Design of Ammonia Production Plant with Membrane Separation Technology

1. Muhammad Asim Shahzad 2008-NUST-BE-Chem-68 2. Muhammad Mujahid 2009-NUST-BE-Chem-14 3. Muhammad Numan 2009-NUST-BE-Chem-16 4. Mubeen Mukhtar 2009-NUST-BE-Chem-67

This work is submitted as a FYP report in partial fulfillment of the requirement for the degree of

(BE in Chemical Engineering)

Project Supervisor: Dr. Arshad Hussain

School of Chemical and Materials Engineering (SCME)

National University of Sciences & Technology (NUST), H-12

Islamabad, Pakistan

June 2013

In the Name of Allah, the Most Gracious, the Most Merciful

CERTIFICATE This is to certify that work in this dissertation/report has been completed by Muhammad Mujahid, Mubeen Mukhtar, Muhammad Numan and Muhammad Asim Shahzad under my supervision at School of Chemical and Materials Engineering, National University of Sciences and Technology, H-12, Islamabad, Pakistan.

Supervisor:

Dr. Arshad Hussain

HoD Chemical Engineering

School of Chemical and Materials Engineering (SCME)

National University of Sciences and Technology (NUST)

Submitted through:

HoD Principal/Dean

Chemical Engineering SCME NUST

Dedicated to our Beloved Parents and Respected Teachers !

ACKNOWLEDGEMENTS

All praises to ALMIGHTY ALLAH, who provided us with the strength to accomplish the final year

Before anybody else we thank our Parents who have always been a source of moral support and

driving force behind whatever we do. We are indebted to our project advisor Dr. Arshad Hussain for

his worthy discussions, encouragement, inspiring guidance, remarkable suggestions, keen interest,

constructive criticism & friendly discussions which enabled us to complete this report. He spared a lot

of his precious time in advising & helping us in writing this report. Without his painstaking tuition,

kind patronization, sincere coaching and continuous consultation, we would not have been able to

complete this arduous task successfully. We are also grateful to Mr. Shahid Hussain Ansari, Mr. Umair Sikandar, Mr. Nauman Ahmed and all

the Teachers for their profound gratitude and superb guidance in connection with the project.

Authors

project. All respects are for His HOLY PROPHET (Alayehis Salam), whose teachings are

true source of knowledge & guidance for whole mankind.

ABSTRACT

Fertilizer Industry which constitutes the backbone of the agriculture, the largest GDP

contributing sector in Pakistan is of strategic importance. There is an ever increasing demand of

fertilizers especially urea. Exceeding expense on the import bills of this crucial commodity

requires expensive foreign exchange. Therefore Pakistan has a huge growth potential for urea

industry. Ammonia production is the first step before manufacturing urea. So ammonia

technology always remains the field of key interest for chemical engineering students.

This project consists of comprehensive design of ammonia plant in generalized manner. It

includes literature review and process selection, creation of PFDs with material and energy

balances, designing of process equipment, cost estimation and analysis of ammonia plant.

This project also includes the design of membrane unit capable of recovering maximum quantity

of H2 from purge gas of ammonia synthesis loop in spite of conventional cryogenic separation

technique being used in major ammonia production plants in Pakistan. Prior to H2 recovery,

NH3 is being recovered from purge gas by water scrubbing.

Table of Contents Certificate.......................................................................................................................... ii

Acknowledgements........................................................................................................... iv

Abstract ............................................................................................................................ v

List of tables...vi

List of Figures ...vi

Table of Contents................................................................................................................vii

1 INTRODUCTION ..1

2 LITERATURE REVIEW 2

2.1 Process Selection and Description ...2

2.2 LTS and HTS design.6

2.3 Carbon Dioxide Removal.15

2.4 Ammonia Convertor.17

2.5 Hydrogen Recovery Unit..23

3 MEMBRANE SEPARATION TECHNOLOGY24

3.1 Membranes for Hydrogen Separation.....24

3.2 Why Polyimide? ..25

3.3 Hollow Fibers Preparation...25

3.4 Hollow Fiber Module...27

3.5 Asymmetric Membrane Transport Mechanism.,.28

3.6 Working of Hollow Fiber Module..30

4 MATERIALS BALANCE ....32

4.1 NG and Recycle Stream Mixing Point32

4.2 Desulfurizer.33

4.3 Feed Gas and Steam Mixing Point..34

4.4 Primary Reformer...35

4.5 Secondary Reformer...36

4.6 High Temperature Shift Converter - HTSC...37

4.7 Low Temperature Shift Converter LTSC...38

4.8 2nd KO Drum.39

4.9 CO2 Absorption.....40

4.10 Methanator...41

4.11 3rd KO Drum41

4.12 1st Stage of Compressor...42

4.13 2nd and 3rd Stage of Compressor...43

4.14 Make up Gas Mixing Point..43

4.15 Separator..44

4.16 Ammonia Synthesis Convertor44

4.17 Purge Gas Exit Point45

4.18 Ammonia Recovery Unit - ARU.46

4.19 Recovered NH3 Mixing Point..46

4.20 Hydrogen Recovery Unit - HRU.47

5 ENERGY BALANCE..48

6 EQUIPMENT DESIGNING51

6.1 High Temperature Shift Converter51

6.2 Low Temperature Shift Converter56

6.3 ABSORPTION COLUMN61

6.4 DESIGN STEPS OF AMMONIA REACTOR.....67

6.5 AMMONIA RECOVERY UNIT..74

6.6 Hydrogen Recovery Unit...78

7 COST ESTIMATION82

7.1 Cost Index Calculation82

7.2 Fixed Costs.82

7.3 Cost of Equipment..84

7.4 Operating Cost: ..86

7.5 Variable Costs: ..86

7.6 Fixed Cost..87

8 INDUSTRIAL APPLICATIONS.89

8.1 Fertilizer.89

8.2 Precursor to Nitrogenous Compounds..89

8.3 Cleaner..89

8.4 Refrigeration R717.89

8.5 Remediation of Gaseous Emissions..89

8.6 Fuel...90

8.7 Stimulant in Sports90

8.8 Textile: .90

8.9 Lifting Gas: ..90

8.10 Wood Working: .91

9 CONCLUSIONS.92

9.1 Recoveries Importance.92

9.2 Membrane Separation Technology..92

10 References.93

List of Tables Table 2.1 Comparison of Carbon Dioxide absorption Processes .15

Table 3.1 Engineering Strategies for Hydrogen Separation membrane24

Table 3.2 Comparison of Membranes for Hydrogen Separation...25

Table 3.3 Influence of Operating Variables...31

Table 9.1 Comparison of Polymer membrane and Cryogenic separation..85

List of Figures Figure 2.1 Process steps of Ammonia Production..2

Figure 2.2 Equilibrium CO concentration in LTS......12

Figure 2.3 Sulfur and Temperature Profile For used LYS Catalyst...13

Figure 2.4 S -300 Ammonia Converter ..............22

Figure 3.1 Experimental set-up for hollow fiber spinning..27

Figure 6.1 Fixed bed Reacor51

Figure 6.2 Reaction Mechanism...52

1

mmonia is a compound of nitrogen and hydrogen with the formula NH3. It is a colorless gas with a characteristic pungent odor. Ammonia contributes significantly to the nutritional needs of terrestrial organisms by serving as a precursor to food and fertilizers.

Ammonia, either directly or indirectly, is also a building block for the synthesis of many pharmaceuticals. Although in wide use, ammonia is both caustic and hazardous. In 2006, worldwide production was estimated at 146.5 million tones. It is used in commercial cleaning products.

Ammonia, as used commercially, is often called anhydrous ammonia. This term emphasizes the absence of water in the material. Because NH3 boils at -33.34 C (-28.012 F), the liquid must be stored under high pressure or at low temperature. Its heat of vaporization is, however, sufficiently high so that NH3 can be readily handled in ordinary beakers, in a fume hood (i.e., if it is already a liquid it will not boil readily). "Household ammonia" or "ammonium hydroxide" is a solution of NH3 in water. The strength of such solutions is measured in units of Baume (density), with 26 degrees Baume (about 30% w/w ammonia at 15.5 C) being the typical high concentration commercial product. Household ammonia ranges in concentration from 5 to 10 weight percent ammonia. Ammonia is an intermediate product in the manufacture of nitrogenous fertilizers. It is also used for direct application to the soil and in aqua condition with solutions of other nitrogenous fertilizers like ammonium nitrate and/or urea. Besides these, ammonia finds application in the production of nitric acid, soda ash, cleaning agents, leather tanning, petroleum refining, pulp & paper industry, textiles, refrigeration, rubber & synthetic resin industries, explosives and food & beverage industries.

A

1 INTRODUCTION

Natural gas. naphtha Fuel oil. coal

~,~

.l!;~g=>~~~

s""~..,~ AirC7>~

c:sQ)::-!t.c:(/)

~,gXl~

g''=s!Ea;f

HT shiftconversion

Shift conversion(CoMo cat.]

Ammonia syntnesis

2

2.1 Process Selection and Description

In the past years large amount of work is done on designing new and energy conservative process for the production of Ammonia. The main drive behind designing of year process for ammonia is to reduce the Energy requirement. Below are mentioned few process developed by different Companies throughout, for steam reforming. Before going through different process let us have a look at the components of an Ammonia plant.

Fig 2.1 Process steps of Ammonia Production

2 LITERATURE REVIEW

3

This process combines the use of excess air (up to 25%) in the secondary reformer with a very active synthesis catalyst. In LCA process the heat generated in the secondary reformer is utilized in primary reformer by direct heat exchange in a tubular Gas Heated Reformer (GHR). The CO shift is performed in a single stage shift reactor at 2500 oC using special copper basic catalyst.

CO2, inserts and excess nitrogen are removed from raw synthesis gas by pressure swing absorption. Ammonia synthesis takes place at low pressure of below 100 kg/cm2g using ICIs highly active cobalt promoted catalyst. Net energy consumption of around 7.2 Gcal/ MT of ammonia has been demonstrated for 450 MT per day plant.

2.1.2 Kellogg Brown and Roots Advanced Ammonia Process:

KAAP uses a high pressure heat exchange based steam reforming process integrated with a low pressure advanced ammonia synthesis process. Raw synthesis gas is produced by steam reforming of hydrocarbons in a heat exchange based system under pressure, based on Kellogg Brown and Root Reforming Exchange System (KRES). KRES also reduces energy consumption and capital cost besides reduced emission and enhanced reliability. After sulfur removal, the auto-thermal reformer and reforming exchanger which operate in parallel, convert 100% feed into raw synthesis gas in the presence of steam using nickel catalyst. In auto-thermal reformer, enriched air supplies nitrogen.

