Thesis Hydro Treating

Chapter #1 Introduction

Hydrotreating of Naphtha 1

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1.1 NAPHTHA

Naphtha is a petroleum fraction invariably consists of c6 to c10 hydrocarbons.

Naphtha is widely used in fertilizer plants and petrochemical industries as a feed stock. It

is a highly volatile product, manufactured from crude oil by direct atmospheric

distillation and by catalytic cracking of heavy residues. There are two types of Naphtha

marketed namely, High Aromatic Naphtha (HAN) and Low Aromatic Naphtha (LAN)

known as Naphtha (Petrochemical). Naphtha essentially consists of paraffin, naphthenic

and aromatic Hydrocarbons. The presence of Aromatic Hydrocarbons in Naphtha is very

critical especially when it is used in fertilizer plants. In fact, the design of a fertilizer plant

may entirely depend upon the composition of Naphtha available or a refinery has to

produce Naphtha according to the needs and specifications demanded by a fertilizer

plant. This is one of the reasons that IS Specifications for Naphtha has been withdrawn.

Naphtha is used as a fuel in fertilizer plant reformers where high temperatures are

required. It is also used as a fuel for steam generation in the plants where reforming is

done with the help of steam. Some gas turbines for power generation have also been

installed recently which will require Naphtha as fuel.

1.2 HYDROTREATING OF NAPHTHA

Hydrotreating processing is commonly used to remove Platforming catalyst poisons from

straight run or cracked naphthas prior to charging to the Platforming Process unit. It can

be seen that the primary function of the naphtha Hydrotreating Process can be

characterized as a ―Clean up‖ Operation. The catalyst used 1n the Naphtha Hydrotreating

Process 1s composed of an alumina base impregnated with compounds of cobalt or nickel

and molybdenum. The catalyst is insensitive to most poisons which affect



dehydrogenation reactions. A relatively high percentage of carbon on the catalyst does

not materially affect its sensitivity or selectivity. Volumetric recoveries of products

depend on the sulfur and olefin contents, but usually are 100% + 2%

The Naphtha Hydrotreating Process 1s a catalytic refining process employing a selected

catalyst and a hydrogen-rich gas stream to decompose organic sulfur, oxygen and

nitrogen compounds contained in hydrocarbon fractions. In addition, hydrotreating

removes organo-metallic compounds and saturates olefinic compounds.

Organo-metallic compounds, notably arsenic and lead compounds, are known to be

permanent poisons to platinum catalysts. "The complete removal of these materials by

Hydrotreating processing gives longer catalyst life in the Platforming unit.

Sulfur, above a critical level, is a temporary poison to Platforming catalysts and causes an

unfavorable change 1n the product distribution. Organic nitrogen is also a temporary

poison to Platforming catalyst. It is an extremely potent one, however, and relatively

small amounts of nitrogen compounds in the Platformer feed can cause large deactivation

effects, as well as the deposition of ammonium chloride salts in the Platforming unit cold

sections.

Oxygen compounds are detrimental to the operation of a Platformer. Any oxygen

compounds which are not removed in the hydrotreater will be converted to water 1n the

Platforming unit, thus affecting the water/ chloride balance of the Platforming catalyst.

Large amounts of olefins contribute to increase coking of the Platforming catalyst. Also,

olefins can polymerize at Platforming operating conditions which can result in exchanger

and reactor fouling.

The Naphtha Hydrotreating Process makes a major contribution to the ease of operation

and economy of Platforming. Much greater flexibility is afforded in choice of allowable

charge stocks to the Platforming unit. Because this unit protects the Platforming catalyst,

it is important to maintain consistently good operation in the Hydrotreating Unit.

In addition to treating naphtha for Platformer feed, there are uses for the UOP Naphtha

Hydrotreating Process in other areas. Naphthas produced from thermal cracking



processes, such as delayed coking and visbreaking, are usually high in olefinic content

and other contaminants, and may not be stable in storage. These naphthas may be

hydrotreated to stabilize the olefins and to remove organic or metallic contaminants, thus

providing a marketable product.

1.3 PROCESS SELECTION

History:

Until the end of World War 2, there was little incentive for the oil industry to pay

significant attention to improving product quality by hydrogen treatment.

However, soon after the war the production of high sulphur crudes increased

significantly, which gave a more stringent demand on the product blending flexibility of

refineries, and the marketing specifications for the products became tighter, largely due to

environmental considerations.

Furthermore, the catalyst used in the Platforming process can only handle sulfur in the

very low ppm level, so hydrotreating of naphtha became a must. The necessity for

hydrotreating of middle distillates (kerosene/gas oil) originates from pressure to reduce

sulfur emissions into the environment. Overall, this situation resulted in an increased

necessity for high sulphur removal capability in many refineries.

REFINING PROCESSES

Today's refinery is a complex combination of interdependent processes. These processes

can be divided into three basic categories:

a. Separation processes

The feed to these processes is separated into two or more components based on some

physical property, usually boiling point. These processes do not otherwise change the

feedstock. The most common separation process in the refinery is distillation.



b. Upgrading processes

These processes improve the quality of a material by using chemical reactions to remove

any compounds present in trace amounts that give the material the undesired quality.

Otherwise, the bulk properties of the feedstock are not changed. The most commonly

used upgrading processes for jet fuel are sweetening, hydrotreating, and clay treatment.

c. Conversion processes

These processes fundamentally change the molecular structure of the feedstock, usually

by "cracking" large molecules into small ones, for example, catalytic cracking and

hydrocracking

“Here we are concerned with upgrading processes for petroleum’’

UPGRADING PROCESSES

Sweetening processes remove a particular class of sulfur-containing compounds called

mercaptans from jet fuel. Mercaptans are undesirable because they are corrosive and also

because of their offensive odor.

Processes for merceptans removal:

Several processes have been developed to remove mercaptans by converting them to

disulfides. These disulfides are not corrosive and their odors are not as strong as the

mercaptans they replace. Sodium plumbite and copper chloride have been used as

catalysts for this conversion in the past.

Merox Process:

In recent years, the Merox (mercaptan oxidation) process, which uses a cobalt-

based catalyst, has almost completely replaced the older technologies.

Most of these chemical sweetening processes do not change the total sulfur

content of the fuel; they merely convert sulfur from one chemical form to another.

http://www.chevron.com/prodserv/fuels/bulletin/aviationfuel/B_about_hydrocarbons.shtm#mercaptans



Some versions of the Merox process extract the disulfides that are formed and

thus lower the total sulfur content

Hydrotreating Process:

The objective of the Hydrotreating processes is to remove sulfur as well as other

unwanted compounds, e.g. unsaturated hydrocarbons, nitrogen, oxygen, organo-metallic

compounds from refinery process streams. It is catalytic hydrogenation process with very

high efficiency, even some plants remove sulphur up to 0.2ppm.

A main representative reaction is shown as under.

PROCESSES FOR HYDROTREATING

For Hydrotreating, two basic processes are applied,

1. The liquid phase (or trickle flow) process for kerosene and heavier

straight-run and cracked distillates up to vacuum gas oil

2. Vapor phase process for light straight-run and cracked fractions.

Both processes use the same basic configuration: the feedstock is mixed with hydrogen-

rich make up gas and recycle gas.

The mixture is heated by heat exchange with reactor effluent and by a furnace and it

enters a reactor loaded with catalyst, in the reactor, the sulphur and nitrogen compounds

present in the feedstock are converted into hydrogen sulphide and ammonia respectively.

The olefins present are saturated with hydrogen to become di-olefins and part of the



aromatics will be hydrogenated. If all aromatics need to be hydrogenated, a higher

pressure is needed in the reactor compared to the conventional operating mode. The

reactor operates at temperatures in the range of 300-380 0C and at a pressure of 10-20

bars for naphtha and kero, as compared with 30-50 bar for gas oil, with excess hydrogen

supplied. The temperature should not exceed 380 0C, as above this temperature cracking

reactions can occur, which deteriorates the color of the final product.

The reaction products leave the reactor and, after having been cooled to a low

temperature, typically 40-50 0C, enter a liquid/gas separation stage.

The hydrogen-rich gas from the high pressure separation is recycled to combine with the

feedstock, and the low pressure off-gas stream rich in hydrogen Sulphide is sent to a gas-

treating unit, where hydrogen Sulphide is removed.

The clean gas is then suitable as fuel for the refinery furnaces. The liquid stream is the

product from hydrotreating. It is normally sent to a stripping column where H2S and other

undesirable components are removed, and finally, in cases where steam is used for

stripping, the product is sent to a vacuum drier for removal of water. Some refiners use a

salt dryer instead of a vacuum drier to remove the water.

The catalyst used is normally cobalt, molybdenum and nickel finely distributed on

alumina extrudates. It slowly becomes choked by coke and must be renewed at regular

intervals (typically 2-3 years). It can be regenerated (by burning off the coke) and reused

typically once or twice before the breakdown of the support's porous structure

unacceptably reduces its activity.



DIFFERENCE BETWEEN HYDROTREATING AND

HYDRODESULPHURIZATION:

A hyrotreater and a hydrodesulphuriser are basically the same process but

A hydrotreater termed is used for treating kerosene or lighter feedstock

A hydrodesulphuriser mainly refers to gas oil treating.

The hydrotreating process is used in every major refinery and is therefore also

termed as the work horse of the refinery as it is the hydrotreater unit that ensures several

significant product quality specifications.

In most countries the Diesel produced is hydrodesulphuriser before it‘s sold.

Sulphur specifications are getting more and more stringent. In Asia, countries such as

Thailand, Singapore and Hong Kong already have a 0.05%S specification and large

hydrodesulphurization units are required to meet such specs.

The by-products obtained from HDT/HDS are light ends formed from small amounts of

cracking and these products are used in the refinery fuel gas pool. The other main by-

product is Hydrogen Sulphide which is oxidized to sulphur and sold to the chemical

industry for further processing.

In combination with temperature, the pressure level (or rather the partial pressure

of hydrogen) generally determines the types of components that can be removed and also

determines the working life of the catalyst. At higher (part ial) pressures, the

desulphurization process is 'easier', however, the unit becomes more expensive for

instance due to larger compressors and heavier reactors. Also, at higher pressure, the

hydrogen consumption of the unit increases, which can be a significant cost factor for the

refinery. The minimum pressure required typically goes up with the required severity of

the unit, i.e. the heavier the feedstock, or the lower levels of sulphur in product required.



1.4 APPLICATIONS OF HYDROTREATING PROCESS:

A more recent development is the application of Hydrotreating for pretreatment of

feedstock for the catalytic cracking process. By utilization of a suitable hydrogenation-

promoting catalyst for conversion of aromatics and nitrogen in potential feedstock, and

selection of severe operating conditions, hydrogen is taken up by the aromatic molecules.

The increased hydrogen content of the feedstock obtained by this treatment leads to

significant conversion advantages in subsequent catalytic cracking, and higher yie ld of

light products can be achieved.

Hydrotreating can also be used for kerosene smoke point improvement (SPI). It closely

resembles the conventional Hydrotreating Process however an aromatic hydrogenation

catalyst consisting of noble metals on a special carrier is used. The reactor operates at

pressure range of 50-70 bar and temperatures of 260-320 0C. To restrict temperature rise

due to the highly exothermic aromatics conversion reactions, quench oil is applied

between the catalysts beds. The catalyst used is very sensitive to traces of sulphur and

nitrogen in the feedstock and therefore pretreatment is normally applied in a conventional

hydrotreater before kerosene is introduced into the SPI unit. The main objective of

Smoke Point Improvement is improvement in burning characteristics as the kerosene

aromatics are converted to naphthenes.

Hydrotreating is also used for production of feedstock for summarization unit from

paralysis gasoline (pygas) which is one of the byproducts of steam cracking of

hydrocarbon fractions such as naphtha and gas oil.

CONCLUSION

It is obvious from economical data of many commercial Plants that the fixed Capital

Investment on Merox sweetening Process is 90% less then Hydrotreating and the

operating Cost is almost 95% less then Hydrotreating, But the efficiency of Hydrotreating

Units are normally above 99% which cannot be achieved by Merox process, the feed



quality requirements of Platformer Section cannot be fulfilled by Merox Process. Further

more hydrotreating also removes many other impurities and saturated some olefins as

well. This is why; Hydrotreating Process is employed as feed preparation unit, where

ever Platformer Plant is to be installed.

Chapter #2 Description of Process Flow


CHAPTER 2

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2.1 PROCESS DESCRIPTION

A typical Naphtha Hydrotreating unit processing a straight run naphtha for Platfonner

feed will have a reactor section and a stripper 'section. In addition, some units have a

prefractionation section upstream of the reactor section.

A. Prefractionation Section

In some special applications, it is desirable to produce a narrow boiling range naphtha cut

for feed to a Platformer. An example of this would be an operation aimed at making

aromatics, where the end point of the feed to the Platformer is limited to about 160°C

(325°F) to concentrate aromatic precursors in the feed. A full boiling range naphtha cut

from the crude unit could be processed through a prefractionation section to accomplish

this task.

The prefractionation section typically consists of two fraction-action columns in series,

with the overhead of the second (rerun) column being the heart cut for processing in the

reactor section of the hydrotreater. The heart cut boiling range is controlled by the

amount of light naphtha taken overhead in the prefractionation column and by the amount

of heart cut taken overhead in the rerun column. For example, if a 38-204°C (100-400°F)

boiling range naphtha is charged to a prefractionation section, the overhead temperature

controller of the first column sets the amount of overhead product, and increasing the

overhead temperature will increase the endpoint and quantity of the overhead product.

This cut is what controls the initial boiling point of the heart cut.

The prefractionator column bottom is charged to the second (rerun) column, where the

desired product is taken overhead, again controlled by an overhead temperature

controller. Increasing the overhead temperature will increase the amount of material



taken overhead and will increase its endpoint. Thus, if a heart cut of 82-174°C (180-

345°F) is desired, it can be obtained • by adjusting the prefractionation column overhead/

temperature to set the initial boiling point, and the rerun column overhead temperature to

set the endpoint.

Usually, the feed to the prefractionator will be heat exchanged with rerun column

bottoms, and a steam heater can be used to provide the remaining heat that is required.

The prefractionator bottom is normally pumped directly to the rerun column without any

reheat. Both columns have reboilers to provide the heat necessary for vaporization of

naphtha so that sufficient reflux can be maintained. The overhead product from the

prefractionator and the rerun bottoms product are sent to storage for blending or further

processing downstream units. A typical prefractionation flow scheme 1s depicted in

Figure IV-1.

B. Reactor Section

Naphtha feed can enter the unit either from intermediate storage or from another process

unit. In the case of feed from storage, the tank must be properly gas blanketed to prevent

oxygen from being dissolved 1n the naphtha. Even trace quantities of oxygen and/or

olefin in the feed can cause polymerization of olefins in the storage tank when stored for

long periods or in the combined feed/reactor effluent exchangers if the feed is not

prestripped. 'This results in fouling and a loss of heat transfer efficiency.

Naphtha feed from the charge pump combines with a-hydrogen-rich gas stream, and this

combined feed enters the combined feed/ reactor effluent exchangers, where the feed is

heated and the reactor effluent is cooled. The combined feed leaving the exchanger 1s all

vapor, and flows to the ch_aj2ge_jTe^tej^here it is heated to the required reaction

temperature. The amount of fuel burned in the heater is controlled by the temperature of

the combined feed leaving the heater and flowing to the reactor. Most reactors are

designed for down flow operation, and contain/ sufficient catalyst to remove

contaminants to the level required.

The reactor effluent flows through the combined feed/reactor effluent exchanger, usually



on the tube side, and then to the product condenser. A water wash injection point is

provided in the reactor effluent line to the product condenser so that any salt buildup in

the line or condenser may be washed out. Reactor effluent flows out of the condenser at a

low enough temperature to ensure complete recovery of the naphtha and enters the

product separator. A mesh blanket coalescer is provided in the separator to ensure

complete separation of gas, hydrocarbon liquid, and water. The product separator is also

provided with a water boot to collect the water injected for salt removal. This water is

usually pressured to a sour water stripper for disposal.

There are alternate methods for providing the required hydrogen-rich gas to the reactor.

Most common is a recycle gas compressor taking suction from the top of the product

separator with the discharge joining the naphtha feed upstream of the combined

feed/reactor effluent exchanger. Since the process consumes hydrogen, a hydrogen-rich

gas stream is brought into the unit as makeup just upstream of the product condenser.

This stream is controlled by the product separator pressure controller, allowing gas to

enter and hold a constant separator pressure. This flow scheme is depicted in Figure IV-2.

In some units, rather than having a recycle gas compressor, a comparable amount of a

hydrogen-rich gas stream is brought Into the unit on flow control, and flows on a once-

through basis through the reactor section to the product separator where it is vented ''on

pressure control. This flow scheme is depicted in Figure IV-3.

The choice between these two flow schemes is made during the design of each unit based

upon the availability of a high pressure hydrogen-rich gas stream, and the cost of

compression for each stream.

Stripping Section

The liquid hydrocarbon in the separator is pressured on level control through the stripper

feed/bottoms exchanger, and thus heated enters near the top of the stripper. A reboiler is

provided to supply the required heat input for generating vapor. This vapor strips

hydrogen sulfide, water, light hydrocarbons and dissolved hydrogen from the feed to the

stripper, which then passes overhead to the overhead condenser and to the overhead



receiver. Normally, no net overhead liquid product is produced, and all of the liquid in

the receiver is pumped back to the stripper as reflux. A reflux/feed ratio of approximately

0.25 is sufficient to strip the light ends and water from the tower. The re flux is pumped

into the stripper on receiver level control. To increase the amount of reflux, the reboiler

heat Input must be Increased to provide more overhead material. The net overhead gas

leaves the receiver on pressure control, usually to amine scrubb ing and then to fuel gas.

The stripper overhead system is equipped with inhibitor addition facilities to prevent

corrosion of the process lines and equipment by the hydrogen sulfide in the overhead

vapor. The corrosion inhibitor is pumped directly from a drum, diluted with a small'

slipstream of reflux, and injected directly into the overhead vapor line at the top of the

stripper.

The stripper bottoms material is pumped through the feed/bottoms exchanger and usually

is charged directly to the Plat forming unit. On many units, a small slipstream of stripper

bottoms is further cooled in a trim cooler and sent to storage for later use as sweet startup

naphtha. This flow scheme is depicted in Figure IV-4.

The dry, stripped naphtha hydrotreating unit product must meet the following

specifications to be acceptable as Plat former feed:

Total Sulfur, wt-ppm0.5 max.

Total Nitrogen, wt-ppm 0.5 max.

EP, °F 400 max.

*Lead, wt-ppb <20 max.

*Arsen1c, wt-ppb 1 max.

*Iron + Chloride, wt-ppm 1 max.

*Copper + Heavy Metals, wt-ppb <25 max.

Additionally, water plus total oxygen must be low enough to produce less than 5 mole



ppm water in the Platformer Recycle Gas with no water injection to that unit.

2.2 CHEMISTRY

As previously stated, the main purpose of the Naphtha Hydrotreating Process is to "clean-

up" a naphtha fraction so that it is suitable as charge to a Platforming unit. There are six

basic types of reactions that occur in the hydrotreating unit.

