Production Of By Methanol Carbonylation: “acetic Acid”

PRODUCTION OF “ACETIC METHANOL CARBONYLATION

ACID”

BY

INTRODUCTION

Acetic acid, also known as ethanoic acid, is an organic chemical compound with the formula CH3COOH best recognized for giving vinegar its sour taste and pungent smell. Pure, water-free acetic acid (glacial acetic acid) is a colourless liquid that attracts water from the environment (hygroscopy), and freezes below 16.7°C (62°F) to a colourless crystalline solid. Acetic acid is corrosive, and its vapour causes irritation to the eyes, a dry and burning nose, sore throat and congestion to the lungs, however, it is considered a weak acid due to its limited ability to dissociate in aqueous solutions. Following is the sunnary of its general properties. • Systematic Names: Acetic Acid Ethanoic Acid • Molecular Formula: CH3COOH • Appearance: Colorless Liquid or Crystals • • • • •

Molecular Weight: 60.05 g/mol Melting Point: 16.5 oC Boiling Point: 118.1 oC pKa: 4.76 EU Classification: Corrosive ( C )

History Vinegar is as old as civilization itself, perhaps older. Acetic acid-producing bacteria are present throughout the world, and any culture practicing the brewing of beer or

wine inevitably discovered vinegar as the natural result of these alcoholic beverages being exposed to air.The use of acetic acid in chemistry extends into antiquity. In the 3rd century BC, the Greek philosopher Theophrastos described how vinegar acted on metals to produce pigments useful in art, including white lead (lead carbonate) and verdigris, a green mixture of copper salts including copper(II) acetate. Ancient Romans boiled soured wine in lead pots to produce a highly sweet syrup called sapa. Sapa was rich in lead acetate, a sweet substance also called sugar of lead or sugar of Saturn, which contributed to lead poisoning among the Roman aristocracy. The 8th century Persian alchemist Jabir Ibn Hayyan (Geber) concentrated acetic acid from vinegar through distillation. In the Renaissance, glacial acetic acid was prepared through the dry distillation of metal acetates. The 16th century German alchemist Andreas Libavius described such a procedure, and he compared the glacial acetic acid produced by this means to vinegar. The presence of water in vinegar has such a profound effect on acetic acid's properties that for centuries many chemists believed that glacial acetic acid and the acid found in vinegar were two different substances. The French chemist Pierre Adet proved them to be identical. In 1847 the German chemist Hermann Kolbe synthesised acetic acid from inorganic materials for the first time. This reaction sequence consisted of chlorination of carbon disulfide to carbon tetrachloride, followed by pyrolysis to tetrachloroethylene

and aqueous chlorination to trichloroacetic acid, and concluded with electrolytic reduction to acetic acid.

Frozen acetic acid FIG 1-1 By 1910 most glacial acetic acid was obtained from the "pyroligneous liquor" from distillation of wood. The acetic acid was isolated from this by treatment with milk of lime, and the resultant calcium acetate was then acidified with sulfuric acid to recover acetic acid. At this time Germany was producing 10,000 tons of glacial acetic acid, around 30% of which was used for the manufacture of indigo dye. Chemical properties Acidity The hydrogen (H) atom in the carboxyl group (−COOH) in carboxylic acids such as acetic acid can be given off as an H+ ion (proton), giving them their acidic character. Acetic acid is a weak, effectively monoprotic acid in aqueous solution, with a pKa value of 4.8. Its conjugate base is acetate (CH3COO−). A 1.0 M solution (about the

concentration of domestic vinegar) has a pH of 2.4, indicating that merely 0.4% of the acetic acid molecules are dissociated.

Solvent Liquid acetic acid is a hydrophilic (polar) protic solvent, similar to ethanol and water. It can dissolve not only polar compounds such as inorganic salts and sugars, but also non-polar compounds such as oils and elements such as sulfur and iodine. It readily mixes with many other polar and non-polar solvents such as water, chloroform, and hexane. This dissolving property and miscibility of acetic acid makes it a widely used industrial chemical. Detection Acetic acid can be detected by its characteristic smell.

Production by region Acetic acid is a very important commodity. Because of it’s versatile uses the production of acetic acid is increasing rapidly.

Production capacity is rapidly increasing in East Asia, also new plants are being set in Asia. The main companies producing acetic acid are Celanese (23 %) and BP (21 %). BP is the fastest growing producer and 2/3 of world uses BP technology. Acetic acid production by region is statistically shown as below.

FIG 1-2 The overall production capacity of world is about 6.4 million tones. Out of which 1.42 million tones is produced in Europe. And 0.57 million tones are produces in UK. Major uses of acetic acid 

Ethanoic anhydride

Used to manufacture cellulose acetate; pharmaceuticals (aspirin);

bleaching agents in detergents; agro-chemicals (herbicides) and dyes.



Vinyl acetate monomer (VAM) -

used to make Polyvinyl alcohol (polyethenol) (PVOH), a safe biodegradable water soluble polymer, polyvinyl acetate and ethylene vinyl acetate/ethylene vinyl alcohol - used in manufacturing/ processing a wide range of paints, adhesives, films, textiles and paper.

Ethanoic acid is one of the world's most important chemicals and serves as an intermediate in the production of a vast range of products. 

Ethanoate ester solvents - used in the production of many paints, dyes adhesives and inks.



Purified terephthalic acid (PTA)- used to make lightweight, recyclable plastics such as packaging films and plastic PET bottles; and in the manufacture of polyester fibres/ films.



Chloroethanoic acid - used in the production of wallpaper adhesives; herbicides; pharmaceuticals and cosmetics.

FIG 1-3 These wide uses are the cause of increasing world demand of acetic acid.

TECHNOLOGY SELECTION

Acetic acid can be produced industrially by a number of methods, some having discrepancies while others are

modern methods. By methanol carbonylation acetic acid is made by two major methods .The purpose of this chapter is to show how the best among them t is selected. The methods or the technologies used to manufacture acetic acid are as follows. 1. Monsanto Process. 2. Cativa Process One by one these methods are discussed below. Methanol carbonylation 1. Monsanto process The carbonylation of methanol produces acetic acid:

CH3OH

150 - 200 oC 30 - 60 atmosphere + CO CH3COOH Rhodium/Iodide Catalyst

This is the second largest industrial homogeneous carbonylation process with over 7 billion pounds of acetic acid produced each year using this technology. Prior to 1970, acetic acid was made using cobalt catalysts (BASF process) requiring rather severe conditions. In 1970 Monsanto commercialized a rhodium carbonyl iodide

catalyst that is commonly called the Monsanto Acetic Acid Process (developed in the late 60’s by James Roth and his research team at the corporate research center in St. Louis). In 1986 Monsanto sold the acetic acid plant and technology to British Petroleum (BP), but it is still commonly referred to as the Monsanto Acetic Acid process.

Merits:  Introduced methanol carbonylation which resulted in huge reduction of waste.  The whole process uses less energy as compared to non methanol carbonylation processes.

 Uses methanol, a cheaper feedstock.  Although methanol is usually manufactured from synthesis gas, produced from oil, it can also be produced from biomass (wood), municipal wastes and sewage. This may eventually lead to the process being no longer dependent oil. Demerits:  Rhodium metal is very expensive more expensive than gold.  Rhodium and iodide form insoluble salts like RhI3. So water level in the reactor vessel has to be kept high in order to prevent this.  A final distillation step has to be added in order to remove water. Adding to the costs and energy demands.  Rhodium is capable of catalyzing a lot of side reactions. 2. CATIVA PROCESS: Since 1997, ethanoic acid is increasingly being produced by the BP Cativa process, also involving the carbonylation of methanol. The difference is that it uses an iridium metal iodide complex ion as a catalyst, with promoters. 190 oC

27 atmospheres CH3OH +

CO

CH3COOH

Iridium Iodide Complex & promoter

Merits:  The mechanism involving iridium is different to that of rhodium as catalyst, as iridium works best under different conditions. The Cativa process also uses ruthenium compounds as promoters in the reaction. These increase the reaction rate by three times, even though ruthenium on its own has negligible catalytic activity in this system.  Iridium costs only about one fifth as much as rhodium.  The process is faster and more effective, requiring less catalyst to be used.

 Iridium is even more selective for methanol, which increases the overall yield and reduces by- products, resulting in lower purification costs and reduced waste.  Iridium complexes are more soluble in the reaction mixture than rhodium complexes. This means that the catalyst is not lost by precipitation and does not need replacing so frequently.  The energy needed at the distillation and purifying stages.  Existing plant can be modified to run the Cativa process at half the cost of building a new plant. This is referred to as retrofitting.  Cativa plants have a higher throughput a single plant can produce up to 75% more ethanoic acid than was previously possible using the Monsanto process.

BECAUSE OF ALL THESE ADVANTAGES WHICH CATIVA PROCESS HAS OVER MONSANTO PROCESS WE SELECTED CATIVA PROCESS.

PROCESS DESCRIPTION

Process for the formation of acetic acid by methanol and carbon monoxide involves carbonylation process. Carbonylation is the addition of carbon monoxide to a molecule. The reaction is given as follows 27 atmospheres, 190 oC

CH3OH

+

CO

CH3COOH -

[Ir (CO) 2 I2] + Ru (CO) 4I2

Figure 1 is a schematic diagram illustrating the process flow. Carbon monoxide is bubbled through a liquid reaction medium in a continuous stirred tank reactor maintained at 190 oC and 27 atmospheres. The reaction medium consists of methanol, acetic acid, iodomethane, methyl acetate, water, and iridium iodocarbonyl catalyst complex and ruthenium iodocarbonyl as a promoter (Ru (CO) 4I2). Side reactions producing byproducts are as follows: 27 atm, 190 oC

1. CH3OH + CH3COOH

CH3COOCH3 + H2O -

[Ir (CO) 2 I2] + Ru (CO) 4I2

27 atm, 190 oC

2. CO + H2O

CO2 + H2 -

[Ir (CO) 2 I2] + Ru (CO) 4I2

27 atm, 190 oC

3. 4CO + 2H2O

3CO2 + CH4 -

[Ir (CO) 2 I2] + Ru (CO) 4I2

Carbon monoxide is the excess reactant and methanol is the limiting reactant. As shown in Figure 1 stream A consists of carbon monoxide entering at pressure of 1 atm and 25 oC into compressor. The Compressor compresses the gas to 28 atm and temperature after the after cooler is 190 oC (stream B). Methanol enters into pump through stream C at 25 oC and 1 atm. The pump increases the pressure to 28 atm and temperature to 28 oC (stream D). All the reactions described above occur at 190 oC and 27 atm. The overall reaction is endothermic. Dowtherm Q is being used as a heating medium in the reactor jacket. Dowtherm Q enters into reactor jacket through stream 1 and leaves through stream 2. Some of the gases are dissolved or entrained in the liquid and leave with the stream of liquid going out of the reactor (stream E).While the rest of gases are used to build up pressure in the reactor and leave from top (stream F) through a control valve.