The heat of combustion of partially reformed gas supplies energy to remaining hydrocarbon feed. The exit gas of auto thermal reformer is fed on the shell side of the KRES and thus heat of combustion is supplied to the reforming reaction taking place inside the tubes. KRES exit gas is cooled in waste heat boiler where high pressure steam is generated. After cooling, the gas is sent to the CO shift converters (high temperature and low temperature). CO2 is removed from the process gas using hot potassium carbonate solution, methyl diethanol amine (MDEA) etc. After CO2 removal, methanation and gas drying processes are carried out. The gas is then compressed and mixed with recycle stream of synthesis loop where gas mixture is sent to converter designed by KAAP.

2.1.3 Krupp Uhde GmbH Ammonia Process:

The Krupp Uhde GmbH process uses conventional steam reforming for synthesis gas generation (front end) and a medium - pressure ammonia synthesis loop. The primary reforming is carried out at pressure 40 bars and temperature range of 800- 850 oC. The steam reformer is top fired and tubes are made of centrifugal high alloy steel which enhances reliability. Process air is added in secondary reformer through nozzles installed in the wall of vessel. This provides proper mixing of the air and reformer gas. Subsequent high pressure steam generation & superheating, guarantee maximum heat usage to achieve energy efficiency. Carbon monoxide is converted to carbon dioxide in HT and LT shift converters.

2.1.1 ICI Leading Concept Ammonia (LCA) Process:

4

The MDEA or Benfield system is used for carbon dioxide removal. The ammonia synthesis loop uses two ammonia converters with three catalyst beds with waste heat boiler located downstream of each reactor. The converters have small grain iron catalyst. The radial flow concept minimizes pressure drop and allows maximum ammonia conversion.

2.1.4 The Linde Ammonia Concept (LAC) Ammonia Process:

The LAC process consists essentially of a modern hydrogen plant, a standard nitrogen unit and a high efficiency ammonia synthesis loop. Secondary reformer, one shift conversion and methanation steps have been eliminated in this process. The primary reformer is top fired and operates at an exit temperature of about 850 oC. The CO shift conversion is carried out at 250 oC in a single stage in the tube cooled isothermal shift converter and gas is sent to pressure swing absorption (PSA) unit wherein the process gas is purified to 99.99 mole % hydrogen. A low temperature air separation in cold box is used to produce pure nitrogen.

BASFs MDEA process has been eliminated in this process used for CO 2 removal. The ammonia synthesis loop is based on Casale axial - radial three bed converter with internal heat exchanger giving a high conversion. The energy consumption (feed + fuel) is 7 Gcal/ MT of ammonia.

2.1.5 Haldor Topsoe A/S Process:

The companys low energy ammonia process uses the conventional sequence of process steps which are optimized by the introduction of improved catalysts, new equipment design and extensive process optimization studies. A pre-reformer containing nickel catalyst has been also provided upstream of primary reformer for converting all the higher hydrocarbons, so that only methane, carbon monoxide, carbon dioxide, hydrogen and steam are present in the product gas. Firing in primary reformer is reduced by 15% due to pre- reformer. Highly active shift catalyst ensures the lowest carbon monoxide at the exit of converters and thus highest utilization of feedstock. New temperature resistant iron free catalyst makes it possible to operate at low steam to carbon ratio at high temperature shift converter. The company has also developed Heat Exchange Reforming Process (HERA). Low energy CO 2 removal processes, such as selexol, MDEA or low heat potassium carbonate, are used. Topsoe has developed new converters especially for high conversion loops. The S- 250 loop features an S- 200 two bed radial flow converter followed by a boiler and S- 50 single bed radial flow converter in series. A new three bed radial flow S- 300 converter has also been developed which is cheaper than S - 250 configuration and conversion is about the same.

Special focus should be paid to the following features:

Improved reformer design New nozzle burner for secondary reformer Modified S-200 ammonia converter design Condensate stripping with medium pressure steam

5

Improved catalysts such as e.g. the shape-optimized secondary reforming catalyst

With regard to recovery of the hydrogen in the purge gas from the ammonia synthesis loop, membrane technology is used today in ammonia plants. The operating principle is that the pressure difference is used as the driving force for the hydrogen to permeate the membrane. In contrast to the cryogenic unit, this means that the hydrogen-rich stream is recovered at a lower pressure. The recovered hydrogen is added to the make-up gas upstream the synthesis gas compressor.

After going through all the above process and having a look at their positive and negative points, we took all the positive points from the above process and tried to develop a new process Haldor Topsoe A/S Process. In this process the new things which we have introduce are that we are using aMDEA for CO2 Removal. In this process we are using a molecular sieve, Methanator and a CO2 absorber in a single train which provides us with extra life for Ammonia converter Catalyst. We have also used radial flow S - 300 Ammonia converter which gives higher conversion.

Looking at the investment cost, the reformer cost goes up more or less proportionally with the number of tubes, and since the reformer cost is a significant part of the overall plant cost, it is important to keep the reformer size small, and thereby minimize the plant cost. From an energy point of view, there is no doubt that this scheme is very energy efficient, and gives the lowest specific energy consumption figures of all currently studied schemes considered for large scale plants.

To increase the amount of Hydrogen in the process gas

A reactor is a vessel used to carry out the desired reaction under a

controlled set of conditions. This unit is said to be the heart of the process industry, and in

fact, all the other units usually act as its auxiliaries.

2.2 LTS and HTS designObjective:

Reactor:

Introduction:

Every industrial chemical process is designed to produce

economically a desired product from a variety of starting materials through a succession of

treatment steps. The raw materials undergo a number of physical treatment steps to put them

in the form in which they can be reacted chemically. In the next step they pass through the

reactor. The products of the reaction must then undergo further physical treatment

separations, purification, etc. For the final desired product to be obtained. Design of

equipment for the physical treatment steps is studied in the unit operations. Economically this

may be an inconsequential unit, perhaps a simple mixing tank. Frequently, however, THE

CHEMICAL TREATMENT STEP IS THE HEART OF THE PROCESS, the thing that

makes or breaks the process economically. Design of the reactor is no routine matter, and

many alternatives can be proposed for a process. In searching for the optimum it is not just

the cost of the reactor that must be minimized. One design may have low reactor cost but the

materials leaving the unit may be such that their treatment requires a much higher cost than

alternative designs. Hence, the economics of the overall process must be considered. Reactor

design uses information, knowledge and experience from a variety of areas Thermodynamics,

Chemical kinetics, Fluid mechanics, Heat transfer, Mass transfer, and Economics. Chemical

reaction engineering is the synthesis of all these factors with the aim of properly designing a

chemical reactor. To find what a reactor is able to do we need to know the Kinetics, the

contacting pattern and the performance equations. In the designing of reactor we want to

know that what size and type of reactor and method of operation are best for a given job.

Because this may require that the conditions in the reactor vary with position as well as time,

this question can only be answered by a proper integration of the rate equation for the

operation. This may pose difficulties because of the temperature and the composition of the

Muhammad MujahidTypewritten Text6

reacting fluid may vary point to point within the reactor, depending on the endothermic or

exothermic character of the reaction, the rate of heat addition or removal from the system,

and the flow pattern of the fluid through the vessels. In effect, then, many factors must be

accounted for in predicting the performance of a reactor.

Classification of chemical reactors can be achieved in two different kinds.

1. Depending upon the nature of reactants and product

(a) Homogeneous Reactors

(b) Heterogeneous Reactors

2. Depending upon the mode of operation

(a) Batch Reactors

(b) Continuous Reactors ( Stirred tank reactor, Tubular flow Reactor)

In Homogeneous Reactors only one phase, usually a gas

or a liquid, is present. If more than one reactant is involved, provision must of course be

made for mixing them together to form a homogeneous whole. Often mixing the reactants is

the way of starting off the reaction, although some times the reactants are mixed and then

brought to the required temperature.

In Heterogeneous Reactors two or possibly three,

phases are involved, common examples are gas liquid, liquid solid, and liquid liquid

systems. In cases where one of the phases is a solid, it is quite often present as a catalyst; gas

solid catalytic reactors particularly form an important class of Heterogeneous catalytic

Reactors. In a gas solid catalytic reactor, the reaction takes place on the surface of catalyst

and hence heterogeneous. However, bubbling a gas through a liquid may serve just to

dissolve the gas in the liquid where it then reacts homogeneously. The reaction is then

homogeneous but the reactor is heterogeneous itself. Generally, a heterogeneous reactor

exhibits a greater variety of configurations and contacting pattern than homogeneous reactors.

Another kind of classification which cuts cross

the homogeneous heterogeneous division is the mode of operation. Batch wise operation

has carried out small scale preparative reactions in the laboratory. There are many situations,

Classification of Reactors:

Homogeneous Reactors:

Heterogeneous Reactors:

Batch and Continuous Reactors:

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however, especially in large scale operations, where considerable advantage occurs by

carrying out the chemical reaction continuously in the flow reactor.

Continuous reactor

Batch reactor

1. It is used for low

residence time

operations.

2. It has low operating cost.

3. Production variation

takes time to attain

steady state.

4. Continuous flow reactors

are smaller than batch

reactors providing the

same productivity.

1. It is used for high

residence time

operations.

2. Its operating cost is high

3. Production can be varied

easily during each batch.

4. Batch reactors are larger

than continuous reactors

providing the same

productivity.

As we have very high flow rate in our process so we are going to use a continuous flow

reactor. Flow the below comparison we will come to know which continuous flow reactor is

best to fulfill our purpose.

There are mainly three types of continuous flow reactors:

1. Fixed Bed reactor.

2. Fluidized bed reactor.

3. CSTR reactor.

Difference between continuous and batch reactor:

Continuous flow reactors:

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Fixed bed reactor

Fluidized bed reactor

CSTR reactor

1. They primarily used for

gas phase solid

catalyzed reaction &

gas solid reactions.

2. Pressure drop is low.

3. High catalyst load per

unit volume of reactor.

4. Catalyst replacement is

relatively hard.

5. Relatively low heat and

mass transfer.

1. It is also mostly

used for gas

phase solid

catalyzed reaction

and gas solid

reactions.

2. Pressure drop is

high.

3. Catalyst is

regenerated

continuously.

4. High heat and

mass transfer.

1.Useful for Liquid-Liquid

system.

2.Useful for slow reactions

requiring large hold up time.

3.Useful where high mixing is

required.

4.High heat and mass transfer

efficiencies.

5.High power consumption per

unit volume of fluid.

The points which should be considered in

case of selection of type of reactor are given below.

1- Conversion.

2- Selectivity.

3- Productivity.

4- Yield.

5- Heat exchange.

6- Mixing.

7- Catalyst Distribution.

8- Hold-up Time.

9- Availability.

10- Compatibility with processing.

11- Energy utilization.

12- Safety.

13- Economics.

Selection Criteria of Rector:

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We have selected fixed bed reactor. Reasons are given below.

