33170313 ammonia production project

8/8/2019 33170313 Ammonia Production Project

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1200 TPD Ammonia production By Vikrant Rana(DCET)

Introduction

Ammonia is an important ni trogenous material used as fert i l izer.

Most of ammonia is made synthetically.It is also obtained as by-product in

some cases.

Ammonia gas is used directly as fertilizer in heat treatment, paper pulping, nitric

acid and nitrate manufacturing, nitric acid esters and nitrocompound manufacture,

explosive of various types and as a refrigerant.

The most important field is of fertilizer. Importance of ammonia in this field is due

to the fact that it is by far the most simple form of actual nitrogen which can be

administrated and the most economical source of other nitrogen chemicals. Hence

ammonia is the backbone of Nitrogen fertilizer processes. Its production is the first

consideration in the initiation of a fertilizer industry. Great strides have been made in the

last few year in the economy and efficiency of ammonia processes.The synthesis of ammonia from Nitrogen and hydrogen was f irst

chemical react ion to be carr ied out under high pressure on commercial

scale. This synthesis reaction has probably been studied more extensively

and is better understood than any other high presure reactions.

The name of Haber, Corl Bosch, Claude, Casale, Fauser, Mont Cenis

s tand out f rom those of ear ly engineers and scient is ts , worked for the

development of first commercial synthesis.

In India, t i l l 1948, the only factory which was producing synthetic

ammonia was small concern, Mysore Chemicals and Fert i l izers Company


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Ltd. at Belegula, Mysore. The Fer t i l izer Factory at Sindr i , Indias f i rs t

major state-owned enterprise, went into production on October 30, 1961. a

gradual expansion followed and soon fert i l izer plants were commissioned

at Nangal (1961), Trombay (1968), Namrup (1969), Durgapur (1974) and

Barauni (1976).

In 7 t h decade of 20 t h century,a revolution in ammonia manufacturing

came about as the reciprocat ing compressors common at that t ime were

replaced by centrifugal units. At that time, the ammonia

process was regarded as highly ef f icient , a mature process , but careful

examination of the process step by step and the severe economic pressure

led to major improvements and cost drop approaching to 50 percent. Some

of the impor tant developments are bet ter catalys ts , bet ter conver ters ,

r eplacement of r ec iprocating compressors to cent ri fugal unit s, lone

pressure techniques, energy conservation, digital computer control process

and the most recent development on which researches have been made and

c om me rc ia l b ac kg ro un ds a re b ei ng s tu die d i s t o c om mi ss io ne d a

Compressor Ammonia Plant . High pressure ammonia synthesis becomes

economically attractive with this technique of compressorless manufacture

process of ammonia. As without application of compressor, high pressure

synthesi s gases are obta ined which have many advantages over low

pressure techniques e.g. at high pressure ammonia condenses at normal

cooling water temperature, reduced amount of recirculat ion thus replacing


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recirculat ing compressure by a stat ic ejector . Trend in again in favour of

high pressure techniques for ammonia manufacture.

Process And Process Selection

Ammonnia manufacturing consists of six phases :-

1. Manufacture of Reactant gases.

2. Purification

3. Compression

4. Catalytic reaction

5. Recovery of ammonia formed

6. Recirculation

RAW MATERIAL-

Raw material requirement consists of a sources of hydrogen, the nitrogen is

obtained from air, the feed stock supplies the hydrogen. The production of ammonia

depend upon cost and operability of available raw material for hydrogen. Different raw

materials for hydrogen can be

Wood

Coal Refinery gas

Electrolysis of water Naphtha

Coke oven gas Fuel oil

Natural gas Crude oil

Liquid petroleum gas

MANUFACTURING PROCEDURES-

The principal manufacturing process that are used for synthetic ammonia

production are steam water gas process, the steam hydrocarbon process, the coke oven

gas process & electrolysis of water.


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Much research work is s t i l l diverted to f ind cheaper manufacturing

methods .Many new plants use ethanol amine to remove carbon dioxide

from gas streams instead of high pressure water.

Nitrogen required for synthesis gas is obtained usual from an air liquefaction

plant. In water gas and natural gas process, nitrogen is obtained in same manufacturing

process as hydrogen. Air is added in secondary reforms of natural gas process and reacts

with carbon monoxide to furnish carbon dioxide and nitrogen.

Cost is greatly influenced by the pressure, temperature, catalysts and

raw material used. Raw material selection depends upon the availabil i ty

and cos t of r aw mater ia l . I f the p lant i s near some ref inery then raw

material can be naphtha, natural gas or LPG, but in case i t is near a coal

mine, one should depend on coal as raw material for hydrogen. There may

be some other cases where electricity is readily available to meet the

r eq ui re me nt s f or e le ct ro lys is th en h yd ro ge n c an b e o bt ai ne d b y

electrolysis of water

Comparsion of processes

Many variation of the original Haber process for the synthesis of ammonia are now used

in commercial practice, some varying to such an extent that they are identified by a name,

often that of group of men developing them. Important among these are modified Haber

Bosch, Claude, Casale, Fauser, and Mont Cenis processes. All of them are fundamentally

the same in that nitrogen is fixed with hydrogen as ammonia in the presence of a

catalyst , but have variation in arrangement and construction of equipment , composition

of catalyst, temperature and pressure used. Table C gives a condensed comparison of

different processes 40% conversion of the gas upon passes through a single converter and

85% conversion after passes through a series of converter . gas is vented after one pass

through the converter.


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Table C : Synthesis Ammonia Systems

Designation Pressure Temp.

0C

Catalyst Recir-

culation

Conver.

%

Haber-Bosch 200-350 550 Doubly

Promoted

iron

Yes 8

Modified

Haber-Bosch

200-300 500-550 -Do- Yes 20-22

Claude 900-1000 500-650 Promoted

iron

No 40-85

Casale 600 500 Promoted

iron

Yes 15-18

Fauser 200 500 Promoted

iron

Yes 12-23

Modified Haber- Bosch Process

The ammonia concentration in the circulation gas leaving the catalyst is 10-11

mole %. In the condenser the ammonia concentration is reduced by condensation to

equilibrium at the exit temperature of the condenser. The condensed ammonia is

separated from the circulating gas, and the gas is boosted in pressure by a circulating

compressor to overcome the pressure drop in the synthesis loop.


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The Claude Process

This process depar t f rom the Haber process more than any of the

o ther ammonia s yn thes es p rocess , t he r es idua l gas i s was ted t o t he

a tmosphere or u ti li zed for i ts heat content. The l arge amount of heat

evolved in operating at space velocit ies of 100,000 with as much as 40%

of the hydrogen n it rogen mixture conver ted to ammonia in one pass ,

called for a special converter design. In his original process, Claude used

-+ from the liquefaction of air.

Among advantages claimed for the Claude process are the following:

1. Greater compactness, simplicity, and case of

construction of the converter , s ince under the high pressure used

the gases have smaller volume.

2 . E li mi na ti on o f t he expensi ve hea t exchanger s r equi red i n

processes operated at lower pressure.

3. Removal of ammonia wi th water cool ing alone, rather than by

ammonia refrigerators or scrubbing processes.

Ci ted agains t these advantages are the shor ter l i fe of conver ters , high

apparatus upkeep in the high-pressure operation, and the efficiency loss in

wasting approximately 20% of the makeup gas , which is unconverted.