A. Reactions

1. Conversion of organic sulfur compounds to hydrogen sulfide

2. Conversion of organic nitrogen compounds to ammonia

3. Conversion of organic oxygen compounds to water

4. Saturation of olefins

5. Conversion of organic halides to hydrogen halides

6. Removal of organo-metallic compounds

B, Discussion

1. Sulfur Removal

For bimetallic Platforming catalysts, the feed naphtha must contain less than 0.5 weight

ppm sulfur to optimize the selectivity and stability characteristics of the catalyst. In

general, sulfur removal in the hydrotreating process is relatively easy, and for the best

operation of a Platformer, the hydrotreated naphtha sulfur content should be maintained

well below the 0.5 weight ppm maximum. Commercial operation at 0.2 weight ppm

sulfur or less in the hydrotreated naphtha is common.

Typical sulfur removal reactions are shown below.

a. (Mercaptan) C-C-C-C-C-C-SH + H2 ———> C-C-C-C-C-C +H2S



b. (Sulfide) C-C-C-S-C-C-C t 2H?—————> 2 C-C-C + H2S

c.(Disulfide) C-C-C-S-S-C-C-C + 3H2———> 2 C-C-C + 2

d.(Cyclic sulfide) C - C-C + 2\\2 ————> C-C-C-C-C + C C-C C

e. (Thiophenic) C — C-C + 4H2 —————> C-C-C-C-C + H2S

C C-C C

It is possible, however, to operate at too high a temperature for maximum sulfur removal.

Recombination of hydrogen sulfide with small amounts of olefins or olefin Intermediates

can then result, producing mercaptans in the product.

C-C-C-C = C-C + H2S——————> C-C-C-C-C -C-

If this reaction is occurring, the reactor temperature must be lowered. Generally,

operation at 315-340°C (600-645°F) reactor Inlet temperature will give acceptable rates

of the desired hydrogenation reactions and will not result in a significant amount of

olefin/hydrogen sulfide recombination. This temperature is dependent upon feedstock

composition, operating pressure, and LHSV.

2. Nitrogen Removal

Nitrogen removal is considerably more difficult than sulfur* removal in naphtha

hydrotreating. The rate of denitrification is only' about one-fifth the rate of

desulphurization. Most straight run naphtha contain much less nitrogen than sulfur, but

attention must be given to ensure that the feed naphtha to a bimetallic Platforming

catalyst contains a maximum of 0.5 weight ppm nitrogen and normally much less. Any

organic nitrogen that does enter the Platformer will react to ammonia and further with the

chloride in the recycle gas and form ammonium chloride. The ammonium chloride then

deposits in the recycle gas circuit or stabilizer overhead system. This problem can be very

annoying and time consuming, but it can be avoided or minimized by maximizing

nitrogen removal in the Naphtha Hydrotreating unit. Nitrogen removal is much more

important when a Naphtha Hydrotreating unit processes some cracked naphtha, since



these feed stocks normally contain much more nitrogen than a straight run naphtha. The

ammonia formed in the denitrification reactions, detailed below, is subsequently removed

in the hydrotreater reactor effluent wash water.

5. Halide Removal

Organic halides can be decomposed 1n the Naphtha Hydrotreating Unit to the

corresponding hydrogen halide, which is either absorbed In the reactor effluent water

wash or taken overhead in the stripper gas. Decomposition of organic halides is much

more difficult than desulphurization. Maximum organic halide" removal is thought to be

about 90 percent, but is much less at operating conditions set for sulfur and nitrogen

removal only. For this reason, periodic analysis of the hydrotreated naphtha for chloride

content should be made, since this chloride level must be used to set the proper

Platformer chloride injection rate. A typical organic chloride decomposition reaction is

shown below.

C-C-C-C-C-C-C1 H2 ———————> HC1 + C-C-C-C-C-C

6. Metal Removal

Most metallic impurities occur at the part per billion (ppb) levels in naphtha. The UOP

Hydrobon catalyst is capable of removing these materials at fairly high concentrations, up

to 5 weight ppm or more, on an intermittent basis at normal operating conditions. Most

metallic Impurities are permanently deposited on the catalyst when removed from the

naphtha. The catalyst loses activity for sulfur removal as higher metal loadings are

reached. Some commonly detected components found on used Hydrobon catalyst are

arsenic, iron, calcium, magnesium, phosphorous, lead, silicon, copper, and sodium.

Removal of metals from the feed normally occurs in plug flow with respect to the catalyst

bed. Iron is found concentrated at the top of catalyst beds as iron sulfides. Arsenic, even

though it is rarely found in excess of 1 weight ppb in straight run naphtha's, is of major

importance, because it is a potent Platinum poison. Arsenic levels of 3 weight percent and

higher have been detected on used Hydrobon catalysts that retain their activity for sulfur

removal / Contamination of storage facilities by leaded gasoline and reprocessing of



leaded gasoline in crude towers are the common sources of lead on used Hydrobon

catalysts. Sodium, calcium and magnesium are apparently due to contact of the feed with

salt water or additives. Improper use of additives to protect fract ionator‘s overhead

systems from corrosion or to control foaming account for the presence of phosphorus and

silicon.

Removal of metals is essentially complete above temperatures of 315°C (600°F) up to a

total metal loading of about 2-3 weight percent on the catalyst:. Above this level, the

catalyst begins approaching the equilibrium saturation level rapidly, and metal

breakthrough is likely to occur. In this regard, mechanical problems inside the reactor,

such as channeling, are especially bad since these results in a substantial overload on a

small portion of the catalyst in the reactor.

I.e Reaction Rates and Heats of Reaction

The approximate relative reaction rates for the three major reaction types are:

Desulphurization 100

Olefin Saturation 80

Denitrification 20

The approximate heats of reaction (1n kJ per kg of feed per cubic meter of hydrogen

consumed) and relative heats of reaction are:

Heat of Reaction Relative Heat of "Reaction

Desulphurization 8.1 1

Olefin Saturation 40.6 5

Denitrification 0.8 0.1

As can be seen from the above summary, desulphurization is the most rapid reaction

taking place, but it is the saturation of olefins which generates the greatest amount of

heat. Certainly, as the feed sulfur level increases, the heat of reaction also increases.



However, for most of the feedstock processed, the heat of reaction will just about balance

the reactor heat loss, such that the Naphtha Hydrotreating reactor inlet and o utlet

temperatures are essentially equal. Conversion of organic chlorides and oxygenated

compounds are about as difficult as denitrification. Consequently, more severe operating

conditions must be used when these compounds are present.

The following table summarizes the physical properties of UOP Hydrobon catalysts.

TABLE I

UOP HYDROBON CATALYSTS FOR NAPHTHA HYDROTREATING SERVICE

Designator S-6* S-9* S-12 S-15 S-16 /

Base Alumina Alumina Alumina Alumina Alumina

Form Sphere Sphere Extrudate Extrudate Extrudate

Size 1/16" 1/16" 1/16" 1/16" 1/16"

ABD

ABD

(lbs/ft3) 36

38

45

45

45

Lbs/Drum

250

275

300

325

300

Metals:

Ni

N1

Ni

Mo

Mo

Mo

Mo

Mo

Co

Co

Co

Regeneration:

Steam/

Steam/

Inert

Inert

Inert

A1r A1r Gas Gas Gas

*Also available In 1/8" spheres designated as S-6 (L) and S-9(L).



2.3 PROCESS VARIABLES

A. Reactor Pressure

The unit pressure is dependent on catalyst life required and feed stock properties. At

higher reactor pressures, the catalyst is generally effective for a longer time and reactions

are brought to a greater degree of completion. For straight run naphtha desulphurization,

20 to 35 kg/cm2g (300 to 500 psig) reactor pressure is normally used, although design

pressure can be higher if feed nitrogen and/or sulfur contents are higher than normal.

Cracked naphtha contain substantially more nitrogen and sulfur than straight run naphtha

and consequently require higher processing pressures, up to 55 kg/cm2g (800 psig).

Similarly, higher operating pressures are necessary to completely remove organic halides.

Halide contamination of naphtha is usually sporadic in occurrence and is normally due to

contamination by crude oil well operators.

The selection of the operating pressure is influenced to a degree by the hydrogen to feed

ratio set in the design, since both of these parameters determine the hydrogen partial

pressure in the reactor. The hydrogen partial pressure can be increased by operation at a

higher ratio of gas to feed at the reactor inlet. The extent of substitution is limited by

economic considerations.

Most units have been designed so that the desulphurization and denitrification reactions

go substantially to completion well below the design reactor temperature, for the design

feedstock. Small variations in pressure or hydrogen gas rate in the unit will not cause

changes great enough to be reflected by significant differences in product quality.

B. Temperature

Temperature has a significant effect in promoting hydrotreating reactions. Its effect,

however, is slightly different for each of the ' reactions that occur. Desulphurization

increases as/ temperature 1s raised. The desulphurization reaction begins to take place at

temperatures even as low as 230°C (450°F) "With the rate of reaction increasing

markedly with temperature. Above 340°C (650°F) there are only slight increases in



further removal of sulfur compounds due to temperature.

The decomposition of chloride compounds in low concentrations (< 10 weight ppm) will

require about the same temperature as the sulfur compounds decomposition.

Olefin saturation behaves somewhat similarly to the desulphurization reaction with

respect to temperature, except that olefin removal may level off at a somewhat higher

temperature. Because this reaction is very exothermic, the olefin content of, the feed must

be monitored and perhaps limited to keep reactor peak temperature within an acceptable

temperature range.

At very high temperatures, an apparent equilibrium condition limits the degree of olefin

saturation. This may even cause the residual olefins in the product to be greater at higher

temperatures than would be the case at lower operating temperatures. In certain cases,

when processing a naphtha with a significant amount of light ends over fresh catalyst, S

can react with these olefins to form mercaptans. In such a case, lowering the reactor

temperature can eliminate residual olefins and thus mercaptan formation.

Decomposition of oxygen and nitrogen compounds requires a somewhat higher

temperature than desulphurization or olefin saturation, and the removal of these

compounds does not appear to level .off in the same way at elevated temperatures. Units

with significant levels pf nitrogen or oxygen must be designed for high pressure and low

LHSV to ensure complete conversion.

The demetalization reactions are not very dependent on temp erature. Above 315°C

(600°F), metals removal is essentially complete. Below this temperature, there may be

some cases where all the metals will not be removed.

The .recommended minimum reactor inlet temperature to ensure a properly prepared

Platformer feed is 315°C (600°F). There are two factors which are important in

determining this minimum temperature; first, below the minimum temperature, reaction

rates for contaminant removal may be too low. Second, the temperature must be

maintained high enough to ensure that the combined feed (recycle or once-through gas

plus naphtha) to the charge heater is all vapor.



Normal reactor design temperatures for both straight run and c racked naphtha (SRN) are

399°C (750°F) maximum. "Actual operating temperatures will vary, depending upon the

feed type, from 285°C (550°F) to 285 °C (650°F). Cracked stocks may require processing

at higher temperatures because of the higher sulfur, nitrogen, and olefin contents. For

these feeds, the reactor delta T will be higher, 1n the range of 10-55°C (20-100°F).

As the catalyst ages, the product quality may degenerate, which may be corrected by

Increasing reactor inlet temperature. If increasing the temperature does not improve the

product quality, a regeneration or change of catalyst will be required, depending on the

history of the operation and catalyst state.

In addition to catalyst deterioration, scale and polymer formation at the top of the bed

may cause high reactor pressure drops which may result in reactor channeling. This may

be corrected by skimming the top of the catalyst bed; and/or unloading, screening and

reloading. High pressure drop problems should be/ corrected as soon as possible to

minimize the possibility of equipment damage and degradation of product quality

C. Feed Quality

For normal operation, daily changes in hydrotreater inlet temperature to accommodate

changes in feed quality should not be necessary. However, in some cases, such as when a

refinery is purchasing outside crude from widely different sources, the naphtha quality

may change significantly, and adjustment of reactor Inlet temperature may be necessary.

The final selection of reactor temperature should be based upon prod uct quality. The

above relations of feed quality and temperature assume operation within the normal

temperature operating ranges given 1n the preceding section.

D. Hydrogen to Hydrocarbon Ratio

The minimum hydrogen to feed ratio (nm3/m3 or SCFB) is dependent on hydrogen

consumption, feed characteristics, and desired product quality.

For straight run naphtha of moderate sulfur content, 40-75 nm3/m3 (250-400 SCFB) is

normally required. Cracked naphtha must be processed at higher H2 ratios [up to 500



nm3/m3 (3000 SCFB)]. As feedstock varies between these limits, the hydrogen to feed

ratio is proportioned between the extremes.

Ratios above 500 nm3/m3 (3000 SCFB) do not contribute to the rate of reactions. The use

of low purity hydrogen as makeup gas is limited by economical operation of the recycle

compressor. Recycle gas with hydrogen sulfide contents up to 10X and with large

quantities of carbon monoxide and nitrogen are not harmful/ to the catalyst, again when

reasonable desulphurization is the only criterion. For nitrogen removal or complete'

sulfur removal, high hydrogen purity (70X minimum) is necessary, and CO may act as a

temporary catalyst poison. The prevention of excessive carbon accumulation on the

catalyst requires maintenance of a minimum H2 partial pressure, so impurities present in

the makeup gas require higher operating pressures.

Lower hydrogen to hydrocarbon ratios can be compensated for by increasing reactor inlet

temperature. The approximate relation for these variables is 10°C (18°F) higher reactor

temperature .requirement for a halving of the hydrogen/feed ratio. This rule assumes

operation above the minimum values of 315°C (600°F) reactor inlet temperature and 40

nm3/M3 (250 SCFB) hydrogen ratio. This relation is approximate, and it should again be

pointed out that product quality should dictate the actual reactor temperature utilized.

E. Space Velocity

The quantity of catalyst per unit of feed will depend upon feedstock properties, operating

conditions, and product quality required. The liquid hourly space velocity (LHSV) is

defined as f ol1ows:

LHSV = volume of charge per hour volume of catalyst

With most charge stocks and product objectives, a simplified kinetic expression based on

sulfur and/or nitrogen removal determines the initial liquid hourly space velocity. This

initial value may be modified due to other considerations, such as size 'of unit, extended

first cycle catalyst service, abnormal levels of feed metals and requirements of other

processing units in the refinery flow scheme. Relative ease of conversion for Hydrobon®

catalysts indicate that olefins react most easily sulfur compounds next, then nitrogen and



oxygen compounds. There is considerable overlap with several reactions occurring

simultaneously and to different degrees. Charge stock variability is so large that only

approximate ranges of space velocities can be indicated for the various feed types. SRN

is processed at 4-12 LHSV and cracked naphtha at 2-8 LHSV.

For daily changes in the LHSV, inlet temperature on the Naphtha Hydrotreating reactor

may be adjusted according to the equation below:

T2 = Ti - 45 in LHSVi (for °F) LHSV2

or

T2 = TI - 25 in LHSVi (for °C) LHSV2

Where T^ = required inlet reactor temperature at LHSVi T2 = " " " " ―LHSV2

The above relation assumes operation between 4 and 12 LHSV and assumes that reactor

temperatures are within the limits discussed in Section II.

F. Catalyst Protection, Aging, and Poisons

The process variables employed affect the catalyst life by their effect on the rate of

carbon deposition on the catalyst. There is a moderate buildup of carbon on the catalyst

during the initial days of operation, but the rate of increase in carbon level soon drops to a

very low figure under normal processing conditions. This desirable control of the carbon-

forming reactions is obtained by maintaining the proper hydrogen to hydrocarbon ratio

and by keeping the catalyst temperature at the proper level.

Temperature is a minor factor in respect to the hydrotreating catalyst life. A higher

catalyst temperature increases somewhat the rate of the carbon-forming reactions, with

other factors being equal. It must be remembered that a combination of high catalyst

temperatures and inadequate hydrogen is very injurious to the catalyst activity.

Catalyst deactivation Is measured by the decrease in relative effectiveness of the catalyst

at fixed processing conditions after a period of catalyst use.



The primary causes of catalyst deactivation are: 1. accumula tion of coke on the active

sites, and 2. chemical combination of contaminants from the feedstock with the catalyst

components. In normal operation, a carbon level above 5 wt-% may be tolerated without

significant decrease in desulphurization although nitrogen removal ability may be

decreased.

Permanent loss of activity requiring catalyst replacement 1s usually caused by the gradual

accumulation of inorganic species picked up from the charge stock, makeup hydrogen or

effluent wash water. Examples of such contaminants are arsenic, lead, calcium, sodium,

silicon and phosphorus. Very low concentrations of these species, ppm and/or ppb, will

cause deactivation over a long period of service because buildup of deposits depends on

the integrated effect of both temperature and time. This effect is important in SRN

processing for Platformer feed.

Apparent catalyst deactivation may be caused by the accumulation of a deposit on top of

the catalyst bed. The flow pattern through the balance of the bed is disturbed and product

quality is diminished. This condition is easily remedied by skimming a portion of the

catalyst, screening and reloading, or replacing with fresh catalyst. The deposit is

generally iron sulfide,

Hydrobon® catalysts exhibit a high tolerance for metals such as arsenic and lead. Total

metals content as high as 2 to 3 wt-% of the catalyst have been observed with the catalyst

still effective. However, if the calculated metals content of the catalyst is 0.5 wt-%, the

frequency of product analyses should be increased to prevent metal breakthrough to the

Platforming catalyst. Organic lead compounds are decomposed by Hydrobon® catalysts

and for the most part deposit in the upper portion of the catalyst bed as lead sulfide.

Metals are not removed from the catalyst during regeneration. When the total metals

content of the catalyst starts to approach 1 to 2 wt-%, consideration should be given to

replacing the catalyst.

The only certain method of minimizing the effect of trace metal contaminants on the

catalyst is to limit their entry to the system. This is done by careful, conscientious feed

analysis and correcting the source of, or conditions, causing the presence of the metal



contaminant.

Dissolved oxygen, though not a catalyst poison, should be eliminated from the feed. With

oxygen in the feed, excessive fouling of equipment, particularly the feed-effluent

exchangers can occur.

Chapter #3 Material and Energy Balance


CCHHAAPPTTEERR 33

MMAATTEERRIIAALL AANNDD EENNEERRGGYY BBAALLAANNCCEE

3.1 MATERIAL BALANCE

Basis

4000 Barrel of Naphtha per stream day to be Hydrotreated

Bbl/hr Hydrotreated = 4000/24

= 166.66 bbl/hr

MW = 109.7

Density = 0.7424 kg/lit

So

Weight of Naphtha

= 166.66 bbl/hr 42/lbbb gallons 3.78 lt/1us gallon 0.7424kg/1lit

= 19539.18 kg/hr

= 42987 lbs/hr

For getting high %age desulphurization we take 400 SCF H2 Per bbl

(from literature)

So

Feed rate of H2 = 166.66 400

= 66664 scf/hr



Taking H2 purity = 0.712

(from analysis of fresh and recycle gas)

H2 lb moles = 66664/379

= 175.9 lb mol

= 79.95 kg moles

Total H2 Stream = 79.95/.712

= 247.05 lb

As avg M.W of H2 Stream = 10.02

So w.t. of H2 Stream = 2475.44 lb

= 1125.2 kg

Balance Around Heat Exchanger and furnace is same as same amount of combined feed

is entering and leaving that is

Wt of comb feed = 42987 + 2475

= 45462 lbs/hr

= 20664.5 kg/hr

BALANCE AROUND REACTOR

As first we know how to calculate the chemical hydrogen consumption in the reactor.