1. Very high Flow rate is required thats why we are using fixed bed reactor.

2. The reactor is adiabatic in nature (no heat transfer media is used) because it aids in the

reaction.

3. Fixed bed reactors have a low pressure drop.

4. There is no requirement of catalyst regeneration.

5. Low operation cost.

Now we turn our attention to the reaction

taking place in the reactor. Water Gas shift reaction is taking place in the reactor. There are

three types of fixed bed reactors in which Water gas shift reaction can be carried out. Water

gas Shift reaction is defined as

Water gas Shift Converter is an reactor in which Carbon monoxide reacts with water to

Produce Carbon dioxide and Hydrogen, at certain temperature and Pressure Condition, in

the presence of catalyst.

CO + H2O CO2 + H2 H298= 41.1kJmol

1

We will go through their theory and type of catalyst that they use one by one. 1. High Temperature Shift Converter.

2. Intermediate temperature Shift Converter.

3. Low Temperature Shift Converter.

In the unreduced state the HTS

catalyst is iron (III) oxide (Fe2O3) containing additionally 5 - 10 % chromic oxide (Cr203).

During operation, it is reduced more or less stoichiometrically to the composition of

magnetite (Fe2O3) 15931 - 15951. The catalyst is active in the temperature range of 300 - 500

"C. Steam surplus is not only necessary for thermodynamic reasons but also to suppress

undesirable side reactions. Decreasing the steam surplus lowers the oxygen to carbon ratio in

the HTS to such an extent that the atmosphere can reduce magnetite partially to metallic iron.

In addition the Koudouard reaction can occur under these conditions. The resulting carbon is

deposited within the catalyst particles causing their disintegration, and iron carbides will be

formed, which are effective Fischer - Tropsch catalysts that lead to the formation of some

High Temperature Shift Converter:

Water Gas Shift reaction:

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methane and higher hydrocarbons. The minimum steam surplus depends on the CO/CO,

ratio, lower values allowing lower steam surplus. Thus in conventional plant operation at 32 -

35 bar with CO/CO, of 1.65 -2.13, an S/C ratio in the primary reformer of around 3.0 seemed

to be the minimum with respect to the HT shift, whereas for concepts with reduced p rimary

reforming and a CO/CO2 ratio of 1.2 and S/C ratio of 2.8 was sufficient. Newly introduced

HTS catalysts with additional copper promotion (e.g. 3 %) suppress this side reaction and are

therefore less sensitive to lower steam-to-gas ratios. The presence of MgO and ZnO should

also effect some reduction of methane formation. A review of catalyst research, surface

science investigations and relevant literature is given in. The function of Cr2O3 is mainly to

prevent the sintering of the iron oxide, which leads to a reduction of surface area. But

attempts to substitute Cr by Ce or Zr [604] - [607] (because in commercial Fe/Cr catalysts a

minor amount of the chromium is still present as Cr6+ resulted in a lower activity of these Cr-

free catalysts. In the production of the HTS catalyst, mixed iron(III)hydroxide and

chromium(III)- hydroxide are precipitated from an aqueous solution of ferrous sulfate and

chromic acid by addition of sodium hydroxide solution while agitating with air to perform the

oxidation of Fe2 to Fe3 and injecting steam for temperature adjustment. The oxide slurry is

washed in alternating agitating and settler operations (to achieve very low residual sulfur

content). The washed hydroxides are spray-dried and the resulting powder is granulated with

an additive which serves as lubricant and binder and then pressed into tablets of the required

size 6 x 6 mm, 9 x 5 m, 9 x 9 mm are standard, but other dimensions are possible, too). Some

manufactures use iron nitrate instead of iron sulfate as raw material. The BET surface area of

commercial Fe-Cr HT shift catalysts is between 30 and 80 m2/g, depending on Cr2O3 content

and calcining temperature. The space velocities in commercial HT reactors are today

approximately 3500 - 4500 h-1. The classical HTS iron catalyst is resistant against sulfur

compounds, but this is of greater importance in partial oxidation processes and less for the

practically sulfur- free steam reforming gas. In some ammonia process schemes operating

without a secondary reformer and applying pressure swing adsorption (PSA) for further

purification (KTI PARC), only a HTS is used.

Relatively new process

concept is the intermediate temperature shift, which performs the reaction in a single step.

The catalyst is based on a copper-zinc alumina formulation and optimized for operating in a

wider temperature range (200 - 350 0C) than the standard LTS catalyst (190 - 275 0C). The

Intermediate Temperature Shift Converter:

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04 jI

Composition of gas at entry.vol-% (dry)

Hz : 59.94N2 . 20.47CO : 3.11CO2: 1591A : 0.26CH4 : 0.32

t 0.3~3~ 0.2

100000

Catalyst temperature 230

r 10000 r2203= E 1000 D 0~ cDc::t.. 210 ---~ ~.a. Sulfur -=:; CD~c/ 100 E

~

10Top Reactor height Bottom

The LTS catalyst, supplied in pellets like the HTS catalyst, consists of 40-55 % copper oxide,

20 - 30 % zinc oxide, the balance being alumina. The catalyst properties are influenced far

more by the formulation and manufacturing procedure than by its chemical composition. It

makes a great difference whether the individually prepared components are just mixed

physically as oxides or are incorporated by co-precipitation. The latter procedure assures a

rather fine distribution of the copper oxide crystals, which are well separated from each other

by zinc oxide and even smaller alumina crystals. This retards sintering of the copper

cristallites in the reduced catalyst during prolonged operation. The co-precipitation step of the

catalyst manufacturing sequence has to be very carefully controlled with respect to

temperature, pH, and agitation to achieve the desired result. Commercial pellet sizes range

from 6.4 x 3.2 mm to 3.5 x 3 mm; the surface area is around 60 - 120 m'/g, and the pore

volume 0.35 - 0.45 ml/g. The copper oxide is reduced in situ with hydrogen and a carrier gas

(usually nitrogen) to form the fine copper crystallites of about lo-' cm on which the activity

depends. Sulfur, usually present as H2S, has to be below 0.1 ppm, but even with such low

concentrations, the catalyst is slowly poisoned. The ZnO adsorbs the sulfur and it finally

transforms into bulk ZnS. When the ZnO is exhausted in a given layer of the catalyst, the H,S

causes deactivation of the copper by sintering. The poisoning process moves through the

catalyst as a relatively sharp front and can be seen in the change of the catalyst temperature

profile over time. The LTS catalyst is protected by a guard bed, formerly loaded with ZnO,

but nowadays usually with LTS catalyst. Changing the guard bed more frequently prolongs

the service life of the main LTS catalyst bed. Without a guard bed the lifetime is normally 2-4

years, depending on gas quality. With an upstream guard bed.

Sulfur and Temperature Profile For used LYS Catalyst

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Changed at appropriate intervals the life time of the main bed can be extended to 6 10

years. Traces of chlorine compounds, which may be introduced with the natural gas or more

often with the process air to the secondary reformer, may also deactivate the LTS catalyst by

accelerating the sintering of the copper particles. Unlike sulfur poisoning, chlorine is more

diffusely distributed over the whole catalyst bed by migration as volatile zinc and copper

chlorides. A chlorine guard catalyst for installation upstream of the LTS is offered by catalyst

vendors. As chemical composition and formulation of the LTS catalyst are very similar to

methanol production catalysts, small quantities of methanol are formed and found in the

process condensate after cooling the LTS effluent. In a consecutive reaction, amines (mainly

methylamine) are formed from the methanol and traces of ammonia originating from the

secondary reformer and the HTS. These pollutants are removed from the process condensate

by steam stripping and ion exchange. Byproduct formation is higher with fresh catalyst and

declines with operating time. New catalyst types with increased activity and higher selectivity

have reduced the problem. The tendency for methanol formation increases with decreasing

steam/gas ratio.

We have used High temperature and Low temperature shift converter in our process due to

following reasons:

1. There not sufficient and reliable data available for intermediate shift converter and it

is not frequently used in the Ammonia plants.

2. In HTS at high temperature we get very high rate of reaction at the cost of low

conversion. In LTS we get a high conversion at the cost of low rate of reaction. The

combination of these two gives us a very high yield of water gas shift reaction.

3. They are help in generation of steam which fulfill the seam requirements for the

ammonia Plant.

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15

Carbon dioxide is an undesirable constituent in the synthesis gas because it poisons the ammonia synthesis catalysts. Commercially available hot potassium carbonate processes are Benfield process, Glycine Vetrocoke process and Catacarb process, while Alkanolamines processes are Monoethanolamine (MEA) and aMDEA.

Table 2.1 Comparison of Carbon Dioxide absorption Processes

MEA Catacarb Benfield Vetrocoke BASF aMDEA

Absorbent Monoethanol amine

Potassium Carbonate

Potassium Carbonate

Potassium Carbonate

Methyl diethanol amine

Activator Amine Guard IV 3

Boric acid DEA Glycine Peprazine

Concentration %

25-30 25-30 20-25 35-50

Corrosion Problems

Yes Yes Yes Yes Nil

Corrosion Inhibitor

V2O5 V2O5 As2O5 Nil

Foaming Problems

High High High High Low

Degradation and Volatility2

Degrades at high temperature, volatile

Non-volatile, No degradation

Non-volatile, No degradation

Non-volatile, No degradation

Non-volatile, No degradation

CO2 Slip in product gas (ppm)1

100 500 300 100

CO2 purity %1 99 98.5 99.4 99.95

Energy Required Kcal/kgmol of CO2 1

25000 19000 16000 9500

Toxicity Toxic corrosion inhibitors

Toxic corrosion inhibitors

Toxic corrosion inhibitors

Toxic corrosion inhibitors

Biodegradable and non-toxic solvent

2.3.1 Carbon Dioxide Absorption Processes

2.3 Carbon Dioxide Removal

16

The main selection criteria are high solubility of carbon dioxide and, equally important, high absorption selectivity of carbon dioxide. Furthermore, easy desorption is highly desirable, as it reduces the necessary regeneration temperature and pressure difference. In order to prevent the loss of solvent, a low vapor pressure and high thermal stability as well as long-term stability are benecial. Additionally, the cost and environmental toxicity of the solvents have to be taken into account, especially when evaporative loss and chemical degradation are taken into account.

BASF aMDEA process is highly energy efficient due to the elevated acid gas loadings achievable with the solvent; this enables using low circulation rates and reduced energy consumption, as well as reducing equipment size. BASF has used piperazine in its activated MDEA formulations. Piperazine is ten times more reactive than MEA.