Modifications of the Claude process include recycling of the gas through

the synthesis converters.


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The Casale Process

Pressure of 500-900 atm are used in this process which is otherwise distinguished by

the method used for controlling catalyst temperature in the specially designed

converter. A method involving recirculating gas around a synthesis loop,

similar to the Haber process, is used. As in the Claude process, the higher

pressures allow liquefaction of the ammonia at temperatures that can be

attained by water cooling. The basis for heat control of the catalyst is to leave 2

or 3% ammonia in the gas to the converter, thereby showing down the rate of

formation of ammonia and eliminating excessive heating of the catalyst.

The Fauser Process

This pressure incorporates some features not previously mentioned. Electrolytichydrogen from Fauser cells and nitrogen from liquid air unit or from a purification unitutilizing tail gases from absorption towers in the ammonia oxidation plant are used. Themixture of hydrogen and nitrogen is compressed to 200-300 atm and, after passingthrough an oil separator goes to an oxygen burner.

In the oxygen burner any oxygen contained in the gas mixture combines with hydrogen

in the presence of a copper catalyst the water formed is condensed out in a cooler and

removed in a water separator.

The Mont Cenis Process

This process was originally developed to use hydrogen separated from coke oven

gas by a liquefaction process, and nitrogen was obtained by the liquefaction of the air.

The essential characteristics of the Mont Cenis process are its operating pressure of 100

atm or less.

Mixed hydrogen and nitrogen, after being compressed to 100 atm and heated to about

3000C in interchangers, passes through a carbon monoxide purifier. In the purifier, while

in contact with a nickel catalyst, carbon monoxide and oxygen contained in small

quantities in the gas react with hydrogen to form methane and water.


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Material Balance

Ammonia Production = 1200 TPD=2941.18 kmol/hr

N2 + 3H2 2 NH3

Let % conversion=20%

Which means, 1 kmol N2 forms 2x0.2=0.4kmol NH3

Assuming losses=8%

Amoonia production=2941.18/0.92=3196.93kmol/hr


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Hence, N2 required=3196.93/0.4=7992.34kmol/hr

Air required=7992.34x100/79=10,117kmol/hr=10117x22.4=226621 Nm3/hr

Siimilarly, H2 required=3x7992.34=23977kmol/hr

Reaction taking place in the gasifier isCnHm + n/2O2 nCO + 2H2OCnHm + nH2O nCO + (m/2+n)H2O

At high temperature,

CnHm + ( n + m/4)O2 nCO2 + m/2H2O

Composition is given in weight percent.

C = 84.98%

H = 12.07%

S = 1%

N = 0.4%

O = 1.5%

Composition of gases coming out of the gasifier is given as-

CO = 42%

H2 = 51.23%

H2S = 0.23%

CO2 = 5.7%

N2 = 0.18%

Ar = 0.02%

CH4 = 0.5%

Material balance across H2S absorber


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Basis: 100 kmoles or Raw gas

CO = 42 kmoleH2 = 51.23 kmole

H2S = 0.23 kmole

CO2 = 5.7 kmole

N2 = 0.18 kmole

Ar = 0.09 kmole

Now, maximum concentration of H2S in H2S free raw gas in limited to 0.3 ppm.

Amount of H2S to be removed = 0.229997 kmole

In H2S free gas

H2 = 51.23 kmole % of H2 = 51.23/99.77x100 = 51.34

CO = 42 kmole % of CO = 42/99.31 X 100 = 42.09

H2S = 0.3 ppm % of H2S = 0.3 ppm

CO2 = 5.77 kmole % of CO2 = 5.77/99.31x 100 = 5.78

CH4 = 0.5 kmole % of CH4 = 0.5/ 99.31x 100 = 0.501N2 = 0.18 kmole % of N2 = 0.18/99.31x 100 = 0.18

Ar = 0.09 kmole % of Ar = 0.09/99.31x 100 = 0.09

Total no. of moles = 51.23 + 42.09 + 0.00003 + 5.77 + 0.5+0.18 + 0.09 = 99.77 kmol

Material balance across CO shift converter

Basis 100 kmole of H2S free gases

Now since 97% of CO is converted in converter

Therefore kmoles of CO2 in product = 5.78 + 42.09 x 0.97

= 5.78 + 40.8 = 46.6 kmoles


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similarly, kmoles of H2 in the

Product = 51.34 + 40.8 = 92.14

Kmole of CO = 42.09 40.4 = 1.26

Kmole of H2S = 0.3ppm

Kmole of CH4 = 0.501 kmole

Kmole of Ar = 0.09 kmole

Kmole of N2 = 0.18 kmole

Total no of moles in the product stream = ( 46.4 + 92.14 + 1,26 + 0.501 + 0.09 + 0.18 )

= 140.77 kmole

H2 in the product = 92.14/ 140.77 x 100 = 65.45%

CO2 in the product = 46.6/140.77 x 100 = 33.1%

H2S = 0.3 ppm

CH4 = 0.501/140.77 x 100 = 0.36%

Ar = 0.09/140.77 x 100 = 0.063%

N2 = 0.18/140.77 x 100 = 0.13%

Material balance across CO2 Absorber

Basis 100 mole of converted gas

Taking efficiency of the absorber to be 92%

CO2 to be removed = 0.92 x 33.1

= 30.45 kmole

CO2 in the product stream = 2.65 kmole

Total no of moles in the product stream


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= ( 2.65 + 65.45 + 0.89 + 0.13 + 0.06 + 0.36 )

= 69.54 kmole

CO2 = 2.65/69.54 X 100 = 3.8

H2 = 65.45/69.45 X 100 = 94.2

CO = 0.89/69.45 X 100 = 1.3

CH4 = 0.36/69.45 X 100 = 0.518

N2 = 0.13/69.45 X 100 = 0.19

Ar = 0.064/69.45 x 100 = 0.09

H2S = 0.3 ppm

Material balance across Adsorber

The main function of absorber is to adsorb H2S, CO & CO2. They are reduced to an

amount which is negligible.

Moles remained = 100 3.8 1.3 = 94.9 kmole

% of gas after adsorption

H2 = 94.2/95 X 100 = 99%

CH4 = 0.54

Ar = 0.095

N2 = 0.19/95 = 0.2

N2 is added in such a way that

H2 : N2 becomes 3 : 1

Suppose x mole of N2 is added

99 = ( 0.18 + x ) x 3


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x = 32.80

Total no of moles = 99 + 32.80 + 0.54 + 0.095

= 132.43 kmole

% of N2 in the product = 32.80/132.43 x 100

= 24.76 %

% of H2 in the product = 99/132.43 x 100

= 74.76%

% of Ar = 0.072%

% of CH4 = 0.54/132.43 X 100 = 0.4%

Material balance across NH3 Separator

Now in NH3 Separator 3 streams are there out of which 2 are going out and one is making

in the stream that moves in F13 & stream that gives main product is F14 & F 15.

F16 is recycle stream moving to compressor.