Stream E is than send to the blowdown drum where the entrained gasses are removed by impingement and escape from top of blowdown drum (stream G). These gases combine with the gases from the top of reactor (stream F) and go to the utility section where there heat is utilized. After giving there heat these gases are scrubbed before venting them to the atmosphere. Liquid stream from the blow down drum (stream H) at temperature of 190 oC and 27 atmospheres is passed through the heat exchanger, where it is cooled down from 190 oC to 110 oC using Dowtherm as cooling media. Dowtherm enters the heat exchanger (stream 3) at 30 oC and leaves (stream 4) at 80 oC. Dowtherm comes to the plant through parallel streams from the utility section at different temperatures as required. The cooled liquid stream at 110 oC and 26.4 atmospheres escapes from the heat exchanger (Stream I) and enters the throttling valve. The throttling valve reduces the pressure and produces a vapor liquid mixture (stream J) which is send to the flash tank operating at 1.4 atmospheres and 110 oC. Catalyst and promoter are highly non volatile and along with some liquid proportion consisting of acetic acid, methyl acetate, water, iodomethane, (Co-catalyst) iridium iodocarbony catalyst and ruthenium promoter are recycled back to the reactor

(stream K) at the temperature of 101 oC and 1.4 atmospheres (this stream is first pumped to the pressure of 28 atmospheres before entering it to the reactor). The vapor stream from the top of flash tank at the temperature of 119 oC through stream L are send to the distillation column operating at 1 atmospheres. Stream L consists of acetic acid, methyl acetate, water and iodomethane (Co-catalyst). Acetic acid is the required product and the heaviest among all other components with the boiling point of about 118 oC. All the other components except water have boiling point very much less than acetic acid. 99 % acetic acid and 1 % water escape from bottom of distillation column through stream M at temperature of 118 oC and 1 atmosphere. This stream is first to be cooled this is done by passing it through a water cooler. Water enters the cooler at 30 oC (Stream 11) and leaves a 55oC (Stream 12). Finally Product is obtained from stream M’ at 35 oC. Steam is used in reboiler of distillation column and enters through stream 5 and leaves through stream 6. And from the top of distillation column methyl acetate, water, iodomethane and little amount of acetic

acid is obtained ( stream N ) at the temperature of 71 o C and 1 atmospheres. Cooling water is used as a coolant in the condenser of distillation column. It enters the condenser through stream 7 and leaves through stream 8. Now we have to obtain our co-catalyst (iodomethane) from the mixture obtained from the top of distillation column and recycle it back to the reactor. Stream N is than send to water cooler. This reduces its temperature to 35 oC. Cooling water enters the cooler at 30 oC (stream 9) and leaves at 50 oC (stream 10). As there is a large density difference between iodomethane and other components methyl acetate, water and little amount of acetic acid which have very similar density. Iodomethane is heaviest in terms of density as compared to others. Therefore the stream from top of distillation column is send to decanter separator operating at 1 atmosphere. As iodomethane has high density so it moves to the bottom of decanter and is obtained in stream O which is recycled back to the reactor.

A mixture of methyl acetate, water and very little amount of acetic acid moves to the top of the decanter and is obtained in stream P.

MATERIAL BALANCE

MATERIAL BALANCE AROUND REACTOR

O

F

D

E

B

K MATERIAL GOING IN: Stream

Component

Mass (Kg/hr)

Mass %

D

CH3OH

2470

100

B

CO

2250

100

O

CH3I

750

100

K

CH3COOH CH3COOCH3 H2O CH3I

915.33 114.16 20.08 45.5

84 10 2 4

Catalyst = 2250 ppm &

promoter = 2300 ppm

TOTAL MASS IN = 6565 Kg/hr

MATERIAL COMING OUT: Stream

Component

Mass (Kg/hr)

Mass %

F

CO CO2 CH4 H2

37 364 1 7

.2 89 8.9 1.8

E

CH3COOH CH3COOCH3 H2O CH3I CO CO2 CH4

3694.1 1256.5 103.4 795.5 15.7 270 21.21

60 20 2 13 0.3 4.4 0.3

Catalyst = 2250 ppm &

promoter = 2300 ppm

TOTAL MASS OUT = 6565 Kg/hr According to the law of conservation of mass MASS GOING IN = MASS COMING OUT 6565 = 6565 Kg/hr Thus verified

MATERIAL BALANCE AROUND SEPARATOR

G

E

H

MATERIAL GOING IN: Stream

Component

Mass %

E

CH3COOH CH3COOCH3 H2O CH3I CO CO2 CH4

60 20 2 13 0.3 4.4 0.3

TOTAL

Catalyst = 2250 ppm &

100

promoter = 2300 ppm

TOTAL MASS IN = 6156.46 Kg/hr

MATERIAL COMING OUT: Stream Component Mass % G

CO CO2 CH4

5.12 87.97 6.91

TOTAL

100

Total = 306.9 Kg/hr

Stream

Components

Mass %

H

CH3COOH CH3COOCH3 H2O CH3I

63.15 21.48 1.77 13.60

TOTAL Catalyst = 2250 ppm &

100 promoter = 2300 ppm

Total = 5849.5 Kg/hr TOTAL MASS OUT = 6156.46 Kg/hr According to the law of conservation of mass MASS GOING IN = MASS COMING OUT 6156.46 Kg/hr = 6156.46 Kg/hr

MATERIAL BALANCE AROUND THE FLASH TANK

L

J

K MATERIAL GOING IN: Stream Component J

CH3COOH CH3COOC H3 H2O CH3I

TOTAL Catalyst = 2250 ppm & Total = 5849.5 Kg/hr

Mass % 63.15 21.48 1.77 13.60

100 promoter = 2300 ppm

TOTAL MASS IN = 5849.5 Kg/hr

MATERIAL COMING OUT: Stream Component Mass % K

CH3COOH CH3COOC H3 H2O CH3I

TOTAL

84 10 2 4

100

Catalyst = 2250 ppm & Total = 1095.06 Kg/hr

promoter = 2300 ppm

Stream

Component

Mass %

L

CH3COOH CH3COOCH3 H2O CH3I

58.4 24.0 1.8 15.8

TOTAL

100

Total = 4754.5 Kg/hr TOTAL MASS OUT = 5849.5 Kg/hr According to the law of conservation of mass MASS GOING IN = MASS COMING OUT 5849.5 Kg/hr = 5849.5 Kg/hr

MATERIAL BALANCE AROUND THE DISTILLATION COLUMN

N

L

M MATERIAL GOING IN: Stream Component L

CH3COOH CH3COOCH3 H2O CH3I

TOTAL

TOTAL MASS IN = 4754.5 Kg/hr

Mass % 58.4 24.0 1.8 15.8 100

MATERIAL COMING OUT: Stream Component N

CH3COOH CH3COOCH3 H2O CH3I

TOTAL

Mass % 0.04 58.05 3.81 38.11 100

Total = 1968.07 Kg/hr

Stream

Component

Mass %

M

CH3COOH H2O

99.7 0.30

TOTAL

100

Total = 2786.42 Kg/hr TOTAL MASS OUT = 4754.5 Kg/hr According to the law of conservation of mass MASS GOING IN = MASS COMING OUT 4754.5 Kg/hr = 4754.5 Kg/hr

MATERIAL BALANCE AROUND DECANTER

P

N

O MATERIAL GOING IN: Stream Component N

CH3COOH CH3COOCH3 H2O CH3I

TOTAL Total = 1968.07 Kg/hr TOTAL MASS IN = 1968.07 Kg/hr

Mass % 0.04 58.05 3.81 38.11 100

MATERIAL COMING OUT: Stream Component Mass %

O

CH3I

TOTAL

100

100

Total = 750 Kg/hr

Stream

Component

Mass %

P

CH3COOH CH3COOCH3 H2O

0.06 93.79 6.15

TOTAL

100

Total = 1218.07 Kg/hr TOTAL MASS OUT = 1968.07 Kg/hr According to the law of conservation of mass MASS GOING IN = MASS COMING OUT 1968.07 Kg/hr = 1968.07 Kg/hr

OVERALL MATERIAL BALANCE AROUND THE PLANT MATERIAL GOING IN: STREAM

COMPONENT

MASS (Kg/hr)

C

CH3OH

2470

A

CO

2250

TOTAL

4720

MATERIAL COMING OUT: STREAM

MASS ( Kg/hr)

Weigh t%

1.008 363.96 36.42 7.204 408.6

.2 89.1 1.8 8.9 100

CO CO2 CH4

15.7 270.0 21.2 306.9

5 88 7 100

H2o CH3COOC H3 CH3COOH

74.94 1142.36 0.75 1218.1

6.2 93.8 0.01 100

CH3COOH H2O

TOTAL

2778 8.42 2786.42

99.7 0.3 100

TOTAL

4720

E

COMPON ENT CO CO2 CH4 H2

TOTAL G

TOTAL P

TOTAL M

SO THE LAW OF CONSERVATION OF MASS IS VERIFIED. THAT IS TOTAL MASS IN = TOTAL MASS OUT 4720 Kg/hr = 4720 Kg /hr

ENERGY BALANCE

ENERGY BALANCE AROUND THE REACTOR

F

O 1 D

E 2

B K ENERGY GOING IN: Stream

Component

Energy ( KJ/hr)

D

CH3OH

18844

B

CO

432506

O

CH3I

4380

K

CH3COOH CH3COOCH3 H2O CH3I

204686

1

Dowtherm Q

34463298

TOTAL ENERGY IN: 35123714 KJ/hr

ENERGY COMING OUT: Stream

Component

Energy (KJ/hr)

