maleic anhydride.pdf

CHAPTER-1

Introduction 1.1 MALEIC ANHYDRIDE

Maleic anhydride is multifunctional chemical intermediates that find applications

in nearly every field of industrial chemistry.

Each molecule contains two acid carbonyl groups and a double bond in the ,

Maleic anhydride is important raw materials used in the manufacture of phthalic-

typealkyd and unsaturated polyester resins, surface coatings, lubricant additives,

plasticizers , copolymers and agricultural chemicals .

1.2 Other names

Both chemicals derive their common names from naturally occurring malic acid .

Other names for maleic anhydride are 2,5-furandione, dihydro-2,5-dioxofuran,

toxilic anhydride, or cis-butenedioic anhydride. Maleic acid is also called (Z)-2-

butenedioic acid, toxilic acid, malenic acid, maleinic acid, or cis-1,2-

ethylenedicarboxylic acid.

1.3 History

Maleic anhydride and the two diacid isomers were first prepared in the 1830s (1)

but commercial manufacture did not begin until a century later. In 1933 the

National Aniline and Chemical Co., Inc.,installed a process for maleic anhydride

based on benzene oxidation using a vanadium oxide catalyst .

1.4 Physical Properties

Maleic anhydride, Maleic acid, and fumaric acid

including solid and solution properties are given in Tables 1,2

Table 1

Production of Maleic Anhydride

3

Table 2

From single crystal x-ray diffraction data , Maleic anhydride is a nearly planar

molecule with the ring oxygen atom lying 0.003 nm out of the molecular plane. A

twofold rotation axis bisects the double bond and passes through the ring oxygen

atom. Figure 1


5

summarizes the bond distance for maleic anhydride . Similar bond distances and

angles for maleic anhydride were obtained using electron diffraction and double

resonance modulation microwave spectroscopic (1 techniques. Values of the

Raman polarizability were reported for single crystals of maleic anhydride .

Density functional theory has been applied to maleic anhydride to give optimized

geometry, harmonicvibrational frequencies, and electron affinity .

1.5 Chemical Properties

The General References and two other reviews provide extensive descriptions of

the chemistry ofmaleic anhydride and its derivatives. The broad industrial

applications for this chemistry derive from thereactivity of the double bond in

conjugation with the two carbonyl oxygens

Acid Chloride Formation.

Monoacid chlorides of maleic and fumaric acid are not known. Treatment of

maleic anhydride or maleic acid with various reagents such as phosgene , phthaloyl

chloride

phosphorus pentachloride , or thionyl chloride gives 5,5-dichloro-2furanone .

Noncyclic maleyl chloride (6) forms in 11% yield at 220C in the reaction ofone

mole of maleic anhydride with six moles of carbon tetrachlorideover an activated

carbon catalyst .

Alkylation.

H bond

activated by-unsaturation or an adjacent aromatic resonance to produce the

following succinic anhydride derivatives.


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Typical reaction conditions are 150 to 300C and up to 2 MPa pressure.

Polyalkenyl succinic anhydridesare prepared under these conditions by the reaction

of polyalkenes in a nonaqueous dispersion of maleic anhydride, mineral oil, and

surfactant .

N-Alkylpyrroles react with maleic anhydride to give the electrophilic substitution

product (7) and not the Diels-Alder addition product found for furan and thiophene

compounds (35). However, the course of this reaction can be altered by

coordination of the pyrrole compound to a metal center.

Amidation.

Reaction of maleic anhydride or its isomeric acids with ammonia, primary amines ,

and secondary amines produces mono- or diamides. The monoamide derivative

from the reaction of ammonia and maleic anhydride is called maleamic acid (8).

Another monoamide derivative formed from the reaction of aniline and maleic

anhydride is maleanilic acid

The reactions of primary amines and maleic anhydride yield amic acids that can be

dehydrated to imides, polyimides , or isoimides depending on the reaction

conditions . However, these products require multistep processes. Pathways with

favorable economics are difficult to achieve. Amines and pyridines decompose

maleic anhydride, often in a violent reaction. Carbon dioxide is a typical end

product for this exothermic reaction .

Maleic hydrazide is one of a number of commercial agricultural chemicals derived

from maleic anhydride. Maleic hydrazide was first prepared in 1895 but about 60

years elapsed before the intermediate products were elucidated .

1.6 Uses

Maleic anhydride is truly a remarkable molecule in that it possesses two types of

chemical functionality making it uniquely useful in chemical synthesis and

applications. Maleic anhydride itself has few, if any, consumer uses but in

derivatized form it is extremely versatile in the consumer uses in which it is found.

The chemical structure of each maleic anhydride derivative of significant

commercial interest.

The distribution of end uses for maleic anhydride is presented in Table 9 for the

year 2000 . The


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majority of the maleic anhydride produced is used in unsaturated polyester

resin

UNSATURATED. Unsaturated polyester resin is then used in both glass-

reinforced applications and in unreinforced applications.

There are many unsaturated polyester resin formulations. A typical unsaturated

polyester resin formulation consists of an aromatic dibasic acid (or anhydride) such

as phthalic anhydride, an unsaturated dibasic acid(or anhydride) such as maleic

anhydride and a glycol such a propylene glycol. The polyester chains are then

cross-linked through the double bond with vinyl cross linking agents such as

styrene. Reinforcement in the form of glass fibers or other reinforcement fibers

may be added to provide the strength requirements of the end product. The exact

unsaturated polyester formulation, its cross linking agent, and reinforcement fiber,

if

any, are selected to optimize the performance of the end product.

Fumaric acid and malic acid are produced from maleic anhydride. The primary use

for fumaric acid is in the manufacture of paper sizing products . Fumaric acid is

also used as a food acidulant, as is malic acid. Malic acid is a particularly desirable

acidulant in certain beverage selections, specifically those sweetened with the

artificial sweetener aspartame Lube oil additives represent another important

market segment for maleic anhydride derivatives, the molecular structure of

importance being adducts of polyalkenyl succinic anhydrides These materials act

as dispersants and corrosion inhibitors One particularly important polyalkenyl

succinic anhydride molecule in this market is polyisobutylene succinic anhydride

(PIBSA) where the polyisobutylene group has a molecular weight of 900 to 1500.

Other polyalkenes are also used. Polyalkenyl succinic anhydride is further

derivatized with various amines to produce both dispersants and corrosion

inhibitors. Another type of dispersant is a polyester produced from a polyalkenyl

succinic anhydride and pentaerythritol .

Maleic anhydride is used in a multitude of applications in which a vinyl copolymer

is produced by the copolymerization of maleic anhydride with other molecules

having a vinyl functionality. Typical copolymers (and their end uses) are styrene-

maleic (engineering thermoplastic, paper treatment chemical, floor polishes,

emulsifiers, protective colloids, antisoil agents, dispersants, stabilizing agent,

adhesives,detergents, cosmetics, and toiletries), diisobutylene-maleic (dispersing

agent), acrylic acid-maleic (detergent ingredient), butadiene-maleic (sizing agent),

and C18 alpha olefin-maleic (emulsification agent and paper coating).

The use of maleic anhydride in the manufacture of agricultural chemicals has

declined in the United States over the last decade. Malathion (S-[1,2-

dicarbethoxyethyl],-dimethyldithiophosphate) and

Difolatan(cis-N-[1,1,2,2-tetrachloroethylthio]-4-cyclohexene-1,2-dicarboximide)

are no longer produced in the United States and Alar (N-dimethylaminosuccinamic


11

acid) volumes have been significantly reduced by intense environmental scrutiny.

Maleic hydrazide, Captan (cis-N- [trichloromethylthio]-4-cyclohexene-1,2-

dicarboximide), Endothall (7-oxabicyclo[2,2,1]-heptane- 2,3-dicarboxylic acid ,

disodium salt), and several other maleic derivatives continue use in a number of

agricultural functions: plant growth regulation, fungicides, insecticides, and

herbicides There are numerous further applications for which maleic anhydride

serves as a raw material. These applications prove the versatility of this molecule.

The popular artificial sweetener aspartame is a dipeptide with one amino acid

which is produced from maleic anhydride as 32 the starting material. Processes

have been reported for production of poly(aspartic acid) with applications for this

biodegradable polymer aimed at detergent builders, water treatment, and

poly(acrylic acid) replacement Alkenylsuccinic anhydrides made from several

linear alpha olefins are used in paper sizing, detergents, and other uses.