Suppose the moles of F 13 stream are as follows

N2 = a

H2 = 3a

NH3 = b

Ar = c

CH4 = d

Now solubility of H2 , N2 , Ar, NH4 in liquid NH3 is given as

Solubility of H2 = 0.0998 cm3 of H2 at NTP/gm of liq NH3

Solubility of N2 = 0.1195 cm3 of N2 at NTP/gm of liq NH3

Solubility of Ar = 0.154 cm3 of Ar at NTP/gm of liq NH3


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Solubility of CH4 = 0.304 cm3 of CH4 at NTP/gm of liq NH3

For F 14 stream

H2 = 0.0998/22414 x 17 x 3a x 22 x 298/273

= 5.45 x 10-3 a

N2 = 0.1195/22414 x 17 x a x 22 x 298/273

= 2.18 x 10-3 a

Ar = 0.154/22414 x 17 x c x 22 x 298/273

= 2.8 x 10-3 c

CH4 = 0.304/22414 x17 x d x 22 x 298/273

= 5.54 x 10-3 d

Now we want 1200 TPD production

Therefore NH3 production in F 14

= 1200 x 1000/(17x24)

= 2941.18 kmole/hr

Now for F15 stream

Equilibrium of NH3 in gas phase with liq. NH3 in separator = 5.9%

Assuming that total inert conc. In stream from the separator = 5%

H2 + N2 = 100 5.9 5 = 89.1%

H2 = x 89.1 = 66.83%

N2 = x 89.1 = 22.28 %

NH3 = 5.9%

Ar + CH4 = 5%

Since CH4 is 2.5 times the Ar.


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Ar = 1.43%

CH4 = 3.57%

Now applying material balance on NH3

separator

F 13 = F 14 + F 15

H2 balance : 3a = 0.6683 F15 + 5.45 x 10-3 a (1)

N2 balance : a = 0.2228 F 15 + 2.18 x 10-3 a (2)

NH3 balance : b = 2450.98 + 0.059 F 15 (3)

CH4 balance : d = 5.54 x 10-3d + 0.0357 F 15 (4)

Ar balance : c = 2.8 x 10-3c + 0.0143 F 15 (5)

From (5)

0.9945d = 0.0357 F15 (6)

d = 0.0359 F15

From eqn. (1)

2.99478 a = 0.6683 F15

a = 0.223 x F15 (7)

From eqn. (4)

0.9972 c = 0.0143 F 15

c = 0.01434 F 15 (8)

Now total moles in product stream F14

= 2941 + ( 2.18 + 5.45 ) x 10-3 x 0.223 x F15 + 2.8 x 10-3 x

0.0143 F 15 + 5.54 x 10-3 x 0.0359 F15

= 2941 + 1.9404 x 10-3 F15


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of NH3 separator = 98.5%

Therfore

2941 = 0.985

2941 + 1.9404 x 10-3 F15

2941= 2414.22 + 1.9113 x 10-3 F 15

F 15 = 2302kmol/hr

0.0019113

From (7)

a = 5133.7 kmol/hr

From (8)

c = 330.12 kmol/hr

From (6)

d = 826.45 kmol/hr

& from (3)

b = 4299.24 kmol/hr

Now for F-13 stream

H2 = 3 x 5133.7 = 15401.1 kmol/hr

N2 = 5133.7 kmol/hr

NH3 = 4299.24 kmol/hr

Ar = 330.12kmol/hr

CH4 = 826.45 kmol/hr

Total no of moles = 25990.61kmol/hr


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H = 15401.1 = 59.26%

25990.61

N2 = 19.75%

NH3 = 16.52%

Ar = 1.28%

CH4 = 3.18%

For F14 stream

NH3 = 2941 kmol/hr

H2 = 5.45 x 10-3 a

= 5.45 x 10-3 x 5133.7

= 27.98kmol/hr

N2 = 2.18 x 10-3 x a

= 2.18 x 10-3 x 5133.7

= 11.19 kmol/hr

CH4 = 5.54 x 10-3 x d

= 5.54 x 10-3 x 826.45

= 4.58 kmol/hr

Ar = 2.8 x 10-3c

= 2.8 x 10-3 x 330.12

= 0.924 kmol/hr

Total no of moles in product

= 2985.67 kmol/hr


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NH3 = 2941 = 98.5%

2985.67

H2 = 27.98 = 0.94%

2905.67

N2 = 11.19 = 0.38%

2985.67

CH4 = 4.58 = 0.15%

2985.67

Ar = 0.924 = 0.03%

2985.67

Material balance across NH3 converter

Suppose in F12 stream

N2 = x kmole

H2 = y kmole

NH3 = z kmole

Assuming 20% conversion

For N2 balance

X [1 0.2] = 5133.7

X = N2 = 6417.125 kmol/hr

H2 = 3 x 6417.125 = 19251.4 kmol/hr

NH3 formed = 2 x 0.2 x 6147.125

= 2566.85kmol/hr

NH3 in F12 = 4299.24 - 2144.48


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= 1732.4 kmol/hr

CH4 in F12 = 826.45 kmol/hr

Ar in F12 = 330.12kmol/hr

Total no of moles in F12 = 28557.5 kmol/hr

N2 = 6417.125 X 1000 = 22.47%

28557.5

H2 = 19251.4 X 1000 = 67.42%

28557.5

NH3 = 1732.4 X 1000 = 6.04%

28557.5

CH4 = 826.45 X 1000 = 2.89%

28557.5

Ar = 330.12 X 1000 = 1.17%

28577.5

Applying material balance across compressor

F11 + F 16 = F 12

F11 + F 16 = 28557.5

Applying H2 balance

0.7476 F11 + 0.6683 F 16 = 0.6742 x 28557.5

0.7476 [285557.5 F16] + 0.6683 F 16 = 0.6742 X 23855.25

Therefore F16 = 26432.79 kmol/hr


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Recycle = 26432.79 kmol/hr

F11 = 28557.5-26432.79

= 2124.79 kmol/hr

Suppose the kmole of feed gases coming from absorber is A

A x 0.99 = 2124.71 x 0.7476

A = 1588.43 kmol/hr

Amt of N2 Added = 2124.71 x 0.2476

= 526.1kmol/hr

Across the absorber applying overall material balance.

H balance,

F10 x 0.942 kmol/hr

F10 = 1422.8025 kmol/hr

Applying material balance across CO2 removal

F8 x 0.6545 = F10 x 0.942

F8 = 1422.8025 x 0.942 = 2047.79 kmol/hr

0.6545

Applying material balance across shift converter

F7 x 0.534 = 2286.17 x 0.6545

F7 = 2914.49 kmol/hr

Applying material balance across H2S removal

F5 x 0.5123 = 2914.49 x 0.5134

F5 = 2920.75 kmol/hr

H2S removed = 2920.75 x 0.0023 2914.49 x 0.3 x 10-4


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= 6.63 kmol/hr

For H2S gas

W = 0.094, TC = 373.K PC =89.63 bar

P =49 bar T = -220C = 2510 K

Tr = 257 = 0.67

373.5

Pr = 49 = 0.55

39.63

B0 = 0.083 - O.42 2 = - 0.59

( 0.67)1.6

kg/m2B = 0,139 - 0.172 = -0.78

( 0.67)4.2

Z = 1 - 0.59 x 0.55 = 0.52

0.67

PV = Z n RT

49 x V = 0.082 x 0.55 x 251 x 5.94

V = 1.297 m3

Solubility of H2S in Methanol

= 60 m3 of H2S at - 22

0C & 49 atm

m3 of methanol

Weight of methanol required


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= 1.534 = 2.56 x 10-2 x 1215

60 = 31.1 kg

kmol of methanol required = 31.1 = 0.97 kmol/hr

32

3% by weight of carbon is converted into coke, therefore only 97% of carbon is present

in raw gas.