F

CO CO2 CH4 H2

205296

E

CH3COOH CH3COOCH3 H2O CH3I CO CO2 CH4

2678755

2

Dowtherm Q

32239663

TOTAL ENERGY OUT: 35123714 KJ/hr ENERGY BALANCE AROUND THE SEPARATOR

G

E

ENERGY GOING IN:

H

Stream

Component

Energy (KJ/hr)

E

CH3COOH CH3COOCH3 H2O CH3I CO CO2 CH4

2678755

TOTAL ENERGY IN: 2678755 KJ/hr ENERGY COMING OUT: Stream

Component Energy ( KJ/hr)

G

CO CO2 CH4

Stream

Components

Energy (KJ/hr)

H

CH3COOH CH3COOCH3 H2O CH3I

2538933.788

139821.45

TOTAL ENERGY OUT: 2678755 KJ/hr

ENERGY BALANCE EXCHANGER

AROUND

THE

4

3

H

I

ENERGY GOING IN: Stream

Components

Energy (KJ/hr)

H

CH3COOH CH3COOCH3 H2O CH3I

2538933.788

Stream

Components

Energy (KJ/hr)

3

Dowtherm Q

112204.12

TOTAL ENERGY IN: 2651137.91 KJ/hr ENERGY COMING OUT: Stream

Components

Energy (KJ/hr)

I

CH3COOH CH3COOCH3 H2O CH3I

1336266.82

HEAT

Stream

Components

Energy (KJ/hr)

4

Dowtherm Q

1314871.09

TOTAL ENERGY OUT: 2651137.91 KJ/hr ENERGY BALANC AROUND THE FLASH TANK:

L

J

K ENERGY GOING IN: Stream Component

J

CH3COOH CH3COOC H3 H2O CH3I

Energy (KJ/hr) 1336266.8

TOTAL ENERGY IN: 1336266.8 KJ/hr

ENERGY COMING OUT: Stream Component Energy (KJ/hr)

K

CH3COOH CH3COOC H3 H2O CH3I

204686

Stream

Component

Energy (KJ/hr)

L

CH3COOH CH3COOCH3 H2O CH3I

1131580.7

TOTAL ENERGY OUT: 1336266.8 KJ/hr

ENERGY BALANCE AROUND THE DISTILLATION COLUMN: N

8

N’’’

7

N’

N’’

10

L

9

N’’’’

6 12 5 M ENERGY GOING IN: Stream Component

L

CH3COOH CH3COOCH 3

H2O CH3I

11

M’

Energy (KJ/hr) 1131580.7

Stream

Component

Energy (KJ/hr)

7

Cooling water

2788767.981

Stream

Component

Energy (KJ/hr)

9

Cooling water

28004

Stream

Component

Energy (KJ/hr)

11

Cooling water

97495.0

Stream

Component

Energy (KJ/hr)

5

Steam

6469834.25

TOTAL ENERGY IN: 10515681.93 KJ/hr

ENEGRY COMING OUT: Stream Component

Energy (KJ/hr)

8

Cooling water

8360969.1

Stream

Component

Energy (KJ/hr)

10

Cooling water

139884.1

Stream

Component

Energy (KJ/hr)

12

Cooling water

584620.2

Stream

Component

Energy (KJ/hr)

6

Steam

1298533.96

Stream

Component

Energy (KJ/hr)

N’’’’

CH3COOH CH3COOCH3 H2O CH3I

31381

Stream

Component

Energy (KJ/hr)

M’

CH3COOH H2O

100292.6

TOTAL ENERGY OUT: 10515681.93 KJ/hr

ENERGY BALANCE AROUND THE DECANTER

P

N’’’’

O ENERGY GOING IN:

Stream

Component

Energy (KJ/hr)

N’’’’

CH3COOH CH3COOCH3 H2O CH3I

31381

TOTAL ENERGY IN: 31381 KJ/hr

ENERGY COMING OUT: Stream Component Energy (KJ/hr)

O

CH3I

4380

Stream

Component

Energy (KJ/hr)

P

CH3COOH CH3COOCH3 H2O

27001

TOTAL ENERGY OUT: 31381 KJ/hr

OVERALL BALANCE AROUND THE PLANT: ENERGY IN:

Stream

Component

Energy (KJ /hr)

Utility flowrate ( Kg / hr )

B

CO

432506

--

D

CH3OH

18844

--

1

Dowtherm Q

34463298

100678

3

Dowthern Q

112204

13437.6

5

Steam

6469834

2378.7

7

Cooling water

2788768

133370

9

Cooling water

28004

1339

11

Cooling water

97495

4663

TOTAL ENERGY IN = 44410953 KJ/hr

ENERGY OUT: Stream

Components

Energy (KJ/hr)

Utility Flowrate (Kg/hr)

F

•CO •CO2 •CH4 •H2

205296

--

G

• CO •CO2 •CH4

139821

--

2

Dowtherm Q

32239663

100678

4

Dowtherm Q

1314871

13438

6

Steam

1298534

2379

8

Cooling Water

8360969

133370

p

• CH3COOCH3 • H2O • CH3COOH

27001

--

M’

• CH3COOH • H2O

100293

--

10

Cooling water

139884

1339

12

Cooling water

584620

4663

TOTAL ENERGY OUT = 44410953 KJ/hr Thus law of conservation of energy is verified. That is Energy going in at steady state = Energy coming out at steady state 44410953 KJ/hr = 44410953 KJ/hr

REACTOR DESIGN

REACTOR DEFINITION: Reactors or more precisely speaking chemical reactors are vessels designed to contain chemical reactions. TYPES OF GAS LIQUID REACTORS: Following are the four main types of gas liquid reactors. • • • •

Gas liquid continuous stirred tank reactor Bubble column Packed column Plate column

REACTOR SELECTION: The selection of the gas liquid reactor to be used is as follows. Reasons for selecting gas liquid continuous stirred tank reactor: o Excellent gas and liquid mixing. o High mass transfer occurs. o High heat transfer occurs. o Good temperature control. o Gas spends more time in liquid phase due to stirring this facilitates reaction in liquid phase. o Stirring greatly decreases the probability of coalescence of bubbles. o Very less pressure drop. o High liquid hold up. Reasons for not selecting Bubble Column: o Non negligible pressure drop because the column is usually high. o Less mass transfer as compared to stirred tank reactor. o Less heat transfer as compared to stirred tank reactor.

o There is more probability of coalescence of bubbles in thus type of reactor. o Less efficient gas liquid mixing as compared to stirred tank reactor o Problem of foaming can occur. Reasons for not selecting Plate Column: o Relatively high capital investment. o Less liquid holdup. o Flooding can occur. o Less efficient mass transfer as compared to stirred tank reactors Reasons for not selecting Packed Column: o Poor heat transfer. o High pressure drop. o Less efficient heat transfer as compared to mechanically stirred tank reactor. o Cost is more than as compared to mechanically stirred tank reactor. I

IMPELLER SELECTION: There are two main types of impellers. 1. Axial flow impellers 2. Radial flow impellers Axial flow impellers: The impellers that generate currents parallel to the shaft of the impeller are called axial flow impellers. Radial flow impellers: The impellers that generate currents which flow tangentially or radially from the impeller blade are termed as radial flow impellers. There are three main categories of impellers for low to moderate viscosity systems. 1. Propellers 2. High efficiency impellers 3. Turbines As the system of this process is of low to moderate viscosity, so we will have to choose from the categories described above. 1. Propellers: A propeller is an axial flow high speed impeller for low viscosity liquids. Small propellers turn at about 1150 to 1750 rpm. And larger ones turn at

bout 400 to 800 rpm. The direction of rotation is usually chosen to force the liquid downward, and the flow currents leaving the impeller continue until deflected by the floor of the vessel.As described above these are recommended for low viscosity liquids, but are not recommended for gas dispersion in liquid. Also note that they are only axial flow. So they are ruled out from the selection. 2. High efficiency impellers: They are designed to provide more uniform axial flow and better mixing. These impellers are widely used for low to moderate viscosity liquids, but are not recommended for very viscous liquids or for gas dispersion. So they are also eliminated from the selection. 3. Turbines: These push the liquid radially and tangentially. The currents they generate travel towards the vessel wall and than flow either upward or downward. Main types of turbines are shown below.

simple straight blade turbine

pitched blade turbine

disk turbine Concave blade CD-6 impeller Out of these only disk turbine and concave blade CD-6 impeller are used for gas dispersion. So the selection is to be made among them. Following is the selection between these two types of turbine impellers. • Disk turbine is useful for gas-liquid dispersion. But efficiency is not high. • Concave blade CD-6 disk turbine impellers are highly efficient. It is different from simple disk turbine in the manner that its blade are in a concave shape which cup the gas and disperse it more efficiently than simple disk turbine. • Concave blade CD-6 impeller show enhanced mass transfer for the same power/volume & superficial gas velocity than simple disk turbine. • There are two processes in gas dispersion which continuously oppose each other, that is dispersion and coalescence. CD-6 impeller decrease the probability of coalescence to a large extent by decreasing the probability of bubbles existing in the same portion of volume. • Also Concave blade CD-6 disk turbine impellers are quite economical • Concave blade CD-6 disk turbine impellers are being widely used in industries which require gas dispersion in liquid. Based upon all only these points Concave Blade CD-6 Impellers are selected.

SELECTION OF HEATING MEDIUM: As the reactions occurring in the reactor are overall endothermic therefore there is a need of a heating medium. Dowtherm Q (transfer fluid contains a mixture of diphenylethane and alkylated aromatics) has been selected as the heating medium. Because of the following reasons. 1. Compared to hot oils, it exhibits better thermal stability, particularly at the upper end of hot oils' use range, and significantly better low-temperature pumpability. 2. Does not corrodes the reactor jacket. 3. Highly efficient. 4. Suitable working temperature range. 5. Considered as economical in the thermal fluid family.

