balance across Benzene Column:
We have assumed reflux ratio :
R = 1.5
Thus in L:
B = 12494 mol
I = 134.25 mol
& in V:
B = 20823.16 mol
I = 223.75 mol
41
So condenser duty will be:
Qc = QL + QD - QV
= -658.55 MJ /hr
Thus requirement of water as utility will be:
Qc = mw * cp * ∆T
mw = 10453 Kg/ hr
Now reboiler duty will be:
QR - QC = QD + QB - QF
QR = 699.06 MJ/hr
Thus steam requirement will be:
ms = 276.09 Kg/hr (at 90 atmp. & 577 K)
Energy Balance across Ethyl Benzene Column:
We have assumed reflux ratio :
R = 1.5
Thus in L:
EB = 1779.75 mol
B = 1.2495 mol
& in V:
B = 2966.13 mol
I = 2.0825 mol
So condenser duty will be:
Qc = QL + QD - QV
= -117.12 MJ/hr
42
Thus requirement of water as utility will be:
Qc = mw * cp * ∆T mw = 1868.38 Kg/ hr
Now reboiler duty will be:
QR - QC = QD + QB - QF
QR = 63.167 MJ/hr
Thus steam requirement will be:
ms = 24.95 Kg/hr
Thus Plant Utilities at a glance:
s.no. Equipment Utility Amount
(Kg/hr)
1 Heater Steam 272.5
2 Reactor Water 2173
3 Cooler Water 5677.73
4 Benzene Column
(Condenser)
Water 10453
5 Benzene Column
(Reboiler)
Steam 276.09
6 EB Column
(Condenser)
Water 1868.38
7 EB Column
(Reboiler)
Steam 24.95
43
Chapter 8 Detailed Equipment Design
44
8. Detailed Equipment Design
8.1 Fixed Bed Catalytic Reactor Process Design:
Feed composition component Kg/hr
Benzene 742.56
Ethylene 33.42
Inerts 2.685
Total 778.665
Feed flow rate in reactor = 778.665 Kg/hr.
Reaction: C6 H6 + C2 H4 C6 H5 – C2 H5
Reaction temperature =270 0 C ( 543 K)
Catalyst: Zeolite
• Density of catalyst particle: 630 kg/m3 � Catalyst porosity : 0.3
� Type of reactor: Shell and tube heat exchanger type in which catalyst is placed inside the tube.
� Reaction is exothermic: It is carried out in isothermal manner. Dowtherm A is to be circulated in liquid form on shell side to maintain the isothermal condition.
� Mass of catalyst required in commercial scale plant = 2541 kg
� Superficial mass velocity of feed gas G=0.55 kg/m2.s
� Capacity of plant =1000 MTA of Ethyl benzene
Let no. of working days per annum=330 days
Production rate of ethyl benzene=(1000*1000)/(330*24)
=126.2626 Kg/h
Mass of catalyst;
W/FA0= ………………………………………………………..(1)
-ra = KrCE/(1+KEB.CEB)
45
Where;
Kr =0.69*106exp[(-6.344*104)/RT]
KEB= -1.5202*10-2exp[(-3.933*103/RT]
Putting all the values in equation (1)
W = 2541 Kg
Volume of solid catalyst = mass of catalyst/ density of catalyst
=2541/630
= 4.03 m3
Porosity of catalyst bed = 0.3
Bulk volume occupied by the catalyst= 4.03/(1-0.3)
= 5.76 m3
Total cross sectional area of catalyst of tubes
= (778.3/3600)/0.55
= 0.3935 m2
MOC of tube = Stainless steel
Tube OD = 50.8mm
Tube ID = 43.28 mm
Total number of tube required
nt = 0.3935/(∏/4)(0.04328)2 =267.48
=268 tubes
Length of tube required
L= net vol. of catalyst/nt (∏/4)di 2
= 4.03 *4/268*3.14*(.04328)2
=10.22m
=10m
Aav = nt ∏d0 L
= 268*3.14*0.0508*10
46
=427.49 m2
Heat duty= ∆HR Kmol/hr of ethylene consumption
Qt =114123*33.42/28
= 37.837 KW
Calculation of fixed bed side film coefficient hi ;
dp=equivalent dia. Of catalyst cylindrical partical,
∏ dp3 / 6 = (∏/4)(1.5)2 * 5*10-9