As there are a number of reactions going on in reactor so the scientist have developed a

formula for calculation of chemical hydrogen consumption that is as follows.

The general formula for the chemical hydrogen consumption applicable to all feed stocks

can be written as



(aS+bN+cB+E) % Wt H2 on feed where in:

S Sulphur content in feed minus. Sulphur contents in product

N Nitrogen contents as above

Br Bromine number of fed g/100gm

E extra consumption

a, b,c coefficient, depending on type of feed stokes

According to our feed specification

a = 0.12 (coefficient accounts for desulphurization)

b = 0.57 (coefficient accounts for denitrification)

c = 0 (coefficient accounts for olefin saturation)

E = 0.042 (coefficient accounts for extra consumption)

Wt% sulfur contents in feeds = 0.10335%

Wt% sulfur contents in removed = .10335 0.9995

= .1033

N2 contents in feed = 0.001%

Wt% N2 contents removed from feeds = 0.001 0.95

= 0.0095 % wt

Extra Cons. = 0.042

Putting values in formula

= 0.12 0.1033 + 0.57 0.0095 + 0 + 0.042

= 0.0598 wt % H2 on feed



So

H2 Consumption = 42987100

0598.0

= 25.71 lbs

= 11.68 kg/hr

Lb moles H2S formed = 32100

429871033.0

= 1.388lbs

= 0.63 kg mol/hr

The reactor effluents passé through H. exchanger & after this some amount of condensate

is added for removing salts coming with Naphtha and for dissolving some NH3 which is

formed in reactor.

Water added = 1750lbs

= 795.45 kg/hr

BALANCE AROUND SEPARATOR

In separator some gases streams are separated in gas phase and most of water added is

separated from boot (The remaining waters evaporated).

So

Combined reactor effluent in = 45462 lbs/hr

Water in = 1750 lbs

Water out from boot = 1736 lbs

Gaseous stream out = 1257 lbs



Naphtha out = 44219 lbs/hr

BALANCE FOR STRIPPING SECTION

In stripping section the remainder gases ion naphtha are removed and hydrotreated

product is obtained

SO

Naphtha in = 44219 lbs/hr

Gases out = 1275 lbs/hr

Pure Naphtha out = 44219-1275

= 42944 lbs/hr

= 19520 kg/hr

Detailed material balance is tabulated with Process flow sheet

3.2 ENERGY BALANCE

BALANCE AROUND EXCHANGER TRAIN E-110

Cold Side

Naphtha flow rate = 42987 lb/hr

= 19539.18 kg/hr

H2 stream flow rate = 1125.2 kg /hr

Naphtha + H2 Stream in at = 120 F = 49 C

Heat Capacity of liquid Naphtha from (120 F) 49 C to 248.0 (4708.4 F)

Cp = (0.388+0.00045T)/(SP.G)½



2

)7424.0(

47800045.0388.

)7424.0(

12000045.0388.2/12/1

Cp1 = 0.605 Kcal/kgm (Equation might subject an error 4% up to)

Heat Capacity of H2 Stream

Taking molar weighted heat capacity and neglecting. Molar heat capacity departure from

weighted value

Cp2 = 0.745 k cal/kg

Heat capacity of Naphtha vapor for the range

= (478.4 F) 248 C to (621 F) 327.2 C

Cpg = (4-Sp) (T+670)/(6450)

Cp =64502

)670621)(7424.4()6704.478)(7424.4(

= 0.616 kcal/kg

Now calculating heat loads

Heat Requirements for heating liquid Naphtha up to boiling point

= m Cp T

= 19539.18 .650 (248-49)

= 2352419.57 kcal/hr



Latent heat requirements for vaporizing Naphtha

= m

= 19539.18 98

= 1,914,839.64 kcal/hr

Hear requirements for superheating Naphtha Vapors up to 327.22 C

= m Cp T

= 19539.18 /616 (327.22-248)

= 953.502.60 kcal/hr

Heat Requirements for heating hydrogen gas stream from 49 C 327.22 C

= m Cp T

= 1125.2 .745 (327.22-49)

= 233,224.59 kcal/hr

Total heat load of exchanger train E1

= 2352419.57 + 1914839.64 + 953502.6 + 233244.59

= 5,453,986.4 kcal/hr

= 5.5 109 cal/hr

Hot side

Inlet stream at = 374 C

Overall Cp = 0.715

Mass flow rate in = 20664.2 kg/hr



Heats required to cool the reactor effluents up to condensation

Temp = mCp T

= 20664.2 0.715 0.715 (374-242)

= 1950287.19Kcal/hr

Latent heat of condensation = 19539.18 98

Neglecting pressure effects)

Change in = 1914839.64 kcal/hr

Heat required to lower the temp of liquid naphtha + gaseous stream up to 115 C.

For Naphtha = 19539.18 .615 (242-110)

= 1586176 kcal/hr

For H2 Stream = 1125.18 0.71 (242-110)

= 105453.74

Total load = 1950287.19 + 1914839 + 15861756 + 105453.74

= 5.5 109 cal/hr

BALANCE AROUND FURNACE E-120

Flow rate of Naphtha + H2 stream = 19539.18 + 1125.2

= 20664.38 kg/hr

Feed in at = 621 F

= 327.2 C

Over all Cp = 0.715



Out Temp = 696 F

= 368.8 C

Heat Requirements in F1 = m Cp T

= 20664.38 0.715 41.68 C

= 615 954.65 kcal/hr

BALANCE AROUND REACTOR R-130

Feed flow rate in = 20664.38 kg/hr

Feed inlet Temp = 368.8 C

H2 Consumed in desulphurization reactions =2

18.19539.

100

1033.012.

= 1.21 kg moles

Overall heat of reaction for H2 consumption in hydrosulphurization reaction

= 1100 kcal/kg mol

So

Heat evolved due to desulphurization = 1.21 11000

= 13310 kcal/hr

As Extra Consumption of H2 is assumed to be due to saturation of Aromatics

So

H2 Extra consumption = 2

18.19539

100

042.0

= 4.103 kg mol/hr



Heat evolved in Saturation of aromatics = 4.103 16000 kcal/kg mol

= 65651.64 kcal/hr

Total heat evolved = 65651.64 + 13310

= 78961.64 kcal/hr

Cp = 0.72 Kcal/kg

Temp rise in reactor

Q = mCp T

Q/mCp = T

T = 72.038.20664

64.78961

= 5.3 C

So reactor out let temp = 368.8 + 5.3

= 374 C

BALANCE AROUND AIR COOLED HEAT EXCHANGER E-140

Inlet Temp = 93.3 C

Outlet temp = 60 C

Heat load = mCp0.A T

( Subscript CpOA stands for Over all Heat capacity)

= 21460 0.62 (93.3-60)



= 443063 Kcal/hr

BALANCE AROUND TRIM COOLER E-150

Inlet temp = 60 C

Outlet Temp = 43.3 C

Heat load = mCp T

= 21460 0.618 (60-433)

= 221480 kcal/hr

BALANCE AROUND

COLUMN FEED/EFFLUENT HEAT EXCHANGER E-185

Inlet temp = 43.3 C

Outlet temp = 173.9 C


= 20098 (173.9-43.3) (0.626)

= 1643156.7 kcal/hr

BALANCE AROUND ACHE E-182

Inlet Temp = 151.66 C

Outlet Temp = 60 C

Heat required for condensation of vapours = m

= 4645 .95



= 441275 kcal/hr

Heat required to Fall the temps of combined mixture upto 60 C

= 5225 0.62 (151.66-60)

= 296932.5 Kcal/hr

Total load of ACHE

= 441275 + 296932.5

= 7.38207 105 kcal/hr

BALANCE AROUND TRIM COOLER E-183

Inlet temp = 60 C

Outlet temp = 43.3 C


= 5225 0.618 (60-43.3)

= 53925.13 kcal/hr



Chapter #4 Equipment Design


CCHHAAPPTTEERR 44

EEQQUUIIPPMMEENNTT DDEESSIIGGNN

4.1 SHELL AND TUBE HEAT EXCHANGER DESIGN

Introduction:

In the majority of chemical processes heat is either given out or absorbed, and fluids must

often be either heated or cooled in a wide range of plant such as furnaces, evaporators,

distillation units, dryers and reaction vessels.

The process of heat exchange between two fluids that are at a different temperature and

are separated by a solid wall occurs in many chemical engineering applications. And the

device used to implement this exchange is known as ‗heat exchanger’.

Definition:

The word ‗exchanger‘ really applies to all type of equipment in which heat is exchanged

but is often used specifically to denote equipment in which heat is exchanged between

two process fluids.

Such as:

Heaters And Coolers:

Exchangers in which a process fluid is heated or cooled by a plant service stream.

Vaporizer:

If the process stream is vaporized the exchanger is termed as a vaporizer.



Reboiler:

If the stream is essentially completely vaporized then the exchanger is a reboiler. It is

associated with a distillation column.

Evaporator:

For the purpose of concentration of a solution the exchanger is called as an evaporator.

Fired exchanger:

It is used for exchangers heated by combustion gases, such as boilers.

Modes Of Heat Transfer

Heat transfer will take place in one or more of three different ways:

Conduction:

In a solid, the flow of heat by conduction is the result of the transfer of vibrational energy

from one molecule to another and in fluids it occurs in addition as a result of the transfer

of kinetic energy. Heat transfer by conduction may also arise from the movement of free

electrons.

Convection:

Heat transfer by convection arises from the mixing of elements of fluid. It occurs as a

result of actual mixing of hotter part of the fluid with the colder part of fluid due to

density variation caused by temperature difference. There are two type of convection:

(a) Natural convection

when convective heat transfer is caused by temperature variation.

(b) Force convection

when convective heat transfer is caused by temperature variation and so me external



source for the mixing purpose.

Radiation:

All the bodies radiate thermal energy in the form of electromagnetic waves at a certain

temperature. These waves pass through vacuum and air and falls on a body then there are

three possibilities either they are:

Transmitted

Reflected

Absorbed

Depending upon the material upon which they fall. Only the absorbed radiations affect

the heat transfer.

In many of the applications of heat transfer in process plants, one or more of the

mechanisms of heat transfer may be involved. For example in the case of heat exchangers

heat passes through a series of different intervening layers before reaching the second

fluid.

TYPES OF HEAT EXCHANGERS:

Following are the ways of classification of heat exchangers:

(1) According to transfer process:

1. Direct contact type

2. Indirect contact type

(a) Direct transfer type

(b) Storage type



(2) According to surface compactness:

1. Compact (Surface density > = 700 m2/m3)

2. Non – compact (Surface density < 700 m2 /m3)

(3) According to construction:

1. Tubular

(a) Double pipe

(b) Shell And Tube

(c) Spiral Plate

2. Plate

(a) Gasketed

(b) Spiral

(c) Lamella

3. Extended Surface

(a) Plate-Fin

(b) Tube-Fin

4. Regenerative

(a) Rotary

(b) Fixed- matrix



(4) According to Flow arrangement:

1. Single Pass

(a) Parallel Flow

(b) Counter Flow

(c) Cross Flow

2. Multi Pass

(5) According to Number of Fluids:

1. Two-Fluid

2. Three-Fluid

3. n-Fluid(n >3)

(6) According to Heat transfer mechanism flow arrangement:

1. Single phase convection on both sides

2. Single phase convection on one side, two phase convection on other side

3. Two phase convection on both sides

4. Radiation heat combined transfer convection

Principal Types Used in Chemical Industry:

The principle types of heat exchanger used in the chemical process and allied industries

are as follows:

1. Double pipe exchangers

2. Shell and tube exchangers

3. Plate and frame exchangers

4. Plate- Fin exchangers

5. Spiral heat exchangers



6. Air cooled: coolers and condensers

7. Direct contact: cooling and quenching

8. Agitated Vessels

9. Fired Heaters

Selection of Heat Exchanger Type:

One of the more important actions taken by the design engineer in arriving at a

satisfactory solution for a specific heat exchange is the careful selection of the heat

exchanger type that should be used.

The selection process include a number of factors, all of which are related to the heat

transfer application. These are as:

1. Thermal requirement

2. Material Compatibility

3. Operational maintenance

4. Environmental, health, and safety considerations and regulations

5. Availability

6. Cost

In the chemical industry the preferred choice has been the shell and tube heat exchanger

due to the fact:

(1) These exchangers give a large surface area in a small volume

(2) Good mechanical layout

(3) Uses well-established fabrication techniques



(4) Can be constructed from a wide range of materials

(5) Easily Cleaned

(6) Well-established design procedures

(7) More than one heat exchanger can be used in a parallel or series arrangement to

meet special heat transfer or physical requirements.

(8) High thermal performance, even with fouled heat transfer fluids.

Shell and Tube Heat Exchanger

___________________________________________

Shell Side Inlet

Temp = t1=120.2 oF=49 o

C

Shell Side Outlet

Temp = t2 =219.2 oF=104 oC

Tube Side Inlet

Temp= T1 = 323.6 oF=162o

C

Tube Side Outlet

Temp=T2= 230 oF=110o

C



Shell Side: (cold) Tube Side: (hot)

Naphtha + H2 Reactor effluents

Temp of naphtha + H2 inlet = t1 = 49 oC = 120.2 oF

Temp of naphtha + H2 outlet = t2 = 104 oC = 219.2 oF

Total Pressure = 451 psia = 30.68 atm

Temp of reactor effluents inlet = T1 = 162oC = 323.6 oF

Temp of reactor effluents outlet = T2 = 110oC = 230 oF

Flow rate of naphtha stream entering = 42987 lb / hr

Flow rate of naphtha stream leaving = 42987 lb / hr

Flow rate of reactor feed and effluent = 2475 lb / hr

Designing Steps:

STEP 1

To Calculate Heat Duty – ‘Q’:

For Exchanger E-1:

For naphtha = mCp ΔT

= 42987 x 0.6 (219.2-120.2)

= 2553427.8 Btu / hr

For H2 Stream = mCp ΔT

= 2475 x 0.73 (219.2-120.2)

= 178868.25 Btu / hr

Total Heat Load (Q) = 2553427.8+ 178868.25



= 2.74 x 106 Btu / hr

STEP 2

Assumed Overall Coefficient:

U = 50 Btu /hr ft2 oF

STEP 3

Log Mean Temperature Difference:

Δ T1 – Δ T2 Δ Tl m =

ln ΔT1

Δ T2 (T1 – t2) - (T2 – t2)

= ln (T1 – t2)

(T2 – t1)

(323.6 – 219.2) - (230 – 120.2) =

ln (323.6 – 219.2) (230 – 120.2)

104.4 - 109.8

= ln (104.4)

(109.8) = 107 oF



For True Temperature Difference:

(T1 – t2) R = (T2 – t1)

323.6 – 230.0 = = 0.94

219.2 – 120.2 (t2 – t1) S =

(T1 – t1) 219.6 – 120.2

= = 0.48 323.6 – 120.2

(Using Figure 12.19 for 1-shell 2-tube pass)

Ft = 0.86

Δ Tm = Ft x Δ Tl m

= 0.86 x 107

= 92.02 oF

STEP 4

Provisional Area:

As

Q = UAΔ Tm

1

= (2.74 x 106) / (50 x 92.02)

= 496 ft2

= 55.1 m2



STEP 5

Heat Exchanger Specifications:

1 – 2 Pull Through Floating Head Type

Tubes

16BWG

Outside Dia 0D = ¾ inch = 0.75 inch = 0.0625 ft = 0.0208

Inside Dia 1D = 0.620 inch = 0.0516 ft (Table 10 Kern)

Length of Tubes L = 12 ft = 3.65 m

Square pitch = 1 inch = 0.0833 ft

Baffles:

25% cut horizontal segmental baffles

Area of single tube = (0.1623 x 12 x 30.48) / 100 = 0.1809m2

No. of tubes = 55.1 / 0.1809= 304.8

(From nearest count table 9)

No. of Tubes = 324

For 2 pass

Bundle Dia:

Db = do (Nt / K1)1/n1

(From Table 12.4 Coulson)

For Square Pitch:



K1 = 0.156

N1 = 2.291

Db = 0.0208 (324 / 0.156) 1/2.291

= 0.58 m = 1.9 ft

Shell:

(From Fig 12.10)

Shell-bundle clearance:

C = shell inside dia – bundle dia = 0.093 m = 0.31 ft

Shell dia inside = 0.093 + 0.58

(Ds) = 0.673 m = 2.21 ft

Baffle Spacing (B):

B =0.6 x shell dia

= 0.6 x 0.673

= 0.4 m = 1.31 ft

No of Baffles:

No. of Baffles = 3.65 / 0.4 = 9.1 9 Baffles

STEP 6

Physical Properties:

Average Temp Shell Side = 169 oF

Average Temp Tube Side = 276.8 oF



API Gravity = 59

k for H2 in feed stream = 0.122 Btu / lb ft2 oF

k H2 in leaving stream = 0.135 Btu / lb ft2 oF

k for naphtha in feed stream = 0.0858 Btu / lb ft2 oF

k for naphtha in leaving stream = 0.0845 Btu / lb ft2 oF

Heat capacity of naphtha in feed stream = 0.545 Btu / lb oF

Heat capacity of naphtha in leaving stream = 0.650 Btu / lb oF

Heat capacity of gases in feed stream = 0.9 Btu / lb oF

Heat capacity of gases in leaving stream = 1.025 Btu / lb oF

for H2 in feed = 0.0099 cP

for H2 in product = 0.011 cP

for naphtha in feed stream = 0.3 cP

for naphtha in leaving stream = 0.2 cP

Mean Properties: (Feed)

(i) Mean Heat Capacity

0.9314 x 2475 + 0.545 x 42987

C = 2475 +42987

= 0.61 Btu / lb oF

(ii) Mean Density

0.635 x 2475+ 46.325 x 42987

= 2475 +42987



= 43.958 lb / ft3

(iii) Mean Viscosity

0.0099 x 2475+ 0.3 x42987

= 2475+42987

= 0.285 cþ

= 0.6897 lb / ft hr.

(iv) Mean Thermal Conductivity

0.122 x 2347.88 + 0.0858 x 42987

k = 2347.88 +42987

= 0.087 Btu / lb ft2 oF

Mean Properties: (Product Stream)

(i) Mean Heat Capacity

0.0099 x 2475 + 0.3 x42944

= 2475 +42944

= 0.669 Btu / lb oF

(ii) Mean Density

1.74 x 2475+ 0.650 x 42944

= 2475 +42944

= 4309 lb / ft3



(iii) Mean Viscosity

0.011 x 2475 + 0.2 x 42944

=

2475 +42944

= 0.19 cP x 2.42

= 0.4598 lb / ft hr.

(iv) Mean (k)

0.027 Btu / lb ft2 oF

STEP 7

Over All Heat Transfer Coefficient:

Shell Side Calculations:

1. Flow Area :

ID x C x B as =

PT

ID = 0.673 m

= 0.673 x 3.281

= 2.2 ft.