GAS LIQUID REACTOR

CONTINUOUS

STIRRED

TANK

DESIGN STEPS:  Calculate gas holdup.  Subtract the gas holdup from 1 to get liquid holdup.  Calculate the volume occupied by the liquid using overall residence time.  As liquid holdup is volume occupied by the liquid divided by total volume. From this calculate the total volume of reactor.  Give allowance of headspace to this total reactor volume  Assume a value of superficial gas velocity.  Divide the volumetric flowrate of entering gas with superficial gas velocity. To get area of reactor.  From this area calculate the diameter of the reactor.  Divide the volume of rector by area of reactor to get the height of reactor.  Now the ratio of height of rector to diameter of reactor should be equal to 1. This provides the right value of superficial gas velocity.  Now calculate the internal dimensions using standard shape factors.  Calculation of power absorbed by the impeller requires first the determination of power absorbed if only liquid phase is present

 For this purpose determine Reynolds number against this Reynolds number see the value of power number.  From this power number determine the value of power absorbed by the impeller.  Now determination of aeration number is done. Against this aeration number see the ratio of power absorbed by the liquid to the power absorbed by the gas.  This ratio is than used to calculate the power absorbed in the gas liquid mixing.  This is than followed by the calculation of gas and liquid holdup.  Than bubble diameter and sparger type is determined. REACTOR JACKET DESIGN STEPS:  Gather all the standard suitable jacket specifications.  Calculate the inside heat transfer coefficient.  Calculate the outside heat transfer coefficient.  Calculate the overall heat transfer coefficient and compare it with standard values.  Determine the pressure drop in the reactor jacket.

DESIGN CALCULATIONS: • Step 1 : Reactor Volume calculations  GAS HOLDUP(ЄG) ЄG = 0.25(QVGNR2/σ)0.45 ( Ref : Hassan and Robinson (1997) ) QVG = Volumetric flowrate of gas entering = 0.031m3/s NR = Rotational speed = 150 rpm = 2.5 rps σ = Surface tension = 0.07 N/m ЄG = 0.40  LIQUID HOLDUP (ЄL): ЄL = 1 – ЄG So putting the values e get ЄL = 0.60

 VOLUME: ЄL = VL/VT Where VL = volume occupied by liquid = Flowrate of liquid entering multiplied (QL) multiplied by overall residence time(tr). As residence time(tr) is as follows tr = 1 hr = 60 minutes = 3600 seconds So VL = QL * tr = 0.00087 * 3600 = 3 m3 VT = Total reactor volume =? Putting the values in the above equation we get VT = 5.2 m3 According to the book “Chemical Process Equipment Design And Economics by H.Silla (2003)  If Volume < 1.9 m3 than 15 % allowance is must  If volume > 1.9 m3 than 10 % allowance is must Because VT > 1.9 m3. So with allowance of 10 % Total Reactor Volume = VR = 6 m3 • Step 2:

Diameter and Height Calculation  SUPERFICIAL GAS VELOCITY(VSG): Assume a superficial gas velocity VSG = 0.01 m/s  AREA (A): A = QVG / VSG Putting the values we get. A = 3 m2  DIAMETER(DR ): A = (π/4) DR2 From this equation we calculate diameter. DR = 2.00 m  HEIGHT(HR): H R = VR / A Putting the values we get. HR = 2.00 m Optimum HR / DR ratio of a gas liquid CSTR is 1. Reference : book Chemical Reactors by Tremobuze and Euzen. And also in my case

HR / DR = 1 • Step 3 : Internal Dimension calculations According to the book “CHEMICAL REACTORS” by Trembouze and Euzen we have.  Distance from the bottom of rector to the bottom of impeller = H1 = DR/3 = 0.67m  Height of the impeller blade = H3 = DR/5 = 0.13 m  Width of the impeller blade = L3 = DR/4 = 0.17m  Diameter of the impeller = DA = DR/3 = 0.67 m  Width of the baffles = DR/12 = 0.17 m  Step 4 : Power absorbed by the impeller  First let us calculate the power absorbed by the liquid stream only = PL  Reynolds number = Re = (ρL*NR*DA2)/µL Nomenclature  ρL = Average liquid density = 1103 Kg/m3  µL = Average liquid viscosity = 0.00065 Kg/m.s  NR = Rotational speed of impeller = 150 rpm = 2.5 rps Substituting the values we get Re = 1882356

Now against this Reynolds number, Power number from figure 4.20 book “Chemical Reactors” by Trembouze and Euzen is. NP = 6.5 Now as power number is NP = PL / (ρL*NR3*DA5) Substituting the values and calculating PL PL = 4.0 hp/m3 Aeration number = NQG = QVG / ( NR*DA3) Putting the values we get NQG = 0.042 Now from figure 4.21 page # 276 book “chemical reactors“ We have against this NQG ratio PLG/PL equal to PLG/PL = 0.58 PLG = 2.5 hp / m3 • Step 7: Liquid holdup LH = 3.14 m3 • Step 8: Gas holdup GH = 2.07 m3

• Step 9: Sparger calculations and selection For this first we will have to calculate bubble diameter According to Jiang et al (1995) bubble diameter is given as below. dB2 = 8.8 * ((VSG*µL/σ)-0.04) * ((ρL* σ3)/gµL4) -0.12) * ((ρL/ρG) 0.22) * (σ/g ρL) Where dB = bubble diameter Substituting the values we get dB = 3 mm According to data given by Trembouze et al if bubble diameter comes this small than sintered metal sparger is used.

Selected Sparger: Sintered metal sparger • Step 10: Reactor jacket calculations

Jacket Specifications

• • • • • •

Jacket type = Spiral baffle jacket Height of jacket = 1.36 m Spacing between reactor and jacket = 75 mm = 0.075m Pitch = 200 mm= 0.2 m Tenetring = 200 oC = 473 K Tleaving = 190 oC = 463 K

• Calculation of overall heat transfer coefficient:  Heat transfer coefficient at the outside wall of the reactor: Using the following equation to calculate the heat transfer coefficient at the outside wall of reactor (from Chemical Engineering by Coulson and Richardson volume 6). Nu = C * Re 0.8 * Pr 0.33 Where Re = Reynolds number Pr = Prandtl number C = constant = 0.023 (Reference: Chemical Engineering by Coulson and Richardson volume 6) Re = (ρ * v * de) / µ

eq (E)

Where ρ = density of Dowtherm Q = 833.1 Kg/m3 v = Velocity with which dowtherm is moving in jacket = 8.7 m/s de= hydraulic mean diameter = 0.109 m µ = viscosity of dowtherm = 0.000323 Pa.s Substituting the values in equation E

We get Re = 2440250 And Pr = Cp * µ / Kf

eq (F)

Where Cp = Specific heat = 2.193 KJ/Kg .K Kf = Thermal conductivity = 0.0974 W/m. K Substituting the values in equation F we get Pr = 7.3 Substituting all these values in eq (D) Nu = 5702 As Nu = ho de / Kf

eq ( G )

Equating the value of Nusselt number in the equation above we get ho = 5091 W / m2 .C  Heat transfer coefficient at the inside wall of reactor Nu = 1.10 * Re .62 * Pr .33 Re = Reynolds number = ρ * NR * DA 2/ µ = 1882356 Pr = Prandtl number = Cp µ / K = 11 Nu = 18646.3

From definition of Nusselt number we have hi = 1448.30 W/m2.oC  OVERALL HEAT TRANSFER COEFFICIENT: 1/U = (1/hi) + (1/ho) + (Xw/K) Where Xw = wall thickness=.01 m K = Thermal conductivity = 22 W/m. oC U = 760 W/m2 oC Also from Chemical Engineering by Coulson and Richardson Volume 6 Figure 12.1 page # 639 U is between 750-1000 W/m2.oC • Calculation of pressure drop: From equation # 12.18 chemical engineering by Coulson and Richardson Volume 6 it is given as. ΔP = 8*Jf*(L/de)*ρ*v2/2 Where Jf = friction factor Putting the values we get ΔP = 17412 Pa = 0.2 atm = 2.5 psi

MECHANICAL DESIGN OF REACTOR

• Step 1: Design Pressure

Pdesign = Poperating + 0.075Poperating + Phydrostatic Pdesign = 29.2 atmospheres = 2958.7 K Pa • Step 2: Design Temperature Tdesign = 509 K • Step 3: Material Selection Following factors must be considered in selecting a suitable material of construction for the reactor. 1. 2. 3. 4.

Mechanical properties Corrosion resistance Availability Cost

Selection: The selected material is Zirconium (Zircadyne 702). The reasons for selection are as follows. a) The criteria of selection that overrides all others from chemical engineering point of view is corrosion. And whenever hot acids especially acetic acid because of its high corrosivity is present than Zirconium is used. b) It has great mechanical properties c) It is now easily available d) Its cost is not very high. It is very similar in price with high Nickel steel.

Selected Material = Zirconium (Zircadyne 702) • Step 4: Design Stress (f) f = 369300000 N/m2 = 369.3 N/mm2 • Step 5: Wall thickness e = PiDi/ (2*f-Pi) Where e = minimum wall thickness Pi= internal pressure Di=internal diameter Substituting the values in the above equation we get e = 0.008 m = 8 mm Giving 2mm corrosion allowance we get efinal = 10 mm = 0.001 m • Step 6: Head thickness Ellipsoidal head is chosen ehead=PiDi/ (2*J*f-0.2*Pi)

J = joint factor = 1 Substituting the values in the above equation we get ehead = 8mm = 0.008 m Observe the similarity between the values of wall thickness and head thickness. This indicates that the choice of ellipsoidal head is correct. • Step 7: Weight of vessel: It is given as Wv = CvπρmDmg(Hv+0.8Dm) t Where Wv = Weight of vessel Cv = A factor to account for manways, internal support etc. = 1.08 Hv = Length of cylindrical section = 2 m g = Acceleration due to gravity = 9.8m/s ρm = Density of material of construction Dm = Mean diameter of the vessel = Di + e = 2.01 m After giving 2 mm corrosion allowance it becomes ehead = 10mm Substituting the values in the above equation we get Wv = 15081.4 N • Step 8: Direct Stress

It is due to the weight of the vessel and its contents σw = W/π* (Di + t)*t Substituting the values we get σw = 1428335 N / m2 = 1.4 N/mm2 • Step 9: Principal stresses σ 1 = (1/2)*(σh+σz+sqrt((σh-σz)^2+4ĩ^2))= longitudinal σ 2 = (1/2)*(σh+σz-sqrt((σh-σz)^2+4ĩ^2))=circumferential σ 3 = 0.5(P)=radial Where ĩ = Torsional shear stress is very small and is usually neglected σz=Total longitudinal stress = σL + σw = 147 N / mm2 Substituting the values in the above equations we get σ1=292 N/mm2 σ2=147 N/mm2 σ3= 1.48N/mm2 Step 10: Vessel support The method used to support a vessel depends on the size, shape and weight of the vessel; the design temperature and pressure; the vessel location and arrangement; and the internal and external fittings and attachments. There are basically three types of support for vessels. 1. Saddle support

2. Skirt support 3. Bracket or lug support 1. Saddle Support: These are used for horizontal vessels. As the reactor is not a horizontal vessel in this case, so this choice is ruled out. 2. Skirt Support: These are used for tall vertical columns. As the H/D ration of reactor is 1 therefore it cannot be supported for this type of support. 3. Bracket Support: These are used for all types of vessels. Therefore as the above two types of supports are not suitable for the reactor, so this type of support must e chosen. Thus the selected support is BRACKET SUPPORT.