Sulfosuccinic acid esters serve as surface active agents. Alkyd resins are used as

surface coatings. Chlorendric anhydride is used as a flame resistant component .

Tetrahydrophthalic acid and hexahydrophthalic anhydride have specialty resin

applications. Gas barrier films made by grafting maleic anhydride to polypropylene

film are used in food packaging Poly(maleic anhydride) is used as a scale

preventer and corrosion inhibitor Maleic anhydride formscopolymers with mono-

O-methyl-oligoethylene glycol vinyl ethers that are partially esterified for

biomedical

and pharmaceutical uses .An important developing use for maleic anhydride is the

production of products in the 1,4-butanediol-

family. Kvaerner Process Technology licenses a process for producing 1,4-

butanediol from maleic anhydride. This technology can be used to produce the

product mix of the threemolecules as needed by the producer. Two plants were in

operation in 1998 using the Kvaerner technology, having a combined capacity of

50 kt/yr 1,4-butanediol. Several other plants using the Kvaerner technology are

under construction or have been announced. SISAS produces 1,4-butanediol from

maleic anhydride in their facility in Feluy, Belgium. DuPont produces

tetrahydrofuran in Spain from maleic anhydride. The DuPont technology oxidizes

butane to maleic anhydride, which is recovered as maleic acid and then reduced to

tetrahydrofuran. BP Amoco have announced a facility in Lima, OH to produce 1,4-

butanediol from maleic anhydride using their own technology.

Chapter-2

Manufacturing processes

2.1 Butane-Based Catalyst Technology.


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The increased importance of the butane-to-maleic anhydride conversion route has

resulted in efforts being made to understand and improve this process. Since 1980,

over 225 U.S. patents have been issued relating to malefic anhydride technology.

The predominant area of research concerns the catalyst because it is at the heart of

this process. The reasons for this statement are twofold. First, there is the

complexity of this reaction: for malefic anhydride to be produced from butane,

eight hydrogen atoms must be abstracted, three oxygen atoms inserted, and a ring

closure performed. This is a 14-electron oxidation that occurs exclusively on the

surface of the catalyst. The second reason for the emphasis placed on the catalyst is

that all of the commercial processes use the same catalyst. This catalyst is

the only commercially viable system that selectively produces malefic anhydride

from butane.

The catalyst used in the production of malefic anhydride from butane is vanadium

oxide (VPO). Several routes may be used to prepare the catalyst ,

but the route favored by industry

2.2 Benzene-Based Catalyst Technology

. The catalyst used for the conversion of benzene to malefic anhydride consists of

supported vanadium oxide . The support is an inert oxide such as kieselguhr,

alumina , or silica, and is of low surface area. Supports with higher surface area

adversely affect conversion of benzene to maleic anhydride. The conversion of

benzene to maleic anhydride is a less complex oxidation than the conversion of

butane, so higher catalyst selectivities are obtained. The vanadium oxide on the

surface of the support is often modified with molybdenum oxides. There is

approximately 70% vanadium oxide and 30% molybdenum oxide in the active

phase for these fixed-bed catalysts. The molybdenum oxide is thought to form

either a solid solution or compound oxide with the vanadium oxide and result in a

more active catalyst .

2.3 Process Technology Evolution

. Maleic anhydride was first commercially produced in the early 1930s by the

vapor-phase oxidation of benzene . The use of benzene as a feedstock for the

production of maleic anhydride was dominant in the world market well into the

1980s. Several processes have been used for the production of maleic anhydride

from benzene with the most common one from Scientific Design. Small amounts

of maleic acid are produced as a by-product in production of phthalic anhydride .

This can be converted to either maleic anhydride or fumaric acid. Benzene,

although easily oxidized to maleic anhydride with high selectivity, is an inherently

inefficient feedstock since two excess carbon atoms are present in the raw material.

Various C4 compounds have been evaluated as raw material substitutes for

benzene in the production of maleic anhydride. Fixed- and fluid-bed processes for

production of maleic anhydride from the butenes present in mixed C 4 streams

have been practiced commercially. None of these processes is currently in

operation.

Rapid increases in the price of benzene and the recognition of benzene as a

hazardous material intensified the search for alternative process technology in the

United States. These factors led to the first commercial production of maleic

anhydride from butane at Monsanto's J. F. Queeny plant in 1974. By the early

1980s, the conversion of the U.S. maleic anhydride manufacturing capacity from

benzene to butane feedstock was well under way using catalysts developed by

Monsanto, Denka, and Halcon. One factor that inhibited the conversion of the

installed benzene-based capacity was that early butane-based catalysts were not

active and selective enough to allow the conversion of benzene-based plant without

significant loss of nameplate capacity. In 1983, Monsanto started up the world's


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first butane-to-maleic anhydride plant, incorporating an energy efficient solvent-

based product collection and refining system. This plant was the world's largest

maleic anhydride production facility in 1983 at 59,000t/yr capacity, and through

rapid

advances in catalyst technology has been debottlenecked to a current capacity of

105,000t/yr (1999).

Advances in catalyst technology, increased regulatory pressures, and continuing

cost advantages of butane over benzene have led to a rapid conversion of benzene-

to butane-based plants. By the mid-1980s in the United States 100% of maleic

anhydride production used butane as the feedstock. Coincident with the rapid

development of the butane-based fixed-bed process, several companies have

developed fluidized-bed processes. Two companies, Badger and Denka,

collaborated on the development of an early fluid-bed reaction system which was

developed through the pilot-plant stage but was never commercialized. Three fluid-

bed, butane-based technologies were commercialized during the latter half of the

1980s by Mitsubishi Kasei, Sohio (British Petroleum), and Alusuisse. A second

fluidized-bed technology for the oxidation of butane to maleic anhydride, known

as transport bed, has been developed by Du Pont. A world-scale plant in Spain for

the production of THF by the hydrogenation of maleic acid using this

technology began production in 1996 . Europe has largely converted from

benzene-based to butane-based maleic anhydride technology with the construction

of several new butane based facilities by CONDEA-Huntsman, Pantochim and

Lonza. Growth in the worldwide maleic anhydride industry is predominantly in the

butane-to-maleic anhydride route, often

at the expense of benzene-based production. Table 4 shows 1993 and 2000

worldwide maleic production

capacity broken down in categories of fixed-bed benzene, fixed-bed butane,

fluidized-bed butane, andphthalic anhydride coproduct. As can be seen from this

table, both fixed- and fluidized-bed butane routes have grown dramatically with

the fixed-bed route adding 336,000t/yr capacity compared to 90,000t/yr for the

fluid-bed process. Only a few newer benzene-based fixed-bed processes have been

built since the early 1980s and these were built where the availability of butane

was limited. The fluidized-bed butane-based process is experiencing some growth,

but based on growth rates from Table 4 (178,179), it does not appear destined to

challenge fixed-bed technology. The announcement from Huntsman Specialty

Chemicals Corp.,

Butane-Based Fixed-Bed Process Technology.

Maleic anhydride is produced by reaction of butane with oxygen using the

vanadium phosphorus oxide heterogeneous catalyst discussed earlier. The butane

oxidation reaction to produce maleic anhydride is very exothermic. The main

reaction by-products are carbon monoxide and carbon dioxide. Stoichiometries and

heats of reaction for the three principal reactions are as

follows:


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21

C4H10 + 3.5 O2 H

C4H10 + 6.5 O2 4 CO2 + 5 H2O H

C4H10 + 4.5 O2 4 CO + 5 H2O H )

Air is compressed to modest pressures, typically 100 to 200 kPa with either a

centrifugal or axial compressor, and mixed with superheated vaporized butane.

Static mixers are normally employed to ensure good mixing. Butane concentrations

are often limited to less than 1.7 mol % to stay below the lower flammable limit of

butane . Operation of the reactor at butane concentrations below the flammable

limit does not eliminate the requirement for combustion venting, and consequently

most processes use rupture disks on both the inlet and exit reactor heads.

The highly exothermic nature of the butane-to-maleic anhydride reaction and the

principal by-product reactions require substantial heat removal from the reactor.