dp=0.002565m
d p/d t =0.059 ( d t = tube diameter)
(hidt /k) =0.813 e-6dp/dt (dp G/µ)0.9
G= 0.55 kg/m2 .s
k= 0.04 W/m.k
µ=0.015mPa
On Calculation ;
hi =31.49w/m2 C
hi dp/k = 3.6(dpG/µε)0.365
hi*0.002565/0.04 =3.6(0.002565*0.55/0.02*10-3*0.3)
hi =457.51W/m2 0C
Let hi=31.49 W/m2 0C (Lesser of two values)
Calculation of shell side heat transfer, h0 ;
Tube pitch , Pt =1.25d0
=1.25*0.0508
= 0.0635 m
Type of arrangement = Equilateral triangular
Equivalent diameter de = 1.1/d0 (Pt – 0.907 d02 )
= 1.1/0.0508[0.06352 – 0.907*0.05082]
47
= 0.03663 m
Shell side mass flow rate m ;
Qt =m*CL∆t
Properties of Dowtherm A at 270 0C
Property Value CL 2.5832 KJ/kg 0C µ 0.135 mPa .s
K 0.098 W/m 0C Let
∆t=2 0C
m=37.837/2.5832*2)
= 7.323 Kg/s
Density of Dowtherm A at 270 0C,
Ρ =0.709 Kg/L
Circulation rate qv =7.323/0.709*(3600/1000)
= 37.18 m3 /hr
Shell side flow area As:
As= (Pt –d0 ) Ds Bs /Pt
Shell inside diameter, Ds :
Db =d0 (Nt /K1 )1/n1
K1 = 0.319 and n1 = 2.142
Db = 50.8(268/0.319)1/2.142
= 1171.3mm
Let clearance between shell internal dia and bundle(Db),
48
Db = 15 mm
Let Ds= 1171.13+15= 1186 mm
Baffle spacing
Bs= 0.4D s = 0.4*1186= 474mm
Shell side flow area As = (Pt-d0)DsBs/Pt
=( 0.0635-0.0508)/0.0635(1.186*.474
= .1124 m2
Gs=m/As
= 7.323/.1124
= 65.1512 kg/m2 s
Reynolds number, Re=de*Gs/µ
=(0.03663*65.1512)/0.135*10-3
=17677.90
Prandtl number;
,Pr = CLµ/k
= (2.5832*0.135*10-3*103)/0.098
=3.558
(h0 de/k)=Jh RePr0.33 (µ/µm )0.14
J h = 0.0046 (Heat transfer factor)
( h0 *0.03663)/0.098 =0.0046*17677*3.5580.33
h 0 =336.50 W/m2 0 C
Overall heat transfer coefficient U0 ;
1/ U0 =1/h0+1/h0d +d0 ln(d0 /di)/2Kw+d0 l/di hid +d0 l/di hi
= 1/336.5 + 1/5000 +{ 0.0508ln(0.0508/0.04328)}/(2*16) + (0.0508/0.04328)*(1/5000) + 0.0508/0.04328
U0 = 24.47 W/m2 0C
Shell side pressure drop
49
∆ps =8 Jf(Ds/de)(L/Bs)(ρsus2/2)(µ/µw)0.14 us= Gs/ρs =65.152/3.831
= 8*0.0046*(1.186/0.03663)(10/0.474)(3.831*172/2) = 17.006
= 13.91 KPa
Tube Side pressure drop
∆pt =[{4f(LN p/di)+ 4Np}(ρum2/2]
Re = ρumdi/µ
= (810*0.0014*0.04328)/0.015*10-3
= 3271.968
f = [1.58 ln Re- 3.28 ]-2
= [ 1.58 ln 3271.968-3.28]-2
= 0.0110
Atp =( πdi2/4)(Nt/2)
= 3.14*(0.04328)2/4*(268/2)
= 0.19 m2
Um = mt/ρt A tp
= 778.665/(810*0.19*3600)
= 0.0014 m/s
∆pt =[{(4*0.0110*10*1)/0.04328 +(4*1)}(810*0.00142)/2
= 0.011240 KPa
50
8.2 Mechanical Design for reactor
Shell and tube type reactor
Data;
(a) Shell side Material carbon steel (Corrosion allowance- 3)
Number of shells - 1
Number of passes - 1
Fluid - Dowtherm A
Working pressure - 0.2 N/mm2
Design pressure - 0.25 N/mm2
Inlet temperature - 25 0C
Outlet temperature - 400C
Segmental baffles (25% cut ) with tie rods and spacers
Head
Crown radius -1200 mm
Knuckle radius -120 mm
(b) Tube Side Tube and tube sheet material - stainless steel Number of tubes - 268 Out side diameter - 50.8 mm Length - 10 m Tube pitch - 0.0635 Fluid -Benzene and ethelene Working pressure - 3.2 N/mm2 Design pressure - 3.6 N/mm2 Inlet temp. – 25 0C Outlet temp. – 270 0C Permissible stress - 100.6 N /mm2 Shell thickness; t s =PD/(2fJ+P) = .25*1200/(2*87*.85)+0.52 = 2.02 mm Including corrosion allowance. Use 8 mm thickness Nozzel thickness ( diameter -75mm) tn=PD/(2fJ-P)