C = 1 – 0.75

= 0.25 inch.

= 0.0208 ft.

B = 0.4 m



= 1.3124 ft

Pt = 1 inch = 0.0833 ft.

2.2 x 0.0208 x 1.3124

as = 0.0833

= 0.72 ft2

2. Equivalent Dia.

For square pitch:

1.27

De = (Pt2 – 0.785 do

2 ) do

do = tube outside dia = 0.625 ft.

Pt = 0.083 ft

De = 0.078 ft

3. Mass Velocity

Gs = Ws / as

= 45506.67 / 0.66

= 68949.5 lb/hr ft2

4. Reynolds No

Re.s = De Gs /

0.078 x 68949.5 =

0.6897

= 7797.68



5. jH Factor

From fig (28)

jH = 47

6. Prandtl No

Pr = (c / k) 1/3

= 1.6475

7. Outside Heat Transfer Coefficient

jH k (c / k) 1/3 ( / w)0.14

ho = De

= 47 x (0.087/0.078) x 1.6475 x 1

ho = 86.36 Btu/ hr. ft2 oF

Tube Side Calculations:

1. Flow Area:

Nt at

at = n

n = 2

Nt = 324

at = 0.3302 in2 (From Table 10)

= 0.0021 ft2



324 x 0.0021 at =

2

= 0.34 ft2

2. De = ID of tube

= 0.62 inches

= 0.0516 ft.

3. Mass Velocity

Gt = Wt / at

= 45506.67 / 0.34

= 133843.14 lb/hr ft2

4. Reynolds No

Re.t = De Gt /

0.0516 x 133843.14 =

0.4598

= 15020.2

5. jH Factor

jH = 70

6. Prandtl No

Pr = (c / k) 1/3

= 1.5228



7. Inside Heat Transfer Coefficient

jH k (c / k) 1/3 ( / w)0.14 hi =

De

= 70 x (0.087 / 0.0516) x 1.5228

= 179.7 Btu/ hr. ft2 oF

8. Inside to Outside Heat Transfer Coefficient

hio = hi x ID/OD

= 179.7 x (0.0516 / 0.0625)

= 148.36 Btu/ hr. ft2 oF

9. Clean Overall Coefficient:

hio x ho Uc= hio + ho

148.36 x 82.69 =

148.36 + 82.69 12267.8 =

231.05

Uc = 53.09 Btu/ hr. ft2 oF

10. Corrected Overall Coefficient:

Q

Ud = A x ΔTm

A = Nt a‖ L

= 324 x 0.1623 x 12

= 631.0 ft2



Corrected Area = 58.6 m2

2.74 x 106 Ud = 631.0 x 92.02

Corrected Overall coefficient = 47 Btu /hr. ft2 0F

Rd = 0.00243

11. Dirt Overall Coefficient:

1 1 Rd =

Ud Uc

Ud = 48 Btu /hr. ft2 0F

STEP 8

Pressure Drop:

Shell Side Pressure Drop: (Δ Ps)

For single phase flow:

f Gs 2 Ds (N + 1)

Δ Ps =

5.22 x 1010 De S s

Re.s = 7797.68

f = from fig (29)

= 0.0023

Ds = 0.642 m

= 2.11 ft.

N+1 = 10

De = 0.078 ft.



s = 0.7424

Gs = 68949.5 lb / hr ft2

0.0023 x (68949.5) 2 x 2.11 x 10

Δ Ps = = 0.076 Psi 5.22 x 1010 x 0.7424 x 0.078

For two phase flow: (Δ PTP)

Δ PTP = lo2 x Δ Plo

Where

lo2 = 1 + (Y2 – 1)[B xo

g (1 - xog) + xo

g2]

` xg = flow quality of liq phase

Or

mass flow rate of liq phase = total mass flow rate

42987 =

4.88

= 0.95

2 B =

Y + 1

B = 0.25

Y = 2/2.25 – 1

= 8-1 = 7

lo2 = 1 + (72 – 1) [0.25 x 0.95 (1 – 0.95) + 0.952]

= 49 [0.01187 + 0.9025]

= 44.8



Δ PTP = 44.8 x 0.076

= 3.42 Psi

Tube Side Pressure Drop: (Δ Pt)

For single phase flow:

f Gt 2 Ds Ln

Δ Pt =

5.22 x 1010 De s t

Re.t = 15020.2

f = from fig (26)

= 0.00025

L = 12 ft.

n = 2

De = 0.0516 ft.

s = 0.7424

Gt = 133843.14 lb/hr ft2

0.00025 x 133843.142 x 12 x 2 Δ Pt = 5.22 x 1010 x 0.0516 x 0.7424

= 0.053 Psi

4n v2 Δ Pr =

s 2g



From Fig (27)

V2 = 0.0023 x 46/144

2g

= 7.3 x 10-4

4 x 2 x 7.3 x 10-4

Δ Pr = 0.7424

= 7.8 x 10-3

Δ PT = Δ Pt + Δ Pr

Δ PT = 0.053 + 0.0078

= 0.0608 psi

For two phase flow:

lo2 = 1 + (Y2 – 1)[B xo

g (1 - xog) + xo

g2]

lo2 = 1+ ( 72 – 1 )[ 0.25 x 0.95( 1- 0.95 ) + 0.952]

= 49 [ 0.001187 +0.9025 ]

= 44.8

Δ PTP = lo2 x Δ PT

= 44.8 x 0.0608

= 2.7 psi



SPECIFICATION SHEET

Identification: Unit Shell And Tube Heat Exchanger

Item No. E-110-A

Type 1-2 Pull Through Floating Head

Function: To heat the reactor feed

Operation:

Heat Duty ‗Q‘ 690480 kcal/hr

Heat Transfer Area ‗A‘ 58.6 m2

Uc 301.5 W/ m2. K

Ud 272.55 W/ m2. K

Rd 0.00243

Shell Side Tube Side

Fluid Circulated Naphtha + H2 Naphtha + H2

Flow Rates 20664.5 kg / hr 20664.5 kg / hr

Temperature Inlet 49 oC

Outlet 104 oC

Inlet 162 oC

Outlet 110 oC

Pressure Drop 0.23 atm 0.18 atm

Material Of Construction Carbon Steel Carbon Steel

Specification I.D 0.67 m

C 0.094 m

B 0.4 m

I.D 0.015 m

O.D 0.019 m

Pt 0.0253 m

L 3.65 m

n 2

Nt 324



4.2 FURNACES

Introduction:

Furnace is a device for generating the control heat with the objective of performing work.

Definition:

A furnace is an enclosed place in which heat is produced by the combustion of fuel, as for

reducing ores or melting metals, for warming a house, for baking pottery, etc.‘ this is

drawn sufficiently wide terms to cover almost all heating operations. The range of

operation and the condition under which those processes must be carried out cover a very

wide field, and the types of furnaces are equally diverse; therefore, no attempt will be

made to describe or figure particular types of furnace.

Furnaces may operate over a range of temperature from 300 F or thereabouts, to upwards

of 3000 F and they may be intermittent or continuous in operation. The capacity may

vary from that of a small pot furnace used for the tempering steel springs, where a few

gallons of oil is heated to about 500 F to a blast furnace producing a thousand tons of pig

iron a day at a temperature of about 3000 F. The number of types are as great as the

number of heating operations, but most comprise a combustion chamber in which the fuel

is burned and a hearth (or its equivalent) on which the charge is heated.

The principle of fuel economy is the same in all furnaces, and involve

a) The complete combustion of fuel

b) The rejection of the products of combustion at the lowest practicable temperature

b) The reduction of external losses by means of suitable insulation

Types of Furnaces

However we have made an attempt to classify the furnaces so only the names are

indicated here, which also shows the purposes for which the furnaces are employed.



(1) Classification on the basis of fuel used

- Fuel heated furnaces

(a) Fuel and charge in contact

(1) Hearth furnaces

Without blast

With blast

(2) Shaft furnaces

Natural draught

Forced blast

(b) Charge heated by flame alone

(1) Reverberatory Furnace

Natural Draught

Forced Draught

(2) Rotary Furnace

(3) Sintering machines

(c) Closed Vessel Furnaces

Charge heated by conduction of heat through the walls of the vessel

(d) Charge containing its own fuel

Solid charge

Liquid Charge



- Electrical furnaces

(2)Classification on the basis of material heating

(a) Boiler furnace (To get steam from water)

Fire Tube Boiler

Water tube Boiler

Longitudinal Drum

Cross Drum Straight tubes

Cross Drum Bend tubes

(b) Refinery furnace (Crude oil cracking)

De Florez circular furnace

Box Type furnaces

Double Radiant section Box Type Furnace

Furnace with overhead convection bank

(c) Metallurgical furnaces

Metal Industries

The metallurgical furnaces are further classified on the basis of following purpose

For Tempering

For Annealing

For Carbonizing

For Forging



For Ensiling

(3)Classification on the basis of nature of transfer of heat

(a) Oven Furnaces

On the basis of method of firing Oven furnaces have four types

Direct Fired

Over Fired

Side Fired

Under Fired

(b) Muffle Furnace

(c) Liquid Bath Furnaces

(d) Recirculating type Furnace

(e) Radiant tube furnace

(4) Classification on the basis of material handling

(a) Batch furnaces

(b) Continuous furnaces

Fired Heaters

Introduction and operation:

Most of the fired heaters used in the petroleum refinery and petrochemical and other

chemical plants is a pipe still heater, which is designed to heat process fluid in tubes

effectively by burning fuel. The function of the heater is similar to that of the steam-

generating boiler except that usually process fluid is heated instead of water.



Basically a pipe still heater consists of a combustion chamber for heat release, surrounded

by tubes through which the process feedstock flows to absorb heat by both radiation and

convection. Using a predetermined air mix ratio, the heat is supplied by the gas or oil

burners provided on the floor or on the walls of combustion chamber.

The feedstock is fed into and passed through tubes inside the heater. If a convection

section is provided, the feedstock is fed to the convection section first, then introduced

into the radiant section. During passage of the feedstock through the fired heater, it is

subjected to both radiant and convection heat.

The inside walls of the heater are refractory lined, to cope with the high temperatures

generated by the firing fuel.

The feed stock is heated to the required temperature at the specified phase and fed to the

next unit in the process sequence; e.g., distillation column, fractionators or reactor etc. the

temperature of the feed stock when leaving the fired heater differs according to the

operating requirements, but is generally with in a range of 250oC to 500oC.

The fired heaters most generally used are the box type and the vertical cylinder type.

Types of Fired Heaters

Fired heaters are classified by their construction and purpose. There are basically two

types of construction, the box type and the vertical cylinder type. These are further

divided by their tube layout, combustion method, purpose and characteristics. Although

there are many types of construction to meet process requirements

Purpose of fired heaters

1) Heating

Raising the temperature of a liquid

Raising the temperature of a gas

Vaporizing a liquid



2) Thermal Cracking

Gas cracking

Liquid cracking

3) Thermal Reforming

Gas reforming

Heat Transfer

The purpose of a fired heater is to transfer heat to the process feedstock at a

predetermined temperature. This is accomplished by burning a fuel or gas, causing large

quantities of flue gas to enter the heater. The heat is transferred to the feedstock by

radiation, conduction and convection.

Heat Transfer in Fired Heaters

In the radiant section, the heat is transmitted to the tubes by heat radiation from the

burner flame and the heater walls. This heat is transferred to the feedstock by conduction

and convection. In the convection section, the heat is transmitted to the tubes by

convection of the hot flue gas and is then transferred to the feedstock by conduction and

convection Mechanical drafts are induced, forced or balanced. All the four draft methods

are used in fired heaters, although a natural draft is more generally applied.

Heater Components

Burners

The following types of burners are used for the combustion of oil or gas, or both in fired

heaters:

1) Premix gas burner

2) Non-Premix gas burner



3) Steam atomizing oil burner

4) Combination gas and oil burner

The burners are designed to produce a uniform flame suitable to the type of firebox

involved, together with the most efficient, safe and complete combustion of the fuel.

Refractory

The following kinds of refractories are used in fired heaters to protect the heater casing

(insulating materials) from hot flue gas:

1) Cast able (aluminum cement +aggregates)

2) Brick (fire bricks & insulation fire bricks)

3) Ceramic fiber (Al2O3 and SiO2)

Heating tube

Heating tube is a kind of container in which high temperature and high pressure process

feed stock is contained and receives the heat of combustion.

In some special heaters the tube metal temperature will be more then 8000C.

The material of heating tube is selected from among carbon steel, low alloy and high

alloy steel depending upon service temperature, corrosiveness of process feed stock and

others.

Generally the heating tube is classified into three types, that is, bare tube, finned tube and

studded tube.

Tube Support

The tube support is literally, the component that supports the tubes.

Tube supports are usually of high alloy casting.



Auxiliary Equipment

There are some auxiliary equipment for fired heaters to achieve higher heater efficiency

and to keep the fired heater in proper condition.

Air Preheater

The flue gas at the exit of the fired heater still contains some available heat which is high

enough to heat the combustion air.

Air Preheater is a kind of heat exchanger and is designed to exchange the heat between

flue gas and combustion air efficiently.

With preheated combustion air the fuel quantity required can be reduced, since preheated

air has more heat than ambient air.

When the flue gas is cooled too much some trouble may occur in the air preheater

elements, fan elements and the refractories in the duct or stack, since flue gas generally

contains sulfur compounds.

Soot Blower

When fuel oil is used as a fuel, a large amount of ash, carbon etc. will be generated and

accumulate onto the convection heating tubes, resulting in low heat transfer.

Furnace Calculation

Step 1

Partial pressure of CO2 and H2O (P) is found from graph which is plotted against excess

air.

Step 2

Emmisivity is found from the graph, it is plotted against PL (product of partial pressure

and flame length) different curves for different temperatures.



Step 3

Exchange factor finally is found by graph, where it is plotted against ratio of refractory

area to cold plane area, (AR/a ACP), different curves for different emmisivities.

What is meant by design of a furnace?

When we talk about furnace design it means we want to find

1. Size require for the given heat duty

2. Number of tubes require

3. Arrangement of tubes

4. Flue gas temperature

5. Amounts of fuel air steam

Methods for designing

There is no universally applicable method for the furnace design for all types of the

furnaces specially fuel used determent the design method applicable , there are four

known design methods

1. Method of Lobo and Even

2. Method of Wilson Lobo and Hattel

3. The Orrak- Hudson equation

4. Wohlenberg Simplified Method

Here we shall consider only method of lobo and evans. This is a trial and error method

which make use of the overall exchange factor ( ) and a Stefan-Boltzmann type equation.

It has a good theoretical basis and is used extensively in refinery furnace design work. It

is also recommended for oil or gas fired boilers.



As in all trial and error solutions, a starting point must be assumed and checked. For

orientation purposes, we shall estimate the number of tubes required in the radiant section

by assuming an average flux ( permissible average radiant rate ) Btu/(hr)(ft2 of

circumferential tube area).

Here we have taken vertical tube cylindrical furnace

Furnace Design Calculation By Method Of Lobo And Evvans

1-Average radiant heat flux:

First of all we shall assume radiant heat flux. In literature, permissible average

radiant rate for different types of feedstocks are available.

From table 19.2 (Kern)

For naphtha hydro treating charge heater

Average radiant heat flux= 3000 Btu/(hr)(ft2 of circumferential tube area)

2-Find Q/ Acp:

=2 x Average flux

=20x3000

=6000 Btu/(hr)(ft2 of circumferential tube area)

Where,

Acp= Equivalent cold plane surface (ft2)

= effectiveness factor

Acp=effective cold plane surface (ft2)

Q= heat transferred to cold surface (Btu/hr)



3- Overall exchange factor ( ) :

Assume overall exchange factor. Normally it is in the range of 0.55 to 0.65

Here lets take = 0.57

4- Actual heat transfer between hot and cold surfaces:

Q/ Acp = 6000/0.6

= 10000(Btu/ hr. ft2)

5- Tube surface temperature (Ts):

It is fixed depending upon the desired temperature of fluid in tubes.

Lets, Ts=800 F

6- Evaluate temperature of the gases leaving the radiant section:

From fig. 19.14 (Kern)

Tg= 1140 F

Or by substituting Q/ Acp and Ts in the following heat transfer equation

Q/ Acp =0.173[(Tg/100) 4-(Ts/100) 4]+7(Tg-Ts)

Where, all the terms have usual meaning as described.

7- Heat balance:

Heat balance is necessary for the solution of heat absorption problem. The heat balance is

as follows:

Q = Qf + Qa + Qr + Qs – Qw – Qg

Where,



Q = total radiant section duty , (Btu/hr)

Qa =Sensible heat above 60F in combustion air, (Btu/hr)

Qf = Heat liberated by fuel, (Btu/hr)

Qr =sensible heat above 60F in recirculated flue gases, (Btu/hr)

Qs =Sensible heat above 60F in steam used for oil atomization, (Btu/hr)

Qw =Heat loss through furnace walls, (Btu/hr)

Qg =Heat leaving the furnace radiant section in the flue gases, (Btu/hr)

A- Total required heat duty (Q):

Q =2.3x 106 Btu/ hr (from overall energy balance)

B- Efficiency of furnace ( ):

Suppose the overall efficiency of furnace = = 0.70

C- Heat liberated by fuel (Qf):

Qf=Q/ = (2.3x106)/ 0.70

=3.28x106 Btu/hr.

D- Lower heating value of fuel (L.H.V):

We have taken refinery gases as fuel which are obtained during distillation , cracking and

other processing of petroleum and its fractions which contain paraffins (e.g. methane

ethane , propane and butane) olefins (e.g. ethylene, propene and butene) and hydrogen

are called refinery gases.



L.H.V of refinery gases = 20500 kcal/ Nm3

=20500x(1/0.252)x(22.4/1)x(1/2.2)

=828282.82 Btu / lb.mol.

E- Amount of fuel consumed (qf):

qf =Qf/L.H.V

=3.28x106/828282.82

=3.96 lb.mols. / hr.

as the composition of refinery gas is:

COMPONENTS COMPOSITION % LBS./LB.MOLS.

Propane 40 =0.4 x 44 17.6

Butane 30 =0.3 x 58 17.4

Ethane 10 =0.1 x 30 0.3

Methane 10 =0.1 x 16 1.6

Hydrogen 10 =0.1 x 2 0.2

=39.8

Lbs of fuel gas = qf =3.96x39.8 =157.6 lbs./ hr.

F- Sensible heat in combustion air (Qa):

(a)- evaluate lb. Air/ lb. Fuel for 20 %excess air

lb. Air / lb. Fuel = 20.67 ( from table 18-10 ( nelson))

(b)- evaluate air required (qa):

qa = qf x (lb.of air / lb.of fuel)



=157.6x20.67 = 3263.43 lbs. / hr.

(c)- air enter at ambient temperature = 77 F

enthalpy of air at this temperature = Ha = 10.78 Btu/lbs.

(d)- Qa = qa x Ha

= 3263.43 x10.78

= 35179.8 Btu/ hr.

G- heat loss through wall ( Qw) :

Qw= 2% of Qf

=0.02 x3.28x 106

= 65714.28 Btu/ hr.