Thus the reaction is carried out in what is effectively a large multitubular heat

exchanger which circulates a mixture of 53% potassium nitrate, KNO3; 40%

sodium nitrite , NaNO2; and 7% sodium nitrate , NaNO3.

facilitate heat removal. Reactor tube lengths are between 3 and 6 meters. The

exothermic heat of reaction is removed from the salt mixture by the production of

steam in an external salt cooler. Reactor temperatures are in the range of 390 to

430C. Despite the rapid circulation of salt on the shell side of the reactor, catalyst

temperatures can be 40 to 60C higher than the salt temperature. The butane to

maleic anhydride reaction typically reaches its maximum efficiency (maximum

yield) at about 85% butane conversion. Reported molar yields are typically 50 to

60%. Efficient utilization of waste heat from a maleic anhydride plant is critical to

the economic viability of the plant. Often site selection is dictated by the presence

of an economic use for by-product steam. The steam can also be used to drive an

air compressor, generate electricity, or both. Alternatively, an energy consuming

process, such as a butanediol plant, can be closely coupled with the maleic

anhydride plant. Several such plants have been announced . Design and integration

of the heat recovery systems for a maleic anhydride plant are very site specific.

Heat is removed from the reaction gas through primary and sometimes secondary

heat exchangers. In addition to the heat recovered from the reactor and process

gasheat exchangers, additional heat can be recovered from the destruction of

unreacted butane, the carbon monoxide by-product, and other by-products which

cannot be vented directly to the atmosphere. This destruction is done typically in a

specially designed thermal oxidizer or a modified boiler. Reactor operation at 80 to

85% butane conversion to produce maximum yields provides an opportunity for

recycle processes to recover the unreacted butane in the stream that is sent to the

oxidation reactor. Patents have been issued on recycle processes both with and

without added oxygen. Pantochim has announced the commercialization of a

partial recycle process . Mitsubishi Chemical Corporation has announced plans to

add butane recovery from the offgas of their fluid bed process through the use of

BOC Gases proprietary selective hydrocarbon separation system (PETROX) .

This technology is particularly well suited to use in fluid bed processes where the

hydrocarbon to air ratio is relatively high and in world areas where butane has a

high value relative to its energy content. Operation of the butane to maleic

anhydride process in a total recycle configuration can produce molar yields that

approach the reaction selectivity which is typically 65 to 75%, significantly higher

than the 50 to 60% molar yields from a single pass, high conversion process. The

Du Pont transport bed process achieves its high reported yields at least partially

through implementation of recycle technology. Recovery of the fuel value of the


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butane in the offgas from a single pass configuration plant reduces the economic

attractiveness of recycle operation.

Butane-Based Fluidized-Bed Process Technology.

Fluidized-bed processes offer the advantage of excellent control of hot spots by

rapid catalyst mixing, simplification of safety issues when operating above the

flammable limit, and a simplified reactor heat-transfer system. Some disadvantages

include the effect of back mixing on the kinetics in the reactor, product destruction

and by-product reactions in the space above the fluidized bed, and vulnerability to

large-scale catalyst releases from explosion venting. Compressed air and butane

are typically introduced separately into the bottom of the fluidized-bed reactor.

Heat from the exothermic reaction is removed from the fluidized bed through

steam coils in direct contact with the bed of fluidized solids. Fluidized-bed reactors

exploit the extremely high heat-transfer coefficient between the bed of fluidized

solids and the steam coils. This high heat-transfer coefficient allows a relatively

small heattransfer area in the fluid-bed process for the removal of the heat of

reaction compared to the fixed-bed process. Gas flow patterns in a commercial

scale fluid-bed reactor are generally backmixed, which can lead

to maleic anhydride destruction. Patents have been issued for mechanical

modifications to the reactorinternals that claim to control backmixing .Other

methods to reduce backmixing include introductionof catalyst fines (small particles

of catalyst) to decrease bubble size and operation of the reactor in the turbulent,

fast fluidization regime in an attempt to minimize bubbling. Fluidized-bed reactors

require a significant amount of space above the catalyst level to allow the solids to

separate from the gases. This exposure of the product to high temperatures at

relatively long residence times can lead to side reactions and product destruction.

Fluidized-bed processes are operated at high butane concentrations but at longer

gas residence times than fixed-bed processes. The product stream contains gases

and solids. The solids are removed by using either cyclones, filters, or both in

combination. Cyclones are devices used to separate solids from fluids using vortex

flow. The product gas stream must be cooled before being sent to the collection

and refining system. The ALMA process uses cyclones as a primary separation

technique with filters employed as a final separation step after the off-gas has been

cooled and before it is sent to the collection and refining system . As in the

fixedbedprocess, the reactor off-gas must be incinerated to destroy unreacted

butane and by-products before being vented to the atmosphere. Fluidized-bed

reaction systems are not normally shut down for changing catalyst. Fresh catalyst

is periodically added to manage catalyst activity and particle size distribution. The

ALMA process includes facilities for adding back both catalyst fines and fresh

catalyst to the reactor.

Benzene-Based Fixed-Bed Process Technology.

The benzene fixed-bed process is very similar to the butane fixed-bed process and,

in fact, the Scientific Design butane process has evolved directly from its benzene

process. Benzene-based processes are easily converted to butane-based processes.

Typically, only a catalyst change, installation of butane handling equipment, and

minor modifications to the recovery process are required. The benzene reaction is a

vapor-phase partial oxidation reaction using a fixed-bed

catalyst of mixed vanadium and molybdenum oxides. The reactors used are the

same multitubular reactors cooled by circulating a molten mixture of

KNO

process. The benzene concentrations used are about 1.5 mol % or just below the

lower flammable limit of benzene in air. Unlike the butane reaction, the reactor

normally operates at conversions greater than 95% and molar yields greater than

70%. The benzene oxidation reaction runs a little cooler than the butane oxidation

reaction with typical reactor temperatures being in the 350 to 400C range. The


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reactor off-gas is cooled by one or more heat exchangers and sent to the collection

and refining section of the plant. Unreacted benzene and by-products are

incinerated.

Recovery and Purification.

All processes for the recovery and refining of maleic anhydride must deal with the

efficient separation of maleic anhydride from the large amount of water produced

in the reaction process. Recovery systems can be separated into two general

categories: aqueous- and nonaqueous-based absorption systems. Solvent-based

systems have a higher recovery of maleic anhydride and are more energy efficient

than water-based systems.

The Huntsman solvent-based collection and refining system will be used as a

generic model for solventbased recovery systems .The reactor exit gas is cooled in

two heat exchangers for energy recovery. The cooled gas product stream is passed

to a solvent absorber where a proprietary solvent is used to absorb, almost

completely, the maleic anhydride contained in the product stream. The solvent

stream, coming from the bottom of the absorber with a high concentration of

maleic anhydride, known as rich oil, is sent to a stripper where the rich oil is

heated and maleic anhydride is vacuum stripped from the solvent. The vacuum-

stripped maleic anhydride is typically greater than 99.8% purity, and is sent to the

purification section of the plant where it is batch distilled to produce extremely

pure maleic anhydride. A small slip stream of the solvent which has had the maleic

anhydride removed by stripping is sent to the solvent purification section of the

plant where impurities are removed. The Scientific Design water-based collection

and refining system is in broad use throughout the world in butane-based and

benzene-based plants . The reactor off-gas is cooled from reaction temperatures in

a gas cooler with generation of steam. The off-gas is then sent to a tempered water-

fed aftercooler where it is cooled below the dew point of maleic anhydride. The

liquid droplets of maleic anhydride are separated from the off-gas by a separator.

The condensed crude is pumped to a crude tank for storage. The maleic anhydride

remaining in the gas stream after partial condensation is removed in a water

scrubber by conversion to maleic acid which accumulates in the acid storage

section at the bottom of the scrubber. The acid solution is converted to crude

maleic anhydride in a dual purpose dehydrator/refiner. Xylene is used as an

azeotropic agent for the conversion of maleic acid to maleic anhydride. Water from

the dehydration step is recycled to the scrubber. When the conversion of the acid

solution to crude maleic anhydride is complete, condensed crude maleic anhydride

is added to the still pot and a batch distillation refining step is conducted. The UCB

collection and refining technology also depends on

partial condensation of maleic anhydride and scrubbing with water to recover the

maleic anhydride present in the reaction off-gas. The UCB process departs

significantly from the Scientific Design process when the maleic acid is dehydrated

to maleic anhydride. In the UCB process the water in the maleic acid solution is

evaporated to concentrate the acid solution. The concentrated acid solution and

condensed crude maleic anhydride is converted to maleic anhydride by a thermal

process in a specially designed reactor. The resulting crude maleic anhydride is

then purified by distillation.