Coke formed = 20556.22 x 0.8489 x 0.03

= 523.5 kmol/hr

Applying mass balance across carbon recovery unit;

F3*=F5*- mass of coke formed (* denotes mass per unit time)

F3*=2920.75x14.6-523.5x12=36361 kg/hr

Applying mass balance across gasifier

F1*=F3*-F2*=36361-462.42x32=21564 kg/hr

Hence, Fuel Oil requirement=21,564 kg/hr

Energy Balance

Data for Specific heats

For Gases Cp/R= A+ BT +CT2 +DT2


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A 103 B 106 C 10-5 D

CH4 1.702 9.081 -2.164

CH 3 OH 2.211 12.216 -3.450

NH3 (g) 3.578 3.020 -0.186

CO 4.376 1.257 - 0.031

CO 2 5.457 1.045 - 1.157

H 2 4.249 0.422 0.083

H2S 3.931 1.490 - 0.231

N2 3.280 0.593 0.040

O2 3.639 0.506 - 0.227

SO2 5.699 0.801 - 1.015

H2S (vap) 3.470 1.450 0.121

For liquids Cp/R =A +BT +CT

2

A 103 B 106 C

NH3 (l) 22.626 - 100.75 192.77

Methanol 41.653 - 210.32 427.20

Water 8.712 1.25 - 0.18

For Solids Cp/R = A +BT + CT2

A 103 B 10-5 C

Carbon (s) 1.771 0.771 - 8.67

S 4.114 - 1.728 - .78


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Since, the assumed conversion=97% , from graph Inlet Temperature=3500C

Reaction is, CO +H2O CO2 +H2

( H)298 = (HCO2 + HH2 ) (HCO +HH2O)

= (-393509)- (- 110525 285830)

= 2846 kj/kmole

Mole, of CO converted =583.6 kmole/hr

( H)298 = 2846x 583.6

= 16.6 x 105 kj/hr

( H)R = -2.27 x 107 kj/hr

For adiabatic condition, HR + HF8 + H298 = 0 Hp =2.17 x 107 kj/hr

460 (T- 298) + 1.17/2 x 10 3 (T2 2982 )

+0.3 x 105 (1/T 1/298) = 1954.07

Solving this equation by heat & trail we get


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T = 622.50 K

i.e.t = 349.50C

In cooler converted gas is cooled from 349.5 to 42.50C, i.e. from

Heat lost by converted gas = 1335.5 x 8.314 {4.6 (622.5-315.5)+ 1.17x 10-3/2 (622.52 315.52 6332 ) 0.3x105 (315.5 622.5)}

=22.9 x 107 kj/hrIf m be amount of water required for cooling

M 4.18 x 25 =22.9 x 10

6

Therefore M =2.19 x 107 kg/hr

Taking refernce temp. =298k for balance around the

Converter . Feed to converter is at 3500C

( H)F12 = 12585.8x 8.314 [4.12(298-923) + 0.8319/2 x


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10-3(298 2 623 2) + 0.0565 x 10 5 (1/623 1/298)]

= - 15.36 x107 kj/hr

NH3 formed =1167.66 kmole/hr

( Hrxn) = -46110 KH/kmole

( Hrxn)298 =700.6 x 46110 =- 3.23 x 107 kj/hr

If product is heated to a tamp . of T1 then

( H)F13 = 8.134 x 7145.79 [3.845 (T- 298) + 1.1299 x 10-3

(T2 2982) + 0.040 x 105 (1/T 1/298)]

For adiabatic condition,

(H)F12 + ( H)F13 + ( Hrxn)298 = 0

( H)F13 =12.45 x107 =3.845(T 298) + 1.1299 x 10-3 /2

+ (T2 2982 ) + 0.040 x 105 (1/T-1/298)

= 12.45 x 107 /7145.79 x 8.314

= 2095.60

Solving this equation by hit and trial wet dot, T= 773 k

Temperature of outlet gas = 500.50C

Product gases are cooled from 500.50C to 4100C while heating

The feed gases from 2580C to 350 0C


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Specification Sheet for Heat Exchanger design

Number required : 5

Function : To cool the gases coming out of shift

Converter

Type : 1- 4 shell and tube with floating head.

Shell side:

Number of shell passes : 1

Shell side fluid : Hot Gases

Shell internal diameter : 1067 mm

Baffle spacing : 534 mm

Entering temp. of gases : 132 0C

Leaving temp. of gases : 400C

Tube side:

Number of passes : 4

Number of tubes : 738

Tube length : 4.88 m

Tube side fluid : Cooling Water


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Tube pitch : 31.75 mm

Outside diameter of tube : 25.4 mm

Inside diameter of tube : 22.1 mm

Entering temperature of Water : 200 C

Leaving temperature of Water : 350 C

Reactor Design

CATALYST

A triply promoted ( K2O-CaO-Al2O3) iron oxide catalyst will be used. The iron oxide

(Fe2O3-FeO) is in the form of nonstoichiometric magnetite. It is made by fusing the

magnetite with the promoters. The catalyst is reduced in situ, and the removal of oxygenyields a highly porous structure of iron with promoters present as interphases between

iron crystals and as porous clusters along the pore walls. The pores range from 500A to

1000A and intraparticle diffusion is diffusion to occur by the bulk mechanism.

Alumina prevents sintering and corresponding loss of surface area and also bonds

the K2O, preventing its loss during use. The K2O and CaO neutralize the acid character of

Al2O3. Both K2O and CaO decrease the electron work function of iron and increase its

ability to chemisorb nitrogen by charge transfer to the nitrogen.

PROPERTIES

Particle size: Granules, in size range 6-10 mm

Bulk density: 1200kg/m3

Particle density: 305 lb/cu ft (4.9 g/cm3)


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Activity loss in service, 30-35% in 3yr depending on severity of operating conditions and

presence of poisons. Catalyst is slowly deactivated at operating temperatures above

9850F.

CATALYSTS POISIONS

In addition these poisons, hydrocarbons such as lubricating oils and olefins can crack

and plug pores. Sulphur, phosphorous and arsenic compounds should not exceed 15ppm.

Though temporary poisons, they cause crystals growth and attendant area decline.

Chlorine compounds form volatile alkali chlorides with promoters.