H- Sensible heat in steam (Qs):

Since it is a gas fuel , no steam is required for atomization , so

qs. = 0 lb.mols. / hr.

Qs = 0 lb.mols. / hr.

I - Heat in the flue gases (Qg):

Qg = Q(N2) +Q(O2) + Q(CO2) +Q(H2O)

a- Mass flow rate of the flue gases

(qg) = qf +qa +qs

= 157.6 + 3263.43 + 0

= 3421.23 lbs.hr.



b- basis : 1 lb.mol. of fuel gases:

O2 Required

(lb.mols.)

CO2 Produced (lb.mols) H2O Produced

(lb.mols.)

H2O + 0.5O2 H2O 0.05 0.1

CH4+2 O2 2 H2O + CO2 0.2 0.1 0.2

C2H6+3.5 O2 2 H2O + CO2 0.35 0.2 0.3

C3H8+5 O2 3 CO2+4 H2O 2 1.2 1.6

C4H10+6.5 O2 4 CO2+5H2O 1.95 1.2 1.5

=4.55 =2.7 =3.7

As, Fuel gas required (qf) =3.9657 lb.mols.

O2 required = 4.55 (lb O2/ lb.mol. Fuel gas) x 3.965 (lb.mol fuel gas)

=18.044 lb.mols.

with 20% excess air O2 fed = 18.044x1.2

= 21.65 lb.mols.

N2 entered = (21.65 x 0.79)/ 0.21 = 81.45 lb.mols. = N2 leaving

O2 consumed = 18.044 lb.mols.

O2 unconverted = 21.65- 18.044 = 3.606 lb.mols. = O2 leaving

CO2 leaving = 2.7 x 3.9657 =10.707 lb.mols.

H2O leaving = 3.7x3.9657 = 14.67 lb.mols.



b- enthalpies of flue gases as flue gas temperature :

COMPONENT OF FLUE GAS ENTHALPY OF COMPONENT AT Tg (Btu/ lb.mol.)

N2 8004

O2 8427

CO2 12200

H2O 9602

So,

Q (N2) = 81.4x8004 = 651925.8 Btu/ hr.

Q (O2) = 3.606x8427 =8430.606 Btu/ hr.

Q (CO2) = 10.207x12200 =124525.4 Btu/ hr.

Q (H2O) = 14.67x9602=140861.34Btu/ hr.

Therefore,

Qg = Q(N2) +Q(O2) + Q(CO2) +Q(H2O)

= 925743.146 Btu/ hr.

So overall heat balance is:


= 3.28 x 106 + 35179.8 + 0+0 - 65714.28 –925743.146

= 2.3 x 106 Btu/ hr.



8- Establish the number and Sizes of tubes:

Fix tube length = l = 22 ft

Fix outer diameter of tubes = 3.5 inch. = 3.5/12 ft.

Area of tubes = xDx l = x (3.5/12)x 22 = 20.158 ft.2

Heat transferred per tube = Average flux x surface area per tube

=3000x20.158

= 60444 Btu/ hr.

Number of tubes = total radiant section duty (Q)/ heat transferred per tube

= 2.3 x106 / 60444

= 38 tubes

9- Arrangement of the tubes:

Height of furnace = 23.5 ft

Center to center distance = (3.5+ 4) /12 = 0.625 ft

Tubes are vertically mounted in a single row along the wall of the cylindrical furnace

about one tube diameter away from wall.

Diameter of furnace (D) =(number of tubes x center to center distance)/

= (38x0.625)/ 3.14

= 8 ft.



Checking the performance of furnace

10- Evaluate effectiveness factor ( ):

For C-C distance / O.D = 2.3

& Arrangement of tubes is single row when only one is present

From fig 19.11 (kern)

= 0.79

11- Evaluate equivalent cold plan surface area (Acp):

Acp = (number of tubes)x (length of each tube) x (C-C distance)

= 38 x 22 x 0.625

= 523.21 ft2

So,

Acp. = 0.79 x 523.21

= 413.33 ft2

12- Evaluate the total area of furnace surface (At):

For cylindrical furnace,

At = xD xH

=3.14x8x23.5

= 590.32 ft2



13- Evaluate effective refractory surface (Ar):

Ar = At - Acp

= 590.32- 413.33

= 176.99 ft2

14- Evaluate (Ar/ Acp):

Ar/ Acp = 179.99/ 413.33

= 0.43 (this would be used to evaluate

exchange factor)

15- evaluate mean beam length (L):

It depends on the dimensions of the furnace & found from any suitable formula from

table (19.1)

Here, for cylindrical furnace whose dimensions are like Dx2D

L = 1x diameter of furnace (D)

= 1x8

= 8 ft

16- evaluate gas emissivity ( g):

a- Evaluate partial pressure of (CO2+ H2O)

At 20% excess air

p(CO2 + H2O) = 0.24 atm. from fig. 1-7 (Evans)



b- Calculate pxL:

pL = p (CO2+ H2O) xL

= 0.24x8

=1.92 atm.ft.

c- At pL =1.92 atm.ft. & Tg = 1140 F

g = 0.44 from fig (1-8) (Evans)

17- Evaluate Overall exchange Factor ( ):

at g = 0.44 & Ar/ Acp = 0.43

= 0.55 (from fig. 19.15 Kern)

18- CHECK:

Check of gas temperature (Tg) required to effect assumed duty on

assumed surface)

a- Calculate Q/ Acp using the above calculated value of

= 0.55

Q = 2.3e6

Acp. = 413.33 ft2

Q/ Acp = 2.3e6 /(413.33x0.55)

= 10117.3 Btu/ hr. ft2

b- Evaluate Tg (actual) at calculated Q/ Acp & Ts



Tg = 1170 F

So, trial indicates that less duty than 2.2 million Btu/ hr. is performed since this duty

does not cool the fuel gases to 1140 F but to only 1170 F , while the flux corresponding

to this duty could be effected by a gas temperature of 1140F.

In short as assumed Tg (1140 F) is quite different from calculated Tg (1170 F), so we

have to repeat the calculations by assuming another value of Tg.

SECOND TRIAL:

Suppose Tg = 1160 in step 6 (basically we have supposed = 0.57 in step 3 so

that Tg will come out to be 1160F after performing step 6)

Qf, Qa, Qw, will remain the same, only Qg (heat taken away by the flue gases) changed.

enthalpies of flue gases as flue gas temperature :

COMPONENT OF FLUE GAS ENTHALPY OF COMPONENT AT Tg (Btu/ lb.mol.)

N2 8774

O2 9251

CO2 13470

H2O 10562

So,

Q (N2) = 81.4x8774 = 714203.6 Btu/ hr.

Q (O2) = 3.606x9251 =33359.1 Btu/ hr.

Q (CO2) = 10.207x13470 =137488.29 Btu/ hr.

Q (H2O) = 14.67x10562=154944.54 Btu/ hr.



Therefore,

Qg = Q(N2) +Q(O2) + Q(CO2) +Q(H2O)

= 1039995.53 Btu/ hr.

So overall heat balance is:


= 3.28x106 + 35179.8 + 0+0 - 65714.28 –1039995.53

= 2.20 x 106 Btu/ hr.

So number of tubes:

Number of tubes Nt = total radiant section heat duty / heat transferred per tube

=(2.0x106 )/60444

=36

(In the previous trial, no of tubes were 38. as there is very small change in the no. of

tubes so we use the previous value (i.e. 38 ) to avoid the repetition of calculations)

Similarly assuming that does not change (actually it will fall slightly)

i.e.

(Calculated) = 0.55

so,

Q/ Acp = 2.3e6 / (413.33x0.55)

= 101173.38 Btu/hr.ft.2



We have,

Tg = 1155F (calculated)

Where as assumed Tg was 1160 that is close enough.

Circumferential flux = Q / ( Dl) (Nt)

=(2.3x106 )/ (3.14x0.29x22x38)

= 3019 Btu/hr.ft2

As compared with the specified flux =3000 Btu/hr.ft2

Such a difference is negligible.

Final Results

Number of tubes = 38

Flue gas temperature = 1160F

Heat duty = 2300000 Btu / hr

Flux calculated 1‘ = 3019 Btu/ hr ft2

Dimension of furnace = 23.5 x 8

% Error = [(3000-3019) / (3000)]x 100 = 0.63 %



SPECIFICATION SHEET

Identification:

Unit Furnace Item No. E-120

Type Vertical Cylinderical

Function: To heat the reactor feed Operation:

Heat Duty ‗Q‘ 2.42*109 J/hr

Furnace Tube

Fluid Circulated Combustion Gases Naphtha + H2

Flow Rates 20664.5 kg / hr

Temperature Inlet 327 oC

Outlet 369 oC

Material Of Construction Carbon Steel Carbon Steel

Refractory material Refractory Brick

Specification

Diameter 2.44 m

Height 6.55 m

B 0.4 m

L.T 6.096 m

O.D 0.088 m

C-C.Dis 0.1905 m



4.3 REACTOR

Types of Reactors

The most common types of Reactors are

1. Fixed bed Reactor

2. Fluidized bed Rector

3. Stirrer tank Reactor

Fixed bed reactor can be further classified on the biases of either heat is supplied during

reaction or not.

o Adiabatic

o Non adiabatic

The reactions taking place within the reactor may be in gas phase or there might a case of

trickle operation.

For gas phase reactions some important reactor configurations are as under.

1. Single adiabatic bed

2. Radial flow

3. Adiabatic beds in series with intermediate cooling or heating

4. Direct- fired non-adiabatic

Except reactor type and configuration some other factors are important like , Distribution

system and Sporting ceramic balls which also serves for uniform distribution of flow as

well.

Our Reactor in this case is non isothermal adiabatic reactor with basket type distribution

system and standard ceramic balls installation . Detailed calculations of distribution

system is given in design calculations.



REACTOR DESIGN CALCULATIONS

Temp =374oC

Pressure = 29.12atm

Sulpher contents = .5ppm

Temp = 369oC

Pressure = 29.4atm

Sulpher Contents = 1033.5ppm

Feed rate = 20664Kg/hr



PLANT SPECIFICATIONS

Feed rate = 4000 bbl/std

Wt % S in Feed = .10335

Wt % S in Product = .00005

Operating Pressure = 430 Psi =28.4atm

Operating Temperature = 700 F = 374oC

Pressure Drop maximum allowable = 10 Psi

Catalyst

Each pellet (dimention 1/8 inch = .3175 cm)

Area = 2*.785*.31752 + .3175 .3175

= .474cm2

Dp = diameter of a sphere of same surface = (.474 )1/ 2 = .388cm

Diameter of a sphere of same volume 2 ( /4)(0.31752)(0.3175) 1/ 3

4 /3

= 0.388cm

Outer surface / unit volume solid = 0.474 /

( /4)(0.31743)

= 18.9 cm2 /cm3

Volume of solid /cu ft bed is (1-e) = 0.6 cu ft



Total Pellet surface in sq ft /cu ft bed = (0.6)(18.9)(30.48)

= 345 ft 2 / ft 3

However, some of surface is blocked where the pellets touch. A better figure is 310ft2

= 28.81m2

Per cu ft bed, obtained by graphical interpolation of the values given by Sherwood and

Pig ford, p. 87.

Pellet surface in sq ft/ cuft of bed

From test data

Kg = Gm/(aPl) pf dy/y = 45.9 ln(.10335/.00005)

(310)(30)(9)

= 0.004186

Feed rate = 4000bbl * 42gal * 3.76 lit * 0.7424 Kg * 2.2 lbs

24hr 1bbl 1 gal 1 lit 1Kg

= 42987 lbs/hr = 19539.5 kg/hr

Moler Feed rate = 42987/109.7 = 391.86 lb mols/hr = 178.1Kg mol/hr

H2 stream Feed rate = 245.7 lb mols /hr =111.68 Kg mol / hr

Mixed feed rate = 391.86 + 245.7 = 637.6 lb mols/hr

Letting S be crossectional area , ft2

L be bed height in ft

As Gm = G/S



So our design equation for calculating bed volume is

Gm p f dy/y = aPKg L

Putting values in above equation according to our conditions

We get

637.6 * ln 0.10335 = 310*29.25*0.004186 L

S 0.00005

Sl = Volume of catalyst bed = 128.23 ft3 =11.91m3

Know we have to calculate height and dia of catalyst bed suitable for our bed volume,

Which is decided on the base of pressure drop across the catalyst bed.

For using equation for pressure drop, which is in Perry chemical engineering handbook

on p. 393, we have to calculate Nre first

Nre = DpG

Nre = 0.364 G

(30.48)(0.0383)(2.42)

= 0.128G

Total average molar flow rate = 637.6

Average molecular weight = 4`2987+``2475

637.6

= 71.3



Gas density at 710 F = 2.31 lbs/cuft =37Kg/ m3

And equation for P

P/L = 2f U2 = 2fG2 = (2fG2) (30.48)

gcDp gc Dp (32.2)(36002)(2.31)(0.364)

= 0.0000001737fG2

G = 45462/S

F = 5Nre –1 + 0.4Nre – 0.1

G Nre 5/Nre 0.4/Nre F P/L S L P dia

1000 128.00 0.039063 0.246229 0.285 0.03 45.46 2.82 0.08 7.61

1500 192.00 0.026042 0.236445 0.262 0.06 30.31 4.23 0.25 6.21

2000 256.00 0.019531 0.22974 0.249 0.10 22.73 5.64 0.56 5.38

2500 320.00 0.015625 0.22467 0.24 0.15 18.18 7.05 1.06 4.81

3000 384.00 0.013021 0.220611 0.234 0.21 15.15 8.46 1.78 4.39

3500 448.00 0.011161 0.217236 0.228 0.28 12.99 9.87 2.76 4.07

4000 512.00 0.009766 0.214355 0.224 0.36 11.37 11.28 4.05 3.80

4500 576.00 0.008681 0.211845 0.221 0.45 10.10 12.69 5.67 3.59

5000 640.00 0.007813 0.209624 0.217 0.54 9.09 14.10 7.67 3.40

5500 704.00 0.007102 0.207636 0.215 0.65 8.27 15.51 10.08 3.24

6000 768.00 0.00651 0.205837 0.212 0.76 7.58 16.92 12.94 3.10



6500 832.00 0.00601 0.204196 0.21 0.89 6.99 18.33 16.28 2.98

7000 896.00 0.00558 0.202689 0.208 1.02 6.49 19.74 20.15 2.87

Reactor design Parameters

0.00

5.00

10.00

15.00

20.00

25.00

0 1000 2000 3000 4000 5000 6000 7000 8000

G (mass velocity)

Heig

ht,

Pre

ssu

re d

rop

, D

ia

Know choose a value of G and calculate Nre, f, P/L, S, L, and P

From above table we suggested value of

Dia of bed = 3.8ft = 1.158m

Height of bed = 11.28ft3.44m

G = 4000lb/ft2.hr = 19560Kg/(m2.hr)

P = 4.05 Psi =0.2755atm

Know we have to select the distribution system for the feed



According to stander procedure ceramic balls are located at both ends of catalyst bed.

Generally the balls used are of 3mm Dia, 6mmdia and 19mmdia.

For a reactor ID of 3.8ft from table V-4

Main inlet distribution baskets Dia = 2 ft = 0.61m

Small distribution baskets dipped in catalyst bed, Dia = 6 inches

= 0.1524m

Small baskets height = 4 ft = 1.219m

Number of small baskets = 7

According to the conventional procedure 60% of small baskets is dipped in catalyst bed,

So increase in catalyst bed height due to dipping of small baskets.

= 7x (0.52) x 4 x 0.6

4 x 13ft2

= 0.217 ft = .0661m

So

Corrected bed height = 11.28 + 0.217 = 11.5 ft

Bed height with ceramic ball = 11.5 +2 = 13.5 ft

Giving 20% vacant space on top and bottom.

Additional Reactor height = .4 x 11.5 = 4.6 ft

Total Reactor Height = 18.1ft = 5.51m

Reactor Dia = catalyst bed Dia = 3.8 ft = 1.158m



SPACIFICATION SHEET REACTOR-130

Identification: Unit Reactor

Item No. R-130 Type Fixed Bed, Catalytic, Adiabatic,

Nonisothermal

Function: Hydrotreating of Naphtha Type of Operation Continuous, Gas phase operation

Reactor Catalyst

Fluid Circulated Naphtha + H2 Stream Name S-7

Feed Rates 20664.5 kg / hr Composition

Temperature

Inlet 369 oC

Outlet 374 oC

P = 0.278atm

Sulphur Contents

Feed 1033.5ppm

Product: 0.5ppm

5% Cobalt Oxide

10% Nicle Oxide

20% Molybdenium oxide

On Silica / Alumina Sport

Material

Material Of Construction Killed Steel

Specification Distribution System

Bed volume 11.9m3

Diameter 1.16 m

Height 5.51 m

Main basket

Small Baskets

No

1

6

Dia

0.61m

0.152m

Height

0.914m

1.21m





4.4 AIR COOLED HEAT EXCHANGER

An ACHE is a device for rejecting heat from a fluid directly to ambient air. This is in

contrast to rejecting heat to water and then rejecting it to air, as with a shell and tube heat

exchanger.

Air cooler has many advantages over water cooling so there is a comparison between air

and water cooling

Air Versus Water Cooling

Air Water

1. Air is available free in

abundant quantity with no preparation

cost.

2. Mechanical design of an air

cooler is very much easy as the process

fluid is always on the tube side.

3. Cleaning and Maintenance

is easy in air coolers.

4. Non corrosive in nature.

1 Water is corrosive and require

treatment to control both scaling and

deposition of dirt.

2 Danger of process fluid

contamination is much greater.

3 Operating cost for water

cooler is high, because of higher

cooling water circulation pumps HP &

water treatment cost.



TYPES OF AIR COOLED HEAT EXCHANGER

1) Forced Draft

2) Induced Draft



An ACHE consists of the following components:

One or more bundles of heat transfer surface.

An air-moving device, such as a fan, blower, or stack.

Unless it is natural draft, a driver and power transmission to mechanically

rotate the fan or blower.

A plenum between the bundle or bundles and the air-moving device.

A support structure high enough to allow air to enter beneath the ACHE at a

reasonable rate.

Optional header and fan maintenance walkways with ladders to grade.

Optional louvers for process outlet temperature control.

Comparison of forced and induced draft Air cooled Heat exchangers

Forced Draft ACHE Induced Draft ACHE

Lower HP requirement if the

effluent air is hot.

Better accessibility for

maintenance.

Easy to work on fan

assembly.

Offered higher heat transfer

coefficient.

Better distribution of air.

The hoods offer protection

from weather.

More difficult to work on fan

assembly, due to heat from the bundle

and due to their location.