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Chapter-3

PROCESS DESCRIPTION

Figure 1 shows a PFD for the overall process. Pure butane, Stream 2, and

compressed air, Stream 3, are mixed and fed to R-101, an adiabatic reactor, where

butane reacts with oxygen to form maleic anhydride. The reaction is exothermic,

therefore, one could consider either a fluidized bed reactor or a packed bed reactor

with heat removal to stay close to isothermal. The reactor effluent is cooled and

sent to T-101, a packed bed absorber, where it is contacted with water, Stream 7, to

remove the light gases and all of the maleic anhydride reacts to form maleic acid.

The vapor effluent, which consists of non-condensables, Stream 8, must be sent to

an after-burner to remove any carbon monoxide prior to venting to the atmosphere.

This is not shown here. The liquid effluent, Stream 9, is then cooled and flashed at

101 kPa and 120C in V-101. The vapor effluent from V-101, Stream 11, is sent to

waste treatment. Stream 12, the liquid 2 effluent, is sent to R-102 where maleic

acid is broken down to maleic anhydride and water. The reactor effluent is then

sent to distillation column, T-102, where maleic anhydride and water are separated.

The distillate, Stream 14, is sent to waste treatment.

Stream 15, the bottoms, consists of 99-wt.% maleic anhydride.

where

C-101 Air Compressor

E-101 Heat Exchanger

E-102 Heat Exchanger

E-103 Condenser

E-104 Reboiler

P-101A/B Reflux Pump

R-101 Packed Bed Reactor

R-102 Maleic Acid Reactor

T-101 Absorbtion Tower

T-102 Distillation Column

V-101 Flash Vessel


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V-102 Reflux Vessel

Chapter-4

MATERIAL BALANCE

4.1- Reactor

Air entering at 25 0C and assume the Humidity is 65% from the

Psychometer chart from appendix figure 5.1

H= 0.018kg water/kg dry air

=0.03 kgmol water/kgmol dry air

Basis

100 ton of maleic anhydride per day.

According to US Patent # 4317778 air is provided in this reaction is

Bu : O2

1 : 8.65 (in mol fraction )

From Encyclopedia

Butane unreacted = 17% of entering

Butane converted to maliec anhydride = 53% of entering

Butane converted to Acrilic acid = 1.1% of entering

Butane converted to formic acid = 1.07% of entering

Butane entering =116 ton/day

= 1829 Kgmol/day

O2 required =1829*8.65

=15820.85 kgmol/day

= 553.6 ton/day

N2= 15820.85*0.79/0.21

=59516.7 kgmol/day

= 1822.26 ton/day

total dry air=75337.55 kgmol/day

H2O with air =75337.55*0.03

=2260.1kgmol/day

=44.48 ton/day

So butane coming out =0.17*1829

=310.83kgmol/day

=19.72 ton/day

Butane is converted to maliec anhydride=0.53*1829

=969.37 kgmol/day

=103.88 ton/day

Butane converted to Acrilic acid =0 .011*1829

= 20.119kgmol/day

=1.584ton/day

Butane converted to formic acid =0.0107*1829

=19.57kgmol/day

=0.9844ton/day

Total butane in-butane consumed=butane coming out of reactor

1829-(969.37+20.119+19.57+x)=310.93

x=amount of butane consume in cox

x=509.01kgmol/day

C4H10+3.5O2 C4H2O3+4H2O ----------- ( 1 )


27

C4H10+5.5O2 CO2+CO+5H2O ----------- ( 2 )

C4H10+3.5O2 C3H4O2+CO2+3H2O ----------- ( 3 )

C4H10+6O2 CH2O+CO2+4H2O ----------- ( 4 )

O2 consumed in reaction (1) = 3.5*969.37

=3392.79 kgmol/day

O2 consumed in reaction (2)= 5.5*509.01

=2799.56 kgmol/day

O2 Balance

O2 consumed in reaction (3) = 3.5*20.119

=70.42 kgmol/day

O2 consumed in reaction (4) = 6*19.57

=117.42 kgmol/day

Total O2 consumed = 6380.19 kgmol/day

O2 leaving unreacted=O2 entering -O2 consumeds

=15820-6380.19

= 9439.81 kg mol/day=330.342 ton/day

CO2Balance

CO2 produced in reaction in (2) =2*509.01

=1018.02 kgmol/day


=20.119 kgmol/day


=58.71 kgmol/day

Total CO2 produced=1096.85kg mol/day

=52.77 ton/day

CO Balance

CO produced in reaction in (2) =2*509.01

=1018.02 kgmol/day

Total CO produced=1018.02kgmol/day

=31.17 ton/day

H2O Balance

H2O produced in reaction (1)=4*969.37

=3777.48 kgmol/day

H2O produced in reaction (2) =5*509.01

=2545.05 kgmol/day


=60.357 kgmol/day


=18.57 kgmol/day

Total water produced =6567.17 kgmol/day

Water with air =2260.1 kgmol/day

Total H2O outlet =8821.29 kgmol/day

=173.63 ton/day

Reactor

C4H10=116 t/d

N2=1822.26 t/d

O2=553.6 t/d

H2O=44.48t/d

C4H10=19.72t/d

O2=330.342 t/d

N2=1822.26t/d

CO2=52.77 t/d

CO=31.17 t/d

H2O=173.63 t/d

C3H4O2=1.548 t/d

CH2O=0.9844 t/d

C4H2O3= 103.88 t/d


29

4.2 Material Balance around absorber

The solubility of Butane in water is 0.0098 kgmol butane/kgmol water at 60 oc and

the solubility of CO,CO2,N2and O2 is negligible.

Water required for absorption = 310.93/0.0098

=31727.55 kgmol/day

= 625 ton/day

Absorber

C4H10=19.72t/d

O2=330.342 t/d

N2=1822.26t/d

CO2=52.77 t/d

CO=31.17 t/d

H2O=173.63 t/d

C3H4O2=1.548 t/d

CH2O=0.9844 t/d

C4H2O3= 103.88 t/d

H2O=625t/d

C4H10=0.3942t/d

O2=330.342 t/d

N2=1822.26t/d

CO2=52.77 t/d

CO=31.17 t/d

H2O=778.788 t/d

C3H4O2=1.548 t/d

CH2O=0.9844 t/d

C4H4O4= 122.96 t/d

C4H10=19.33t/d

4.3 Balance around flash vessel

Stream entering in flash vessel is given below

H2O=778.788 t/d

C3H4O2=1.548 t/d

CH2O=0.9844 t/d

C4H4O4= 103.88 t/d

C4H10=19.33t/d

Total moles entering =40880.53 kgmol/day

Zj=mol fraction in feed

Xj=mol fraction in liquid in outlet

Yj=mol fraction in vapour in outlet

F

V=vapour to feed ratio

Pj=vapour pressure of component supposed

P=Total pressure

By hit and trail method

At 120 oC and V/F=0.95

Components Zj Pj

KPa P

Pj

11P

P

F

V

ZX

i

j

j jj

XP

PYi

C4H10 0.0076 2068.5 20.41 0.000391 0.00797

CH2O 0.00047 179.27 1.769 0.000272 0.00048

C3H4O2 0.0005 51.7125 0.51 0.000935 0.000477

C4H4O3 0.024 0.06895 0.00068 0.4738 0.000322

H2O 0.967 193.06 1.91 0.5198 0.9905


31

Total 1.000 0.996 0.999

V=0.95*F

=0.95*40880.53

=38836.5kgmol/day

L=0.05*F

= 0.05*40880.53

=2044.03kgmol/day

components Liquid stream

kgmol

L*Xj

vapour stream

kgmol

V*Yj

C4H10 0.799 303.91

CH2O 0.55 19.02

C3H4O2 1.83 18.291

C4H4O4 969 0.37

H2O 1072.34 38494.42

Flash Vessel

H2O=757.67 t/d

C3H4O2=1.44 t/d

CH2O=0.9568 t/d

C4H4O4= 0.05 t/d

C4H10=19.26t/d

H2O=21.11 t/d

C3H4O2=0.144 t/d

CH2O=0.0276 t/d

C4H4O4= 122.91 t/d

C4H10=0.051t/d

H2O=778.788 t/d

C3H4O2=1.548 t/d

CH2O=0.9844 t/d

C4H4O4= 122.96 t/d

C4H10=19.33t/d


33

4.4 Balance around maleic acid reactor

By the ref

At 135 oc maliec acid rapidly decomposes into maliec acid and water according to

the reaction

C4H4O4 C4H2O3+H2O

So H2O produced =969 kgmol /day

Total H2O =969+1072.34=2041.34 kgmol /day=40.18ton/day

Maliec acid produced =969 kgmol/day=103.841ton/day

4.5 Balance around distillation column

As Butane, Formic acid and Acrylic acid are so small that they can be neglected.