51
= 0.25*75/9\(2*87-0.25) = 0.1080 mm Adding corrosion allowance tn = 4mm Head thickness(th) = PCrW/2fJ W=1/4(3+(Rc/R1)
1/2 ) = 1/4(3+(1200/120)1/2) = 1.54 J=1 th=2.431mm adding corrosion allowance th=5.431mm Using thickness same as for shell i.e 8 mm Transverse Baffles Spacing between baffles= 0.4Ds = 0.4*1200 =480 mm Thickness of baffls = 5mm Tube Side Thickness of tube; tf =PD0/(2fJ+P) J=1 (seamless tube) =3.6*50.8/(2*87*0.85)+0.36 = 1.233 mm No corrosion allowance , since the tubes are of stainless steel. Use a thickness of 2 mm
Design of Gasket and Bolt Size
Gasket factor, m= 2.00
Minimum design stress =11.2 N/mm2
Basic gasket seating width – b0
Internal dia. Of Gasket – 1200 mm
External dia. – 1240mm
b0 = ½(1240-1200)/2 = 10mm
Effective Gasket seating width, b= 2.5b01/2
=7.90mm
Minimum bolt load at atm. Condition, Wm1 = 3.14*b*G*Ya
= 3.14*7.9*(1240+1200)/2*(11.2)
52
= 338949.184
At operating condition,
Wm2=3.14(2b)GMP+3.14G2P
=3.14*2*7.9*(1240+1200)/2(0.25)+3.14/4(1240+1200)2/2(0.25)
= 322361.82
Cross section area of bolt
Am1=338949.184/5870 = 57.74 Cm2
Am2= 322361.82/5450= 59.14 Cm2
Number of bolts= 1220/(2.5*10)
= 49 bolts
Diameter of bolts= [(Am2/no. of bolts)*3.14/4]1/2
= [(59/49)/49*(4/3.14)]1/2
= 1.533 cm
Bolt area ,Ab = 2*3.14*YaGN/fa
= (2*3.14*1220**11.2*20)/58
= 292 Cm2
N- Width of Gasket
Ya –Gasket seating stress
G – dia. Of gasket lode reaction
fa – permissible stress.
pitch of bolts= 4.75*18=85.6 mm
pitch circle dia.=(85.6*49)/3.14
=1335mm
Flange Thickness, tf = G(p/kf)1/2+c ……………..(2)
K= 1/[0.3+1.5WmhG/HG]
= 1/[0.3+(1.5*322361.82*(1335-1220)/2]/(3.14/4)*12202*0.25*1220]
= 3.17
53
putting all the values in equation (2)
tf = 38.15
Tube Side:
Thickness of tube;
tf =PD0/(2fJ+P) J=1 (seamless tube)
=3.6*50.8/(2*87*0.85)+0.36
= 1.233 mm
No corrosion allowance , since the tubes are of stainless steel.
Use a thickness of 2 mm
Thickness of nozzles
tn= pD/(2fJ-p)
D = 75 mm ( inlet and outlet)
J= 1 ( seamless pipe)
tn= 3.6*75/(2*95-3.6)
=1.448 mm
Add corrosion allowace of 3mm
tn= 5mm
54
8.3 Process Design of Benzene Distillation Column
In Benzene column, a product stream from reactor is coming which consists of benzene, ethyl benzene & some amount of di-ethyl benzene. Since it is more than two component case, hence we will perform multicomponent distillation in order to separate benzene from mixture and recycle it back to reactor.
Here Benzene is light key component and EB is heavy key component. The column is operating at 1 atm.
Benzene Column
Benzene 8330 mol
EB 1186.5 mol
DEB 3.5 mol
Benzene 4.15 mol
EB 1186.5 mol
DEB 3.5 mol
Benzene 8325.85 mol
EB 0.6 mol
55
Now first of all we will calculate bubble point & dew point temperature.
Bubble point calculation:
We have taken a temperature 100 0 c (373 K).
Component Pi’ mm Hg Ki = Pi’/ 760 xi yi = kixi B 1344.933 1.771 0.875 1.55 EB 261.38 0.344 0.124 0.043 DEB 54.42 0.072 0.001 0.000072 Total 1 More than 1
Now will do trail & error untill we get a value of temperature for which yi = 1.