CHEMISTRY AND KINETICS

The overall stoichiometric equation is

1/2N2 + 3/2H2 NH3Extensive studies of ammonia synthesis on iron catalysts

suggest that the reaction occur through the following steps

N2 (g) 2N (ads)

H2 (g) 2H (ads)

N (ads) + H (ads) NH (ads)

NH (ads) + H (ads) NH2 (ads)

NH2 (ads) + H (ads) NH3 (ads)

Nh3 (ads) NH3 (g)

A rate equation based on nitrogen adsorption as the slow step is the most

commonly used although other forms have been development that also correlates the data

Since the effectiveness factor of ammonia catalyst is less that unity in a commercial

size pellets, it is desirafinely ground catalyst and employ and effectiveness factor

correction for other sizes. Ammonia synthesis is another example of an old reaction with

sufficient data existence to make this procedure feasible. The following equation in term

of activity has been recommended.Design of shift converter

F12=28557.5kmol/hr F13=25990.61 kmol/hrPlug Flow Reactor


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Step -1 Reaction occurring in PFR

N2 + 3H2k 2NH3

Or A + 3B k C

Reactor Kinetics

The only reaction that needs to be considered is the formation of ammonia from H2 and

N2,

N2 + 3H2 2NH3

A common rate equation for ammonia synthesis is the Temkin-Pyzhev equation given

where R is the rate of nitrogen consumption per unit volume of catalyst, f is a catalystactivity factor, and Pi is the partial pressure of component i in the gas. Values for K1 andK2:

K1 = ko1exp(-E1/RT)K2=ko2exp(-E2/RT)

Ko1 = 1.78954 x 104 kgmol/m3-hr-atm1.5E1 = 20,800 kcal/kgmol

Ko2 = 2.5714 x 1016 kgmol-atm0.5/m3-hrE2 = 47,400 kcal/kgmol

One limitation of this rate equation is that the rate will be infinite if the amount ofammonia in the gas is zero, and this may occur in a reactor being fed with fresh make-upgas. To avoid this numerical problem, the rate equation may be multiplied by a factorK3PNH3/(1+K3PNH3) ; this will avoid the approach to infinity at low PNH3 while having littleeffect at high ammonia pressures. This modification re-casts the rate equation into the


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Langmuir-Hinshelwood-Haougen-Watson (LHHW) form, which is one of the options forthe RPLUG reactor block in Aspen:

Assume that f = 1.0 and K3 = 2 atm-1.

Assume that the bulk density of the catalyst is 1200 kg/m3 and that the catalyst costs$12/kg.

Now, reactor temperature is 450 degree Celsius

Therefore, k1= 0.00873 kmol/m3-hr-atm1.5

K2=106.3 kmol-atm0.5/m3-hr

Also, pi=pio(1-Xi)/(1+EiXi) ; where i=A,B,C

EA=(2-4)/2=-0.5 & XB=3XA,XC=2XA

Using pAO=0.2247*200=44.94atm,pBO=0.6742*200=134.84atm,pCO=0.06*200=12atm

Finally, we get

pA=44.94(1-XA)/(1-0.5XA)

pB=134.84(1-3XA)/(1-1.5XA)

pC=12(1-2XA)/(1-XA)

1228.6(1-XA)(1-3XA)1.5 9.78(1-2XA)2(1-1.5XA)1.5(1-0.5XA)(1-1.5XA)

1.5 (1-XA)2(1-3XA)

1.5

(-rA) =24(1-2XA)

1 +(1-XA)


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Now plug flow reactor design equation

dxA = V

(-rA) FAO

Where FAO -- feed rate of A

From

V = dxA ( limits from 0 to 0.2)

FAO (-rA)

XA 0 0.04 0.08 0.12 0.16 0.20-rA 48.7 25 24.5 23.9 23.5 23.23

1/-rA 0.02 0.04 0.04 0.04 0.04 0.04


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1/(-rA) vs XA

0

0.005

0.01

0.015

0.02

0.025

0.03

0.035

0.04

0.045

0 0.05 0.1 0.15 0.2 0.25

XA

1/(-r

ASeries1

Area under curve = 30x0.005x0.05=7.6 x 10-3

V = FAO x (R+1)x 7.6 x 10-3

V = 28557.5 x 10.4 x 7.6 x 10-3

V = 16.362m3

Take factor of safety = 10%

New volume = 1.1 x 16.362

= 18 m3

Assume L/D = 4

Volume = /4 x D2 x L

18 = /4 x D2 x (4D)

D3 = 5.729

D = 1.212 m

Hence length (L) = 4 x 1.212

= 4.848 m

Now catalyst is divided in 2 beds in ratios 1:2.5

(as upper bed : lower bed)

Volume of catalyst in 1st bed = 18/3.5


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= 5.14 m3

Volume of catalyst in 2nd bed = (18 5.14)

= 12.86 m3

Bulk density of catalyst = 1200 kg/ m3

Weight of catalyst in 2nd bed = 12.86 x 1200

= 15.432 tonne

Weight of catalyst in 1st bed = 5.14 x 1200

= 6.168 tonne

Now giving allowance for space for gas movement upward and downward and insulation

be 0.5 m

Hence diameter of reactor becomes = 1.212 + 0.5

= 1.712 m

Pressure at which reactor works = 200 kg/m2

Let factor of safety = 20%

Design pressure = 1.2 x 200

= 240 kg/m2

We use carbon steel with internal lining of titanium

Allowable stress = 66,000 psi {hesse & ruston}

= 4494 kgf/cm2

Now Allowable stress = PD(1/k2 1) (from Hesse)

Where PD = design pressure

4494 = 240(1/k2 1)

k = 1.0337

Do = kDi

Do = 1.0337 x 1.212

= 1.252 m

Thickness of shell = (1.252 1.212)/2=20mm


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Specification sheet for Ammonia Convertor

Operating pressure : 200 kg/m2

Design pressure : 240 kg/m2

Temperature : 300 520oC

Allowable stress : 4494 kgf/cm2

Material of construction : Cr Mo steel

Diameter inside of converter : 1.212 m

Bulk density of catalyst : 1200 kg/m3

Total volume of catalyst : 18 m3

Diameter of catalyst bed : 1.252 m

Thickness of the shell : 20mm

Volume of 1st catalyst bed : 5.14 m3

Volume of 2nd catalyst bed : 12.86m3


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Weight of 1st catalyst bed : 6.168 tonne

Weight of 1st catalyst bed : 15.432 tonne

MECHANICAL DESIGN OF SHELL & TUBEHEAT EXCHANGER

Carbon steel(corrosion allowance=3mm)Allowable stress=11,000/14.7=749atm

[ref.PED by hesse & ruston,pg.60,table3.1]SHELL SIDE

No. of pass=1Fluids in shell are hydrogen ,nitrogen ,ammonia & inerts like argon& methaneDesign pressure=51 kgf/sq. cm=51atmShell diameter=1067mm

THICKNESS OF SHELL= [pD/2fJ-p]+c

Where;

P=design pressureD=shell ID


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F=allowable stressJ=joint efficiciencyC=corrosion allowance

Shell thickness= [51x1067/(2x749x.85-51)]+3=48mm

SHELL HEAD : assuming dished heads(because Pressure


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TUBE THICKNESS= pD/(2fJ-p)J=1(seamless tube)

Working pressure=1atm

Design pressure=1.1 times 1atm=1.1atmTube ID=22.1mmTube thickness=0.02mm( required)Tube thickness=(OD-ID)/2=1.65mm(actual)Hence, tube design is safe.