GAS LIQUID CONTINOUOUS STIRRED TANK REACTOR MECHANICAL DIAGRAM SHOWING DIMENSIONS:

10 mm 2.01 m

Ellipsoidal head

2m

Weight = 15081.4 N

0.17 m

0.67 m

1.36 m

Bracket Support 0.13 m

SPECIFICATION SHEET

2m

Equipment Name

Gas Liquid continuous stirred tank reactor

Mode Of Operation

Continuous

Operating Pressure

2735.7 KPa

Operating Temperature

463 K

Volume

6 m3

Height

2m

Diameter

2m

Area

3 m2

Superficial gas velocity

0.01 m / s

Power absorbed

2 hp/m3

Liquid Holdup

3.14 m3

Gas Holdup

2.07 m3

Bubble diameter

3 mm

Sparger type

Sintered metal sparger

Jacket type

Spiral baffled sparger

Overall heat transfer coefficient

760 W/m2.oC

Jacket pressure drop

17412 Pa

Design pressure

2958.7 K Pa

Design temperature

509 K

Material of construction

Zirconium (Zircadyne 702)

Design stress

369.3 N / mm2

Wall thickness

10 mm

Head type

Ellipsoidal

Head thickness

10 mm

Weight of vessel

15081.4 N

Support type

Bracket support

FLASH TANK DESIGN

Introduction If we say that Chemical Engineering is nothing but the combination of art and science to design and control the separation equipment, it won’t be a lie. In a chemical industry, more than the 70% of total capital investment is incurred on separation and purification equipment. These stats might highlight the importance of separation equipment in chemical industry.

Defining the problem:In Cativa Process, one of the product streams is coming out from the reactor. This stream contains the Acetic Acid; which is our sole product, and the Iridium Catalyst Complex. We have to maintain some liquid level in the reactor as well so that we might use this liquid as the solvent for the incoming feed stream. The catalyst has to be recycled back to reactor for further utilization. So we need equipment that might separate out the product (not essentially all of it) and recycle back some fraction of Acetic Acid along with the catalyst. A little amount of water should also be maintained in the reactor as this is the requirement of the technology (Cativa Process) we are using. So up till now, we have successfully defined our problem. Let’s look for a solution to it.

Looking for the solution:Now there are a number of equipments that are available to us for this purpose. We need to have a look at the physical conditions of the stream. All the components are in liquid state at 110 oC and 27 atm pressure.

We need to recycle some of the Acetic Acid and the catalyst back to reactor. Both of these are required to be there for further conversion. The feed mixture is in homogenous phase. This makes our choice quite simple. We can eliminate the possibility of a phase separator. One thing that must be kept in mind is that the solution has to be economical and quite effective. If we have a look at various industries; we find that most industries generate a second phase from this feed and recycle successfully some of the desired components in liquid state. This is quite an energy efficient process. Now let’s have a look at the possible choices that we have at our hand.

Possible Choices Available:We have our feed in liquid state in which catalyst is homogenously dissolved. We want some of the Acetic Acid, little amount of water and the catalyst recycled back to reactor. We’ll make use of equipment that can generate the vapor phase without expenditure of much of external energy and then successfully recycle the desired components back to reactor. One choice looks obvious. It’s the Flash Drum. There are other possible alternatives available to us, likewise Knockout Drum, Horizontal Flash Drum or the spherical one. All have their own characteristics and are used in specific situations. We’ll make use of Vertical Flash Drum.

Construction of a Flash Drum:When feed is flashed in a Flash Drum, vapor and liquid mixture is generated. As this mixture enters the drum, the surface area is increased, due to which pressure drop is generated. Right at eh entrance of the feed, there’s a splash plate in the drum. This splash plate directs the vapor and liquid flow downwards. This way the effect of gravity is enhanced. The liquid settles down at the bottom while the vapors with little momentum, change their path and rise up the vessel. At the top of the vessel, there’s a mist eliminator. Actually when vapors rise up the vessel, small liquid droplets also accompany them. The phenomenon of splashing is avoided by the use of splash plate. So our splash plate is serving two major purposes. First it helps us to avoid the splashing of

liquid. Secondly, it directs the vapor liquid mixture downwards which in turn enhances the effect of gravity. Due to this effect, liquid is separated out of vapor. One thing should be kept in mind is that most of the impaction process takes place at the splash plate. So it has to be mechanically sound so that it can handle all the impact. Now there are two kinds of mist eliminators. 

Vane type Mist Eliminator



Mesh Eliminators

Vane type mist eliminator consists of metallic plates arranged closely to each other. Vapors with small liquid droplets rise. The plates are arranged in such a manner that they provide a zigzag path to the incoming vapor and liquid droplets. Droplets due to inertia and large momentum strike the plates and are captured at the surface while the vapors change their path accordingly and escape the eliminator. The phenomenon is referred to as Impaction and the size increase of droplets is called as Coalescence. Hence vapors are collected at the top of the vessel. A vane type mist eliminator is shown in the following figure.

Now in mesh mist eliminators, a metallic or plastic wire mesh with a diameter ranging 0.006 to 0.011 in is used. The phenomenon is the same; impaction on the wire and then captured. Mist escapes the wire while droplets are captured at the surface where they coalesce and fall down as large drops. A mesh mist eliminator is shown in the following figure:

There’s a radial vane vortex breaker shown at the bottom of the vessel. The purpose of this vortex breaker is to avoid the phenomenon of Vortex Formation. There are a couple of causes that induce the vortex formation in the drum. The first one is the earth’s

rotational speed. Due to the earth’s rotational speed, anticlockwise vortex is observed in Northern Hemisphere while a clockwise motion is observed in Southern Hemisphere. Second reason is the introduction of feed in the vessel tangentially. Whenever feed is entered tangentially, vortexes are formed. Third reason is the vapors. Whenever there’s a two phase mixture and they differ in their velocity; then the fluid with lesser velocity and high density would start the rotational motion (Vortex Formation). In our case, we are handling a vapor-liquid mixture. Vapors are at a higher speed in the vessel while the liquid are a bit slower due to the impaction with the splash plate. So the vapors would induce the vortex to the liquid. The formation of vortexes brings some disadvantages to the system. Our system with vortex formed, experiences: 

Loss of valuable vapors



Downstream equipment damage



Loss of flow



Erroneous liquid level readings resulting in poor control



Vibrations caused by unsteady two phase flow.

The formation of vortexes is shown in the following figure:

To avoid the vortex formation, we should avoid the usage of a tangential feed line. Secondly, we can use a vortex breaker to get rid of vortexes. Following types of vortex breaker are usually used in the industry: 

Flat plate vortex breaker



Crosses



Radial vane or gratings

We are using a Radial Vane Vortex Breaker. A vortex breaker is stationary and it doesn’t move. If it starts the motion with the vortex then it wouldn’t break the vortex rather it would just weaken it. To break the vortex and get rid of it, we’ll have to fix the vortex breaker and make it stationary.

Why use Vertical Flash Drum?

Let’s carry out the process of elimination to justify our choice. We can simply rub aside the choice of Knockout Drum as it is used wherever there’s gas in the feed stream. In our stream there are no gases. We have only liquid phase. So we will not go for the Knockout Drum. Now we are left with Horizontal, Spherical and Vertical Flash Drums. Horizontal Drums are used when we have to handle a large liquid flow rate. But in our case we’ll see that the liquid flow rates wouldn’t be that huge. Instead we’ll have to deal with a high amount of vapor flow rate. Also Walas carried out a survey and in his book “Chemical Process Equipment Selection and Design” writes that out of every ten chemical industries; seven are making use of Vertical Flash Drums. The choice is made due to the economy and the ease with which we can handle the flow rates. A design engineer is required to start designing a Vertical Flash Drum by default and then after the design is complete we have a look at the L/D (length to Diameter Ratio) to decide which configuration to use. So we’ll follow the same procedure. We’ll design a Vertical Flash Tank and then would analyze the L/D ratio obtained to determine which configuration to use. Just remember one rule of thumb; for large liquid flow rates, we’ll use Horizontal Flash Drum and for small liquid flow rate, you’ll go for a vertical configuration. You can start designing any one of these and then the final decision would rest upon the L/D ratio of the drum. So don’t bother. Just start your computer software and begin designing any configuration. Let’s start the design of Vertical Flash Drum. Before the process of designing, we’ll see what exactly flashing is.

Throttling:When a fluid (liquid or a liq/vapor mixture) at high temperature and high pressure experiences sudden reduction in pressure, then some of the liquid is vaporized

and the phenomenon is referred to as Throttling. During the process the temperature of the feed stream doesn’t change that much and in such a case the process is called as Adiabatic Flashing. Actually for an ideal gas or a fluid behaving likewise an ideal gas, there’s no temperature drop. But in real fluids, little temperature drops have been observed. These temperature drops are due to the Joule-Thomson Effect and the frictional loss. Since there’s no appreciable change in the kinetic and potential energy; and also there’s no shaft work or heat transferred, therefore the eq: Δ (H + u2/2 + gz) = Q + Ws reduces to ΔH = 0. We know that the enthalpy depends upon the temperature of fluids. Since there’s no change in enthalpy so theoretically there will be no change in the temperature of the fluid stream. Usually for real fluids, a very little temperature drop is observed. In our case, the feed is at 110 oC and the pressure is 27 atm. We’ll suddenly reduce the pressure of the liquid stream and this would ultimately generate a vapor phase without the expenditure of any external energy. There will be ignorable temperature drop. All the beauty of equipment lies in this phenomenon. We are generating a second phase without expanding any external energy. But we know that energy is always conserved. We have generated the vapors on the expense of the pressure of the incoming feed So although the process of throttling makes us lose some of the energy contents of the feed stream, yet we get more benefits. Now the problem comes out to be the selection of the valve.