After some calculations we got a value of T = 84 0 c (357 K), for which:
Component Pi’ mm Hg Ki = Pi’/ 760 xi yi = kixi B 851.79 1.12 0.875 0.98 EB 148.84 0.196 0.124 0.02 DEB 27.61 0.036 0.001 0.00003 (=0) Total 1 1
Thus bubble point will be 357 K.
Dew point calculations:
Similarly we can calculate dew point as follows:
At T= 112 0 c ( 385 K)
Component Pi’ mm Hg Ki = Pi’/ 760 yi yi/ ki B 1171.84 1.54 0.875 0.556 EB 220.60 0.29 0.124 0.427 DEB 44.35 0.058 0.001 0.017 Total 1 1
The values of Pi can be determined by following relation between T & P;
ln P = A – B/ (T + C) (where P is in mm Hg & T is in Kelvin.)
The various values of A,B,C are:
Component A B C B 15.9008 2788.51 -52.36 EB 16.0195 3272.47 -59.95 DEB 16.1140 3757.22 -71.18 *All the values of Antonie’s cofficient of A,B,C are taken from Coulson & Ricardson Vol. 6, Appendix D, page no. 947 - 967.
56
Calculation of Minimum reflux ratio (R m):-
For a liquid feed, q = 1
∑[(αi xid)/( αi – φ)] = 0
(αi can be calculaed as αi = ki/khk which is 5.714, 1, 0.1836 correspondingly for B, EB, DEB.)
Component Xf Xd K i αi
B 0.8750 0.9999 1.12 5.714 EB 0.1246 0.0001 0.196 1.000 DEB 0.0004 0.0000 0.036 0.1836
From trial & error,
Φ = 1.52
∑[(αi xid)/( αi –φ )] = Rm + 1
Rm + 1 = 1.242 (putting the value of φ = 1.52 in above equation.)
Rm = 0.242
Calculation of Operating Reflux Ratio:-
Now operating reflux will be :
Ro = 1.5 Rm = 0.363
Calculation of number of ideal plates at operating reflux:
By Gilliland correlation: (Fig 5 in appendix)
Now to calculate Nmin :-
57
Nmin = [ln{(xdi/xbi)/(xdj/xbj)}] / [ln ( αav)]
Putting all the values ,
Nmin = 8
Putting this value of Nmin in (a),
We have,
N = 19
Thus we need 19 stages ideally.
Flow Rates:
Average molar mass of feed.
Mav = ∑xi Mi
= (78*0.875) + (106*0.124) + (134*0.001) = 81.528 kg/kmol
F = 9.52 kmol/hr
D = 8.33 kmol/hr
B = 1.194 kmol/hr
Now molar flow rates of vapor & liquid at top in enriching section:
L = R * D = 0.363 * 8.33 = 3.0225 kmol/hr
V = (R+1) * D = (0.363 +1) * 8.33 = 11.349 kmol/hr
Molar flow rates of vapor & liquid in stripping section:
L’ = L + F * q = 3.0225 + 9.52 * 1 = 12.543 kmol/hr
V’ = F*(q-1) + V = 9.52*(1-1) + 11.349 = 11.349 kmol/h
58
Calculation of tower diameter:
(a) Tower diameter required at top:
Operating pressure at the top of column = 1 atm = 101.325 kpa
V = 11.349 kmol/hr
L = 3.0225 kmol/hr
Here total condenser is used,
hence Lw/Vw = L/V = 3.0225/11.349 = 0.266 m
Density of vapor:
ρv = (p* Mav)/(R*T) = (78*273)/(357*22.414) = 2.66 kg/m3 ( Mav = ∑xi Mi = 78)
Density of liquid at top:
ρl = 1/∑(wi/ρi) = 1/ {(7.2*10-5)/867 +(0.9999)/879} = 878.9 = 879 kg/m3
Liquid –vapor flow factor at top:
Flv = (Lw/Vw)*(ρv/ ρl)0.5 = 0.226*(2.66/879)0.5 = 0.146
Tray spacing = 0.3 m (assumed)
Corresponding Cf = 0.06
(from Introduction to Process Engineering & Design by S.B. Thakore & B.I Bhatt, Chapter 6, Process Design of distillation column, Page no. 448)
Now flooding velocity:
vf = Cf * (σ/0.02)0.2 * {( ρl – ρv)/ρv}0.5 (where σ = surface tension of liquid, N/m = ∑σixi)
= 0.06 * (22.267*10-3/0.02)0.2 * {(879-2.66)/2.66}0.5
= 1.126 m/sec.