MECHANICAL DESIGN OF REACTOR

Material used is Carbon SteelThickness of reactor=pD/2fJ-p + C

Treactor=(1.1 x 200)x 1.212/(2x 4494 x 0.85-200)+ 3 mm= 39 mm

therefore, OD of Reactor=ID+2 x t=1.212+2x .039=1.29m

let head be the dished headsknucke radius=0.06x OD of reactor = 0.06x 1.29

= 77.4 mm

Crown radius, L= ID of reactor-6inches= 1.212-6x .0254= 1.06 m


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Head thickness= 0.833pL/fJ= 0.833x 1.1x 200x 1.06/0.85x 4494= 51 mm

h= depth of head=L-(L2-D2/4)1/2 =190 mm

e ffective length = 4.848+2x 0.190= 5.228 m

Surface area = 6.28Lh= 6.28x 1.06x 0.190

= 1.265 m2

Cost Analysis

Cost estimation is a specialized subject and profession in its own right. Chemical plants

are built to make a profit. An estimate of the investment required and the cost of

production needed before the profitability of project can be assessed.

Cost in plant can be classified as-

1. Fixed capital2. Working capital

(1) Fixed capital:-

Fixed capital is the cost of the plant ready to start up. It includes

1. Design

2. Equipment and interaction

3. Piping instrumentation and control system

4. Building and structure

5. Auxiliary facilities such as utilities

(2) Working capital

Working capital is additional investment needed over and above fixed capital to start up

the plant. It includes

1. Start up


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2. Catalyst cost

3. Raw material

4. Finished product inventories

These costs are required to be corrected to the specific

capacity and index price of the given year.

Correction of index price

Cost in yr. A = Cost in yr. B x Cost index of yr. A

Cost index of yr. BCorrection of capacity

C2 = C1 (S2/S1)n

C2 = Capital cost at capacity S2

C1 = Capital cost at capacity S1

n = power factor

(From Peter Timmer Haus, page 186)

Ammonia plant capacity = 100,000 ton/yr.

Fixed capital investment = 24 x 106 $ (Based on 1979)

Power factor for ammonia = 0.55

Our plant capacity = 1200TPD

Equation (2) C2 = C1 (s2 )n

s1

Taking that our plant runs for 335 days /yr. deducting the

time for normal shut down period.

Capacity = 1200 x 335 tons/yr. = 402000 tons/yr.

C2 = 24 x 106 x (402000) 0.55

100000

= 51.59 x 106$

Cost index = 1200

From chemical engineering journal


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CI =585.9(based on 1979)

Cost in yr. 2008 = cost in yr. 1979 x C.I. 2008

C.I.1979Thus,

Cost in yr. 2008 = (51.59 x 106) x 1200

585.9

=105.663 x 106 $

From Peter Timer Haus -

Working capital = 15% [fixed + working capital]

W.C. = F.C. x 15

85

W.C. = 105.663 x 15 x 106

85

= 18.646 x 106 $

Therefore total capital investment

= 105.663 x 106 + 18.646 x 106

= 124.31 x 106 $

Taking currency exchange rate as 1 $ = 44 Rs.

Total capital investment = 546.964 crore

Total capital investment = 124.31 x 106 x44

=5469.64 x 106

=Rs.546.64 crore

DIRECT COST

Component Given range Assumed range(%FCI) Cost (Rs.crores)Purchased equip. 15-40 25 116.25

Installation 06-14 09 41.85


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Instrumentation 02-18 07 32.55

Piping 03-20 08 37.20

Electrical 02-10 05 23.25

B

uilding 03-18 05 23.25

Yarest 02-05 02 9.3

Service facilities 08-20 15 69.75

Land 01-02 01 4.65

Total direct cost = Rs. 358.05 crore

Indirect costEngineer and supervision 04-25 10 46.5

Construction expenses 04-16 12 55.8

Contractor fee 02-06 02 9.3

Contingency 05-15 08 37.20

Total Indirect cost = Rs. 148.8 crore

Direct + Indirect cost = Rs 506.85 crore

COST OF REACTOR

Volume of cylinder=3.14*l*(OD2-ID2)/4

OD=1.290m

ID= 1.212m


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Length,l=4.848m

Therefore, vol of cyl. Material = 0.743m3

Density of Carbon Steel=7820 kg/m3 , Price of carbon steel= Rs 45/kg

Therefore, Cost of Reactor material = 0.743 x 7820 x 45=2.62 lakh

CATALYST

Volume of catalyst = 18 m3

Price of catalyst = $12/kg

Density of catalyst = 1200 kg/m3

Therefore, Cost of catalyst = 18 x 1200 x 12 x 44 = 1.141 Crores

Now, taking fabrication cost = 70% of total cost of equipment= 2.26 x 0.7/0.3 = 5.273 lakh

Hence, total cost of reactor (including catalyst) = 1.141+.05273+.0262= 1.22 Crores

COST OF HEAT EXCHANGER

Volume of shell material = 3.14 x length x (OD2-ID2)/4

OD = 1.115mID = 1.067m

Length=16 ft=16 x 12 x .0254m


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Volume of shell material = 0.4012 m3

Volume of one tube = 3.14 x length x (OD2-ID2)/4

OD=.0254mID = .0221m

Length= 16 ftVolume of single tube = 6 x 10-4 m3

Total vol. of tubes= 6 x 10-4 x 730= 0.4383 m3

Density of material(carbon steel) = 7820 kg/m3

Cost of material = Rs 45/kg

Therefore, Cost of heat exchanger material = (0.4012+0.4383) x 7820 x 45

= 2.954 lakh

total cost of heat exchanger(including fabrication)= 2.954 x 1/0.3= 9.847 lakh

Safety And Hazards

FIRE AND EXPLOSION HAZARDS

Due to high vapour pressure of ammonia, container for ammonia should not be exposed

to the Sun and other sources of heat and should be properly insulated.


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Ammonia vapours are flammable in certain percentages in air, but since these high

concentrations are seldom encountered, the relative fire and explosion hazard is small.

Experience has shown that ammonia is hard to ignite, so it is generally treated as a non

inflammable compressed gas. However, certain safe practices should be followed.

Electrical apparatus and fixtures should be vapour proof and plugged in at

locations free from ammonia gas.

Equipment, tanks and lines that contain ammonia should not be welded without

first being thoroughly washed or steamed. Aqueous fire extinguishers are

satisfactory for ammonia fires.

Since ammonia is readily soluble in water, hose streams can be used effectively to

remove ammonia from the air in contaminated areas.

CLOTHING AND PROTECTIVE EQUIPMENT

Persons subject to exposure should be provided with gas masks of ammonia the type

approved by the Bureau of Mines, rubber hats, suits, gloves and boots. Garment worth

beneath rubber clothing should be cotton. Eye gogg1es or full face masks should be

provided. Easily accessible safety showers and bubble fountains for washing the eye

is .recommended.

FIRE PREVENTION STEPS IN AMMONIA

1. Fire fighting extinguishers

2. Water and foam tenders

3. Breathing apparatus, gas mask, ear plugs, aluminized suits for rescue purpose,

pressure bags for lifting heavy equipments

4. High expansion foam generator5. On site emergency plant

Apart from the above, fire alarms, smoke detectors and fire hydrants are also provided in

plant.

CORROSION


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Since ammonia is corrosive to alloys of copper and zinc, these materials must not be used

in ammonia service- iron or steel should usually be the only metal used for ammonia

storage tank, piping and fittings.

Cleaning and Repairs of Equipments

Wherever possible, ammonia tanks and other equipment should be cleaned and repaired

from the outside. When it is necessary to enter an ammonia vessel all lines leading to the

vessel should be disconnected and provided with blind flanges. Danger signs should be

placed at appropriate points to wary other work men and all electric apparatus should be

turned off. One person keeps close watch outside the tanks and has dosed at hand the

proper emergency equipment, respirator, and fire outing wishes. The tank should 'be

filled with water and drained before being entered and should purge with air and tested

for explosive atmospheres.