D

Maliec acid

reactor

H2O=21.11 t/d

C3H4O2=0.144 t/d

CH2O=0.0276 t/d

C4H4O4= 122.91 t/d

C4H10=0.051t/d

H2O=40.18 t/d

C3H4O2=0.144 t/d

CH2O=0.0276 t/d

C4H2O3= 103.8411 t/d

C4H10=0.051t/d

F

W

It is assumed that 98% of water is in distillate and 99% maliec acid in

bottom.

Feed stream H2O= 2041.35 kgmol/day

C4H2O3=969 kgmol/day

C4H10=0.85kgmol/day

CH2O=0.546 kgmo/day

C3H4O2=1.829kgmol/day

Overall balance

F=D+W --------------1

D+W=3010.35

Maliec acid balance

0.02*D+0.98*W=966.32---------2

By solving equation 1&2

D= 2063.93 kgmol/day

W=946.42 kgmol/day

So in distillate

Maliec acid = 0.98*946.42=927.48kgmol/day=100ton/day

H2O=0.02*946.42=18.93 kgmol/day

Acrylic acid=0.9148 kgmol/day

Formic acid=0.011 kgmol/day

In bottom


35

Maliec acid = 41.32 kgmol/day

H2O= 2025.43kgmol/day

Acrylic acid= 0.9145kgmol/day

Formic acid=05377 kgmol/day

Butane=0.8047 kgmol/day

Chapter-5

Distillation

column

H2O=0.3726 t/d

C3H4O2=0.072/d

CH2O=0.000552 t/d

C4H2O3= 100t/d

H2O=40.18 t/d

C3H4O2=0.144 t/d

CH2O=0.0276 t/d

C4H2O3= 103.8411 t/d

C4H10=0.051t/d

H2O=39.886 t/d

C3H4O2=0.072 t/d

CH2O=0.027t/d

C4H2O3= 4.429 t/d

C4H10=0.051t/d

ENERGY BALANCE

5.1 Around air compressor

Inlet flow rate =77597.65 Kmol/day= 0.898 kmol/s

Inlet volumetric flowrate

P

nRTV

Where

n=0.898kmol/s

R=0.0821 m3atm/kmol K

Air

P2= 300Kpa

T2=?

Air

P1= 101.325Kpa

T1=235oC


37

P= 1 atm

T=298 K

So

V=21.97 m3/s

For fig 3.6 coulson vol 6 For this flow rate centrifugal compressor would be used

with efficiency EP=75%

Outlet temperature

m

1

2

12P

PTT

where

T1=25 oC

P1=101.325Kpa

P2=300 Kpa

PE

1-m

=1.4 (for air)

so

T2 =177.14 oC

Work per kmol

W =

1P

P

1-n

nR T Z

n1n

1

2

11

Where

m-1

1n =1.61

Z1=1 (at 25 oC and 1 atm from perry )

R=8.314 kJ/KmolK

Putting values

W=3326.56KJ/Kmol

Power requirement

P

E

kmol/s W Power

=3.983 Mwatt

5.2 Around butane compressor

Butane

P2= 300Kpa

T2=?

Butane

P1= 101.325Kpa

T1=235oC


39

Inlet flow rate =1829 Kmol/day= 0.021 kmol/s

Inlet volumetric flowrate

P

nRTV

Where

n=0.021kmol/s

R=0.0821 m3atm/kmol K

P= 1 atm

T=298 K

So

V=0.51 m3/s

For fig 3.6 coulson vol 6 For this flow rate centrifugal compressor would be used

with efficiency EP=67%

Outlet temperature

m

1

2

12P

PTT

where

T1=25 oC

P1=101.325Kpa

P2=300 Kpa

PE

1-m

=1.135 (for butane from perry)

so

T2 =88.32 oC

Work per kmol

W =

1P

P

1-n

nR T Z

n1n

1

2

11

Where

m-1

1n =1.216

Since

Tc=425.1 K, Pc=37.96 bar

So

Tr=Tc/T1=0.701 , Pr=P1/Pc=0.0263

So

From fig 3.8 coulson vol 6 , Z1= 0.98

R=8.314 kJ/Kmol oK

Putting values

W=2906.72KJ/Kmol

Power requirement


41

P

E

kmol/s W Power

=0.091 Mwatt

5.3 MIXING TEE

dTCnQiPi

2

1

T

T

1

dTCnQiPi2

3

1

T

T

Where

ni = no. of moles of ith component

Cpi = heat capacity of ith component

T1 = reference temperature=25 oC

T2 = 177.14oC

T3 = 88.32o C

Qin = Q1 +Q2

Air

T2=177.14oC

T4=?

Butane

T3=88.32oC

So

Qin= 3.6* 108 kJ/day

dTCnQ4

1

i

T

T

Piout

Qin = Q out

So

T4= 155 o C

RP2 ff HHQ

TC

QW

p

5.4 Around reactor:-

Inlet feed temperature = T1 =155 o C

Out let product temperature T2= 410 o C

Inlet cooling water temperature=t1= 25 o C

Outlet cooling water temperature=t2= 70 o C

Sensible energy

Is given by eq.

dTCnQiPi

1

ref

T

T

1

where



43


Tref = reference temperatiure=25 oC

dTCnQiPi2

2

ref

T

T

where



T1 = reference temperatiure=25 oC

Q = Q2 -Q1

So

Q= 6.5* 108 kJ/day

Heat of reaction

is given by following equation

RP2 ff HHQ

= -3.19*109 +7.75 * 10 8

= -2.4*109 kJ/day

Q evolved = Q + Q2

= -1.76*109 kj/day

Where ve sign shows that heat is evolved

Amount of water required:-

tC

QW

p

2

where

Cp = 4.184 KJ/kg oC

W= 9.3* 106 KJ / day

5.5 Around heat exchanger (E-101)

Inlet temperature =T1=410oC

Out let temperature =T2=90 oC

dTCnQiPi

2

1

T

T

Q=8.2*108KJ/day

5.6 Around absorber

Feed inlet temperature=T2 =90oC

Water inlet temperature=25 oC

Outlet temperature =T=?

dTCnQiPi

2

1

T

T

IN

where

T1= ref temperature =25oC

=1.59*108 kJ/day

dTCnQT

T

Piout

1

i

where

T1= ref temperature =25oC


45

since

QIN=QOUT

So

T=90oC




dTCnQiPi

2

1

T

T

Q=4.9*107KJ/day=574 kJ/s




Reaction temperature T2=135 oC

3

2

i

2

1

i

T

T

PiR

T

T

PidTCnHdTCnQ

=1.2*106+3.4*107+2.86*106

=1.7*108KJ/day

Q=2019.35KJ/s

5.9 Around distillation column

Condenser

energy balance equation of condenser is

H1Vn=[Lnhd+Dhd]+Qc

ni1i1

yHH at Ti

= 0.979*2886.91+0.0199*1878.34

=2864 KJ/Kmol

didid

xhh

Distillate (D)

Td=106 oC

Xd

hd=?