Selection of Valve:No ordinary valve would be used for this purpose. We need such a valve that would handle a feed stream with such a high temperature and pressure and allow it to expand suddenly. The valve would allow only one sided flow of the stream. There are a

number of options open to us. Globe Valve, Gate Valve, Butterfly Valve, Ball Valve etc are all at our disposal. But none of these is manufactured for the purpose of throttling. As we look for the best choice, we come to know that there’s a valve that is manufactured keeping in mind the sole idea of throttling. This is Lever sealed Plug Cock. The valve operates up to a temperature of 260 oC. It has plastic lining that makes it corrosion resistant. It has a tapered plug that is moved by a lever. The plug contains perforations just like a ball valve. As the feed stream passes through it, pressure drops from 27 atm to 1.4 atm. The temperature change is negligible. So after getting flashed, vapors are generated. The temperature of the stream remains more or less the same.

Determination of Flash Temperature:To determine the Flash temperature, we’ll have to determine the Dew Point and the Bubble Point Temperatures. To calculate the Bubble Point Temperature, we assume that all of our feed is saturated liquid. We assume a temperature and at that temperature, the K value for the component is determined. Then by multiplying this K value with the liquid weight fraction, we get the vapor fractions. The sum of these fractions should be unity in order to have the correct Bubble Point Temperature. So we see that it’s a hit-and-trial method. Similarly, we assume a temperature and at that temperature we determine the K value for the component. Then we assume that all of our feed is saturated vapor. So dividing these fractions with the K values, we get liquid fractions; whose sum should be unity. If the sum of liquid fractions is unity then our assumed Dew Point is correct. Taking the arithmetic average of this Bubble Point and Dew Point Temperatures, we get the Flash Temperature. K values for Iodomethane and Acetic Acid has been determined directly from the Himmelblau Software. The equation

that this software uses is V.P = A- {B/(T + C)} Here A,B and C are empirical constants while T is the assumed temperature. By dividing this Vapor Pressure (V.P) by the total pressure, we get the K value. K values for Water and Methyl Acetate have been determined by using the empirical relation given in Perry’s Chemical Engineering Handbook. The relation is: V.P = exp [C1+C2/T + (C3*lnT) + (C4*TC5)] * 9.869233E-06 atm Here C1, C2, C3, C4 and C5 all are empirical constants and there value is given in Chemical Engineer’s Handbook by Perry. The Vapor Pressure thus obtained is divided by the total pressure to get the K value at the assumed temperature. The process of calculating Bubble Point and the Dew Point Temperature is given below:

Compound

Xi

K at 101 oC

K*Xi

Acetic Acid

0.632

0.416

0.263

Methyl Acetate

0.215

2.673

0.574

Iodomethane

0.018

3.682

0.065

Water

0.136

0.736

0.100

1.000

1.002

This employs that our Bubble Point Temperature is 101 oC.

Compound

Yi

K at 119 oC

Yi/K

Acetic Acid

0.632

0.739

0.855

Methyl Acetate

0.215

4.163

0.052

Iodomethane

0.018

5.469

0.003

Water

0.136

1.348

0.101

1.000

1.011

This determines our Dew Point Temperature which comes out to be 119 oC. Now: Flash Temperature = (101 + 119)/2 = 110 oC

Determining the Vapor and Liquid Flows:Determination of Vapors going out and the liquid draining the drum is a result of some lethal calculations. These calculations are explained over here. First we make a material balance for a single component. It yields: Fxfi = Vyi + Lxi ………….Eq. I From Henry’s Law, we have: xi = yi /K Putting this value in Eq.I, we get: Fxfi = Vyi + L (yi/K)……….Eq. II 

yi = Fxfi / (V + L/K)

Since F = V + L which employs that V = F – L, therefore; yi = Fxfi / (F- L + L/K) yi = xfi / {1 – L/F (1 – 1/K)} Also from Eq. II, we can write that yiV = Fxfi / (1 + L/VK)………..Eq. III Which employs that:

yi = (Fxfi / V) / (1 + L/KV)……..Eq. IV

Now after determining yi’s, we can calculate xi’s by using the K values from the expression:

yi = Ki xi where Ki is determined by using the relation Ki = V.P/P

V.P stands for Vapor Pressure at the specified Flash Temperature. Eq. III can be written in the form as:

i=c

i=c

Σi=1 (yi V) = Σi=1 {Fxfi / (1 + L/KV)}……..Eq. V

Procedure to be followed for Flash Calculations:-

So simplifying all the procedure, we come to know that if we are to calculate V, L, yi’s and xi’s then we’ll have to follow these steps: 1.

Assume V.

2.

Calculate L = F – V

3.

Calculate L / V

4.

Look up for K values at Flash Temperature and Total Vessel Pressure

5.

Substitute values in; i=c

V = Σi=1 {Fxfi / (1 + L/KV)} If equality is obtained between the assumed V and the calculated V, then the assumed value is satisfactory. 6.

Calculate yi’s from Eq. IV

7.

Calculate xi’s from yi = Ki xi

Now using this procedure the values of V, L, yi’s and xi’s have been calculated for the Flash Drum. The calculations are given below.

Getting started with Design of Flash Drum:Since our calculations are based upon an hour of operation, so we have the following amount of vapor and liquid flow rates; FL = 1480.196 kg/hr

pL = 961.55 kg/m3

Fv = 4664.308 kg/hr

pv = 2.654 kg/m3

Vapor liquid separation factor, which is equal to (F L/Fv) / (pv/pL) ½; comes out to be 0.017. Using the graph, we notice that the Vapor Velocity Factor is equal to 0.35 m/sec.

Maximum design vapor velocity is obtained by multiplying the vapor velocity factor with {(pL – pv)/ pL} 0.5. The value of velocity comes out to be Uv = Kv* {(pL – pv)/pL} 0.5 Uv = 6.653 m/sec If we divide the vapor mass flow rate by the density, we get the volumetric flow rate. So VL = 0.488 m3/sec

Dividing volumetric flow rate by the vapor velocity, we get the minimum cross sectional area of the drum. Hence Amin = VL/Uv Amin = 0.073 m2 From this minimum cross sectional area, we can calculate the minimum diameter for the vessel. The minimum diameter is: Dmin = 0.306 m Actual internal diameter is obtained by adding 6in to this minimum diameter. Therefore D = 0.458 m For a vertical Flash Drum the surge time is in the range of 4 to 7 min and that for a horizontal vessel, it ranges between 7 to 12 min. Flash Drum used in Cativa Process has a surge time of 5 min. So multiplying this time with the liquid volumetric flow rate, we get the liquid volume held in the flash drum. Liquid Volume = VL * 300 Liquid Volume = 0.128 m3 Since the vessel is cylindrical, therefore its volume is equal to 3.145*(radius) 2*height. Using this relation, we can determine the liquid height in the vessel. The liquid height comes out to be: Liquid Height = 0.779 m Now both H. Silla and Coulson have suggested the following formula for determining the vapor height in the vessel. This formula is: Vapor Height = 1.5* D + 0.4 Vapor Height = 1.087 m

Now by adding the liquid and vapor heights, we can determine the total internal height of the vessel. Thus Total Height = 1.866 m The L/D ratio for the Flash Drum comes out to be 4.072 which is a satisfactory value. This ratio actually determines the type of vessel. It tells us that whether we should go for a horizontal vessel or a vertical one. If the value of L/D ratio is between 3 and 5, then a vertical flash drum is used. If its value exceeds 5, then a horizontal vessel should be employed.

Material of construction:Though material of construction is the part of mechanical design of the equipment but we can predict about it. Since we are dealing with acidic, corrosive fluid; therefore we’ll have to look for a material that is corrosion resistant. We come across two important choices that are corrosion resistant as well as economical. The flash drum can either be manufactured from Stainless Steel or we may make use of Aluminium. We can use either of the materials. Both have good mechanical strength, quite resistant to corrosion and are also cheap. Most of the heat transfer equipment in industry is made up from Aluminium Alloys. We are not that concerned with the heat transfer over here, so stainless steel is recommended as the priority material of construction.

DESIGN OF DISTILLATION COLUMN

Introduction:

The separation of liquid mixtures

into their various components is one of the major operations in the process industries, and distillation, the most widely used method of achieving this end, is the key operations in any oil refinery. In processing the demand for purer products, coupled with the need for greater efficiency, has promoted continues research into techniques of distillation. This process of getting pure products is accomplished by partial vaporization condensation.

and

subsequent

Distillation: “Process in which a liquid or vapour mixture of two or more substances is separated into its component fractions of desired purity, by the application and removal of heat”

TYPES OF DISTILLATION COLUMNS;

There are basically two types of distillation columns used in industries.  Batch columns  Continuous columns There selection criteria depends upon total number of stages and reflux ratio. As it is shown that when a large number of plates are used, then continuous distillation has the lowest reflux requirements and hence operating costs. If a smaller number of plates are used and high purity product is not required, then batch distillation is probably more attractive.

Batch Columns: In batch distillation the more volatile component is evaporated from the still which therefore becomes

progressively

richer in

the less

volatile

constituent. Distillation is continued, either until the residue of the still contains a material with an acceptably low content of the volatile material, or until the distillate is no

longer sufficiently pure in respect of volatile content. In batch operation, the feed to the column is introduced batchwise. That is, the column is charged with a 'batch' and then the distillation process is carried out. When the desired task is achieved, a next batch of feed is introduced. Most distillation processes operate in a continuous fashion, but there is a growing interest in batch distillation, particularly in the food, pharmaceutical, and biotechnology industries. The advantage of this separation process is that the distillation unit can be used repeatedly, after cleaning, to separate a variety of products. The unit generally is quite simple, but because concentration are continuously changing, the process becomes more difficult to control.

Continuous Distillation: In contrast to batch columns, a continuous feed is given to the column. No interruptions occur unless there is a problem with the column or surrounding process units. They are capable of handling

high throughputs and are the more common used. I will put light only on this type of distillation column.

CHOICE BETWEEN PACKED COLUMN

PLATE

AND

The choice between use of tray column or a packed column for a given mass transfer operation should, theoretically, be based on a detail cost analysis for the two types of contactors. However, the decision can be made on the basis of a qualitative analysis of relative advantages and disadvantages, eliminating the need for a detailed cost comparison. Which are:

(1)

Because of

liquid dispersion

difficulties

in

packed columns, the design of tray column is considerably more reliable.