FIRST AID

A person who has been overcome of burned by ammonia should be placed under the case

of physician at once. The patient should be removed from contact with ammonia as soon

as possible and kept warm and quiet.

Physiological response to various concentration of ammonia

ppm

Least detectable odor - 53

Least amount causing immediateirritations to eyes- 698

Least amount causing immediateirritations to the throat- 408

Least amount causing coughing- 1720Max. Conc. allowable for


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Prolonged exposure- 100

HAZARDS OF AMMONIA

1. Toxic hazards

In general ammonia gas under high concentration causes irritation to eyes.

2. Explosive Hazards

Ammonia is explosive with air in the range of 16 to 25% by volume. Its auto ignition

temperature is 651C since such temperature are not encountered in practice; the

chance of fire and explosive hazards due to ammonia is relatively remote.

The presence of oil in mixture of ammonia with other combustible materials will increase

the /fire hazard. The explosive range of ammonia is;

1. Temperature and pressure higher than atm conditions.

2. Presence of chlorine higher than ammonia causes chlorine to react with ammonia

and form a violently explosive compound.

SAFETY IN AMMONIA PLANTS

Wear ear plugs

wear approved respirator & protective clothing

Wear safety helmet, goggles, hand glows, safety shoes.

Keep away from heat, flames, spark & oxidizing material

Keep area ventilated & report the leakage

Keep people away from hazardous area.

Pollution Control

1) Ash Slurry treatment Scheme

The ash slurry from steam generation plant is pumped by the existing sat of pumps to the

proposed ash ponds. The ash is settled in the ponds by gravity settling. The clear

overflow from the ponds is discharged into sump from where the clear water is pumped

to steam generation plant for de-ashing purpose or can be drained to the river sump

directly.


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2) Oily Effluent Treatment Scheme

Two nos. of oil separators are provided to receive oil effluent. Each separator has about

75 m3 capacity and both put together can receive 150m3 of effluent per day. Oil

separators are to be operated alternately. They are building with stilling, inlet chamber,

and deoiler pipe, baffle for retaining the floating oil and overflow weir. Oil separates andfloats to the surface by gravity separation.

3) Reuse of Process condensate as boiler feed water make up

In ammonia plant, the main source of liquid effluents is process condensate. It is well

known that steam is used in excess of stoichioimetric requirements for reforming and

shift conversion in ammonia plant. This excess steam is subsequently removed as process

condensate which contains impurities like carbon dioxide, ammonia, methanol etc. This

condensate cannot be directly used anywhere due to the presence of impurities.

4) Non-Chromate based cooling water treatment

Cooling water treatment is of vital importance for all the chemical process industry.

When the original plant was commissioned in 1974, chromate based cooling water

treatment was adopted. About 15-20 ppm chromate was maintained in the circulating

cooling water. The chromate is toxic and harmful to the environment. This hexavalent

chromate was converted into trivalent chromate and separated out as sludge.

To overcome chromate sludge storage problem and to adopt environmental friendly

system, non-chromate based cooling water treatment was adopted in Apirl,1998.

Non-chromate treatment programme are mainly based on o-phosphate and zinc with

different combinations. o- phosphate concentration of 6 ppm minimum and

Zincconcentration of 1 ppm maximum is maintained as corrosion inhibitor. Polymer

based bio-dispersant and boicides are also added.

Three different biocides are shock dosed alternatively in addition to chlorination for

micro biological growth control as per requirement. Dispersants are added to keep the

salts of circulating water in dispersion condition. All the chemicals used for non chromate

treatment are biodegradable chemicals.

5) Condensate Stripper Overhead vapour condensation


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Overhead vapours about 6000 kg/hr from condensate stripper were vented to atmosphere.

the overhead vapours contain ammonia, methanol, amines and CO2. Ammonia present in

the overhead vapours is about 5000 ppm. The limit specified by Gujarat Pollution Control

Board (GPCB) for ammonia is 100 ppm. To avoid the venting of ammonia to

atmosphere,

the overhead vapours from the top of the condensate stripper are diverted to the vent

condenser (172-C). All the vapours are condensed and non-condensables mainly CO2 is

vented to atmosphere. The condensate from the vent condenser can either be sent to

Ammonia recovery and stripping system in ammonia plant or to hydrolyser stripper

system in urea plant by condensate transfer pump. The system is shown as annexure - VI.

The system has reduced the 720 kg/d ammonia to atmosphere and pollution norms has

been also met.

6) Biological Treatment of Effluent

After giving treatment, the effluent is further bio-degraded in the effluent ponds having a

holding capacity of 1, 00,000 m3 before final discharge to the river Damodar.

Factory effluent mixed with sanitary waste methanol water and is treated with the help of

bacteria. The bacterial actions consist of hydrolysis, nitrification and denitrification.

HYDROLYSIS

In the process of metabolism through heterotrophic bacteria.complex carbohydrates,

proteins and insoluble fats present in sanitary waste are hydrolyzed into soluble sugar,

amino acids and fatty acids.

Hydrolysis of urea takes place and converts organic nitrogen to ammonium carbonate.

(NH2)2C03 +2H2O (NH4)2 C03

Biochemical oxygen demand (BOD) is reduced by hydrolysis under aerobic condition in

presence of baceotrophic bacteria.

NITRIFICATION

Ammonical nitrogen is converted to nitrites and then to nitrate form in a presence of

autotrophic bacteria under aerobic condition.

N H4++ 3/2 O2 NO2

- + H2+ + H20


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N02- + 1/2 O2 - N03

-

NH4+ +2O2 - NO3H + H2O

DENTRIFICATION:

Heterotrophic bacteria in anaerobic condition reduce nitrite (N02) and Nitrate (N03) to

gaseous N2. An organic carbon, as methanol, acetic acid, acetone, ethanol or sugar is

reduced to act as a hydrogen donor to supply carbon for biological synthesis. The

required BOD is supplied by waste methanol (6000 ppm) which serves as an energy

source for survival and growth of bacteria.

First methanol reduces:

3O2 + 2 CH3OH - 2 CO2 + H2O

Then bacterial reduction of N02 & N03 takes place.

6 NO3- + 5 CH3OH - 3 N2 + 5CO2 + 7 H2O + 6OH

-

2 NO2- + CH3OH N2 + CO + H 2O+ 2 OH

-

Aeration is helpful for the removal of gaseous product and increase dissolved oxygen in

the effluent. In the step of hydrolysis urea (500 ppm) degrades to 50 ppm level.

Instrumentation

The even operation of a process is dependant upon the control of process variables. These

are defined as the conditions in the process material and equipments which are

subjected to change. There may be several materials and several pertinent operating

factors which may change in the simplest of process; the maintenance of control over an

entire process is an important aspect of process design. In Ammonia manufacture

temperature and pressure are two process variables on which operation of whole plant


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depends upon. Instrumentation of each equipment is done separately.

1. REACTOR

Feed to the converter is at 750C which is made to pass through the internal heat

exchangers so that it may gain heat from outgoing gases to attain reaction initiation temp.