Now actual velocity:
v = 0.85 * vf = 0.85 * 1.1126 = 0.946 m/sec.
Volumetric flow rate of liquid at top:
Q’ = (V* M av)/ρv = (11.349*78)/2.66 = 332.79 m3/hr = 0.0924 m3/sec
Net area required at top:
An = Q’/v = 0.0924/0.946 = 0.0973 m2
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Let downcomer area Ad,
Ad = 0.12* Ac (Ac = internal cross-sectional area of tower)
& A n = Ac – Ad = Ac – 0.12*Ac = 0.88 Ac
Ac = An/0.88 = 0.0973/0.88 = 0.11 m2
Inside diameter of column required at top:
Di = {(4*A c)/∏} 0.5 = 0.375 m
(a) Tower diameter required at bottom:
Operating pressure at the top of column = 1 atm + ∆pt
Where ∆pt = total pressure drop in sieve tray tower
Assuming tray efficiency (η) = 0.5
Actual no. of trays =19/0.5 = 38
∆pt = Actual no. of trays*ρ*g*h t (where pressure drop, ht =100mm of LC)
= 32.767 kpa
Thus operating pressure= 101.325 + 32.767 = 134.092 kpa
Temp. at bottom = 112 0 c
Molar flow rates:
V’ = 11.349 kmol/hr
L’ = 12.543 kmol/hr
Here total condenser is used,
hence Lw/Vw = L/V = 12.543/11.349 = 1.105 m
Density of vapor:
ρv = (p* Mav)/(R*T) = {(134.092*106)*(273)}/(385*22.414*101.325) = 4.438 kg/m3
Density of liquid at bottom:
ρl = 1/∑(wi/ρi) = 879 kg/m3
Liquid –vapor flow factor at bottom:
Flv = (Lw/Vw)*(ρv/ ρl)0.5 = 1.105*(4.438/879)0.5 = 0.0785
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Tray spacing = 0.3 m (assumed)
Corresponding Cf = 0.06
Now flooding velocity:
vf = Cf * (σ/0.02)0.2 * {( ρl – ρv)/ρv}0.5 (where σ = surface tension of liquid, N/m = ∑σixi)
= 0.06 * (16.57*10-3/0.02)0.2 * {(879-4.438)/4.438}0.5
= 0.811 m/sec.
Now actual velocity:
v = 0.85 * vf = 0.85 * 0.811 = 0.69 m/sec.
Volumetric flow rate of liquid at bottom:
Q’ = (V* M av)/ρv = (11.349*106)/4.438 = 0.0752 m3/sec
Net area required at bottom:
An = Q’/v = 0.0752/0.69 = 0.109 m2
Let downcomer area Ad,
Ad = 0.12* Ac (Ac = internal cross-sectional area of tower)
& A n = Ac – Ad = Ac – 0.12*Ac = 0.88 Ac
Ac = An/0.88 = 0.109/0.88 = 0.124 m2
Inside diameter of column required at bottom:
Di = {(4*A c)/∏} 0.5 = 0.400 m
Checking for weeping: Minimum vapor velocity through holes to avoid the weeping given by following equation: vh, min = [{K – 0.9(25.4 – dh)}/( ρv)
0.5] K constant can be obtained from Fig 8.19 (Introduction to Process Engineering & Design by S.B. Thakore & B.I Bhatt, Chapter 6, Process Design of distillation column, Page no. 449) or Fig. 4 of appendix. ,is a function of (hw + how), where weir height (hw) = 50 mm hole diameter(dh) = 5 mm Plate thickness (t) = 5mm
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(a) For enriching section : height of liquid crest over the weir how = 750 (Lm/ρl*l w) Lm = 0.7*L*M av = 0.7*(11.349/3600)*78 = 0.172 kg/sec lw = 0.77* Di = 0.77* 0.375 = 0.2885 Now putting all the values , how = 6.075 mm Thus (hw + how) = 56.075 mm corresponding K value from Fig 8.19: K = 30.2 Thus finally vh,min = 4.62 m/sec Actual vapor velocity holes at actual vapor flow rate : vh,a = (0.7* Qv)/ Ah ……………………………………………………………………………………………..(a)
Now Ad = 0.12 * Ac = 0.12 * 0.11 = 0.0132 m2 Active area Aa = Ac – 2Ad = 0.11 – 2(0.12)(0.11) = 0.0836 m2
Thus hole area, Ah= 0.00836 m2
Now putting it in equation (a) we get,
vh,a = 7.74 m/sec
Since vh,a >> vh,min
Thus in enriching section minimum operating rate is well above weep point.