(2)

Tray columns can be designed to handle wide ranges liquid rates without flooding.

(3)

If the operation involves liquids that contain dispersed solids, use of a tray column is preferred

because the plates are more accessible for cleaning.

(4)

For non-foaming systems the plate column is preferred.

(5)

If periodic cleaning is required, man holes

will

be provided for cleaning. In packed columns packing must be removed before cleaning.

(6)

For large column heights, weight of the packed column is more than plate column.

(7)

Design information for plate column is more readily available and more reliable than that for packed column.

(8)

Inter stage cooling can be provided to remove heat of reaction or solution in plate column.

(9)

When temperature change is involved, packing may be damaged.

(10)

Random-packed columns generally are not designed with diameters larger than 1.5 m, and diameters of commercial tray column are seldom less than 0.67m.

As my system is non foaming and diameter calculated is larger than 1.5m so I am going to use tray column. Also as average temperature calculated for my distillation column is higher that is approximately equal to 98oc. So I prefer Tray column.

PLATE CONTACTORS: Cross flow plate are the most commonly used plate contactor in distillation. In which liquid flows downward and vapours flow upward. The liquid move from plate to plate via down comer. A certain level of liquid is maintained on the plates by weir. Other types of plate are used which have no down comer (non-cross flow) the liquid showering down the column through large opening in the plates (called shower plates). Used when low pressure drop is required. Three basic types of cross flow trays used are (1) Sieve Plate (Perforated Plate) (2) Bubble Cap Plates (3) Valve plates (floating cap plates) I prefer sieve plate because:

(1) Their fundamentals are well established,

entailing

low risk. (2) The trays are low in cost relative to many other types of trays. (3) They can easily handle wide variations in flow rates. (4) They are lighter in weight. It is easier and cheaper to install. (5) Pressure drop is low as compared to bubble cap trays. (6) Peak efficiency is generally high. (7) Maintenance cost is reduced due to the ease of cleaning.

Label.Diagram Down comer And weir

Man Way

Calming zone

Plate support ring

Major Beam

Real picture of Sieve Tray

Factors Affecting Selection of Trays:  Relative Cost of plate will depend upon material of construction used. For mild steel, the ratio of cost between plates is Sieve plate 3.0

: valve plate : : 1.5 :

bubble-cap plate 1.0

 There is little difference in Capacity Rating of the three types (the column diameter required for a given flow rate). Sieve tray > valve tray > bubble-cap tray  Operating Range means the range of liquid and vapour flow rates which must be above the weeping conditions and below the flooding conditions. Operating range flexibility comparison is. Bubble cape tray > Valve tray > Sieve tray Sieve plate depends on the vapours flow through the holes to hold the liquid on the plate, and cannot operate at very low vapour flow rates. But with good design, sieve plate gives satisfactory operating range.  The Plate pressure drop will depends on the detailed design of plate but, in general, sieve plate gives the lowest pressure drop, followed by valves, with bubble-caps giving the highest.

Operation of Typical distillation Column: The operation of typical distillation column may by followed by figure. The column consists of a cylindrical structure divided into sections by a series of perforated trays which permit the upward flow of vapour. The liquid reflux flows across each tray, over

a weir and down a down comer to the tray below. The vapour rising from the top tray passes to condenser and then through an accumulator or reflux drum and a reflux divider, where part is withdrawn as the overhead product D and the remainder is returned to the top tray as reflux R. In the bottom there is reboiler which is used to give heat to the system. Liquid from the bottom of distillation column is fed to the reboiler which vaporises the in coming liquid. These vapours in turn move towards the bottom plate interact with the liquid over that plate. Due to which partial condensation of vapours occur. Also partial vaporization of liquid occurs too. That is less volatile component condensed first and more volatile component vaporizes first. This phenomenon occurs on each plate. Causing enrichment on each plate. A schematic of a typical distillation unit with a single feed and two product streams is shown below.

FACTORS AFFECTING DISTILLATION COLUMN OPERATION Vapour Flow Conditions Adverse vapour flow conditions can cause:  Foaming  Entrainment  Weeping/dumping  Flooding

 Foaming Foaming refers to the expansion of liquid due to passage of vapour or gas. Although it provides high interfacial liquidvapour contact, excessive foaming often leads to liquid build-up on trays. In some cases, foaming may be so bad that the foam mixes with liquid on the tray above. Whether foaming will occur depends primarily on physical properties of the liquid mixtures, but is sometimes due to

tray designs and condition. Whatever the cause, separation efficiency is always reduced.  Entrainment Entrainment refers to the liquid carried by vapour up to the tray above and is again caused by high vapour flow rates. It is detrimental because tray efficiency is reduced: lower volatile material is carried to a plate holding liquid of higher volatility. It could also contaminate high purity distillate. Excessive entrainment can lead to flooding.

 Weeping/Dumping This phenomenon is caused by low vapour flow. The pressure exerted by the vapour is insufficient to hold up the liquid on the tray. Therefore, liquid starts to leak through perforations. Excessive weeping will lead to dumping. That is the liquid on all trays will crash (dump) through to the base of the column (via a domino effect) and the column will have to be re-started. Weeping is indicated by a sharp

pressure drop in the column and reduced separation efficiency.  Flooding Flooding is brought about by excessive vapour flow, causing liquid to be entrained in the vapour up the column. The increased pressure from excessive vapour also backs up the liquid in the down comer, causing an increase in liquid hold-up on the plate above.  Depending on the degree of flooding, the maximum capacity of the column may be severely reduced. Flooding is detected by sharp increases in column differential pressure and significant decrease in separation efficiency.

Reflux Conditions: Minimum trays are required under total reflux conditions, i.e. there is no withdrawal of distillate. On the other hand, as reflux is decreased, more and more trays are required.

Feed Conditions: The state of the feed mixture and feed composition affects the operating lines and hence the number of stages required for separation. It also affects the location of feed tray.

State of Trays: Remember that the actual number of trays required for a particular separation duty is determined by the efficiency of the plate. Thus, any factors that cause a decrease

in tray efficiency will also change

the

performance of the column. Tray efficiencies are affected by fouling, wear and tear and corrosion, and the rates at which these occur depends on the properties of the liquids being processed. Thus appropriate materials should be specified for tray construction.

Column Diameter: Vapour flow velocity is dependent on column diameter. Weeping determines the minimum

vapour flow required while flooding determines the maximum vapour flow allowed, hence column capacity. Thus, if the column diameter is not sized properly, the column will not perform well.

DESIGNING STEPS OF DISTILLATION COLUMN  Calculation of Minimum number of stages.Nmin  Calculation of Minimum Reflux Ratio Rm.  Calculation of Actual Reflux Ratio.  Calculation of theoretical number of stages.  Calculation of actual number of stages.  Calculation of diameter of the column.  Calculation of weeping point.  Calculation of pressure drop.  Calculation of thickness of the shell.

 Calculation of the height of the column.

Condenser

REFLUX DRUM

PUMP

(1) Methyl Iodide = 0.212 (2) Acetic Acid = 0.0005 (3)Methyl Acetate = 0.62

FEED (1) Methyl Iodide = 0.074 (2) Acetic Acid = 0.65 (3)Methyl Acetate = 0.215 (4) Water = 0.065

(4) Water = 0.167

Reboiler

(1)Acetic Acid = 0.99

Process Design:

(2)Water = 0.01

Using Vapors Liquid Equilibrium Data Temperature of feed = 119 o C Temperature of top product =71 o C Temperature of bottom product = 118 o C From Material Balance: Feed Component

Fraction xf

Bottom

Top

Fraction Fraction xb

xd

0.074

0

0.212

0.65

0.99

0.0005

(3)Methyl Acetate

0.215

0

0.62

(4) Water

0.065

0.01

0.167

(1) Methyl Iodide (2) Acetic Acid

Heavy Key Component = Acetic Acid Light Key Component = Water

Calculation of Minimum no. of Plates: The minimum no. of stages Nmin is obtained from Fenske relation which is,

Nmin =

LN[(xLK/xHK)D(xHK /xLK)B]

LN (αLK/HK) average To find average geometric relative volatility of light key to heavy key:  Lk / Hk  avg   Lk / Hk  D  Lk / Hk  B 0.5

Average geometric relative volatility = 1.53 So, Nmin = 24 (reboiler is excluded)

Calculation of Minimum Reflux Ratio Rm: Using Underwood equation,

α A xfA α B xfB  1 q αA  θ αB  θ As feed is entering as saturated vapors so, q = 0 By trial,  = 1.68, Using equation of minimum reflux ratio,

α A xDA α B xDB   Rm 1 αA  θ αB  θ Putting all values we get, Rm = 4.154

Actual Reflux Ratio: The rule of thumb is: R = (1.2 ------- 1.5) R min

R = 1.5 R min R = 6.23

Theoretical no. of Plates: Gilliland related the number of equilibrium stages and the minimum reflux ratio and the no. of equilibrium stages with a plot that was transformed by Eduljee into the relation; N  N min

0.566     0.751   R  Rmin N 1 R  1    

From which the theoretical no. of stages to be,

N= 39

Calculation of actual number of stages: Overall Tray Efficiency: Using O'Connell's Correlation overall tray efficiency can calculated using average viscosity and relative volatility evaluated at average temperature.

Eduljee has expressed the O' Connell graphical method to a mathematical relation,

   Eo  51  32.5log  avg . avg     α avg =average relative volatility of light key component = 1.75 μ avg = molar average liquid viscosity of feed evaluated at average temperature of column

Average temperature of column = (118+71)/2 = 95 oC Feed viscosity at average temperature = avg = 0.3850 mNs/m2 So, Eo = 56.60% So, No. of actual trays = 39/0.566 = 68

Location of feed Plate: The Kirk bride method is used to determine the ratio of trays above and below the feed point.