Vapors

Vn=21.5 kmol/hr

Ti=110oC

H1=?

yn

Reflux

Ln=135.3kmol/hr

hd=?


47

=0.979*2564.78+0.0199*1568.64

= 2542.14 KJ/Kmol

so

Qc=71392 KJ/hr =19.83 kJ/s

latent heat of water == 40683KJ/Kmol

m=0.979*221.5*40683

=8822047.5 KJ/hr

=2450.6Kwatt

latent heat of malice anhydride == 54800KJ/Kmol

m=0.0199*221.5*57800

=241550.2 KJ/hr

=67Kwatt

Qact=Qc+m+m

watt

Around reboiler

Temperature =185 oC (isothermal)

Vapour required =m=96Kmol/hr

Q=m=1452.23KJ/s

Since the feed enters and leaves isothermally at 185 oC

So sensible heat is zero

5.10 Condenser after flash vessel

mdTCnQ2

1

i

T

T

Pit

where

T1= inlet temperature =120oC

T2=outlet temperature = 100oC

= latent heat of vapourization

Qt = 1.69*109 KJ/day

Chapter 6


49

EQUIPMENT DESIGN

6.1 FIXED BED CATALYTIC REACTORS

INTRODUCTION

Fixed-bed catalytic reactors have been aptly characterized as the

workhorses of me process industries. For economical production of large amounts

of product, they are usually the first choice, particularly for gas-phase reactions.

Many catalyzed gaseous reactions are amenable to long catalyst life (1-10 years);

and as the time between catalyst change outs increases, annualized replacement

costs decline dramatically, largely due to savings in shutdown costs. It is not

surprising, therefore, that fixed-bed reactors now dominate the scene in large-scale

chemical-product manufacture.

TYPES OF FIXED BED REACTOR

Fixed-bed reactors fall into one of two major categories:

Adiabatic or

Non-adiabatic.

A number of reactor configurations have evolved to fit the unique

requirements of specific types of reactions and conditions. Some of the more

common ones used for gas-phase reactions are summarized in Table(4.1) and the

accompanying illustrations. The table can be used for initial selection of a given

reaction system, particularly by comparing it with the known systems indicated.

Fixed-Bed Reactor Configurations for Gas-Phase Reactions

Classification Use Typical Applications

Single adiabatic bed Moderately exothermic

or

endothermic non-

equilibrium

limited

Mild hydrogenation

Radial flow Where low AP is

essential

and useful where

change

in moles is large

Styrene from

ethylbenzene

Adiabatic beds in series

with intermediate

cooling or heating

High conversion,

equilibrium

limited reactions

SO2 oxidation

Catalytic reforming

Ammonia synthesis

Hydrocracking Styrene

from ethylbenzene

Multi-tabular

non-adiabatic

Highly endothermic or

exothermic reactions

requiring

close temperature

control to

ensure high selectivity

Many hydrogenations

Ethylene oxidation to

ethylene oxide,

formaldehyde

by methanol oxidation,

phthalic anhydride

production

Direct-fired

non-adiabatic

Highly endothermic,

high temperature

reactions

Steam reforming


51

SELECTION OF REACTOR TYPE

After analyzing different configuration of fixed bed reactors we have

concluded that for our system the most suitable reactors is multi tube fixed bed

reactor. Because oxidation of butane is highly exothermic reaction, so cooling will

be required otherwise the temperature of reactor will rise and due to rise in

temperature the catalyst activity and selectivity will be affected and in turn, the

formation of by-products will increase which is direct loss of productions.

As reaction temperature is already high 410 oC if we keep the process

adiabatic temperature of reactor will rise and the structure of the catalyst will be

changed and catalyst will be damaged. For such a situation the best reactor is

multi-tube fixed bed reactor

CONSTRUCTION AND OPERATION OF

MULTI-TUBE FIXED BED REACTOR

Because of the necessity of removing or adding heat, it may not be possible

to use a single large-diameter tube packed with catalyst. In this event the reactor

may be built up of a number of tubes encased in a single body, as illustrated in Fig.

The energy exchange with the surroundings is obtained by circulating, or perhaps

boiling, a fluid in the space between the tubes. If the heat effect is large, each

catalyst tube must be small (tubes as small as 1.0-in. diameter have been used) in

order to prevent excessive temperatures within the reaction mixture. The problem

of deciding how large the tube diameter should be, and thus how many tubes are

necessary, to

Feed Stream

achieve a given production forms an important problem in the design of such

reactors.

A disadvantage of this method of cooling is that the rate of heat transfer to

the fluid surrounding the tubes is about the same all along the tube length, but the

major share of the reaction usually takes place near the entrance. For example, in

an exothermic reaction the rate will be relatively large at the entrance to the reactor

tube owing to the high concentrations of reactants existing there. It will become

even higher as the reaction mixture moves a short distance into the tube, because

the heat liberated by the high rate of reaction is greater than that which can be

transferred to the cooling fluid. Hence the temperature of the reaction mixture will

rise, causing an increase in the rate of reaction. This continues as the mixture

moves up the tube, until the disappearance of reactants has a larger effect on the

rate than the increase in temperature. Farther along the tube the rate will decrease.

The smaller amount of heat can now be removed through the wall with the result

that the temperature decreases. This situation leads to a maximum in the curve of

temperature versus reactor-tube length.

Cooling

(or Heating)

fluid out

Feed Stream


53

Multi-tubular fixed bed reactor

EFFECT OF VARIABLES ON MULTI-TUBE FIXED

BED REACTOR

Particle Diameter

The overall heat transfer coefficient declines with decrease in particle size in

the usual practical range. Redial gradients increase markedly with decrease in

particle size. Small size, however, may improve rate or selectivity in some case by

making catalyst inner surface more accessible.

Tube Diameter

Reducing tube diameter reduces the radial profile. Heat transfer area per unit

volume is inversely proportion al to the tube diameter and reaction temperature is

affected by a change in this area.

Outside Wall Coefficient

Improvement up to the point where this resistance becomes negligible is

worthwhile. Boiling liquids are advantageous because of the high heat transfer

coefficient.

Heat of Reaction and Activation Energy

Accurate values should be used since calculated temp. is sensitive to

both of these, particularly to the value of energy of activation. This roust be

determined carefully over the range of interests, but calculated results should be

obtained based on different activation energies over the probable range of

accuracy for the data so that final equipment sizing can be done with a feel for

uncertainties.


55

Particle Thermal Conductivity

One of the mechanisms of radial heat transfer in a bed, conduction through

the solid packing which must quite logically depend on the thermal conductivity of

the bed, can be reasoned to have some dependence on the thermal conductivity of

the solid. But since it only affects one of the several mechanisms, the

proportionally cannot be direct. Differences in effective conductivity and the wall

heat transfer coefficient h between beds of packing having high and low solid

conductivity may be in the range of a factor of

2-3. The largest difference will occur at lower Reynolds numbers. Most catalyst

carriers have low conductivities, but some such as carbides have high

conductivities.

6.2 DESIGN PROCEDURE FOR MULTI TUBE FIXED

BED REACTOR

To calculate weight of catalyst required

2

1AoF

W A

A

X

X A

A

r

dX

If space time is know then space time = rate flow Volumetric

reactor of Volume

By the knowledge of bulk density of catalyst and weight of catalyst

Calculate volume of reactor

Volume of reactor = catalyst ofdensity bulk

catalyst ofweight

Decide the dimensions of tube; keeping in mind that

particlecatalyst of Dia

tubeof Dia > 10

Calculate volume of one tube and then number of tubes required

No. of tubes = tubeone of Volume

Reactor of Volume

Calculate the shell dia

NT =

2t

431st22

1s

P1.223

knkkDPk4

kD

Calculate pressure drop


57

G

D

1

CD

G

1

L

P

p1fp

Calculate heat transfer co-efficient

i) Shell side

ho =

2.0

8.0b

D

V0.011t1150

ii) Tube side

ddp

4.60.7

ppe

Gd3.50

k

dh

iii) Calculate overall heat transfer coefficient

Calculate area required for heat transfer.

Calculate area available for heat transfer.