DESIGN CALCULATIONS OF AIR COOLER

Mass flow rate of naphtha = 42987lb/hr

Mass flow rate of water = 1750lb/hr

Total mass flow rate = 42987+2475+1750

= 47212.2 lb/hr

Feed Temperature:

Inlet temperature = 2000F

Outlet temperature = 1400F

Q = mCpΔT

= 47212.2x0.74x(200-140)

= 2096221.68Btu/hr

Assuming,

U = 75 Btu/hr.ft2. 0F

Calculation of temperature difference:

T1 = Process fluid inlet temperature = 2000F

T2 = process fluid outlet temperature = 1400F

Ta1= Air inlet temperature = 900F

Fins Selection:

Circular fins of aluminum

Height of fins = 5/8in



Thickness = 0.017in

8fins/inch

Number of tube rows = 4

Ft2 bare tube area /ft2 face area = 5.04

Tube OD = 1 in

Air velocity employed = 650ft/min

For calculating face area required of bundle

Through factor calculation and then through graph

UAt / KVf = 75x5.04/1.08x650

= 0.5384

From graph

KVfAf/wCp = 1.8

So,

Af = 1.8x47212.2x0.74/1.08x650

= 90ft2

So a bundle of standard size

4x24 ft2 face area should be selected

A safety factor of 6.6% is provided

Air temperature rise for 90ft2 face area = Q/KAfVf

= 2096221/1.08x90x650



= 33.170F

Air outlet temperature from bundle = 90+33

= 1230F

ΔTlm = (200-123)-(140-90)/ln (200-123)/ (140-90)

= 62.54 0F

P = t2-t1/T1-t1

= 123-90/200-90

= 0.3

R = T1-T2/t2-T1

= 60/33.1

= 1.81

From graph we found the correction factor for log mean temperature difference

For two pass flow,

FT = 0.99

ΔTlm = 62.54x0.99

= 61.90F

As,

Q = U A LMTDc

Required surface area = 2096221/61.9x75

= 451.41 ft2



Surface area from 90 ft2 face,

Area = 90x5.04

= 453.1 ft2

This is close enough so, selection of U is right

Tube area of 4x24 bundle face for 4 row bundle,

Area = 96x5.04

= 483.84 ft2

Number of tubes /Row = Total bare area/No. of row x Length x Tube bare area /ft of tube

23 = 483.84/4x24x0.2260

Total no of tubes = 90 tubes

Two pass flow of tube side process stream

Air face velocity = 650 ft /min

Density of air =ρair = 0.073 lb/ft3

Gross free area = 12(2.375-1)

= 16.5 in2

Fin blockage = 12x8x2x0.017x0.625

= 2.041 in2

Net free area = 14.4 m2

Free area /face area = 14.46/2.375x12

= 0.508



Gm = 650x0.073x60/0.508

= 5604.3 lb/ft2.hr

From figure,

h = 9.8 Btu/hr ft2 0F

For fin efficiency for alluminium curve

(2h/kt) 1/2xL = (2x9.8/120x0.017x12)1/2 x0.625/12

= 0.56

Dfo/Dt = 2.25/1

= 2.25

From above two factors using graph

Fin Efficiency = 86%

Fin area = 12x8[(2.25)2-12]x2xπ/4 + πx2.25x0.017

= 624in2

= 4.33ft2

Tube bare area per length feet of tube

= πdo(1-nt) L

= 3.14x1x(1-8x.017)x12/144

= 0.2261 ft2

Ratio of fin area to tube area = 4.33/0.2261

= 19.15



Air side heat transfer coefficient

Based on outside diameter = 9.8(1915x0.86+1)

ho = 171. Btu/hr ft2 0F

Now calculating inside heat transfer coefficient ‘hi’

Cross sectional area = 4.5 x π x (0.0695)2/4

= 0.1707ft2

G = Mass flow rate /flow area

= 47212/0.1707

= 276578.8lb/ft2hr

Re = DG/µ

= 0.0695x276578.8/0.46

= 41787.44

hi = 0.023xk/dx(Re)0.8 x(cµ/k)1/3

=0.023 x 0.06/0.0695 x (41787.4)0.8 x (0.71x0.46/0.061)1/3

= 175.59 Btu/hr ft2 0F

For air cooled heat exchanger

1/U = 1/ho + (Do/2Kw) lnDo/Di + (1/hi) Do/Di + Re

= 1/171.19 + 0.083/2 x 30 ln0.083/0.0695 + 1/175 x 0.083/0.0695 + 0.0004

= 0.013308

U = 75.15



So from above it is proved that selection of U is right so the area selected is right too

Calculation Of Air Side Pressure Drop

ΔP = 18.93 (GmDr/µ)-0.316 (Pt/Dr)-0.927 (Pt/Pl)

0.515 (Gm2 n/gcρ)

Where,

Gm = mass velocity at minimum cross section through the rows of the tube

normal to the flow

Dr = root diameter of tube

gc = acceleration of gravity 4.18x108 ft/hr hr

ρ = density of gas

Pl = longitudinal pitch between adjacent tubes in different rows measured on

the diagonal, in

ΔP = 18.93(4836 x 0.0729/0.018 x 2.42) (2.375/0.875) -0.927 (2.375/2) 0.515

(4836x4/4.18x108x0.063)

= 0.17lbf/ft2

= 0.00471psi

= 0.129in H20

ΔP static for 4 rows = 0.516 in H20

ΔP dynamic for 650 ft/min and 4 rows

From graph 0.4 in H20

Total ΔP = 0.916 in



Fans

As fan area is 40 to 50% of bundle face area, fan must be 6in apart from the bundle wall

So,

6 fans of diameter 3ft will be suitable

For a bundle of 4 x 24 ft2

Total fan area = π/4x 32x 6

= 42.41 ft2

Motor Hp = actual ft3 /min (at fan) – total pressure drop /6356

-fan (system efficiency)- (speed reducer efficiency)

= 650 x 42.41-0.916/6356 – 80 - 95

= 4.46Hp for 6 fans = 0.743 Hp for one fan



AIR COOLED HEAT EXCHANGER SPECIFICATION SHEET

PERFORMANCE DATA

TUBE SIDE AIR SIDE

MASS FLOW RATE Kg/hr 21460 AIR VELOCITY m/min 198.12

INLET TEMPERATURE 0C 93.33 MASS VELOCITY Kg/m

2 hr 23635

OUTLET TEMPERATURE 0C 60 INLET TEMPERATURE

0C 32.2

PRESSURE atm 26.4 OUTLET TEMPERATURE 0C 50.7

HEAT CAPACITY KJ/Kg0C 0.0035 PRESSURE DROP atm 0.0306

VISCOSITY Kg/m hr 0.72

HEAT LOAD KJ/hr 2.09×106

ALLOWABLE PRESSURE DROP atm

0.916in H2O

DESIGN PRESSURE DROP atm

CONSTRUCTION

TUBE FIN

MATERIAL

Killed

steel MATERIAL ALUMINUM

OUTER DIA m 0.0254 TYPE CIRCULAR

INNER DIA m 0.0211 HEIGHT m 0.0158

NO. OF TUBES 90 NO. OF FIN /m 315

∆ PITCH m 0.06 THICKNESS m 0.0004

NO. OF PASSES 2

NO. OF ROW 4

MECHANICAL EQUIPMENT

FAN

NO. OF UNIT 6

DIAMETER OF FAN m 0.9146

NO. OF BLADE 4

BLADE MATERIAL PLASTIC

FAN MATERIAL CAST IRON

POWER KW 1



4.5 DESIGN OF SEPARATOR

The separator used in that process is actually three phase separator in which we are

adding water to remove solid particles from naphtha stream .so we have here

Separated gaseous stream

Liquid stream

Water

Here the vessel used is horizontal because

Handling high capacity

Water has to be separated from the stream

DATA:

Gas:

In soluble phase = 1275 lbs/hr

In gaseous phase = 1257.2 lbs/hr

Liquid:

Liquid Naphtha = 42987 lb/hr

Total liquid flow rate = 42987 + 1750 + 2475.52 - 1257.2

= 45955.32 lbs/hr

Total gas flow rate = 1257.2 lbs/hr

= 0.349 lb/sec



Moles of gas

= 1257.2/6.4

= 196.43 lbmoles / hr

Volume of gas:

Pressurjre of gas = 370psi

Temperature of gas = 110F = 570R

Number of moles of gas = 196.43 lbmoles / hr

Gas constant = R = 10.72psi.ft3 / lbmole.oF

Volume of gas = ?

PV = nRT

V = nRT / P

= 196.43 x 10.72 x 570/370

= 3245.8 ft3

Density of gas (ρv)

= 1257.2 / 3245.8

= 0.387 lb/ ft3

Density of liquid (ρ l)

= 46.9 lbs/hr



Selection of Lv/Dv

The most economical length to diameter ratio depends upon operating

pressure here from 290 to 507.64 psi we use Lv / Dv = 4 and the operating pressure of

that process is 370psi

Selection of liquid height ‘hv’

hv = Dv / 2

Fraction of total area occupied by the vapor ‘fv’

Fv = 0.5

Settling velocity of liquid droplets

Ut = 0.07 [(ρL– ρv) / ρv]1/2

= 0.767 ft / sec

Here the separator without demister pad is tried so for that

Ua = 0.15 x Ut

= 0.115 ft / sec

Vapor volumetric flow rate ‘Qv’

= mass flow rate / density

= 0.349 / 0.387

= 0.901 ft3 /sec



Cross sectional area for vapor flow ‘Av’

= π Dv2 / 4 x 0.5

= 0.392 Dv2

Vapor velocity ‘Uv’

= Qv / Av

= 0.901 / 0.392 Dv2

= 2.3 Dv-2

Vapor residence time required for the liquid droplets to settle on the liquid surface:

= hv / Ua

= 0.5 Dv / 0.115

= 4.347 Dv

Actual residence time:

= vessel length /vapor velocity

= Lv / Uv

= 4 Dv / Uv

= 4 Dv / 2.3 Dv-2

= 1.74 Dv3



For satisfactory separation:

Required residence time = actual residence time

4.347 Dv = 1.74 Dv3

Dv = 1.6ft

Liquid hold up time:

Liquid volumetric flow rate = 45955.32 lbs / hr

= 12.76 lbs / sec

= 12.76 / 46.9

= 0.272 ft3/ sec

Liquid cross-sectional area = π Dv2 / 4 x 0.5

= π (1.6)2 /4 x 0.5

= 1 ft2

Length ‗Lv‘

Lv = 4 Dv

= 4x1.6

= 6.4 ft

Hold up volume = area x length

= 1 x 6.4

= 6.4 ft3



Hold up time = liquid volume / liquid flow rate

= 6.4/0.272

= 23.52 sec

= 0.392 min

This is unsatisfactory, 3 minutes minimum required. Need to increase the liquid volume

this is best done by increasing vessel diameter the diameter must be increased by the

factor of roughly (3/0.392)1/2 = 2.76

New Dv = 1.6 x 2.76

= 4.42 ft

New Lv = 4 x Dv

= 16 ft

New liquid volume = cross-sectional area x Length

= (π (4)2 / 4 x 0.5) x (4x 4)

= 100.5 ft3

Liquid residence time = volume/flow rate

= 100.5 / 0.272

= 369.5 sec

= 6.15 min



BBOOOOTT DDEESSIIGGNN::

Residence time = 6.15 min

Water flow rate = 1750 ft3 / hr

Keeping interphase on 50% level of leg

Amount of water = 1750 x 6.15 / 60

= 179.4 lbs

Volume of water hold up in leg = 179.4 / (2.2 x 0.987)

= 82.6 L

= 2.92 ft3

Lv/Dv = 5

As interphase is on 50%,

Total volume of leg = area x length

= π D2 / 4 x 5 x D

Dv = 1.23 ft

Lv = 1.23 x 5

= 6.2 ft



44..66 DDEESSIIGGNN OOFF DDIISSTTIILLLLAATTIIOONN CCOOLLUUMMNN

In industry it is common practice to separate a liquid mixture by distillating the

components, which have lower boiling points when they are in pure condition from those

having higher boiling points. This process is accomplished by partial vaporization and

subsequent condensation.

CHOICE BETWEEN PLATE AND PACKED COLUMN

Vapour liquid mass transfer operation may be carried either in plate column or packed

column. These two types of operations are quite different. A selection scheme

considering the factors under four headings.

i) Factors that depend on the system i.e. scale, foaming, fouling factors,

corrosive systems, heat evolution, pressure drop, liquid holdup.

ii) Factors that depend on the fluid flow moment.

iii) Factors that depends upon the physical characteristics of the column and its

internals i.e. maintenance, weight, side stream, size and cost.

iv) Factors that depend upon mode of operation i.e. batch distillation, continuous

distillation, turndown, intermittent distillation.

The relative merits of plate over packed column are as follows:

i) Plate column are designed to handle wide range of liquid flow rates without

flooding.

ii) If a system contains solid contents, it will be handled in plate column, because

solid will accumulate in the voids, coating the packing materials and making it

ineffective.

iii) Dispersion difficulties are handled in plate column when flow rate of liquid

are low as compared to gases.



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

column.

v) If periodic cleaning is required, man holes will be provided for cleaning. In

packed columns packing must be removed before cleaning.

vi) For non-foaming systems the plate column is preferred.

vii) Design information for plate column are more readily available and more

reliable than that for packed column.

viii) Inter stage cooling can be provide to remove heat of reaction or solution in

plate column.

ix) When temperature change is involved, packing may be damaged.

For this particular process, ―Acetaldehyde, ethyl alcohol and water system‖, I have

selected plate column because:

i) System is non-foaming.

ii) Temperature is high (91o C).

CHOICE OF PLATE TYPE

There are four main tray types, the bubble cap, sieve tray, ballast or valve trays and the

counter flow trays. I have selected sieve tray because:

i) They are lighter in weight and less expensive. It is easier and cheaper to

install.

ii) Pressure drop is low as compared to bubble cap trays.

iii) Peak efficiency is generally high.

iv) Maintenance cost is reduced due to the ease of cleaning.



DESIGNING STEPS OF DISTILLATION COLUMN

Calculation of Minimum Reflux Ratio Rm.

Calculation of optimum reflux ratio.

Calculation of theoretical number of stages.

Calculation of actual number of stages.

Calculation of diameter of the column.

Calculation of weeping point.

Calculation of pressure drop.

Calculation of thickness of the shell.

Calculation of the height of the column.



Colburn’s Method for Minimum Reflux Rm

Rf = ratio of key components in the liquid part of feed.

Vf = XfB / XfA = 0.4345 / 0.0693 = 6.26

Xnl = Pinch composition of light key component.

fhif

fnl

xr

rX

11

738.0195.0126.61

26.6

r

Xnl = Pinch composition of heavy key component.

Feed

20104.9 kg/hr

Distillate

580 kg/hr

D-180

Bottom Product

19524.9 kg/hr

Q-181



= Xnl / rf = 0.738 / 0.26 = 0.118

nH

dBLH

nl

dA

LHm

X

X

X

XR

1

1

738.0

0078.0776.1

118.0

02448.0

1776.0

1mR

= 0.31

Fenske equation

AB

Ba

b

Db

a

m

X

X

X

X

Nlog

log

1

Nm + 1 =1.298/0.249 = 5.16

Nm = 4.16

11 R

RR

N

NN mm

18

3.08

1

16.4

N

N

85.085.016.4 NN

N = 33.4

Viscosity 355.0,776.1,24.0 LHLH

From figure 11.57 Coulson and Richardson, vol.2

75.0



So,

Actual no. of plates = 33.4 / 0.75 = 44.5 = 44 plates

Maximum vapor flow rate in rectifying section = Vn = 25710 lbs

Maximum liquid flow rate in rectifying section = Ln = 24434 lbs

Maximum vapor flow rate in stripping section = Vm = 47193 lbs

Maximum liquid flow rate in stripping section = Lm = 90132 lbs

Plate spacing initial estimate = 0.5m = 18in

Calculation of column diameter based on flooding velocity

Calculate FLV = liquid vapor flow factor

L

V

W

WW

V

LF

LW = liquid mass flow rate kg/s

VW = vapor mass flow rate, kg/s

4.700

9

25710

24434WTopF

= 0.1076

4.742

26.10

47193

90132BottomLVF

= 0.22

From figure 11.27 Coulson and Richardson vol.6



t1 = a constant obtained from fig 11.27

K1 Top = 0.08 K2 Bottom = 0.07

Uf = flooding velocity

V

VLf KU 1

9

94.70008.0TopfU

= 0.700 m/s

2.10

2.104.74207.0BottomfU

= 0.6 m/s

Based on 80% flooding velocity

Superficial Vapor Velocity

48.08.06.0ˆbaseU m/s

56.08.0700.0ˆ,topvU m/s

Maximum volumetric flow rate

Top 36.0360092.2

25710 m3/sec

Bottom 58.0360026.102.2

47193



Net Area Required

Top = 0.36/0.56 = 0.643

Bottom = 0.58/0.48 = 1.2

As first trial take downcomer area as 12% of the total.

Column cross sectional area

Base = 0.643/0.88 = 0.73

Top = 1.2/0.86 = 1.36

Column Diameter

Top 473.0

, Bottom 436.1

= 0.96 m = 1.31 m

Maximum liquid rate (kg/sec)

Top = 11.38

Bottom = 3.08

For bottom column diameter = 1.31m

Column Area Ac 2

4d

Ac = 1.33 m2

Downcomer area Ad 33.112.0

= 0.159 m2



Net area An = Ac – Ad

= 1.33 – 0.159

= 1.171 m2

active area Aa = Ac – 2Ad

= 1.33 – 2(0.159)

= 1.012

Hole area Ah take 10% Aa as first trial = 0.1012 m2

Weir length (from figure 11.31) 76.03.1

= 0.988 m

Take weir height = 50 mm

Hole diameter = 5 mm

Plate thickness = 5 mm

Check Weeping

Maximum liquid rate = 11.38 kg/sec

Minimum liquid rate at 70% turn down 38.117.0

= 7.966 kg/sec

how = weir crust

Maximum

4/5

159.04.742

38.11750owh

= 40.29 mm liquid



Minimum

3/2

98.04.742

966.7750owh

= 37.86 mm liquid

at minimum hw + how = 50 + 37.86

= 87.86 mm liquid

from fig 11.30, Coulson and Richardson Vol.6

K2 = 30.8

2/12

min26.10

4.259.0 naKU

2/1min26.10

54.259.08.30U

= 3.883 m/s

Actual minimum vapour velocity

hA

ratevapour minimum

1012.0

0.580.7

= 4.01 m/s

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



Plate Pressure Drop

Dry Plate Drop

Max. vapour velocity through holes

73.51012.0

0.58ˆhU m/s

from fig. 11.34 for plate thickness/hole dia = 1

and 1.0a

h

p

h

A

A

A

A

lo = 0.84

L

V

o

hd

l

Uh

2ˆ

51

4.742

26.10

84.0

73.551

2

dh = 32.8 mm liquid

Residual Head

83.164.742

105.12 3

rh mm liquid

Total pressure drop = 32.8 + (50 + 40.20) + 16.83

ht = 139.92 mm liquid



Downcomer Liquid Backup

Take hap = hw – 10 = 40 mm

Area under apion 31040988.0

= 0.03952 m2

As this is less than Ad use Aap in eq. 11.92 i.e,

2

166nL

wddc

A

lh

2

039.04.742

38.11166dch

= 25.64 mm ~ 26 mm

Backup in downcomer

hb = 139.92 + 40.29 + 25.64

= 205.85 mm

= 0.20585 m

0.205 < ½ (Tray spacing + weir height)

So tray spacing is acceptable

Check Residence Time

966.71

4.742205.0159.0rt



= 3.03 sec

> 3 sec satisfactory

Check Entrainment

UV = 0.58 / 1.171 = 0.495 m/s

Percent flooding = 0.495/0.6 = 0.082%

FLV = 0.22 from fig. 11.29 = 0.018 well below 0.1

Satisfactory

Trial Lay Out

Use cartridge type construction. Allow 50 mm imperforated strip round plate edge; 50

mm wide calming zeros.