Owing to the low vapor pressure of MDEA and the activator, there are no losses of the active solvent components. The carbon dioxide binds much less strongly to MDEA than to MEA, and the solvent character is more like a hybrid between a strong chemical and a purely physical solvent. On account of the relative weak binding forces, a substantial amount of carbon dioxide can be recovered simply by flashing to low pressure, and only a small amount has to be recovered by stripping. The process is very versatile: increasing the activator concentration shifts the character of the solvent more to the chemical side and vice versa. In BASF aMDEA process, the gas is treated in a two stage absorber using partially regenerated solution in the bottom stage and completely regenerated solution in the top stage. Rich solution from the bottom of the absorber passes through a hydraulic turbine for energy recovery and is then flashed in the high pressure flash unit where most of the dissolved inert gases are released. The rich solution then flows to the low pressure flash unit, which operates close to atmospheric pressure. Hot overhead gas from the thermal stripping column is passed through the solution in the low pressure flash tank to improve the efficiency of CO 2 removal in this vessel. A significant portion of the CO2 contained in the rich solution is stripped in the low pressure flash tank. Partially regenerated solution from the low pressure tank is split into two portions. The larger portion is fed to the bottom stage of the absorber while the balance flows to the re-boiled stripping column. Completely regenerated solution from the striping column is fed to the top absorption stage, completing the cycle.

BASF aMDEA process is highly energy efficient due to the elevated acid gas loadings achievable with the solvent; this enables using low circulation rates and reduced energy consumption, as well as reducing equipment size. BASF has used piperazine in its activated MDEA formulations. Piperazine is 10 times more reactive than MEA.

2.3.2 BASF aMDEA Process:

Basically reactor is a vessel which is used to carry out

the specific set of reactions under the specific conditions. Normally reactors are also

classified on the basis of phases which they deal. There are two general types of phases:

Homogenous phase.

Heterogeneous phase.

In case of gas-solid phase reactions. Heterogeneous phase reactors are considered.

Heterogeneous catalytic reactors are the most important single class of reactors utilized by

the chemical industry. Whether their importance is measured by the wholesale value of the

goods produced, the processing capacity, or the overall investment in the reactors and

associated peripheral equipment, there is no doubt as to the prime economic role that reactors

of this type play in modern technological society.

Heterogeneous reactors

are further classified on the basis of catalyst position or motion. Generally divided into two

types:

1. Reactors in which the solid catalyst particles remain in a fixed position relative to one

another (fixed bed, trickle bed, and moving bed reactors).

2. Reactors in which the particles are suspended in a fluid and are constantly moving

about (fluidized bed and slurry reactors).

As in our case, our reaction medium is gas-solid, in which gases reacts on the surface of solid

so first type of reactor is to be considered here. Fixed bed reactors are named as fixed because

the catalyst bed is fixed. Fixed bed reactor is further classifies on the basis of the type of

fixed bed which is given below:

A single large bed.

Multiple horizontal beds supported on trays arranged in a vertical stack.

Multiple parallel packed tubes in a single shell.

Multiple beds each in their own shell.

The use of multiple catalyst sections usually arises because of the need to maintain adequate

temperature control within the system. Other constraints leading to the use of multiple beds

Key Features of Heterogeneous Reactors:

Ammonia Converter

INTRODUCTION:

Muhammad MujahidTypewritten Text17

include those of pressure drop or adequate fluid distribution. In addition to the shell and tube

configuration, some of the possibilities for

heat transfer to or from fixed bed reactors include the use of internal heat exchangers, annular

cooling spaces or cooling thimbles, and circulation of a portion of the reacting gases through

an external heat exchanger.

The packing itself may consist of spherical, cylindrical, or randomly shaped pellets, wire

screens or gauzes, crushed particles, or a variety of other physical configurations. The

structure of the catalyst pellets is such that the internal surface area far exceeds the superficial

(external) surface area, so that the contact area is, in principle, independent of pellet size. To

make effective use of the internal surface area, one must use a pellet size that minimizes

diffusion resistance within the catalyst pellet but that also gives rise to an appropriate

pressure drop across the catalyst bed. The most commonly used direction of reactant flow is

downward through the bed. This approach gives a stable bed that will not fluidize, dance, or

lift out of the reactor. This approach minimizes catalyst attrition and potential entrainment of

catalyst fines.

1. Low costs for construction and quite simple, operation, and maintenance relative to

moving bed or fluidized bed operation.

2. It requires a minimum of auxiliary equipment.

3. Appropriate for use in small commercial units.

4. Supporting facilities would be economically prohibitive.

5. Another major advantage of this mode of operation is implicit in the use of the term

"fixed bed reactor"; i.e., there are no problems in separating the catalyst from the

reactor effluent stream.

6. Another important attribute of fixed bed reactors is the wide variation in space times

at

which they can be operated.

Catalysts change the reaction mechanism and therefore the rate of the

reaction. If the reaction rate increases, the reaction volume will decrease, reducing the cost of

the reactor. Many chemical syntheses are impractical without using a catalyst. From the early

Advantages of Fixed Bed Reactor:

Catalyst:

Muhammad MujahidTypewritten Text18

A typical composition of anindustrial ammonia-synthesiscatalyst

Composition in Activated form(%)Fe203 1.1- 1.7FeO 1....3-1 ....6Fe 79.7- 81.6CaO 0.1-0.2Si02 0.1- 0.7l\IIgO 0.3 - 0.6AI203 1.5- 2.1K20 0.2-0.5Porosity 40-50

days of ammonia production to the present, the only catalysts that have been used have been

iron catalysts promoted with nonreducible oxides. Recently, a ruthenium-based catalyst

promoted with rubidium has found industrial application. The basic composition of iron

catalysts is still very similar to that of the first catalyst developed by BASF.

Freshly reduced commercial iron catalysts that contain aluminum, potassium, and calcium

oxides as basic promoters consist of approximately 30-nm primary crystallites; the spaces

between them form an interconnecting system of pores.

Selected Catalyst:

Triply promoted Iron oxide catalyst is used

The iron catalyst consists of magnetite (Fe3O4)

which is enriched

(Promoted) most frequently with Al- and K- (or

Ca, Mg, Si)-oxides.

Catalyst is produced by fusing (melting) magnetite

ore with the other

Promoters at 1700oC and pouring melt into water to

form fine particles.

Iron being a transition metal with partially occupied

d-bands represents a surface suitable for adsorption

and dissociation of N2 molecules.

Others catalysis are:

Os, Ru are equally active with Fe.

Ru is most active in ammonia decomposition.

Mo, U and Mn show quite high activity in ammonia synthesis

However, they reduce the inner surface or

lower the temperature stability and the resistance to oxygen-containing catalyst poisons and

the synthesis gas purity. Promoters can be classified in different groups according to the

specific action of the metal oxides:

Influence of the Promoters:

Muhammad MujahidTypewritten Text19

Structural stabilizers, such as Al2O3 produce a high inner surface during reduction and

stabilize it under thermal stress by restraining iron crystallite growth. The ability of the

various metal oxides to create a high specific surface decreases in the following order:

Al2O3 > TiO2 > Cr203 > MgO > MnO = CaO > SiO2 > BeO

So-called electronic promoters, such as the alkali oxides, enhance the specific activity (based

on a unit surface) of iron - alumina catalysts.

Effect of Potassium:

Potassium increases intrinsic activity (increasing desorption rate).

Potassium increases the binding energy of the molecular N2 precursor on Fe and

thereby assists the formation of atomic nitrogen.

Potassium acts as an electron donor enhancing N2 reduction.

Potassium on Fe surface prevents the S poisoning on Fe.

Potassium enhances the reduction through an amide intermediate formation.

A promoter oxide of aluminum (3%) and potassium (1%) prevents sintering.

The activity of an ammonia synthesis catalyst may be lowered

by certain substances, commonly referred to as poisons. These substances can be minor

gaseous constituents of the synthesis gas or solids introduced into the catalysts during the

manufacturing procedure, derived from impurities in the natural magnetite from which the

catalyst is made. In contrast to temporary poisons, permanent poisons can be detected on the

catalyst by chemical analysis. Oxygen-containing compounds such as H2O, CO, CO2, and O2,

are the most common temporary poisons encountered in ammonia.

Catalyst Poisons:

Muhammad MujahidTypewritten Text20

Tllpuill 81 IlIdlal

~iIf

Ammonia converters are classified by

flow type

1. Radial Flow.

2. Axial Flow.

3. Cross Flow.

Cooling method

1. Quench cooling.

2. Indirect cooling

One design uses 2 radial beds with quench gas injection

between them.

A similar radial flow design uses an inter-bed heat

exchanger in the first catalyst bed.

Cold ammonia synthesis gas is introduced from the

bottom of the converter through the second catalyst bed

then through the heat exchanger in the first catalyst bed.

The cold gas flow through the second bed also provides

indirect heat exchange.

An additional heat exchanger is located at the bottom of

the reactor to cool the reacted gases.

This reactor is a cross-flow converter

design where gas flows through the

catalyst bed perpendicular to the axis of

the vessel.

It is available in both quench and

indirectly cooled versions.

AMMONIA CONVERTER CLASSIFICATION:

Topsoe radial flow converter:

Kellogg horizontal converter:

Muhammad MujahidTypewritten Text21

A t1AIN GAS INLETB INLET FOR GAS TO

LOWER THE TUBE SIDEC COLD BY-PASS INLETD GAS OUTLET

PRESSURE SHELL2 OUTER ANNULUS3 OUTER BASKET SHELL4 BASKET INSULATl ON5 BASKET COVER6 INTERBED HEAT EXCHANGER c i HE)7 TRANSFER PIPE8 SCREEN PANELS9 1st CATALYST BED

10 CENTRE SCREEN11 COVER PLATE12 CATALYST SUPPORT13 2nd BED SUPPORT FLANGE14 2nd CATALYST BED15 3rd CATALYST BED

This reactor consists of 4 catalyst beds held on

separate grids.

Quench gas is introduced in the spaces between the

beds.

A heat exchanger is located at the top of the vessel.

With the S-300 converter, the catalyst volume can be reduced by approx. 20%

100% radial flow through the catalyst beds to obtain low pressure and high

conversion with a small size catalyst Particle.

Indirect cooling of the gas in the heat exchangers between the catalysts.

Total converter feed flow passes through all beds .stable operation with great

flexibility in operating range simple temperature control.

S-300 Ammonia Converter

Kellogg axial flow converter:

Topsoe radial flow converter S-300:

Muhammad MujahidTypewritten Text22

23

Old ammonia plants do not use recovery units. Mostly, the new plants are using cryogenic separation technology for hydrogen recovery.

Cryogenic separation technique is a complex system that requires great operator attention and large space to accommodate. It can cause health and safety problems because of the very low temperatures and explosion.

As membrane technology for gas processing becomes wider spread and more accepted, there will undoubtedly be more applications. There will always be special applications in which membrane technology will uniquely fit. Membrane technology is a new emerging technique for hydrogen recovery from purge gas that offers the following advantages:

It is simple, straightforward and, therefore, easy to operate and maintain. It can operate under fluctuating feed conditions It offers unique operating flexibility when planned or unexpected process changes occur.