Available area should be greater than required area

DESIGN CALCULATIONS OF MULTI-TUBULAR

FIXED BED REACTOR

FEED C4H10=116 t/d

N2=1822.26 t/d

O2=553.6 t/d

H2O=44.48t/d

PRODUCT C4H10=19.72t/d

O2=330.342 t/d

N2=1822.26t/d

CO2=52.77 t/d

CO=31.17 t/d

H2O=173.63 t/d

C3H4O2=1.548 t/d

CH2O=0.9844 t/d

C4H2O3= 103.88 t/d

Cooling Water in

Cooling Water Out


59

Volume of Reactor

Volumetric flow rate of feed to reactor = Vo = 19.04 m3/s

Space time = = 0.15s

Ref [US Patent #4317778]

Volume of reactor =V= Vo =2.86 m3

Type and volume of Catalyst

Vanadium Phosphorus Oxide (VPO) catalyst of 0.48 cm in the form of pellet

is used

Bed void fraction = = 0.4

From the appendix table 1.1

volume of catalyst = (1- V= 1.72 m3

weight of catalyst

Bulk density of VPO =c=2836 Kg/m3

Ref [www.chemistry periodic table .htm.]

Weight of catalyst = Vcc

= 4103.92 Kg

Tube length and diameter

Take length and diameter of tube to prevent deviation from plug flow assumption.

From appendix table 1.2

Dt/Dp > 10

Where

Dt = diameter of tube

Dp = diameter of particle

L/Dp>100

Where

L=length of tube

Take L=250 cm Dt=5cm

Dt/Dp=10.48

And L/Dp=520.83

Satisfactory

Number of Tubes

Volume of one tube = Vt = /4 Dt2 L

Dt = 0.05m L = 2.5 m

Vt = 0.006 m3

Total number of tubes = Nt=V/Vt

V = reactor volume =2.86 m3

Nt = 584

Diameter of Shell

Tube layout

Tube layout is triangular.

P=1.25Do

Where P= tube pitch

Do=outside tube diameter

Do= 6 cm

P = 7.5

Number of tubes at bundle diameter


61

2

1

D3

14N

t

N

where ND = number of tubes at bundle diameter

Nt= total number of tubes = 584

So ND =28

1NPDDi

Where P= tube pitch

ND = number of tubes at bundle diameter

Di = shell diameter

So Di = 2.1 m

Shell Height

Length of tube = 2.5 m

Leaving 20 % spacing above and below

So height of shell = 2 (0.2 2.5) + 2.5

= 3.5 m

PRESSURE DROP

Tube side pressure drop

Using Eurgen equation.

1.75G

D

1150

gD

G

1

L

P

PcfP

3

= bed void fraction = 0.4

DP = particle diameter = 4.8 mm = 0.48 cm

f= feed density = .00269 g/cm3

G = mass velocity = 0.274 g/cm2 Sec

= viscosity of feed = 0.00022 g/cm. Sec

gc = 980.67 cm/sec2

L = length = 2.5 m = 250 cm

Putting values in above eq. gives

P = 267.5 gm/cm2

And 1033.074 g/cm2 = 1 atm

So P = 0.25 atm

Shell side pressure drop

Mass flow rate = mw=107 kg/s = 235.4 lb/s

Flow area =Ac= 2ott2s DND4

Where Ds=shell inside diameter= 82.656 in

Nt=total number of tubes = 584

Dot=tube outside diameter= 2.375 in

Ac= 2778.65 in2= 1.79 m2

Wetted primeter = Nt Dot

= 4357.4 in

De= (4*flow area)/wetted primeter

Where De= Equivalent diameter= 2.55 in = 0.21 ft

G= mw/Ac

Where G= mass velocity =19.2 lb/ft2/s

e

e

GDR

where Re= Reynold number

De=equivalent diameter=0.21ft

viscosity=0.000403 lb/ft/s


63

Re=10004.96

From appendix fig.1.1

Friction factor for tube side = f = 0.00025

t

10

2

s

sDe5.22x10

LnfGP

where Ps= pressure drop

Gs= shell side mass velocity= 43920 lb/ft2/hr

L= length of tube = 8.2ft n=

number of passes=1

De=Equivalent diameter=0.21ft

S= specific gravity=1

14.0

w

s

= for water neglecting

Ps=0.00036 psi

Negligible

Calculations of Heat Transfer Co-efficients

Shell Side

Using Eagle and Ferguson equation.

0.2e

0.8

b

oD

V0.011t1150h

tb = average water temperature; oF

= 117.5 oF

De=Equivalent Diameter, in

De = 2.55 in

Now to calculate V = velocity of water in fps

Mass velocity = G = 19.92 lb/ft2/s

Water density =w=62.3lb/ft3

V=w

G

= 0.31fps

so ho

2.08.0

0.31

2.19117.50.0111150

= 184.09 Btu/ hr. ft2 oF

Heat transfer coefficient for wall

mw

i

DK

D

X=tube wall thickness=0.154 in

Di=tube outside diameter =2.375 in

Dm=tube mean diameter=2.22 in

Kw=Thermal conductivity=25 Btu/ hr. ft2 oF

mw

i

DK

D=0.00066 Btu/ hr. ft2 oF

Tube Side

An equation proposed by LEVA to find heat transfer co-efficient inside the

tubes filled with catalyst particles.

D

D4.6

0.7

pi

p

e

GD3.5

k

h D

G = tube side mass velocity=2017 lb/hr. ft2

= viscosity of tube side fluid=0.073 lb/hr. ft

k = 0.0265 Btu/hr. ft oF

Dp = diameter of particle = 0.0157 ft


65

D = diameter of tube = 0.164 ft

Putting values in above equation

hi = 25.51 Btu/hr. ft2 oF

Inside dirt coefficient From appendix table 1.3 for air

hid = 500 Btu/hr. ft2 oF

Outside dirt coefficient

From appendix table 1.3 for water

hid = 333.33 Btu/hr. ft2 oF

Over all H.T. Coefficient

oodiD h

1

h

1

h

1

U

1

mw

o

id

o

i

o

dk

Xd

h

d

d

d

do= tube outside diameter =2.375 in

dm= tube mean diameter

UD=overall heat transfer

By putting the values

UD=16.99 Btu/hr. ft2 oF

Area required for Heat Transfer

Q = 6.8*106 Btu/hr

LMTD = 389o F

A = 38916.99

10*6.8

LMTDU

Q 6

D = 1028.88 ft2 = 95.63 m2

Area Available for Heat Transfer

Length of tube = Lt = 2.5 m

Outer Dia of tube = Dot = 0.06 m

Surface area of one tube = totLD

= 3.14 0.06 2.5

= 0.47 m2

Total surface area available = 584 0.47

= 274.48 m2

so sufficient area is available for heat transfer


67

SPECIFICATION SHEET

Identification

Item Reactor

Item No. R-1

No. required 1

Function: Production of malice anhydride via butane

Operation: Continuous

Type: Catalytic

Multi tube, fixed bed

Chemical Reaction:

Catalyst:

Shape: Spherical

Size: 4.8 mm

Tube side:

Material handled Feed Product

(kg/hr) (kg/hr)

C2H5OH 86326 432.58

H2O 45.44 214.35

CH3CHO ----- 412.8

O2 635.28 484.96

N2 2090.82 2090.82

Temp (oC) 550 550

Tubes:

No. 709

Length 2.438 m

O. D 63.5 mm

Pitch 79.37 mm pattern Material of construction = copper

Shell side

Fluid handled = cooling water

Temperature 25oC to 45oC

Shell

Dia = 2.66 m

Material of construction = Carbon

steel

Heat transfer area required = 77.67 m2

6.2 DESIGN OF ABSORBER

ABSORPTIONS

The removal of one or more component from the mixture of gases by using a

suitable solvent is second major operation of Chemical Engineering that based on

mass transfer.

In gas absorption a soluble vapours are more or less absorbed in the solvent

from its mixture with inert gas. The 'purpose of such gas scrubbing operations may

be any of the following;

a) For Separation of component having the economic value.

b) As a stage in the preparation of some compound.

c) For removing of undesired component (pollution).

TYPES OF ABSORPTION

1) Physical absorption,

2) Chemical Absorption.

Physical Absorption

In physical absorption mass transfer take place purely by diffusion and

physical absorption is governed by the physical equilibria.