This temperature of gas at the inlet of catalyst bed is taken into account by TIC (temp.

indicator controller) put in line with the start up heater.

In case temperature goes below the reaction initiation temperature.

Then automatically start up heater works and temperature of product gas is increased

which in turn heats the incoming gases to the required reaction initiation temperature.

But if the temperature goes high then a portion of by pass feed controls the temperature.

Temperature control is necessary because in first case, reaction would not start till therequired temperature is attained. And in the later case, temperature goes too high then

catalyst may get spoiled. Pressure inside the reactor is 300 atm. and any change in

pressure will effect the conversion therefore pressure control is also necessary so we will

have to employ pressure gauges and pressure controller valves.

Temperature indicator devices are put on both inlet and outlet streams and if temperature

of outgoing gases goes below 162C the water flow in the heat exchanger (waste heat

boiler) is reduced with the help of a hand controller valve.

2. AMMONIA COOLED CONDENSER

A level indicator controller is installed in the ammonia refrigeration loop so as tomaintain a constant level which is necessary to get required refrigeration to bring downthe temperature of gas to -S C where required condensation of ammonia is acting.

3. AMMONIA SEPARATOR

Here also a level indicator controller is required to maintain a particular level of liquid

ammonia in it so no gas but only liquid product is obtained.

4. AMMONIA STORAGE

Ammonia is stored at a pressure of 40 psig and temperature 25.8 0 F. However, efficient

may be the insulation. Some of heat enters the tank from ambient and vapors of ammonia


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are formed. Pressure inside the tank increases so there must be some arrangement to

reduce this pressure. A safety valve may be provided at the top of 'Horton Sphere'.

Whenever pressure inside the vessel rises some of the vapors are vented. These vented

gases either released at. 6 ft. high stack or can be compressed and condensed in a close

loop to reduce atmospheric pollution. In a closed loop system vent are stored in a small

vessel from where those gases are compressed, condensed and then fed to the storage

tank.

TEMPERATURE MEASUREMENT

Temperature measurement is done by using thermocouples. Since ammonia corrodes

metals like Cu and Zn. The thermocouple selected is chromel (Ni 90%, Cr 10%)

Alumd (Ni 94%, Mn 3%, Al 2%). It can be used in temperature ranging from 600 to 2100

of vertically in the converter. Slight variations of temperature inside the converter are

transmitted through the transmitter to the error is passed on to a diaphragm motor valve

which controls the flow of gases to the converter.

INSTRUMENT AND FITTINGS

Gasket material most suitable for anhydrous ammonia service is hard finished. Such as

flange facings where the gasket material is retained in a groove. And on gasket material is

retained in a groove. And on gasket bodies, lead may be used. Aluminum has been

applied satisfactorily as a gasket material for flat faced and grooved flanged joints but. is

not. Generally recommended Rubber rings cannot. be used. All ammonia piping should

be extra heavy (sch. 80). Steel piping may be used where joints are welded. Galvanized

pipe should never be used. Welded pipes should be used wherever possible. Welded pipe

flanges may be used on all pipe joints of 1.25" or larger. Under no circumstances brazed

joint be used, as they will deteriorate rapidly.

Industrial Applications of Ammmonia


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Explosive Industry

The explosive industry is one of the largest consumers of ammonia.

It is used principally in the form of Nitric acid for the manufacture of compounds such as

TNT, Nitroglycerine, nitrocellulose, nitro starch, Pentaerythrital t1tranitrate, tetraethyl

and ammonium nitrate.

Textile Industry

In textile industries, ammonia is used in the production ff synthetic fibres such as

cupraammonium rayon and nylon. In the production of rayon, the ammonia is used in the

preparation of ammonical copper hydroxide solution (Schweiter's reagent) for dissolving

the cotton linters.

In the rayon process ammonia is used in the production of hexamethylene diamine, which

is considered with-adipic acid to form the monomer, which in turn is polymerized and

speen into polyamide fibres.

Fertilizer Industry

The fertilizer industry is a heavy consumer of ammonia and its compounds. The main

nitrogenous fertilizers are

- Anhydrous ammonia

- Ammonia liquor

- Ammonium Nitrate

- Urea

- Ammonium Sulphate

- Calcium Cynamide

- Sodium Ammonium Nitrate.

Ammoniation of phosphatic fertilizer produces fertilizer mixture with improved physical

properties.

Pharmaceutical Industry


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Ammonia is an important ingredient in the manufacture of sulfa drugs such as

sulfanilamide sulphath1azole and sulfapyridine. It. is also used in the manufacture of

vitamins and antimalarials.

Petroleum Industry

In the refining of crude oils, ammonia is used as neutralizing agent to prevent corrosion

in acid condensers, heat exchangers etc. of the distillation equipment.

Ammonia is also used to neutralize final traces of acid in acid-treated lubricating oil

stocks and pressure distillates,

Rubber Industry

Ammonia is used to raw latex to prevent coagulation during transportation from the

rubber plantation to Ammonia is also used in the nuconization the factory process for the

manufacture of fabrics, boots, shoes etc.

Refrigeration

Ammonia is most commonly used refrigerant, especially for large industrial installations.

It can be used in both compression and absorption system for

1) Production of ice

2) Refrigeration cold-storage plants

3) Quick freezing units

4) Food-lockers

5) Air conditioning large industrial plants

Certain characteristics of ammonia such as high latent heat of vaporization low vapour

density. Chemical stability and low corrosion to compressor parts, make operating costs

per ton lower for ammonia than for any other refrigerants used in industrial systems.

Inorganic Chemicals Industry

A large number of ammonium salts having numerous industrial applications are produced

by direct neutralization of their respective acid with ammonia.

Important industrial compounds produced are ammonium nitrate, ammonium sulphate

mono and diammonium phosphate and ammonium acetate.


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Acid Manufacture

In the lead chamber process for sulfuric acid manufacture, Ammonia is oxidized to give

Nitrogen oxide to provide required oxygen required for conversion of SO2 to sulphuric

acid.

Pulp and Paper Industry A recent development is the substitution of ammonia for

calcium in the bisulphate process for pulping wood. The ammonia base is said to improve

the yield and Quality of pulp.

Leather Industry

Ammonia is used as slime and mold preventive in tanning liquors.and as a. curing agent

in making leather.


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Bibliography

1. Perry, R.H., "Perry's Chemical Engineers Handbook", 6th

Ed., McGraw Hill Company, 1984.

2. McCabe Warren L., Smith Julian C., Harriot Peter, "UnitOperations of Chemical Engineering", 4th Ed., McGraw Hill,

Inc., 1981.

3. Levenspiel Octave, "Chemical Reaction Engineering 2nd

Ed., Wiley Eastern Ltd., 1991.

4. Peters M.S., Timmerhaus K.D., "Plant Design and

Economics for Chemical Engineers", 4th Ed., McGraw Hill,

1991.

5. Treybal R.E., "Mass Transfer Operations", 3rd Ed., McGraw

Hill Company, 1981.

6. Hesse, Herman C. and Rushton, J. Henry "Process

Equipment Design" Affiliated East-West Press Pvt. Ltd.,

New Delhi.

7. Richardson & coulson Vol 6

8. Operating Manual, "Ammonia Plant" of National Fertilizers ltd. Bathinda

9. Elementary Principles of Chemical Processes by Felder

33170313 ammonia production project

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