Increment in capacity requirements can be handled by the simple addition of more separators. Some turndown is accommodated automatically.

It can be easily fitted into small or crowded plant areas because of the simplicity and compact size of the skid mounted system.

It offers low capital cost and energy requirements. It is environment friendly.

2.5 Hydrogen Recovery Unit

24

Broadly, hydrogen selective membranes (based on the materials used) can be categorized into two types: organic (polymer) and inorganic (metallic, carbon, and ceramic). High-purity hydrogen (up to 99.99%) could be available through dense metallic membranes and especially through Pd and its alloys, but the commercial application of metallic membranes faces today important limitations due to drawbacks like:

Poisoning effect that hydrogen sulfide (H2S) and other feed stocks have on the hydrogen transport mechanism.

High cost for the preparation of Pd membranes. Mechanical stability.

The key advantages of polymer membranes are the ability to cope with high pressure drops and low cost. Polymers also provide a range of properties including permeability, mechanical stability, and ease of process ability that are important for gas separation. Therefore, the separation of H2 by polymeric membranes has become an attractive existing technology.

Another property of membranes used for subdivision, is the membrane structure and connected to this way in which transport through the membrane takes place. Porous membranes enable transport through their pores, whereas dense membranes allow transport through the bulk of the material.

Finally, the morphology or structural make-up of membranes can also be used for classification. Symmetrical membranes have a homogeneous structure. Asymmetric membranes consist of several layers with different characteristics. There can also be a gradual transition from a dense membrane to a porous support. Membranes consisting of different layers of different materials are called composite membranes.

Table 3.1 Engineering Strategies for Hydrogen Separation membrane

Selectivity General H2 Selectivity H2 Rejection DH2/Dgas >>> 1 >>>> 1 >> 1 SH2/Sgas > 1 < 1

D is the Diffusivity coefficient S is the Sorption Coefficient P is Permeability

3 MEMBRANE SEPARATION TECHNOLOGY 3.1 Membranes for Hydrogen Separation

25

Table 3.2 Comparison of Membranes for Hydrogen Separation

Dense Polymer

Micro porous Ceramic

Dense Metallic

Porous Carbon

Dense Ceramic

Temperature Range

< 100 oC 200-600 oC 300-600 oC 500-900 oC 600-900 oC

H2 Selectivity

Low 5-139 > 1000 4-20 > 1000

Poisoning Issues

HCl, Sox H2S, HCl, CO

Strong adsorbing vapors, organics

H2S

Materials Polymers Silica, alumina, titania, zeolites

Palladium alloy

Carbon Proton conducting ceramics

Transport Mechanism

Solution diffusion

Molecular sieving

Solution diffusion

Surface diffusion, molecular sieving

Solution diffusion

Development Status

Commercial by Air Products, Linde, Air Liquid

Prototype tubular silica membranes

Commercial by Johnson Mathey

Small membrane modules commercial

Small samples available for testing

While choosing a membrane, there are four important parameters that must be considered:

1. Permselectivity for the required component 2. Mechanical strength 3. Chemical resistance 4. Thermal stability

After investigating permselectivity for hydrogen gas, mechanical strength to withstand such high pressure of 122 atm, chemical resistance to incoming gases feed i.e. no chemical reaction occurs with the gases, from the literature, polyimide was our first priority.

3.3 Hollow Fibers Preparation

Modern membrane fabrication is based on the LoebSourirajan technique, which allows the generation of asymmetric membrane morphologies. Membranes prepared according to this

3.2 Why Polyimide?

26

formation principle, consist of an extremely thin layer, which determines the separation properties and a thicker porous support layer, which provides for the required mechanical strength of the whole membrane.

The process of making asymmetric hollow fiber membranes is called spinning. The polymer solution used in making the membranes is commonly called a dope. The dope typically consists of the polymer, a solvent, and a non-solvent.

Three main routes are known to prepare hollow fiber membranes;

1. Melt spinning 2. Dry spinning 3. Wet spinning (or dry/wet spinning).

The latter is the most important technique for the preparation of industrial hollow fiber membranes. The majority of the hollow fibers employed in technical membrane processes are spun by a wet spinning (or dry/wet spinning) technique. Any type of membrane morphology can be obtained with this technique since many parameters involved can be varied.

In this process, the polymer solution is extruded into a non-solvent bath where de-mixing occurs because of exchange of solvent and non-solvent, a high polymer concentration is created near the outer surface, which is necessary to form a dense (gas selective) membrane top layer. Between the spinneret and non-solvent bath there is an air gap where in fact the membrane formation starts. This implies that a good control of this phase is a first requirement. The membrane formation process can be controlled by the choice of the non-solvent which is extruded through the outer opening as a thin film adhering to the polymer solution. Both the asymmetric microfiltration and ultrafiltration hollow fibers can be prepared as well as integrally skinned hollow fibers with a defect-free top layer suitable for gas separation and pervaporation.

Vers and Smolders distinguished two different de-mixing processes:

1. Delayed de-mixing: Where the ratio of solvent outflow and non-solvent inflow is relatively large, which results in a certain time interval between immersion of the polymer solution in the non-solvent bath and the onset of de-mixing.

2. Instantaneous de-mixing: Where liquid-liquid de-mixing takes place immediately after the polymer solution is in contact with the non-solvent.

Vant Hof et al. showed, around 1992, that asymmetric membrane could be made using a dual coagulation bath method, The first bath consists of a non-solvent with reduced non-solvency power or reduced miscibility with solvent so that delayed de-mixing occurs (via increased solvent outflow compared to non-solvent inflow), to form a dense region on the outside. The second bath contains a strong non-solvent which promotes instantaneous de-mixing and complete phase separation.

727

There is an air gap between the spinneret and non-solvent and this air gap has a large impact on the ultimate membrane properties. In order to control the processes in the air gap, a new type of spinneret has been developed where the extrusion mouth has three openings. This new triple orifice allows a much better control of the conditions applicable for the spinning of all types of hollow fibers. For the preparation of commercial membranes the phase inversion process is generally applied. In this process, a polymer solution is immersed into a non-solvent bath where phase separation (liquid-liquid de-mixing) occurs as a result of the exchange of solvent and non-solvent. The first bath consists of a non-solvent with reduced non-solvency power or reduced miscibility with solvent in air gap so that delayed de-mixing occurs (via increased solvent outflow compared to non-solvent inflow), to form a dense region on the outside. The second bath contains a strong non-solvent which promotes instantaneous de-mixing and complete phase separation. After solidification of the polymer an asymmetric membrane with a relatively dense top layer remains. The fibers obtained were dried in air after rinsing in running water for two days followed by immersion in ethanol for 8 hours and then in hexane for another 8 hours.

Figure 3.1 Experimental set-up for hollow fiber spinning (1: spinning dope reservoir, 2: bore liquid reservoir, 3: tube-in-orifice spinneret, 4: chimney (air gap), 5: coagulation bath (water), 6: fiber guiding wheel, 7: pulling wheel, 8: hollow fiber membrane, 9: fiber collecting reservoir (water))

3.4 Hollow Fiber Module

A module is the smallest practical unit containing a set of membrane area and any supporting structures. For each membrane certain module shapes are commercially available. Hollow fiber is preferred commercially for high pressure gas separation.

productivi ty = p. A /).pI

28

Hollow fibers are essentially self-supporting and resistant to collapse in high pressure and environmentally difficult situations (Paul, 1994). The support is a porous substructure that ideally has negligible resistance and lies underneath the outer skin layer, thereby providing the mechanical strength to the fiber. Hollow fiber modules are preferably used when the feed stream is relatively clean. Hollow fiber membranes can withstand very high pressures from the outside, but are limited on pressure exerted from the inside of the fiber, therefore backwash rates are limited to around twice the normal permeate rate. The feed fluid is applied on the outside of the fibers and permeate is removed down the fiber bore.

Asymmetric hollow fibers provide high fluxes required for productive separations due to the ability to reduce the separating layer to a thin integral skin on the outer surface of the membrane. The high trans-membrane pressure differences capable of being handled, and the high surface area to volume ratios and high packing densities achieved from such cylindrical morphologies make this type of membrane structure very favorable industrially. The productivity of a membrane can be characterized by the amount of penetrant that permeates through the membrane within a given time period and is inversely proportional to the thickness of the separating layer according to equation

Where

P is the intrinsic membrane permeability (a property of the material) A is the membrane surface area for permeation p is the trans-membrane pressure difference driving force l is the membranes effective thickness, which can be controlled morphologically

For this reason, it is clear that a thinner skin would allow use of smaller (lighter weight) modules and reduce capital costs for a given membrane.

3.5 Asymmetric Membrane Transport Mechanism

Generally, gas transport in non-porous (dense) membranes occurs via the solution-diffusion mechanism, comprised of three main steps:

1. Sorption of the gaseous penetrants at the upstream side of the membrane, 2. Diffusion of the penetrants across the membrane, 3. Desorption of the penetrants at the downstream side of the membrane

29

3.5.1 Solution-diffusion model

Permeates dissolve in the membrane material and then diffuse through membrane down concentration gradient. The permeates are separated because of the differences in the solubility of the materials in the membrane and the differences in the rates at which the materials diffuse through the membrane.

The solution diffusion mechanism is driven by a difference in the thermodynamic activities existing at the upstream and downstream faces of the membrane as well as the interacting force working between the molecules that constitute the membrane material and the permeate molecules. The activity difference causes a concentration difference that leads to diffusion in the direction of decreasing activity.

3.5.1.1 Diffusion

The basis of the solution-diffusion model is the process by which matter is transported from one part of system to another by a concentration gradient. The individual molecules in the membrane medium are in constant random molecular motion, but in an isotropic medium, individual molecules have no preferred direction of motion. Although the average displacement of an individual molecule from its starting point can be calculated, after a period of time nothing can be said about the direction in which any individual molecule will move. However, if concentration gradient of permeate molecules is formed in the medium, simple statistics show that a net transport of matter will occur from the high concentration to low concentration region.

This concept was first recognized by Fick theoretically and experimentally in 1855. Fick formulated his results as the equation now called Ficks law of diffusion, which states

Where

Ji is the rate of transfer of component i or flux (g/cm s) dci/dx is the concentration gradient of component i Di is called the diffusion coefficient (cm/s) and is a measure of the mobility of the

individual molecules.

The minus sign shows that direction of diffusion is down the concentration gradient. Diffusion is an inherently slow process. In practical diffusion-controlled separation processes, useful fluxes across the membrane are achieved by making the membranes very thin and creating large concentration gradients in the membrane.