(b) For Stripping section: height of liquid crest over the weir how = 750 (Lm/ρl*l w) Lm = 0.7*L*M av = 0.7*(12.543/3600)*106 = 0.259 kg/sec lw = 0.77* Di = 0.77* 0.400 = 0.308
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ρl = 879 kg/m3 Now putting all the values , how = 23.19 mm Thus (hw + how) = 75.19 mm corresponding K value from Fig 8.19: K = 30.68 Thus finally vh,min = 5.8533 m/sec Actual vapor velocity holes at actual vapor flow rate : vh,a = (0.7* Qv)/ Ah ……………………………………………………………………………………………..(b )
Now Ad = 0.12 * Ac = 0.01488 m2 Active area Aa = Ac – 2Ad = 0.09424 m2
Thus hole area, Ah= 0.009424 m2
Now putting it in equation (b) we get,
vh,a = 6.3 m/sec
Since vh,a >> vh,min
Thus in stripping section minimum operating rate is well above weep point.
Tray pressure drop: (a) For enriching section:
Dry plate pressure drop: hd = 51(vh/C0)
2(ρv/ρl) ………………………………………………………………………………………..(c) vh = Q’v/Ah = 0.0924/0.00836 = 11.053 m/sec From fig. 8.20 Plate thickness / Plate Area = 1 Ah/Ap = Ah/Aa =0.1 ( Ap is perforated area which is slightly less than active area.) Thus corresponding C0 = 0.8422 (from fig. 1 of appendix) Now putting all the values in equation (c)
hd = 26.58 mm
Maximum height of liquid of crest over the weir:
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how = 750 (Lm/ρl*l w) = 750{(0.172/0.7)/(867*0.288)} = 23.36 mm Residual Pressure drop: hr = (12.5*103)/ρl = (12500/867) = 14.41 mm Total tray pressure drop: ht = hd + hw + how + hr = 26.58 + 50 + 23.36 +14.41 = 114.16 mm
(b) For stripping section: Dry plate pressure drop: hd = 51(vh/C0)
2(ρv/ρl) ………………………………………………………………………………………..(c) vh = Q’v/Ah = 0.0752/0.009424 = 7.979 m/sec From fig. 8.20 Plate thickness / Plate Area = 1 Ah/Ap = Ah/Aa =0.1 ( Ap is perforated area which is slightly less than active area.) Thus corresponding C0 = 0.8422 Now putting all the values in equation (c)
hd = 23.43 mm
Maximum height of liquid of crest over the weir:
how = 750 (Lm/ρl*l w) = 750{(7.979/0.7)/(879*0.308)} = 9.74 mm Residual Pressure drop: hr = (12.5*103)/ρl = (12500/879) = 14.22 mm Total tray pressure drop: ht = hd + hw + how + hr = 23.43 + 50 + 9.74 +14.22 = 97.39 mm
Checking of downcomer design:
Type of downcomer: Straight & segmental downcomerarea, Ad = 0.12Ac (for both sections)
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(a) For enriching section: hdc = 166 (Lmd/ρl*A m) where Lmd = Liquid flow rate through downcomer, kg/sec = L*Mav
= (3.0225*78)/3600 = 0.065 kg/sec ρl = 867 kg/m3 Am = Ad or Aap whichever is smaller Aap = hap*l w = (hw – 10)*lw = (50 -10)*0.288 = 0.1152 m2 Ad = 0.0836 m2
Since Aap<< Ad Thus Am = Aap = 0.0836 m2 hdc = 166{0.065/(867*0.01152)}2 = 7.03* 10-3 mm Liquid back in downcomer: hb = hw + how + ht + hdc = 50 + 6.075 + 114.16 + 7.03*10-3 = 170.243 mm Tray spacing (lt) = 0.3 m =300 mm Now, (l t + hw)/2 = (300 + 50)/2 = 175 mm Since hb < (lt + hw)/2 Thus downcomer area & spacing are acceptable. Checking for residence time : θr = (Ad*hb*ρl)/Lmd = (0.0836*0.170*867)/0.065 = 189.56 sec which is greater than 3 sec. thus it is satisfactory.