 

2    x LK  B N x       D B HK log  .206 log        x HK  D   D x LK   NB  

From which,

Number of Plates above the feed tray = ND = 47

Number of Plates below the feed tray = NB = 21

Determination of the Column Diameter: Top Conditions

Bottom Conditions

Ln =155.13 Kgmol/hr

Lm = 226.8 Kgmol/hr

Vn = 180.02 Kgmol/hr

Vm = 180.02 Kgmol/h r

T = 71 0C

T = 118 0C

ρV = 3.469 Kg/m3

ρV = 1.87 Kg/m3

ρL = 1123.93 Kg/m3

ρL = 938.85 Kg/m3

Because liquid flow rates are greater at top so based upon bottom liquid flow rates.

Flow Parameter: FLV

L   n  Vn

 ρ v   ρ L

  

0.5

FLV = Liquid Vapor Factor = 0.0562

Capacity Parameter: Assumed tray spacing = 18 inch (0.5 m)

From

Fig

(15-5)

Plant

Design

and

Economics

for

Chemical Engineering, sieve tray flooding capacity, Csb = 0.0760 m/Sec Surface tension of Mixture = σ = 18.35 dynes/Cm   V nf  C sb    20 

0.2

 l  v    v  

0.5

Vnf=1.6740 m/sec Assume 90% of flooding then Vn=0.9Vnf So, actual vapor velocity, Vn=1.507 m/sec Net column area used in separation is An =mv/Vn Volumetric flow rate of vapors = mv mv = (mass vapor flow rate /(3600) vapor density) mv = 2.1184m3/sec Now, net area = mv/Vn = 1.4061m2 Assume that downcommers occupies 15%of cross sectional Area (Ac) of column thus: Ac = A n + A d Where, Ad = downcommer area Ac = An + 0.15(Ac)

Ac = An / 0.85 Ac=1.6542 m2 So Diameter of Column Is Ac =π/4D2 D = (4Ac/π) D = 1.4513 meter = 5ft (based upon bottom conditions)

Liquid flow arrangement: In order to find liquid flow arrangement first find maximum liquid volumetric flow rate So liquid flow rate = (Liquid mass rate)/ (3600) (Liquid

density)

Max Liquid Rate Is At the bottom of column so using "L m" values So Maximum liquid flow rate = 0.0050m3/sec So from fig11.28 cross flow single pass plate is selected

Provisional Plate Design: Column Diameter Dc= 1.4513 m Column Cross-sectional Area(Ac)= 1.6542 m2

Down comer area Ad = 0.15Ac = 0.2481 m2 Net Area (An) = Ac - Ad =1.406 m2

Active area Aa=Ac-2Ad = 1.1580 m2 Hole area Ah take 10% Aa = 0.1 × 1.1580 = 0.0462 m2 Weir length

Ad / Ac = 0.248 / 1.654 = 0.15 (From figure 11.31 volume 6) Coulson and Richardson’s Lw / dc Lw

=

0.80

= 1.452*0.80

= 0.733 m Weir length should be 60 to 85% of column diameter which is satisfactory Take weir height, hw

=

50 mm

Hole diameter, dh

=

5 mm

Plate thickness

=

5 mm

Check Weeping:  K 2  0.9 25.4  d h  U  min    v  1 / 2 where Umin is the minimum design vapor velocity. The vapor velocity at weeping point is the minmum velocity for the stable operation.

In order to have K2 value from fig11.30 we have to 1st find how(depth of the crest of liquid over the weir) Where how is calculated by following formula:

how=750{[Lm/lw*ρ]2/3} Maximum liquid rate

Lm = 4.73 kg/sec

Minimum Liquid Rate At 70% turn down ratio = 3.3Kg/sec At Maximum rate ( how)= 19.95 mm Liquid At Minimum rate (how) = 15.72 mm Liquid hw + how = 50 + 15.7245 = 66 mm Liquid From fig 11.30, Coulson and Richardson Vol.6 K2 = 30.50 So, U (min) = 8.89 m/sec Now maximum volumetric flow rate (vapors) Base = 2.12 m3/sec Top = 1.14 m3/sec At 70% turn down ratio Actual minimum vapor velocity =minimum vapor rate / Ah = 12.81 m/sec

So minimum vapor rate will be well above the weep point.

Plate Pressure Drop (P.D): Consist of dry plate P.D (orifice loss), P.D due to static head of liquid and residual P.D (bubbles formation result in energy loss + froth formed in operating plates) Dry Plate Drop: Max. Vapor velocity through holes (Uh) = Max Volumetric Flow Rate / Hole Area = 18.30 m/sec Perforated area Ap (active area) = 1.158 m2 Ah/Ap = 0.100 From fig. 11.34(for plate thickness/hole diameter) = 1.00 We get,

Co = 0.840 2

Uˆ h   V hd  51   C o   L This equation is derived for orifice meter P.D hd= 48.1 mm Liquid Residual Head (hr): hr = (12.5*103 / ρL) = 13.312 mm Liquid

So,

ht  hd  (hw  how )  hr Total pressure drop = 48.1 + (50 + 19.95) + 13.32 ht = 131.35 mm liquid Total column pressure drop Pa, (N/m2) = (9.81*10-3) htρLN =

82771.6 Pa

Down comer Liquid Backup: Caused by P.D over the plate and resistance to flow in the downcomer it self. hb = (hw+ how) + ht + hdc where hdc is equal to to head loss in downcomer. The main resistance to flow in downcomer will be caused by constriction in the downcomer outlet, and head loss in the down comer can be estimated using the equation given as,  l  hdc  166 wd    L Aap 

2

where Lwd is the liquid flowrate in downcomer, kg/sec and Aap is the clearance area under the downcomer, m2 Aap =hapLw

Where hap the height of bottom edge of apron above the plate. hap = hw – (5 to 10 mm) hap = 40.00 mm so, Area under apron “Aap” = .0464 m2 As this is less than area of downcomer Ad so using A ap values in above formula. So, hdc = 1.95 mm As a result, hb = 203.24 mm = 0.203 m hb < ½ (Tray spacing + weir height) 0.203 < 0.25 So tray spacing is acceptable

Check Residence Time: Sufficient residence time should be allowed in the downcomer for the entrained vapors to disengage from liquid stream to prevent aerated liquid being carried under the downcomer. tr =Ad hbc ρL/L(max)

tr = 10.02 sec It should be > 3 sec. so, result is satisfactory

Check Entrainment: (un) actual velocity (based on net area) = (maximum volumetric flow rate at base Vm / net area An) (un) actual velocity = 1.51 m/sec Velocity at flooding condition Uf = 1.67 m/sec So Percent flooding =un/uf = 0.90 = 90% Liquid flow factor FLV = 0.0562 From fig. 11.29 Coulson vol.6 fractional entrainment

ψ can be

found out. Fractional entrainment (ψ) = 0.0750 Well below the upper limit of (ψ) which is 0.1. Below this the effect of entrainment on efficiency is small.

No of Holes: Area of 1 Hole = (π/4) Dhole2 = 0.0000196 m2 Area of N Holes = 0.1158 m2 So, Number OF Holes = 5900

Height of Distillation Column: Height of column Hc= (Nact-1) Hs+ ∆H No. of plates = 68 Tray spacing Hs = 0.50 m ∆H= (1-1.5% of total height) for liquid hold up and vapor disengagement ∆H=0.55 m Total thickness of trays = 0.005*68 = 0.34 m So, Height of column = (68-1)*.50+ 0.55+0.34 = 35 meters

Chemical Process Control

Introduction A chemical plant is an arrangement of processing units (reactors, heat exchanger, pumps, distillation column, absorber, evaporators, tanks, etc), integrated with one another in a systematic manner. The plant’s overall objective is to convert certain raw materials into desired

products using available sources of energy, in the most economical way. During its operation, a chemical plant must satisfy several requirement imposed by its designers and general technical, economic, and social conditions in the presence of everchanging external influences (disturbances). Among such requirements are the following:(1) Safety: The safe operation of a chemical process is a primary requirements foe the well-being of the people in the plant. Thus the operating pressures, temperatures, concentrations of chemicals, and so on, should always be within allowable limits. (2) Production specifications: A plant should produce the desired amounts and quality of the final products. For example, we may require the production of 2 million pounds of ethylene per day, of 99% purity. Therefore, a control system is needed to ensure that the production level (2 million pounds per day) and purity specifications (99.5% ethylene) are satisfied. (3) Environmental regulations: Various federal and state laws may specify that the temperatures, concentrations of chemicals, and flow rates of the effluent from a plant be within certain limits. Such regulations exist, for example, on the amounts of SO2 that a plant can reject to the atmosphere, and on the quality of water returned to a river or a lake.

(4) Operational constraints: The various types of equipment used in a chemical plant have constraints inherent to their operation. Such constraints should be satisfied throughout the operation of plant. Foe example, pumps should maintain a certain net positive suction heads; distillation columns should not flooded; the temperature in a catalytic reactor should not exceed an upper limit since the catalyst will be destroyed. (5) Economics: The operation of plant must confirm with the market conditions, that is, the availability of raw material and the demand of final products. Thus it is required that the operating conditions are controlled at given optimum levels of minimum operating cost, maximum profit, and so on. All the requirements listed above dictate the need for the continuous monitoring of the operation of chemical plant and external intervention (control) to guarantee the satisfaction of operational objectives. This is accomplished of a rational arrangement of equipment (measuring devices, valves, controllers, computers) and human intervention (plant designers, plant operators), which together constitute control system.

Control Over Continuous Stir Tank Reactor: Temperature Control:

For temperature control we employed cascade control configuration. In a cascade control configuration we have one manipulated variable and more than one measurement. The reaction is endothermic and heat is supplied by dowtherm, which flows in the jacket around the tank. The control objective is to keep the temperature of the reacting mixture, T, constant at the desired value. Possible disturbances to the reactor include the feed temperature T f. and the dowtherm temperature Th. The only manipulated variable is the dowtherm flow rate Fh. Configuration:

We control the reaction temperature by measuring T h and taking control action before its effect has been felt by the reacting mixture. Thus if Th goes down, increase the flow rate of dowtherm to give the same amount of heat. Decrease the flow rate when Th increases. Disturbances arising within the secondary loop are corrected by the

secondary controller before they can affect the value of the primary controlled output.

Control Over Heat Exchanger: The control objective is to keep the exit temperature at 110 oC for our shell and tube heat exchanger. The possible disturbances are: (1) Offset of temperature value from its desired value of 110oC. (2) Variation in temperature of dowtherm used as a coolant media.

Control over Distillation Column: Cascade control is usually employed to regulate the temperature (and consequently the concentration) at the bottom or top of a distillation column.