C = KDP +11/ c~bp1+ bp

30

3.5.1.2 Sorption

The sorption coefficient is the term linking the concentration of a component in the fluid phase with its concentration in the membrane polymer phase. Because sorption is an equilibrium term, conventional thermodynamics can be used to calculate solubility's of gases in polymers to within a factor of two or three. In fact, sorption coefficients of gases in polymers are relatively constant for a wide range of chemically different polymers. Sorption coefficients of gases in polymers remain relatively constant, as sorption in polymers behaves as though the polymers were ideal fluids.

According to dual-sorption model, gas sorption in a polymer (cm) occurs in two types of sites.

1. The first type is filled by gas molecules dissolved in the equilibrium free volume portion of material. Henrys law occurs in the equilibrium free volume portion of the polymer.

2. The other fraction is assumed to be sorbed into the excess free volume elements, which are limited, so sorption will cease when all the sites are filled. Sorption in these sites is best approximated by a Langmuir-type absorption isotherm.

Hence, the total Sorption is given as:

3.6 Working of Hollow Fiber Module

A closed bundle of fibers is contained in a pressure vessel. The feed is entered into the shell side at a high pressure of 122 atm; permeate passes through fiber wall and exits through open fiber ends. Permeate is introduced into the syn compressor. Pressure at permeate side is maintained at 31 atm while the reject is collected from the top of the module and is used for burning purposes.

Membranes consist of an extremely thin layer, which determines the separation properties and a thicker porous support layer, which provides for the required mechanical strength of the whole membrane. For modeling of membrane separation processes, it is often assumed that the total membrane resistance is determined only by the thin active layer of the membrane

Stage cut is an important factor in determining the performance of the system. It is the ratio of permeate flow to the feed flow rate. To achieve the same output results i.e. same recovery and purity of hydrogen gas, stage cut of the system must not be changed.

Feed

Permeate

Cl

55

--5 -22r 56

32

4 MATERIALS BALANCE Calculated by simultaneously solving a set of 124 equations, by using MS Excel.

4.1 NG and Recycle Stream Mixing Point

Components 5-NG feed 22-recycle from syn compressor 6-desulphurizer

inlet mole% kmol/hr mole% kmol/hr mole% kmol/hr

H2 0.00% 0.000 70.21% 85.195 4.55% 85.195 N2 18.00% 315.281 28.34% 34.393 18.67% 349.674

CO2 8.00% 140.125 0.00% 0.000 7.48% 140.125 Ar 0.00% 0.000 0.26% 0.315 0.02% 0.315

CH4 73.80% 1292.651 1.19% 1.447 69.10% 1294.098 C2H6 0.20% 3.503 0.00% 0.000 0.19% 3.503 Total 100.00% 1751.560 100.00% 121.350 100.00% 1872.910

Equations

H2: fH2,22 = fH2,6 fH2,6 fH2,22 = 0

N2: fN2,5 + fN2,22 = fN2,6 fN2,6 fN2,22 = 18 CO2: fC02,5 = f C02,6 f C02,6 = 8 Ar: fAr,22 = fAr,6 fAr,6 fAr,22 = 0 CH4 : fCH4,5 + fCH4,22= fCH4,6 fCH4,6 fCH4,22 = 73.8 C2H6: fC2H6,5 = fC2H6,6 fC2H6,6 = 0.2

33

4.2 Desulfurizer

Desulfurizer

S 6

S 7 Equations

H2: fH2,6 = fH2,7 fH2,6 fH2,7 = 0 N2: fN2,6 fN2,7 = 0

CO2: fCO2,6 fCO2,7 = 0

Ar: fAr,6 fAr,7 = 0

CH4: fCH4,6 fCH4,7 = 0

C2H6: fC2H6,6 fC2H6,7 = 0

Sulfur content is reduced from 1-2 ppm to 1000 then 10 ppb.

Components 7-desulphurizer outlet

mole % kmol/hr H2 4.55% 85.195 N2 18.67% 349.674

CO2 7.48% 140.125 Ar 0.02% 0.315

CH4 69.10% 1294.098 C2H6 0.19% 3.503 Total 100.00% 1872.910

s7 ~1~S-8----~S9

34

4.3 Feed Gas and Steam Mixing Point

Equations

H2: fH2,7 = fH2,9 fH2,7 fH2,9 = 0 N2: fN2,7 fN2,9 = 0

CO2: fCO2,7 fCO2,9 = 0

Ar: fAr,7 fAr,9 = 0

CH4: fCH4,7 fCH4,9 = 0

C2H6: fC2H6,7 fC2H6,9 = 0

H2O: fH2O,8 = fH2O,9 fCH4 - fH2O,9 = 0 As, steam / C ratio = 2.7

H2O,84,7 + 226,7 = 2.7 2.7 fCH4,7 + 5.4 fC2H6,7 fCH4 = 0

Components 8-process steam 9-primary

reformer inlet mole % kmol/hr mole % kmol/hr

H2 0.00% 0.000 1.58% 85.195 N2 0.00% 0.000 6.49% 349.674

CO2 0.00% 0.000 2.60% 140.125 Ar 0.00% 0.000 0.59% 0.315

CH4 0.00% 0.000 24.03% 1294.098 C2H6 0.00% 0.000 6.50% 3.503 H2O 100.00% 3512.982 65.23% 3512.982 Total 100.00% 3512.982 100.00% 5385.892

,--------i-

L '--- -

35

4.4 Primary Reformer

S 9

S 10

Primary Reformer

Equations

The general overall reaction for steam reforming of hydrocarbons can be formulated as:

CnH(2n+2) + nH2O nCO + (2n+1) H2

Specifically for methane:

CH4 + H2O CO + 3 H2 H0 298 = 206 kJ/ mol

Simultaneous with this reaction water gas shift reaction proceeds as:

CO + H2O CO2 + H2 H0 298 = -41 kJ/ mol

When this is included following overall reaction can be written:

CH4 + 2H2O CO2 + 4 H2 H0 298 = 165 kJ/ mol

At the given conditions,

Conversion of methane = 60 % Yield of CO = 60 % Yield of CO2 = 40 %

Let conversion of ethane is 100 % and completely gives CO

H2: fH2,9 + 0.6(0.4)(4) fCH4,9 + 0.6(0.6)(7)fCH4,9 + 5fC2H6,9 = fH2,10

fH2,9 + H2.04 fCH4,9 + 5fC2H6,9 - fH2,10 = 0 N2: fN2,9 = fN2,10 fN2,9 - fN2,10 = 0 CO: 0.6(o.6) fCH4,9 + 2 fC2H6,9 = fCO,10 0.36 fCH4,9 + 2 fC2H6,9 - fCO,10 = 0

Components 10-primary

reformer outlet mole % kmol/hr

H2 39.45% 2742.671 N2 5.03% 349.674 CO 6.80% 472.882 CO2 6.48% 450.708 Ar 0.00% 0.315

CH4 7.45% 517.639 H2O 34.79% 2418.934 Total 100.00% 6952.823

Sl1

36

CO2: fCO2,9 + 0.6(0.4) fCH4,7 = fCO2,10 fCO2,9 + 0.24 fCH4,7 - fCO2,10 = 0 Ar: fAr,9 = fAr,10 fAr,9 - fAr,10 = 0 CH4: fCH4,9 0.6 fCH4,9 = fCH4,10 0.4fCH4,9 fCH4,10 = 0 H2O: fH2O,9 0.6(.6) fCH4,9 0.6(0.4)(2) fCH4,9 2 fC2H6,9 = fH2O,10

-0.84 fCH4,9 2 fC2H6,9 + fH2O,9 fH2O,10 = 0 4.5 Secondary Reformer

Equations

Combustion reactions (upper portion):

CH4 + 2O2 CO2 + 2 H2O

H2 + O2 H2O

CO + O2 CO2

Reforming reaction (lower portion):

CH4 + H2O CO + 3 H2

CH4 + 2H2O CO2 + 4 H2

At the given conditions,

Conversion of methane = 90 % Yield of CO = 90 % Yield of CO2 = 10 %

Components 11-process air 12-2ndry reformer

outlet mole % kmol/hr mole % kmol/hr

H2 0.00% 0.000 37.21% 3449.942 N2 78.04% 1369.963 18.55% 1719.636 O2 20.99% 368.471 0.00% 0.000 CO 0.00% 0.000 9.62% 892.169 CO2 0.09% 1.580 5.38% 498.876 Ar 0.88% 15.448 0.17% 15.763

CH4 0.00% 0.000 0.56% 51.764 H2O 0.00% 0.000 28.51% 2643.414 Total 100.00% 1755.462 100.00% 9271.565

37

Let only H2 combust with O2

H2: fH2,10 2(0.2099) f11 + 0.9(0.9) (1) fCH4,10 + 0.9(0.1)(4) fCH4,10 = fH2,12

fH2,10 + 2.79 fCH4,10 0.4198 f11 fH2,12 = 0 N2: fN2,10 + 0.7804 f11 = fN2,12 fN2,10 + 0.7804 f11 fN2,12 = 0 CO: fCO,10 + 0.9(0.9) fCH4,10 = fCO,12 fCO,10 + 081 fCH4,10 - fCO,12 = 0 CO2: fCO2,10 + 0.9 (0.1) fCH4,10 +0.0009 f11 = fCO2,12

fCO2,10 + 0.09 fCH4,10 +0.0009 f11 fCO2,12 = 0 Ar: fAr,10 + 0.0088 f11 = fAr,12 fAr,10 + 0.0088 f11 - fAr,12 = 0 CH4: fCH4,10 0.9 fCH4,10 = fCH4,12 0.1 fCH4,10 - fCH4,12 = 0 H2O: fH2O,10 + 2(o.2099) f11 0.9(.9) fCH4,10 0.9(0.1)(2) fCH4,10 = fH2O,12

0.99 fCH4,10 + fH2O,10 + 0.4198 f11 fH2O,12 = 0 4.6 High Temperature Shift Converter - HTSC

HTS Convertor

S 12

S 13

Equations

CO + H2O CO2 + H2 H0 298 = -41.2 kJ/ mol

At given conditions, conversion = 90 %

H2: fH2,13 + 0.9 fCO,12 = fH2,13 fH2,13 + 0.9 fCO,12 - fH2,13 = 0 N2: fN2,12 = fN2,13 fN2,12 fN2,13 = 0 CO: fCO,12 0.9 fCO,12 = fCO,13 fCO,12 0.9 fCO,12 - fCO,13 = 0

Components 13-HTSC outlet

mole % kmol/hr H2 45.87% 4252.894 N2 18.55% 1719.636 CO 0.96% 89.217 CO2 14.04% 1301.828 Ar 0.17% 15.763

CH4 0.56% 51.764 H2O 19.85% 1840.461 Total 100.00% 9271.565

38

CO2: fCO2,12 + 0.9 fCO,12 = fCO2,13 fCO2,12 + 0.9 fCO,12 - fCO2,13 = 0 Ar: fAr,12 = fAr,13 fAr,12 - fAr,13 = 0 CH4: fCH4,12 = fCH4,13 fCH4,12 - fCH4,13 = 0 H2O: fH2O,12 0.9 fCO,12 = fH2O,13 fH2O,12 0.9 fCO,12 - fH2O,13 = 0 4.7 Low Temperature Shift Converter LTSC