(b) For stripping section: hdc = 166 (Lmd/ρl*A m) where Lmd = Liquid flow rate through downcomer, kg/sec = L*Mav
= 0.259 kg/sec ρl = 879 kg/m3 Am = Ad or Aap whichever is smaller Aap = hap*l w = (hw – 10)*lw = 0.01232 m2 Ad = 0.09424 m2
Since Aap<< Ad Thus Am = Aap = 0.01232 m2
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hdc = 166{(0.259/0.7)/(867*0.01152)}2 = 5.67 mm Liquid back in downcomer: hb = hw + how + ht + hdc = 50 + 9.74 + 97.39 + 5.67 = 162.80 mm Tray spacing (lt) = 0.3 m =300 mm Now, (l t + hw)/2 = (300 + 50)/2 = 175 mm Since hb < (lt + hw)/2 Thus downcomer area & spacing are acceptable. Checking for residence time : θr = (Ad*hb*ρl)/Lmd = (0.09424*0.162*879)/(0.259/0.7) = 189.56 sec which is greater than 3 sec. thus it is satisfactory.
Checking for entrainment:
(a) For enriching section:
Vapor velocity based on net area (vn) = Q’/ An = 0.0924 / 0.973 = 0.95 m/sec. % Flooding = (vn/vf)*100 = 85.35% (from fig 2 of appendix ) Flv = 0.146
From fig 8.18:
ψ = 0.11 or 11% Since % ψ> 10% , thus higher efficiency will be obtained.
(b) For stripping section: % Flooding = 85% Flv = 0.0785
From fig 8.18:
ψ = 0.053 or 5.43% Since % ψ< 10% , thus higher efficiency will be obtained
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8.4 Mechanical Design of Benzene Distillation Column used in Ethyl Benzene Production
Shell Thickness:
Diameter = 400mm (taken for designing)
Operating pressure = 760mmHg
Taking 10% allowance Pd = 684mm Hg
Operating average temperature = 100 0C
Design temperature = 1500 C
Material used is carbon steel
Specific gravity of carbon steel = 7.7
Permissible tensile stress = 950 kg/cm2 (up to 250 0C)
Insulation material is mineral wood , 75mm thick , density = 130kg/m3
Shell minimum thickness ts
ts = Pd Di/(2f tJ-Pd) + C take C = 3mm
ts = (0.9x400/2x950x1 -0.9) +3 = 3.20
but for high vessels under external pressure take shell thickness ts= 8mm
Head:-
Torispherical heads
Let thickness = 8mm same as shell
t = P dRc Cs/[2f tJ + Pd(Cs-0.2)]
Cs = ¼[3 + (Rc/Rk)0.5]
Take Rc = 400 mm ; J = 1
On solving Cs = 71.24
Therefore Rk = 5.026x10-3mm
But Rk > 0.06Rc ; therefore Rk = 24mm
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Shell Thickness at different Height:-
At a distance X m from the top of shell the stress are
(a) Axial stress:-
fap = Pd Di/4(ts-C) on substituting values
fap = 18 kg/cm2
(b) Stresses due to dead load :
(i) Compressive stress due to weight of shell (fds):-
= 7.7x10 -6X kg/cm2
(ii) Compressive stress due to weight of insulation(fdi):-
D m = Dins = (0.4 + 0.375)/2 = 0.3875 m
fdi = 3.22x10 -5X kg/cm2
(iii) Compressive stress due to liq. In the column up to height X
fd,liq= 2.134x10 -3X kg/cm2
(iv) stress due to attachement (f d, att)
wt. of attachements = 150 kg/m
therefore fd = 11.46x10 -3X kg/cm2
(c) Stress due to wind load:- fwx
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wind pressure = 125 kg/m2 = 125x10 -6 kg/mm2
fwx = 1.34x10 -6X2 kg/cm2
Neglecting seisemic load
equating all the stresses to zero
fwx – ( fdx + fd,liq + fdi + fds ) – fap = 0
solving for X ;
X = 17.6
Hence thickness taken as 8mm is sufficient as column ht. is 11.7m
Support:-
(a) stress due to dead weight:
Skirt diameter = 400 mm (Ds)
Dead weight attachments = 46000 kg
(b) stress due to wind load Mw = 0.7PwDoX2
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70
Chapter 9 Cost Estimation
71
9. Cost Estimation 9.1 Capital Cost Estimation & Distribution: Cost of producing ethyl benzene per annum in 1965 = $ 5.5x106
Chemical plant index for the year 1965 = 76.2 Chemical plant index for the year 2008 = 740 Therefore cost of plant in 2008 = Cost in 1965x {(cost index in 2008)/ (cost index in1965)} = 5.5x106 (740/76.2) = $ 53.41x105
Or Rs. 26.1709 Crores (1$ = Rs 49.0 ) Therefore fixed capital cost = FCC = Rs 26.1709 crores Total capital investment = TCI = FCC + Working capital Working capital = 25% of TCI Therefore working capital = Rs 32.714 crores Distribution of capital cost: Direct cost % of FCC Cost (Crores of Rs) Purchased Equipment 20 5.234 Installation of equipment 9 2.355 Instrumentation (installed) 3 0 .785 Piping 15 3.926 Electrical (installed) 8 2.094 Building 12 3.140 Yard improvement 2 0.523 Service facilities 12 3.140 Land 1 0 .261 Total 82 21.46 Indirect Cost % of FCC Cost (Crores of Rs) Engineering supervisions 4 1.047 Construction expenses 6 1.570 Contractor fees 3 0.785 Contingencies 5 1.308 Total 18 4.71
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9.2 Estimation of total product cost: Let X be the total product cost Distribution of total product cost
Fixed charge = 16% of TPC (let) 20.1082 + 0.07125X = 0.16X X = 22.657 crores General Expenses:
Total product cost = manufacturing cost + general expenses = (30.1082 + 0.76875X) + 0.2375X = 30.1082 + 1.00625X X = 22.657 crores Therefore general expenses = 5.34 crores Therefore manufacturing cost = 20.42 crores or Direct production cost.