69

Chemical Absorption

In this type of absorption as soon as a particular component comes in contact

with the absorbing liquid a chemical reaction take place. Then by reducing the

concentration of component in the liquid phase, which enhances the rate of

diffusion.

TYPES OF ABSOR5SRS

There are two major types of absorbers which are used for absorption purposes:

Packed column

Plate column

COMPARISON BETWEEN PACKED AND PLATE

COLUMN

1) The packed column provides continuous contact between vapour and liquid

phases while the plate column brings the two phases into contact on stage

wise basis.

2) SCALE: For column diameter of less than approximately 3 ft. It is more

usual to employ packed towers because of high fabrication cost of small

trays. But if the column is very large then the liquid distribution is problem

and large volume of packing and its weight is problem.

3) PRESSURE DROP: Pressure drop in packed column is less than the plate

column. In plate column there is additional friction generated as the vapour

passes through the liquid on each tray. If there are large No. of Plates in the

tower, this pressure drop may be quite high and the use of packed column

could effect considerable saving.

4) LIQUID HOLD UP: Because of the liquid on each plate there may be a

Urge quantity of the liquid in plate column, whereas in a packed tower the

liquid flows as a thin film over the packing.

5) SIZE AND COST: For diameters of less than 3 ft. packed tower require

lower fabrication and material costs than plate tower with regard to height, a

packed column is usually shorter than the equivalent plate column.

From the above consideration packed column is selected as the absorber,

because in our case the diameter of the column is approximately 0.8 meter which is

less than 3 ft. As the solubility is infinity so the liquid will absorb as much gases as

it remain in contact with gases so packed tower provide more contact. It is easy to

operate.

PACKING

The packing is the most important component of the system. The packing

provides sufficient area for intimate contact between phases. The efficiency of the

packing with respect to both HTU and flow capacity determines to a significance

extent the overall size of the tower. The economics of the installation is therefore

tied up with packing choice.

The packings are divided into those types which are dumped at random into

the tower and these which must be stacked by hand. Dumped packing consists of

unit 1/4 lo 2 inches in major dimension and are used roost in the smaller columns.


71

The units in stacked packing are 2 to about 8 inches in size, they are used only in

the larger towers.

The Principal Requirement of a Tower packing are:

1) It must be chemically inert to the fluids in the tower.

2) It must be strong without excessive weight.

3) It must contain adequate passages for both streams without excessive

liquid hold up or pressure drop.

4) It must provide good contact between liquid and gas.

5) It must be reasonable in cost.

Thus most packing are made of cheap, inert, fairly light materials such as

clay, porcelain, or graphite. Thin-walled metal rings of steel or aluminum are some

limes used.

Common Packings are:

a) Berl Saddle.

b) Intalox Saddle.

c) Rasching rings.

d) Lessing rings.

e) Cross-partition rings.

f) Single spiral ring.

g) Double - Spiral ring.

h) Triple - Spiral ring.

DESIGN CALCULATIONS OF PACKED ABSORPTION TOWER

98 % of the n-Butane (entering the tower at 12.948 Kgmol/h), 100 % of the

Maleic Anhydride (entering the tower at 40.388 Kgmol/h), 100 % of the Formic

Acid (entering the tower at 0.826 Kgmol/h), and 100% of the Acrylic Acid

(entering the tower at 0.833 Kgmol/h) is to be absorbed. The absorption takes place

at 170 kN/m2 and 373K.

Basis: 1 hour of operation

Compositions of Components in Gas Mixture at Entrance

Compositions of Components in Liquid Mixture at Exit

Component Molec. For. Molec. Wt. Kg/h Kgmol/h Mol%

Acrylic Acid C3H4O2 72 60 0.833 0.024

N-Butane C4H10 58 751 12.948 0.4

Carbon Dioxide CO2 44 2110 47.954 1.42

Carbon Monoxide CO 28 1188 42.428 1.25

Formic Acid HCOOH 46 38 0.826 0.024

Maleic Anhydride C4H2O3 98 3958 40.388 1.200

Nitrogen N2 28 69436 2480.000 73.237

Oxygen O2 32 12587 393.344 11.616

Water H2O 18 6616 367.556 10.829


Acrylic Acid C3H4O2 72 60 0.833 0.04

N-Butane C4H10 58 736 12.689 0.60


Carbon Monoxide CO 28 0.000 0.00


73

Compositions of Components in Gas Mixture at Exit

STEP 1) SELECTION OF SOLVENT

The solubility data of these compounds shows that Formic acid, Acrylic and

Maleic Anhydride acid are very soluble in water. The least soluble component of

the four components, which are to be absorbed, is n-Butane. Therefore, we based

the design of our packed absorption tower on the solubility of n-Butane in water.


Maleic Acid C4H2O3 98 3958 40.388 1.91

Nitrogen N2 28 69436 0.000 0.00

Oxygen O2 32 12587 0.000 0.00

Water H2O 18 37120 2062.222 97.41


Acrylic Acid C3H4O2 72 0 0.000 0.00

N-Butane C4H10 58 15 0.260 0.01


Carbon Monoxide CO 28 1188 42.428 1.43


Maleic Anhydride C4H2O3 98 0 0.000 0.00

Nitrogen N2 28 69436 2480.000 83.66

Oxygen O2 32 12587 393.344 13.28

Water H2O 18 0 0.000 0.00

STEP 2) SELECTION OF PACKING

Our system, is corrosive, therefore, ceramic material is needed. The mass transfer

efficiency of Intalox Saddles is more than the Raching Rings and Berl Saddles. We

have selected this packing for the efficient operation. It is expensive than Raching

Rings and Berl Saddles. To avoid environmental pollution and to recycle n-Butane

to the reactor we have to absorb n-Butane as much as possible. Intalox Saddles (2

inch) of ceramic material is the best choice for our required conditions.

STEP 3) CALCULATION OF THE COLUMN DIAMETER

Most methods for determining the size of randomly packed towers are derived

from the Sherwood correlation, which is used here to find out the diameter of the

absorber.

The physical properties of gas can be taken as that of air at 60 0C and 170kN/m2

because concentration of n-Butane is very small in gas mixture and average

molecular weight of gas mixture = 29. The flow rate of water to absorb the n-

Butane has been optimized in the context of calculation of height of absorber,

which is 31230 Kg/h.

The abscissa of Fig.2.1 =

21

L

V

G

L

L = Flow rate of water 31230 Kg/h

G = Flow rate of gas 85956 Kg/h

V = Density of gas at 60 0C and 170kN/m2

L= Density of Water at 60 0C and 170kN/m2

V = PM / RT

P = 170kPs = 170kN/m2 = 1.7atm


75

M = 29 Kg / Kgmol

R = 0.082(atmKgmol/m3K)

T = 600C = 333K

V = (1.7atm* 29 Kg/Kgmol) / (0.082 atmKgmol/m3K)* 333K

V = 2 Kg/m3

L = 988Kg/m3 at 60 0C

Therefore abscissa of the Fig.2.1

=

21

L

V

G

L

= 0.02

For 42 mmH2O/m of packing height, from Fig.2.1

K4 = 2

From table 2.1, Fp = 22.3m-1

L is viscosity of water at 60 0C = 0.5 cp

21

0.1

L

L

p

VLV4*

13.1F

KG

G* = 7.6 (Kg/m2s)

G = Flow rate of gas 85956 Kg/h

G = Flow rate of gas 23.90 Kg/s

A = Cross-sectional area of the column

A = G / G*

A = 23.90 / 7.6

= 3.14 m2

D = [4*A/

m

STEP 4) CALCULATION OF THE COLUMN HEIGHT

CALCULATION OF NUMBER OF TRANSFER UNITS.

The solubility data of these compounds shows that Formic acid, Acrylic and

Maleic Anhydride acid are very soluble in water. The least soluble component of

the four components, which are to be absorbed, is n-Butane. Therefore, we based

the design of our packed absorption tower on the solubility of n-Butane in water.

Assumptions:

1) Absorption takes place isothermally at 373K

2) Carbondioxide, Carbonmonoxide, Oxygen, and Nitrogen are insoluble in water

at this temperature

3) n-Butane is our key component

A) EQUILIBRIUM CURVE

As the concentration of solute is very small, the flow of gas and the liquid will be

essentially cons

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