S 13

LTS Convertor

S 14

Equations

CO + H2O CO2 + H2 H0 298 = -41.2 kJ/ mol

At given conditions, conversion = 80 %

H2: fH2,13 0.8 fCO,13 = fH2,14 fH2,13 0.8 fCO,13 - fH2,14 = 0 N2: fN2,13 = fN2,14 fN2,13 - fN2,14 = 0 CO: fCO,13 - 0.8 fCO,13 = fCO,14 0.2 fCO,13 - fCO,14 = 0 CO2: fCO2,13 + 0.8 fCO,13 = fCO2,14 0.8 fCO,13 +fCO2,13 fCO2,14 = 0 Ar: fAr,13 = fAr,14 fAr,13 - fAr,14 = 0 CH4: fCH4,13 = fCH4,14 fCH4,13 - fCH4,14 = 0 H2O: fH2O,13 0.8 fCO,13 = fH2O,14 0.9 fCO,13 - fH2O,13 + fH2O,14 = 0

Components 14-LTSC outlet

mole % kmol/hr H2 46.64% 4324.268 N2 18.55% 1719.636 CO 0.19% 17.843 CO2 14.81% 1373.202 Ar 0.17% 15.763

CH4 0.56% 51.764 H2O 19.08% 1769.088 Total 100.00% 9271.565

39

4.8 2nd KO Drum

S 16

KO Drum

S 16

S 14

Equations

Let 95% of H2O is removed

H2: fH2,14 = fH2,16 fH2,14 fH2,16 = 0 N2: fN2,14 - fN2,16 = 0

CO: fCO,14 - fCO,16 = 0

CO2: fCO2,14 fCO2,16 = 0

Ar: fAr,14 - fAr,16 = 0

CH4: fCH4,14 - fCH4,16 = 0

H2O: fH2O,14 = f15 + fH2O,16 fH2O,14 - f15 fH2O,16 = 0 As, f15 = 0.95 fH2O,14 0.95 fH2O,14 - f15 = 0

Components 15-condensate 16-KO drum outlet

mole % kmol/hr mole % kmol/hr H2 0.00% 0.000 56.97% 4324.268 N2 0.00% 0.000 22.65% 1719.636 CO 0.00% 0.000 0.24% 17.843 CO2 0.00% 0.000 18.09% 1373.202 Ar 0.00% 0.000 0.21% 15.763

CH4 0.00% 0.000 0.68% 51.764 H2O 100.00% 1680.633 1.17% 88.454 Total 100.00% 1680.633 100.00% 7590.931

517

C02AbsorptionColumn

516

Column

40

4.9 CO2 Absorption

Equations

Let CO2 removed = 99.8 %

H2: fH2,16 = fH2,18 fH2,16 fH2,18 = 0 N2: f3,16 - f3,18 = 0

CO: fCO,16 - fCO,18 = 0

CO2: fCO2,16 = f17 + fCO2,18 fCO2,16 - f17 fCO2,18 = 0 Ar: fAr,16 - fAr,18 = 0

CH4: fCH4,16 - fCH4,18 = 0

H2O: fH2O,16 fH2O,18 = 0

As, f17 = 0.998 fCO2,16 0.998 fCO2,16 - f17 = 0

Components 17-CO2 to urea

plant 18-CO2 absorber

outlet mole % kmol/hr mole % kmol/hr

H2 0.00% 0.000 69.52% 4324.268 N2 0.00% 0.000 27.64% 1719.636 CO 0.00% 0.000 0.29% 17.843 CO2 100.00% 1370.456 0.04% 2.746 Ar 0.00% 0.000 0.25% 15.763

CH4 0.00% 0.000 0.83% 51.764 H2O 0.00% 0.000 1.42% 88.454 Total 100.00% 1370.456 100.00% 6220.476

41

4.10 Methanator

Methanator

S 18

S 19

Equations

Let 100 % conversion of CO and CO2 into CH4

CO + 3 H2 CH4 + H2O

CO2 + 4 H2 CH4 + 2H2O

H2: fH2,18 3 fCO,18 - 4 fCO2,18 = fH2,19 fH2,18 3 fCO,18 - 4 fCO2,18 - fH2,19 = 0 N2: f3,18 = f3,19 f3,18 - f3,19 = 0 Ar: fAr,18 - fAr,19 = 0

CH4: fCH4,18 + fCO,18 + fCO2,18 = fCH4,19 fCH4,18 + fCO,18 + fCO2,18 - fCH4,19 = 0 H2O: fH2O,18 + fCO,18 + 2 fCO2,18 = fH2O,19 fCO,18 + 2fCO2,18 +fH2O,18 - fH2O,19 = 0

4.11 3rd KO Drum

KO DrumS 19

S 20

S 21

Components 19-methanator

outlet mole % kmol/hr

H2 68.94% 4259.752 N2 27.83% 1719.636 Ar 0.26% 15.763

CH4 1.17% 72.354 H2O 1.81% 111.791 Total 100.00% 6179.296

Components 20-condensate 21-KO drum outlet

mole % kmol/hr mole % kmol/hr H2 0.00% 0.000 70.21% 4259.752 N2 0.00% 0.000 28.34% 1719.636 Ar 0.00% 0.000 0.26% 15.763

CH4 0.00% 0.000 1.19% 72.354 H2O 100.00% 111.791 0.00% 0.000 Total 100.00% 111.791 100.00% 6067.506

42

Equations

Let 100 % of H2O is removed.

H2: fH2,19 = f H2, 21 fH2,19 fH2,21 = 0 N2: f3,19 - f3,21 = 0

Ar: fAr,19 - fAr,21 = 0

CH4: fCH4,19 - fCH4,21 = 0

H2O: fH2O,19 = f20 fH2O,19 - f20 = 0 4.12 1st Stage of Compressor

S 21

S 23

S 22

Equations

Let 2% of stream 21 is recycled back through stream 22

H2: fH2,21 = fH2,22 + fH2,23 fH2,21 - fH2,22 - fH2,23 = 0 fH2,22 =0.02 fH2,21 0.02 fH2,21 - fH2,22 = 0

N2: fN2,21 - fN2,22 - fN2,23 = 0

0.02 fN2,21 - fN2,22 = 0

Ar: fAr,21 - fAr,22 fAr,23 = 0

0.02 f,Ar,21 - fAr,22 = 0

CH4: fCH4,21 - fCH4,22 fCH4,23 = 0

Components 23-compressor 1st

stage outlet mole % kmol/hr

H2 70.21% 4174.557 N2 28.34% 1685.244 Ar 0.26% 15.448

CH4 1.19% 70.907 Total 100.00% 5946.156

529

524

525

43

0.02 fCH4,21 - fCH4,22 = 0

4.13 2nd and 3rd Stage of Compressor

S 34

S 23

S 24

Equations

H2: fH2,23 + 0.9328 f34 = fH2,24 fH2,23 - fH2,24+ 0.9328 f34 = 0 N2: fN2,23 + 0.0411 f34 = fN2,24 fN2,23 + 0.0411 f34 - fN2,24 = 0 Ar: fAr,23 + 0.0111 f34 = fAr,24 fAr,23 + 0.0111 f34 fAr,24 = 0 CH4: fCH4,23 + 0.0122f34 = fCH4,24 fCH4,23 + 0.0122f34 - fCH4,24 = 0 4.14 Make up Gas Mixing Point

Equations

NH3: fNH3,29 = fNH3,25 fNH3,29 - fNH3,25 = 0 H2: fH2,24 + fH2,29 = fH2,25 fH2,24 + fH2,29 - fH2,25 = 0 N2: fN2,24 + fN2,29 = fN2,25 fN2,24 + fN2,29 - fN2,25 = 0

Components 24-compressor 2ndstage outlet

mole % kmol/hr H2 73.81% 5189.948 N2 24.60% 1729.983 Ar 0.39% 27.531

CH4 1.20% 84.187 Total 100.00% 7031.649

Components 25-separator inlet

mole % kmol/hr NH3 7.65% 3162.066 H2 65.21% 26960.770 N2 21.74% 8986.923 Ar 1.33% 550.618

CH4 4.07% 1683.737 Total 100.00% 41344.114

S27 1I I

Ammonia SynthesisConvertor

I S28

44

Ar: fAr,24 + fAr,29 = fAr,25 fAr,24 + fAr,29 - fAr,25 = 0 CH4: fCH4,24 + fCH4,29 = fCH4,25 fCH4,24 + fCH4,29 - fCH4,25 = 0 4.15 Separator

Ammonia Separator

S 25

S 26

S 27

Equations

NH3: fNH3,25 = f26 + fNH3,27 fNH3,25 - f26 - fNH3,27 = 0 H2: fH2,25 = fH2,27 fH2,25 - fH2,27 = 0 N2: fN2,25 = fN2,27 fN2,25 - fN2,27 = 0 Ar: fAr,25 = fAr,27 fAr,25 - fAr,27 = 0 CH4: fCH4,25 = fCH4,27 fCH4,25 - fCH4,27 = 0 4.16 Ammonia Synthesis Convertor

Equations

3 H2 + 2 N2 2 NH3 H0 298 = - 92.44 kJ/mol

Components 26-separator liquid

outlet 27-recycle to NH3

covertor inlet mole % kmol/hr mole % kmol/hr

NH3 100.00% 2529.653 1.63% 632.413 H2 0.00% 0.000 69.46% 26960.770 N2 0.00% 0.000 23.15% 8986.923 Ar 0.00% 0.000 1.42% 550.618

CH4 0.00% 0.000 4.34% 1683.737 Total 100.00% 2529.653 100.00% 38814.461

Components 28-NH3 convertor

outlet mole % kmol/hr

NH3 9.22% 3328.490 H2 63.45% 22916.655 N2 21.15% 7638.885 Ar 1.52% 550.618

CH4 4.66% 1683.737 Total 100.00% 36118.384

529528

45

At given conditions single pass conversion is 15 %

NH3: fNH3,27 + 0.15(2)fN2,27 = fNH3,28 fNH3,27 + 0.30fN2,27 - fNH3,28 = 0 H2: fH2,27 -0.15 fH2,27 = fH2,28 0.85 fH2,27 - fH2,28 = 0 N2: fN2,27 -0.15 f2,27 = fN2,28 0.85 fN2,27 - fN2,28 = 0 Ar: fAr,27 = fAr,28 fAr,27 - fAr,28 = 0 CH4 : fCH4,27 = fCH4,28 fCH4,27 - fCH4,28 = 0 4.17 Purge Gas Exit Point