From fig. 11.32

Lw/Dc = 0.76

QL = 99

Angle subtended at plate edge by imperforated strip = 180 – 99 = 81o

Mean length, unpeeforoted edge strip 76.1180

8110503.1 3rt

Area of unpeeforated edge strip 088.076.150 m2

Mean length of calming zone 95.0sin10503.12

993 m

Area of calming zone 095.0105095.02 3

Total area of peeforations, Ap = 1.012 – 0.095 = 0.917 m2



11.0917.0

1012.0

p

h

A

A

From fig. 11.33 lp/dh = 2.6, satisfactory within 2.5 = 4.0

No of Holes

Area of one hole 510964.1

No. of holes 510964.1

1012.0

= 5152.74

TOP DIAMETER

Max. volume liquid flow rate = 24434 kg/hr

Max. vapor liquid flow rate = 25710 kg/hr

9V

4.700L

Liquid flow rate = 6.787 kg/sec

Vapour flow rate = 7.141 kg/sec

For above feed plate

Column dia = 0.96 m



Ac = Column Area = 72.04

2d m2

0864.072.012.0dA

0864.072.0dcn AAA m2

= 0.6336 m2

Aa = Ac – 2Ad = 0.5472 m2

Hole area Ah take 10% Aa = 0.05472 m2

Weir length 1210072.0

0864.0

76.0c

w

D

l

7296.096.076.0wl

Take weir height = 50 mm

Hole diameter = 5 mm

Plate thickness = 5 mm

Check Weeping

Maximum liquid flow rte = 6.78 kg/sec

Minimum liquid. Rate at 70% turn down 76.478.67.0 kg/sec

Maximum

4/5

0864.0700

78.6750owh



= 48.65 mm liquid

Minimum

3/2

729.0700

76.4750owh

= 33.24 mm liq.

At min. rate = 50 + 33.24

= 83.24 mm liq.

From fig. 11.30

K2 = 30.8

146.49

54.259.08.30ˆ2/1(min)hU m/s

Actual min. vapor velocity

hA

ratevap..min

054.0

36.07.0

= 4.66

So well above weep point

Plate Pressure Drop

Dry Plate Drop

Max. vap. Velocity strength holes

95.13054.0

9/78.6ˆhU m/sec



From fig. 11.34 for plate thickness/hole diameter = 1

and Ah / Ap = 0.1

700

9

84.0

95.1351

2

dA

= 21.5 mm liq.

Residual Head

ht = 21.5 + 50 + 48.65 + 17.85

= 138 mm liq.

Down comer liquid back up

Take hap = hw – 10 = 40 mm

Area under apron 31040729.0apA

= 0.0291 m2

2

029.0700

78.6166dch

= 18.51 mm liq.

Backup in down comer

hb = 50 + 48.65 + 138 + 18.51

= 255.16 mm

= 0.255 m

0.255 < ½ (Tray spacing + weir height) So, tray spacing is acceptable.



Check Residence Time

76.4

700255.00864.0rt = 3.24 sec

> 3 is Satisfactory

Check Entrainment

56.06336.0

36.036.0

hv

AU m/sec

Percent flooding 81.07.0

56.0%

FLV = 0.107, = 0.05 well below 0.1

From fig. 11.29

Trial Layout

Use cartridge type construction. Allow 50mm upperforated strip vannd plate edge; 50mm

wide calming zone.

From fig. 11.32

76.0c

wD

l

QC = 99o

Angle subtended at plate edge by unperforated strip = 180 – 99 = 81o

Mean length, unperforated edge strip 28.1180/8105096.0 3 m

Area of unperforted edge strip 692.0sin105096.02

993 m



Area of calming zone 0692.01050692.02 3 m2

Total area of perforations, Ap = 0.5472 – 0.0692 = 0.478 m2

11.0478.0

0547.0

p

h

A

A

From fig. 11.33 lp/dh = 2.6 satisfactory

No. of Holes

Area of one hole 510964.1 m2

No. of holes 13.278510764.1

0547.05



SPECIFICATION SHEET

Identification: Unit Distillation Column

Item No. D-180

Type Sieve Tray Column

Function: Seperation of Light Components Type of Operation Continuous

DESIGN DATA

Trays design Hole design

No of Trays : 43 Weir Height: 1in

Tray Spacing: 0.4m Weir Length: 0.988m

Diameter: 1.3m Hole area: 0.1012m2

Efficiency: 75% Area of one hole: 0.0000196m2

Pressure drop per plate: 139.92mm liq Hole diameter: 0.00635in

No. of holes: 5152

Down comer area: 0.159 m2

Fraction entrainment: 0.018

Chapter #5 Mechanical Design


CCHHAAPPTTEERR 55

MMEECCHHAANNIICCAALL DDEESSIIGGNN

5.1 Shell & Tube heat exchanger:

Shell side:

Material – carbon steel

Working pressure – 0.1N / mm 2

Design pressure – 0.11N / mm 2

Permissible stress for carbon steel – 95 N / mm 2

Dia of shell = 673mm

Tube side:

Working pressure = 0.5N / mm 2

Design pressure = 0.55N / mm 2

Shell thickness:

ts = PD/2f J+P = (0.11x 673) / {(2 x 95 x 0.85) + 0.11}

= 0.45mm

Minimum thickness of shell must be 6.3 mm

Including corrosion allowance, ts = 8mm.



Head thickness:

Shallow dished & torispherical head

t h = PRcW / 2 fJ

Rc – crown radius

W – stress intensification factor

W= 1/4 [Rc / Rk] 0.5

Rk = 6% Rc

W= 1 / 4 [3+ (1 / 6) 0.5]

J= 1

th = (0.11 x 1.77 x 673) / 2 x 95

= 0.689 mm.

Use thickness as it for shell i.e. 8 mm

Segmental baffles:

Baffle spacing = 0.4 x 673 = 269.2 mm

Thickness of baffles = 6 mm

Tie rods and spaces:

Diameter of tie rod = 10 mm

Number of tie rods = 6



Flanges:

Shell thickness = go = 8 mm

Flange material – IS: 2004 – 1962 class 2

Gasket material – asbestos composition

Bolting steel = 5% Cr Mo steel

Allowable stress of flange material – 100 MN / m2

Allowable stress of bolting material, Sg – 138 MN/m2

Outside diameter = B = 673 + (2 x 8)

= 689 mm

Gasket width:

do / di = [(y- pm)/ (y- p{m+1})] 0.5

m – gasket factor – 2.75

y – min design seating stress – 25.5 MN / m2

Gasket thickness = 1.6 mm

Thus,

do / di = 1.002

Let di of the gasket equal 683 mm [10 mm greater than shell dia]

do = 0.683 x 1.002.= 0.684m

Mean gasket width = (0.684 – 0.683) /2

= 6.83 x 10-4



Taking gasket width of 12 mm,

do = 0.696 m

Basic gasket seating width, bo = 5mm

Diameter of location of gasket load reaction is,

G = di + N

= 0.683 + 0.012

= 0.695m

5.2 Estimation of bolt loads:

Load due to design pressure:

H = G2 P / 4 = (3.14 x 0.695 x 0.11) / 4

= 0.06004 MN

Load to keep joint tight under operation:

Hp = G(26)mp

=3.14 x0.695 x 2 x 5 x 10 -3 x 2.75 x 0.11

= 6.6 x10-3 MN

Total operating load:

Wo = H + Hp

= 0.066MN



Load to seat gasket under bolting up condition:

Wg = Gby

= x 0.695 x 0.005 x 25.5

= 0.2783 MN

Controlling load = 0.2783 MN

Minimum bolting area= Am = Wg/Sg

= 0.2783/138

= 2.02 x 10-3 m2

Take Bolt size – M 18 x 2

Actual number of bolts – 44

R = 0.027m

g1 = go/0.707 = 1.415 go for weld leg

go = 8mm

Bolt circle diameter, C = B +2(g1+R)

= 0.689 + 2 (1.415 x 0.008 + 0.027)

= 0.76564 m

Using 66 mm bolt spacing,

C = 44 x 0.066 /

= 0.9243 m

Bolt circle diameter, C = 0.93 m



Flange outside diameter

A = C + bolt diameter + 0.02 m (minimum)

= 0.93 + 0.018 + 0.02

= 0.968 = 0.97m

Check of gasket width

AbSg / GN = ( 1.56 x 10-4 x 44 x 138) / x 0.012 x 0.4 x 0.475

= 50.43 < 2y.It is satisfied

Flange moment computation:

For operating condition:

Wo = W1 + W2 + W3

W1 = (B2 / 4) P

/ 4 (0.689)2 0.11

= 0.0410

W2 = H-W1

= 0.06004 – 0.0410

= 1.9 x 10 -3

W3 = Wo-H = Hp (gasket load)

= 6.6 x 10 -3 MN

Total flange moment, Mo = W1a1 + W2a2 + W3a3

a1 = (C-B) / 2 = 0.93-0.689/2 = 0.241



a3 = (C-B) / 2 = 0.93-0.689 / 2 = 0.241

a2 = (a1+a2) /2 = 0.241

Mo = 1.1 x 10 -2

For bolting up condition

Mg = W. a3

W = (Am +Ab)/(2). Sg

Ab = area of bolt

= 44 x 1.56 x 10 -4

= 6.76 x 10 -3 m 2

Am = Minimum bolt area. =1.38 x 10 -3 m 2

Sg = 138N/mm 2

W = 0.562 MN

a3 = 0.241

Mg = 0.135 MN-m

Mg is controlling moment

Flange thickness:

t 2 = (MCfY)/(BSt) = (MCfY/BSfo)

K= (A/B)

= (0.97/0.689)

= 1.407



Assume, Cf =1

From the graph, Y = 3

M = 0.1275MN-m

St = Allowable stress

=100MN / m 2

t 2 =(0.1275 x 3) / (0.689 x 100)

= 0.0055

t = 0.074m

Tube sheet thickness:

tts. = F x G [0.025 x P / 95] 1/2

= 1 x 0.695 [0.025 x 0.55 / 95] 1/2

= 8.36 mm

tts = 11.3 mm including corrosion allowance

Channel and channel cover:

th = Gc [KP/95]1/2

= 0.695 [1.407x0.55/95]1/2

= 6.27mm

th =10mm including corrosion allowance.

Nozzle:

Thickness of nozzle = PD/2fJ-P



Inlet & outlet dia – 100 mm

Vent – 50 mm

Drain – 50 mm

Opening for relief value – 75 mm

tn = 0.55 x 100/2 x 95 x (1-0.55)

= 0.293

Corrosion allowance 3 mm

tn = 4 mm

Considering the size of the nozzle & the pressure rating, it is necessary to provide for a

reinforcing pad on the channel cover.

Area required to be compensated for each nozzle

A = d x th = 100 x 10 = 1000 mm 2.

Saddle Support:

Material- low carbon steel

Diameter = 454 mm

Length of the shell, L = 3.8 m

Knuckle radius = 6% of diameter

= 27.24 mm

Total depth of head = [Dx r /2]1/2

= [454x27.4 / 2 ]1/2



H= 78.63mm

Weight of vessel & contents, W = 11943 kg.

Distance of saddle centerline from shell end,

A = 0.5 x R = 113.5 mm

Longitudianl bending moments:

M1 = QA [ {1- ( 1- A/L + {R2 – H2 }/2AL)}/ 1+ 4L/ 3L]

Q = Load carried by each symmetrical support

= W/2 ( L + 4H/3)

= 11943/2 ( 3.05 +4x0.078 /3)

=18834.1 Kg

M2 = QL/4 [{{ 1+2 {(R2 – H2) / L2}/ {1+ 4H / 3L}} – 4A / L]

So, M1 = 12.778 Kg.m

M2 = 10218 Kg.m

Stresses in shell at the saddle

1.At the topmost fibre of the cross section.

F1= M1 / k1 R2 t

K1 = 1

t = thickness of the shell

f1 = 12.778 / ( 3.14 x 0.008 x 0.2272)

= 0.9865 Kg / cm2



2.At the bottom most fibre of the cross – section

F2= M1 / k2 R2 t

K2 = 1

F2= 0.9865 Kg/cm2

Stresses are well within the permissible values.

Stresses in the shell at mid – span:

The stress at the span is,

F3= M2 / R2 t

= 789.46Kg / cm2

Axial stress is the shell due to internal pressure :

Fd = P Di / 4 t

= 1.12 x 673 / 4 x 8 = 23.55 Kg / cm2

f3 + fp = 813.015 kg / cm2

Stresses are well within the permissible values.

Chapter #6 Pump Selection


CCHHAAPPTTEERR 66

PPUUMMPP SSEELLEECCTTIIOONN

6.1 FACTORS AFFECTING CHOICE OF A PUMP

1) Many different factors can influence the final choice of a pump for a particular

operation. The following list indicates the major factors that govern pump selection.

2) The amount of fluid that must be pumped. This factor determines the size of

pump (or pumps) necessary.

3) The properties of the fluid. The density and the viscosity; of the fluid influence

the power requirement for a given set of operating conditions, corrosive properties of the

fluid determine the acceptable materials of construction. If solid particles are suspended

in the fluid, this factor dictates the amount of clearance necessary and may eliminate the

possibility of using certain types of pumps.

4) The increase in pressure of the fluid due to the work input of the pumps. The head

change across the pump is influenced by the inlet and downstream reservoir pressures,

the change in vertical height of the delivery line, and frictional effects. This factor is a

major item in determining the power requirements.

5) Type of flow distribution. If nonpulsating flow is required, certain types of

pumps, such as simplex reciprocating pumps, may be unsatisfactory. Similarly, if

operation is intermittent, a self-priming pump may be desirable, and corrosion difficulties

may be increased.

6) Type of power supply. Rotary positive-displacement pumps and centrifugal

pumps are readily adaptable for use with electric-motor or internal-combustion-engine

drives; reciprocating pumps can be used with steam or gas drives.

7) Cost and mechanical efficiency of the pump.



6.2 PUMP SELECTION OF P-111

Flow rate = 4000 bbl / std

= 4000 bbl / std * 1std / 24hr * 1hr / 60min * 42 U.S . gallons / bbl

= 116.6 gallons / min

Inlet Pressure = 15 Psi

Outlet Pressure = 461Psi

Density = 46.25 Lb/ft3

Developed Pressure = 446 Psi = 64224 Psf

Developed Head = 1388.62 ft

Eff = 80 – 0.2855 F + 3.78 x 10 - 4 FG – 2.38 x 10 -7FG2 + 5.39 x 10 – 4F2

-6.39 x 10 -7 F2 G + 4 x 10 -10 F2 G2

Where

Eff = Pump % age efficiency

F = Developed Head, ft

G = Flow rate, GPM

So putting values in the equation we get

Eff = 76.5%

H.P = GPM (ΔP) / 1715* eff



Where

GPM = flow rate in Gallon Per min.

ΔP = Developed pressure, Psi

eff = Efficiency in fraction

Pump Horsepower = H.P = 35.67 h.

Chapter #7 Instrumentation and control


CCHHAAPPTTEERR 77

IINNSSTTRRUUMMEENNTTAATTIIOONN AANNDD CCOONNTTRROOLL

The important feature common to all process is that a process in never in a state of static

equilibrium except for a very short period of time and process is a dynamic entity subject

to continual upset or disturbance which' tend to drive it away from the desired state of

equilibrium the process must then be manipulated upon or corrected to derive some

disturbance bring about only transient effect in the process behavior. These passes away

and the never occur again. Others may apply periodic or cycle forces which may make

the process respond in a cyclic or periodic fashion. Most disturbances are completely

random with respect to time a show no repetitive pattern. Thus their occurrence may be

expected hut cannot be predicated at any particular time. If a process is to operate

efficiently, disturbances in the process must be controlled.

A process is designed for a particular objective or output and is then found. Sometimes

by trail and error and sometimes by referring from the previous, experience that control

of a particular variable associated with some stages of the process is necessary to achieve

the desired efficiency.

Each process will have associated with it number of variables which are independent of

the process and/ or its operation and which are likely to change at random. Each such

change will lead to changes in the dependent variables of the process one of which is

selected as bring indicative of successfully operation. One of the input variable will be

manipulated to cause further changes in the output variable will be manipulated to cause

further changes in the output variable the original conditions, Process may controlled

more precisely to give more uniform and higher quality products by the application of

automatic control, often leading to higher profits additionally, process which response too

rapidly to be controlled by human operators can be controlled automatically. Automatic

control is also beneficial in certain remote, hazardous or routine operations. After a

period of experimentation, computers are now being used to operate automatically

control processing systems, which may too large and too complex for effective direct



human control.

Since process profit is usually the most important benefit to obtained by applying

automatic control. The quality of control and its cost should be compared with the

economic return expected and the process technical objective. The economic return

includes reduced operating costs, maintenance and of the specification product along with

improved process operability and increased throughout.

7.1 COMPONENTS OF THE CONTROL SYSTEM

Process

Any operation of series of operations that produce a desired final result is a process. In

this discussion the process is the purification of natural

Measuring Means

As all the parts of the control system, measuring element, is perhaps the most important.

If the measurements are not made properly the remainder of the system cannot operate

satisfactorily. The measured variable is chosen to represent the desired condition in the

process.

7.2 ANALYSIS OF MEASUREMENT

Variables to he Measured

a. Pressure Measurement

b. Temperature Measurement

c. Flow Rate Measurement

d. Level Measurement



Variables to be Recorded

Indicated temperature, composition, pressure etc.

7.3 CONTROLLER

The controller is the mechanism that responds to any error indicated by the error

detecting mechanism. The output of the controller is some predetermined function of the

error. There arc three types of controllers.

1. Proportion action which moves the control valve indirect proportion to the

magnitude of the error.

2. Integral action (reset) which moves the control valve based on the time integral of the

error and the purpose of integral actions is to drive the process back to .its set point when

it has been disturbed.

3. Ideal derivative action and its purpose are to anticipate where the process is

heading by cooking at the time a rate of change of error. The final control element

receives the signal from the controller and by some predetermined relationship changes

the energy input to the process.

CHARACTERISTICS OF CONTROLLER

In general the process controllers can be classified as

a. Pneumatic controllers

b. Electronic controllers

c. Hydraulic controllers

While dealing with the gases, the controller and the final control element may be

pneumatically operated due to the following reasons.

i. The pneumatic controller is very rugged and almost free of



maintenance. The maintenance men have not had sufficient training and background in

electronics, so pneumatic equipment is simple.

ii. pneumatic controller appears to be safer in a potentially explosive atmosphere which

is often present in the industry.

iii. Transmissions distances are short pneumatic and electronic

transmissions system are generally equal up to about 200 to 300 feet. Above this distance

electronic system beings to offer savings.