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9.3 Break Even Analysis: Cost price of ethyl benzene per kg = 22.657x107/1x106= 226.57 or Rs. 227 Assuming a profit margin of 20% so selling price of the product = Rs. 273 Gross annual earnings = total annual sales – total annual product cost = Rs. 4.6 crores Net annual earnings = gross annual earnings – income tax = 23 – 40%of 46 = 2.76 Crores Payback period = Total capital investment Net annual earnings = 32.714/ 2.76 = 11.85 years (12 year approx) Rate of return = Net profit FCC = 2.76/26.17 = 10.54%
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Chapter 10 Conclusion
75
Conclusion:
Nearly all commercial ethylbenzene is produced by alkylation of benzene with ethylene. Earlier processes were based on liquid phase alkylation using an aluminum chloride catalyst but this route required disposal of aluminum chloride waste.
In the early 1980s, Mobil/Badger developed an alternative zeolite-based process using vapor phase alkylation, offering higher yields and purity. More recently, liquid phase processes using zeolite catalysts have been introduced. These latest technologies offer low benzene-to-ethylene ratios, which reduces the size of equipment, and lowers the production of byproducts.
Nearly all the ethylbenzene (EB) produced is used in the manufacture of styrene monomer (SM) with the remainder, at less than 1%, used in solvent applications. In addition, most of the EB is used captively, leaving a small merchant market for the product. Hence, EB demand runs in parallel to that of styrene.
Here we have designed a plant for 1000 MTA, with a techno-economic feasibility report which is started with a need, demand & supply analysis and by going through a process of mass, energy balances and detailed design of equipments in the process.
76
References
77
References:
• Chapter 22, Introduction to Multicomponent Distillation, Unit Operations of Chemical Engineering, by McCabe-Smith-Harriott, 6th Edition, Published by McGraw . Hill International Edition, Chemical Engineering Series.
• Chapter 8, Process Design of distillation Column, Introduction to Process Engineering & Design, Second
reprint 2009, Published by, Tata-McGraw-Hill Publishing Company limited, New Delhi.
• Chapter 6, Costing & Project Evaluation, Coulson & Richardson’s volume 6, Third Edition, Chemical Engineering Design, By R.K.Sinnot, Publisher by Butterworth- Heinemann Publications.
• Chapter 11, Separation Columns, Coulson & Richardson’s volume 6, Third Edition, Chemical Engineering
Design, By R.K.Sinnot, Publisher by Butterworth-Heinemann Publications.
• Chapter 10, Process Design of Reactors, Introduction to Process Engineering & Design, Second reprint 2009, Published by, Tata-McGraw-Hill Publishing Company limited, New Delhi.
• Process Equipment Design, By M.V. Joshi & V.V. Mahajani, Third edition, published by McMillan India
Limited.
• Kirk – Othmer Encyclopedia of chemical Technology 4th Edition.
• Ullmann’s Encyclopedia, Industrial Organic Chemicals, Volume – 4
• Chemical Weekly Buyer’s Guide 2005.
• SAX’s Dangerous Properties of Industrial Materials.
• Hydrocarbon Processing.
• Basic Principals & Calculations in Chemical Engineering, David M. Himelblau, 6th Edition, PHI Publication.
• Physical, thermo physical & thermo chemical properties:
Yaws C.L., Physical properties. Perry’s Chemical Engineer’s Handbook , McGraw-Hill Publications.
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Appendices
79
(A) Figures:
Figure 1
80
Figure 2
81
Figure 3
Figure 4
82
Figure 5
83
(B) MSDS:
84
85
86
87
88