Equations

5% of stream 28 is purged out

NH3: fNH3,28 = fNH3,29 + fNH3,30 fNH3,28 - fNH3,29 - fNH3,30 = 0 fNH3,30= 0.05 fNH3,28 0.05 fNH3,28 - fNH3,30 =0

H2: fH2,28 = fH2,29 + fH2,30 fH2,28 - fH2,29 - fH2,30 = 0 fH2,30 = 0.05fH2,28 fH2,30 - 0.05fH2,28 = 0

N2: fN2,28 = fN2,29 + fN2,30 fN2,28 - fN2,29 - fN2,30 = 0 fN2,30 = 0.05 fN2,28 0.05 fN2,28 - fN2,30 = 0

Ar: fAr,28 = fAr,29 + fAr,30 fAr,28 - fAr,29 - fAr,30 = 0 0.05fAr,28 = fAr,30 0.05fAr,28 - fAr,30 = 0

CH4: fCH4,28 = fCH4,29 + fCH4,30 fCH4,28 - fCH4,29 - fCH4,30 = 0

Components 29 30-purge gas to

ARU mole % kmol/hr mole % kmol/hr

NH3 9.22% 3162.066 9.22% 166.425 H2 63.45% 21770.822 63.45% 1145.833 N2 21.15% 7256.941 21.15% 381.944 Ar 1.52% 523.087 1.52% 27.531

CH4 4.66% 1599.550 4.66% 84.187 Total 100.00% 34312.465 100.00% 1805.919

1531

52_6 ~_~ ~~ 532

46

0.05fCH4,28 = fCH4,30 fCH4,30 0.05 fCH4,28 = 0 4.18 Ammonia Recovery Unit - ARU

Ammonia Recovery Unit

S 30

S 31

S 33

Equations

NH3: fNH3,30 = f31 fNH3,30 - f31 = 0 H2: fH2,30 = fH2,33 fH2,30 fH2,33 = 0 N2: fN2,30 = fN2,33 fN2,30 - fN2,33 = 0 Ar: fAr,30 = fAr,33 fAr,30 - fAr,33 = 0 CH4: fCH4,30 = fCH4,33 fCH4,30 - fCH4,33 = 0 4.19 Recovered NH3 Mixing Point

Equations

NH3: f26 + f31 = f32 f26 + f31 - f32 = 0

Components 31-recovered NH3 33-ARU outlet

mole % kmol/hr mole % kmol/hr NH3 100.00% 166.425 0.00% 0.000 H2 0.00% 0.000 69.89% 1145.833 N2 0.00% 0.000 23.30% 381.944 Ar 0.00% 0.000 1.68% 27.531

CH4 0.00% 0.000 5.13% 84.187 Total 100.00% 166.425 100.00% 1639.495

Components 32-product NH3

mole % kmol/hr NH3 100.00% 2696.077 Total 100.00% 2696.077

47

4.20 Hydrogen Recovery Unit - HRU

Hydrogen Recovery Unit

S 34

S 33 S 35

Equations

H2: fH2,33 = 0.9328f34 + fH2,35 fH2,33 - 0.9328f34 - fH2,35 = 0 As 95% of H2 is recovered

0.9328 f34 = 0.95 fH2,33 0.9328f34 - 0.95 fH2,33 = 0 N2: fN2,33 = 0.411f34 + fN2,35 fN2,33 - 0.411f34 - fN2,35 = 0 Ar: fAr,33 = 0.0111f34 + fAr,35 fAr,33 - 0.0111f34 - fAr,35 = 0 CH4: fCH4,33 = 0. 0122 f34 + fCH4,35 fCH4,33 - 0.0122f34 fCH4,35 = 0

Components 34-recoverd H2

(permeate) 35-boiler fuel gas

(raffinate) mole % kmol/hr mole % kmol/hr

H2 93.28% 1088.541 12.02% 57.292 N2 4.11% 47.962 70.05% 333.982 Ar 1.11% 12.953 3.06% 14.578

CH4 1.22% 13.237 14.88% 70.950 Total 100.00% 1166.961 100.00% 476.802

48

Calculated by using MS Excel.

Formulae =

= Where = + + 2 + 3

= [ ( ) + 2

(2 2) + 3

(3 3) + 4

(4 4) ] = ( ) + 2 (2 2) + 3 (3 3) + 4 (4 4)

Reference/Base Conditions

Temperature = 298.15 K Pressure = 1 atm

Cp Constants

Components a B c d NH3 27.568 2.56E-02 9.91E-06 -6.69E-09 H2 29.11 -1.92E-03 4.00E-06 -8.70E-10 N2 28.9 -1.57E-03 8.08E-06 -2.87E-09 O2 25.48 1.52E-02 -7.16E-06 1.31E-09 CO 28.16 1.68E-03 5.37E-06 -2.22E-09 CO2 22.26 5.98E-02 -3.50E-05 7.47E-09 Ar 28.11 1.97E-03 4.80E-06 -1.97E-09

CH4 19.89 5.02E-02 1.27E-05 -1.10E-08 C2H6 6.9 1.73E-01 -6.41E-05 7.29E-09 H2O 32.24 1.92E-03 1.06E-05 -3.60E-09

5 ENERGY BALANCE

49

Calculations

Following table was formed in Excel using above mentioned formula and cp constants in order to calculate the enthalpy of NG feed stream (S5). Enthalpies of other streams were calculated by same method.

x (mole fraction) n (kmol/hr)

Tr (K) T (K) a B C d

Cp (kJ/ kmol K) x*Cp H (kJ/h)

NH3 0.000 0.000 298.15 308.15 27.568 2.56E-02 9.91E-06 -6.69E-09 36.062 0.000 0.00 H2 0.000 0.000 298.15 308.15 29.11 -1.92E-03 4.00E-06 -8.70E-10 28.873 0.000 0.00 N2 0.180 315.281 298.15 308.15 28.9 -1.57E-03 8.08E-06 -2.87E-09 29.086 5.236 91703.83 O2 0.000 0.000 298.15 308.15 25.48 1.52E-02 -7.16E-06 1.31E-09 29.467 0.000 0.00 CO 0.000 0.000 298.15 308.15 28.16 1.68E-03 5.37E-06 -2.22E-09 29.100 0.000 0.00 CO2 0.080 140.125 298.15 308.15 22.26 5.98E-02 -3.50E-05 7.47E-09 37.382 2.991 52381.22 Ar 0.000 0.000 298.15 308.15 28.11 1.97E-03 4.80E-06 -1.97E-09 29.093 0.000 0.00

CH4 0.738 1292.651 298.15 308.15 19.89 5.02E-02 1.27E-05 -1.10E-08 35.980 26.553 465092.78 C2 H6 0.002 3.503 298.15 308.15 6.9 1.73E-01 -6.41E-05 7.29E-09 53.569 0.107 1876.59 H2 O 0.000 0.000 298.15 308.15 32.24 1.92E-03 1.06E-05 -3.60E-09 33.692 0.000 0.00 Total 1.000 1751.560 34.886 6.11E+05

Following calculations were done in Excel to find outlet temperature of HTSC. Similar calculations were done for LTSC, methanator and ammonia synthesis convertor.

As Hout = - Hrxn + Hin (kJ/h)

At given conditions, molar Hrxn = -4.00E+04 kJ/kmol

Therefore, Hrxn = -4.00E+04 kJ/kmol * 802.952 kmol/h = -3.21E+07 kJ/h

Hout = -(-3.21E+07) + 5.67E+07 = 8.88E+07 kJ/h

As, Hout = n cp (T - Tr)

For calculating cp , following table was formed in Excel.

x (mole

fraction) n

(kmol/hr) Tr (K) T (K) a B C d Cp (kJ/kmol

K) x*Cp

NH3 0.000 0.000 298.15 493.15 27.568 2.56E-02 9.91E-06 -6.69E-09 38.851 0.000 H2 0.459 4252.894 298.15 493.15 29.11 -1.92E-03 4.00E-06 -8.70E-10 28.934 13.272 N2 0.185 1719.636 298.15 493.15 28.9 -1.57E-03 8.08E-06 -2.87E-09 29.380 5.449 O2 0.000 0.000 298.15 493.15 25.48 1.52E-02 -7.16E-06 1.31E-09 30.437 0.000 CO 0.010 89.217 298.15 493.15 28.16 1.68E-03 5.37E-06 -2.22E-09 29.535 0.284 CO2 0.140 1301.828 298.15 493.15 22.26 5.98E-02 -3.50E-05 7.47E-09 40.823 5.732 Ar 0.002 15.763 298.15 493.15 28.11 1.97E-03 4.80E-06 -1.97E-09 29.526 0.050

CH4 0.006 51.764 298.15 493.15 19.89 5.02E-02 1.27E-05 -1.10E-08 41.071 0.229 C2 H6 0.000 0.000 298.15 493.15 6.9 1.73E-01 -6.41E-05 7.29E-09 65.476 0.000 H2 O 0.199 1840.461 298.15 493.15 32.24 1.92E-03 1.06E-05 -3.60E-09 34.450 6.838 Total 1.000 9271.565 31.856

50

Therefore, temperature at HTSC outlet, T = 598.83 K or 325.68 C

Results

Stream No. Temperature (C) Pressure

(kg/cm2 g) Flow rate (kmol/h)

Enthalpy (kJ/h)

5-NG feed 35 40 1751.560 6.11E+05 6-desulphurizer inlet 400 40 1872.910 2.95E+07 7-desulphurizer outlet 390 39 1872.910 2.86E+07 8-process steam 380 38 3512.982 4.39E+07 9-mixer outlet 370 37 5385.892 6.94E+07 9-primary reformer inlet 600 35 5385.892 4.96E+07 10-2ndry reformer inlet 810 32 6952.823 1.95E+08 11-compressed air 35 175 1755.462 5.12E+05 11-preheated compressed air 600 32 1755.462 3.09E+07 12-2ndry reformer outlet 945 31 9271.565 2.89E+08 12-HTSC inlet 220 30 9271.565 5.67E+07 13-HTSC outlet 325 29 9271.565 8.88E+07 13-LTSC inlet 210 29 9271.565 5.46E+07 14-LTSC outlet 220 28.5 9271.565 5.74E+07 14-KO drum inlet 75 28 9271.565 1.45E+07 16-CO2 absorber inlet 75 28 7590.931 1.17E+07 17-CO2 to urea plant 40 0.2 1370.456 7.71E+05 18-stripper outlet 50 26.6 6220.476 4.52E+06 18-methanator inlet 320 26.5 6220.476 5.39E+07 19-methanator outlet 343.5 26.2 6179.296 5.92E+07 19-KO drum inlet 40 26 6179.296 2.70E+06 21-KO drum outlet 40 26 6067.506 2.64E+06 22-recycle to desulphuriser 45 40 121.350 7.04E+04 24-fresh make up gas 42 13