MODES OF CONTROL

The various types of control are called modes, and they determine type of response

obtained. In other words these describe the action of controller that is the relationship of

output of output signal to the input or error signal. It must be noted that is error that

achieve the controller. The four basic mode of control are:

1. On-off control

2. Integral control

3. Proportional control

4. Rate or derivative control

In industry purely integral, proportional or derivative modes seldom occur alone in the

control system. The on-off controller is the controller with very high gain. In this case the

error signal at once off the valve or any other parameter upon which it sites or completely

sets system.

7.4 ALARMS AND SAFETY TRIPS

Alarms are used to alert operators of serious and potentially hazardous, deviations in

process conditions, key instruments are fitted with switches and relays to operate audible

and visual alarms on the control panels.



The basic components of an automatic trip system are

1. A sensor to monitor the control variable and provide and output signal when a preset

value is exceeded (the instrument).

2. A link to transfer the signal to the activator, usually consisting of a system of

pneumatic or electric relays.

3. An activator to carry out the required action close or open a valve, switch off a motor.

7.5 CONTROL LOOPS

For instrumentation and control of different sections and equipments of plants, following

control loops are most often used.

1. Feed backward control loop

2. Feed forward control loop

3. Ratio control loop

4. Auctioneering control loop

5. Split range control loop

6. Cascade control loop

Here is given a short outline of these control schemes, so that to justify our selection of a

control loop for specified equipment.

FEED BACK CONTROL LOOP

A method of control in which a measured value of a process variable is compared with

the desired value of the process variable and any necessary action is taken. Feed back

control is considered as the basic control loops system. Its disadvantage lies in its

operational procedure. For example if a certain quantity is entering in a process, then a

monitor will be there at the process to note its value. Any changes from the set point will



be sent to the final control element through the controller so that to adjust the incoming

quantity according to desired value (set point). But in fact change has already occurred

and only corrective action can be taken while using feed back-control system.

FEED FORWARD CONTROL LOOP

A method of control in which the value of a disturbance is measured, and action is taken

to prevent the disturbance by changing the value of a process variable. This is a control

method designed to prevent errors from occurring in a process variable. This control

system is better than feed back control because it anticipates the change in the process

variable before it enters the process takes the preventive action. While in feed back enter

system action is taken after the chanee has occurred.

RATIO CONTROL

A control loop in which, the controlling element maintains a predetermined ratio of one

variable to another. Usually this control loop is attached to such as system where two

different streams enter a vessel for reaction that may be of any kind. To maintain the

stoichiometic quantities of different streams this loop is used so that to ensure proper

process going on in the process vessel.

AUCTIONEERING CONTROL LOOP

This type of control loop is normally used for a huge vessel where, readings of a single

variable may be different at different locations. This type of control loop ensures safe

operation because it employs all the readings of different locations simultaneously, and

compares them with the set point, if any of those readings is deviating from the set point

then the controller sends appropriate signal to final control element.

SPLIT RANGE LOOP

In this loop controller is per set with different values corresponding to different action to

be take at different conditions. The advantage of this loop is to maintain the proper

conditions and avoid abnormalities at very differential levels.



CASCADE CONTROL LOOP

This is a control in which two or more control loops are arranged so that the output of,

one controlling element adjusts the set point of another controlling element. This control

loop is used where proper and quick control is difficult by simple feed forward or feed

backward control. Normally first loop is a feed back control loop. We have selected a

cascade control loop for our heat exchanger in order to get quick on proper control.

7.6 INTERLOCKS

Where it is necessary to follow a fixed sequence of operations for example, during a plant

start-up and shut-down, or in batch operations. Interlocks are includes to prevent

operators departing from the required sequence. They may be incorporated in the control

system design, as pneumatic or electric relays or may be mechanical interlocks.

CONTROL OF HEAT EXCHANGER

The Normal Way

The normal method of controlling a heat exchanger is to measure exit temperature of

process fluid and adjust input of heating or cooling medium to hold the desired

temperature.

To stabilize this feed back control, in almost all cases the control must have a wide

proportional band (i.e, wide range of exit temperature change operates the control valve

through full stroke). The proportional band is determined by gain of other components in

the control loop by process considerations. It is an cxccniion when the usual combination

of conventional control elements permits use of narrow band control mechanism. .

Since heat-exchanger control require a wide proportional band for stabilization, reset

response (rate of change of heating medium How proportional to exit temperature.

deviation from controller set point is normally required to correct for off set in the

controlled variable (temperature). It there are process load change and reset response can

be eliminated in cases where disturbance such as heating fluid header pressure, product



flow rate or inlet temperature changes have small effects relative to desired tolerance on

the controlled variable.

When throughout to a heat exchanger is changed rapidly a short-term error in control

temperature results. The magnitude and duration of this error can normally be reduced by

a factor of two by adding derivative response to the control mechanism and adjusting it

properly. In derivative responses, heating fluid flow rate is proportional to rate or change

of temperature derivation from the set point.

A Pressure Cascade Control

A pressure cascade control system cascades output of a standard three action temperature

controller into the set point of a pressure controller. It achieves a more rapid recovery to

process load disturbances in a shell-and-tube exchanger than can be obtained without the

pressure controller. Heating fluid to the heater is regulated by the pressure controller

which is normally provided with proportional and reset responses. Load change is rapidly

sensed by a change is shell pressure which is compensated for by the pressure controller.

The temperature control system senses the residual error and resets the pressure control

set point.

Bypass Improves Control of Slow-Response Exchanger

In certain cascade, the time response characteristic of heat exchanger is too slow to hold

temperature deviations resulting from load changes within desired tolerances. In some of

these cases, the transient characteristic of the heat exchanger can be circumvented by by-

passing the heater with a parallel line and bledding cold process fluid with hot fluid from

the heater. In the by-pass system care must be taken in sizing valves to obtain the-desired

flow sprit with adequate flow versus steam travel characteristics. Thermal elements

response time is particularly important since this tie constant is a major factor influencing

performance of the system.



Flow Controllers

These are used to control tin- feed rate into a process unit Orifice plates are by far the

most type of How-rate sensor. Normally orifice plates arc designed to give pressure drops

in the range of 20 to 200 inch of water Venture tubes amand turbine meter are also used.

Temperature controller

Thermocouples are the most commonly used temperature sensing device. The two

dissimilar wires produce a millivolt emf that varies with the ―hot- functions‖ temperature.

Iron constant to thermocouples are commonly used over the 0 to 1300 F. temperature

range.

Pressure Controller

Bourdon tubes, bellows and diaphragms are used to sense pressure

and differential pressure. For example, in mechanical system the process pressure force is

balanced by the movement of a spring. The spring positing can be related to process

pressure.

Level Indicator

Liquid levels are detected in a variety of ways. The three common are

1 . The following the position of a float that is lighter than the fluid.

2. Measuring one apparent-weight of a heavy cylinder as it is buoyed up more or

less by the liquid (they are called displacement meters).

3. Measuring the difference in static pressure between two fixed elevations, one in

the vapour above the liquid and the other under the liquid surface. The differential

pressure between the two level taps is directly related to the liquid level in the vessel.

Transmitter

The transmitter is the interface between the process and its control system.The Job of the



transmitter is to convert the sensor signal (millivolts, mechanical movement, pressure

difference etc.) into a control signal 3 to 15 psig air pressure signal, 1 to 5 10 to 50 milli

ampere electrical signal etc.

Control Valves

The interface with the process at the other end of the control loop is made by the final

control element in an automatic control valves control the flow of heating. fluid the open

or close and orifice opening as the system is raised or lowered.

7.7 FEED BACK CONTROL LOOP OF HEAT EXCHANGER E-150

Control Scheme of

Trim Cooler E-150

T

Temperature Recorder

& Indicator(Measuring Instrument)

Controller

Electric

Signal

Control Valve(Final Control Element)

Pneumatic

Signal

Chapter #8 Cost Estimation and Economics of Plant

Location


CCHHAAPPTTEERR 88

CCOOSSTT EESSTTIIMMAATTIIOONN AANNDD EECCOONNOOMMIICCSS OOFF PPLLAANNTT

LLOOCCAATTIIOONN

8.1 PLANT COST ESTIMATION

As the final process-design stage is Complete, it becomes possible to make accurate cost

estimation because detailed equipment specification and definite plant facility

information are available. Direct price quotation based on detailed specification can the n

be obtained from various manufacturers. However o design project should proceed to the

final stages before costs are considered and cost estimate should be made through out all

the early stages of the design when complete specifications are not available. Evaluation

of costs in the preliminary design is said predesign cost estimation. Such estimation

should be capable of providing a basis for company management to decide if further

capital should be invested in the project.

Evaluation of costs in the preliminary design phase is some time called guess estimations.

A plant design obviously must present a process that is capable of operating under

condition which will yield a profit.

A capital investments is required to any industrial process, and determination of the

necessary investment is an important part of a plant design project. The total investment

for any process consists of the physical . equipment and facilities in the plant plus the

working capital for money which must be available to pay salaries keep raw materials

and products on hand and handle other special items requiring a direct cast out lay.

8.2 CAPITAL INVESTMENTS

Before an industrial plant can be put into operation, large amount of -money must be

supplied to purchase and install the necessary machinery and equipment, land and service

facilities must be obtained and the plant-must be erected. Complete with all pipe controls


Location


inn services. In addition it is necessary to have money available for payment of expenses

involved in the plant operation.

The capital needed to supply the necessary manufacturing and plant facilities is called the

fixed capital investment while the necessary for the operation of the plant is termed as the

working capital investment.

1. Working Capital Investment

The capital which is necessary lor the operation of the plant is called working capital

investment.

2. Fixed Capital Investment

The capital needed to supply flu- necessary maMiif'acttirini1 and plant facilities is called

fixed capital investment.

The fixed capital investment classified in to two sub divisions,

i. Direct Cost

ii. Indirect Cost

DIRECT COST

The direct cost items arc incurred in the construction of the plant in addition to the cost of

equipment.

1. Purchased Equipment

2. Purchased Equipment Installation

3. Instrumentation and Control

4. Piping

5. Electrical Equipment and Materials


Location


6. Building (Including Services)

7. Yard Improvement

8. Services Facilities

9. Land

INDIRECT COST

1. Design and Engineering

2. Contractor's Expenses

3. Contractor's Fee

4. Contingency

8.3 METHODS OF CAPITAL INVESTMENT

Various methods are employed for estimating capital investment. The choice of any

method depends on the foil owing-factors,

a. Amount of detailed information available

b. Accuracy Desired

Seven methods of estimating capital investments are outlined, estimate

1. Detailed item estimate

2. Unit estimate

3. Percentage of delivered equipment cost

4. ―Lang‖ factor approximation of capacity ratio

5. Investment cost per capacity


Location


The accuracy of an estimate depends on the amount of design detail available; and the

accuracy of the cost data available; and the time spent on preparing the estimate. In the

early stages of a project only an approximate estimate will be required an justified by the

amount of information by then developed.

PERCENTAGE DELIVERED EQUIPMENT

This method for estimating total investment requires the determination

of the delivered equipment cost. The cost of purchased equipment is the basis of several

pre design methods for estimating capital investment.The most accurate methods for

determining process equipment costs is to obtain firm bids from fabricators or suppliers.

Percentage of delivered equipment cost is the method used for estimating the fixed or

total capital investment requires determination of the delivered equipment cost. The other

items included in the total direct plant cost are then estimated as percentage of the

delivered equipment The addition components of the capital investment are based on

average percentage of total direct plant cost total direct and indirect plant costs or total

capital investment.

Estimating by percentage of delivered equipment cost is commonly used for preliminary

and study estimates. It yield most accurate results when applied to a project similar in

configuration to recently constructed plants.


Location


DIRECT COST

PURCHASED Equipment Cost =E

COMPONENTS % AGES OF E COST ($)

Purchased equipment installation 47% E a

Instrumentation (installed) 12%E b

Piping (installed) 66% E c

Electrical (installed) 11 % E d

Building (including Service) 18% E e

Yard improvement 10% E f

Service facilities 70%. E g

Land 6% E h

Total direct cost D

Total direct cost = D

INDIRECT COST

Engineering and supervision 33%E

Construction Expenses 41%E

Total indirect Cost I

Total direct and indirect cost D+I

Contractor's fee 5%(D+I)=y


Location


Contigency 10%(D+I)= x

E/fixed Capital investment D+I+x+y

Working Capital investment W.C.I

W.C.I 15% total capital


Location


8.4 COST ESTIMATION OF OUR PLANT

EQUIPMENT PURCHASE COST

Heat exchanger train E-110 = 68, 40, 000 Rs

Furnace E-120 = 1710000 Rs

Reactor R-130 = 5700000 Rs

ACHE E-140 = 1180000Rs

Trim cooler E-150 = 1220000 Rs

Three phases Separator H-160 = 1300000 Rs

Column Feed/Effluent exchanger = 2230000 Rs

Distillation column D-180 = 5759280 Rs

CHE-187 = 2145454 Rs

Trim cooler E-183 = 770000Rs Rs

Phase Separator H-184 = 860000 Rs

Total Purchase = 3, 74, 54734 Rs

Direct Cost (Rs)

Purchased equipment cost = 3, 74, 54734 Rs

Purchased equipment installation = 0.47 3, 74, 54734 = Rs.17603709

Instrumentation & Process Control = 0.12 3, 74, 54734 = Rs. 4494564

Piping (installed) = 0.66 3, 74, 54734 = Rs. 24720100


Location


Building (Including Services) = 0.18 3, 74, 54734 = Rs.6741846

Yard improvements = 0.1 3, 74, 54734 = Rs. 3745473

Service facilities (installed) = 0.7 3, 74, 54734 = Rs. 26218290

Land = 0.06 x 3, 74, 54734 = Rs. 2247282

Total direct plant cost = Rs. 123225961

INDIRECT COST

Engg & Supervision = 0.33 3, 74, 54734 = Rs.12360051

Construction expenses = 0.41 3, 74, 54734 = Rs.15356427

Total Indirect Cost = Rs. 27716478

Total Direct & Indirect Cost = Rs150942439

Contractor‘s fee = 0.05 150942439 = Rs. 7547121.95

Contingency = 0.1 150942439 = Rs. 15094243.9

FIXED CAPITAL INVESTMEN

Fixed Capital Investment = Total direct + indirect cost + contingency +

Contractor‘s fee

= Rs. 173583804.9


Location


Total capital investment = F.C.I + W.C

Now

W.C = 0.15 (T.C.I)

= 0.15 (173583804.9 + WC)

= 85.0

74.26037570

= 30632436.15 Rs

Total Capital Investment = T.C.I = 20,42,16,241 Rs.

(Twenty caroor Fourty two lakes sixteen thousands ,two hundred and fourty one rupees

only)

8.5 ECONOMICS OF PLANT LOCATION

The final choice of the plant site usually involves a, presentation ol/the economic factors

for several equally attractive sites. He exact type of economic study of plant locations

will vary with each company making a study. It should include the following.

INVESTMENT

Plant

New Money

Existing facilities

Working capital

Annual sales

Cost


Location


Manufacturing

Distributing

Selling

Research

Annual Earnings

Operative

Net after taxes

Net annual return

On total investment

The limitations of preliminary plan! location cost studies should be recognized pointed

out a management. No matter how carefully a survey is prepared, future trends such as

population and marketing shifts, development of competitive processes and the advent of

new industries. Services and transportation facilities cannot be reliably predicated.

PLANT LOCATION AND SITE SELECTION

The location of plant has a crucial effect on the profitability of project and the scope

for future expansion. Many factors are considered when selecting a suitable site.

A brief explanation of each factor is given below.

i. Raw Materials Supply

Probably the location of the raw materials of an industry contributes more towards the

choice of a plant site than any other factor. This is especially noticeable in those

industries in which the raw material is inexpensive and bulky and is made more compact

and obtains a high bulk value during the process of manufacturing.


Location


ii. Marketing Area

For materials that are produced in bulk quantities, such as cement, minerals acids and

fertilizers, where the cost of e product per ton is relatively low and cost of transportation

has a significant fraction of the sale price. The plant should be located closed to the

primary market. This consideration will be less important for low volume production,

high price product such as pharmaceuticals.

iii. Transportation Facilities

The Transport of material and products too and from the plant will be over riding

consideration in site selection.

If practicable, a site should be selected that is closed to at least two major forms of

transport, road, rail, water way (canal or river) or a sea port. Road transport is being

increasingly used and is suitable for local distribution from a central ware house. Rail

transportation will be cheaper for long distance transport of bulk chemicals.

Air transport is convenient and efficient for the movement of pe rsonnel and essential

equipment and supplies and the proximity of the site to a major airport should be

considered.

iv. Sources of Power

Power for chemical industry is primarily from coal, water and oil; these fuels supply (he

most flexible and economical sources, in as much as they provide for generation of steam

both for processing and for electricity production power can be economically developed

as a by-product in the most chemical plants. If the needs are great enough, since the

process requirements generally call for low-pressure steam. The'turbines of engines used

to generate electricity can be operated non-condensing and supply exhaust steam for

processing purposes.


Location


v. Availability of Labour

Labour will be needed for construction of the plant and its operation. Skilled construction

workers will usually be brought in from outside the site area, but here should be an

adequate pool of unskilled labour available locally; and labour suitable for training to

operate plant. Skilled tradesmen will be needed for plant maintenance. Local trade union

customs and restrictive practices will have to be considered when assessing the

availability and suitability of the local labour for recruitment and training.

vi. Water Supply

Water for industrial purpose can be obtained from one of two general sources: the plant's

own source or municipal supply. If the demand for water is larger, it is more economical

for the industry to supply its own water. Such a supply may be obtained from drilled

wells, rives, lakes, dammed streams or other impounded supplies. Before a company

enters upon any project, it must ensure itself of a sufficient supply of water for all

industrial, sanitary and fire demands, both present and future.

vii. Effluent Disposal

All industrial process produce waste products and full consideration must be given to the

difficulties and cost of their disposal. The disposal of toxic and harmful effluents will be

covered by local regulations and appropriate authorities must be consulted during the

initial site survey to determine the standards that must be met

viii. Local Community Considerations

The proposed plant must fit in with and be acceptable to the local community. Full

consideration must be given to the safe location of the plant so that it dies not impose a

significant additional risk to the community.

On a new site, the local community must be able to provide adequate facilities for, the

plant personnel: school, banks, housing and recreational and cultural facilities.


Location


ix. Land Considerations

Sufficient suitable land must be available for the proposed pant and for future expansion.

The land should ideally be flat, well drained and have suitable load bearing

characteristics. A full site evaluation should be made to determine the need for piling or

other special foundation.

x. Climate

Adverse climatic conditions at a site will increase costs. Abnormally low temperature

will require the provision of additional insulation and special heating for equipment and

pipe runs. Stronger structures will be need at locations subjected to strong winds (cyclone

hurricane areas) or earthquakes.

xi. Political and Strategic Considerations

Capital grants, tax concessions, and other inducements are often given by government's

direct new investment to preferred locations such as areas of high unemployment. The

availability of such grants can be over-riding consideration site selection.

Chapter #9 Hazop Study


CCHHAAPPTTEERR 99

HHAAZZOOPP SSTTUUDDYY

9.1 HYDROGEN SULFIDE POISONING

Hydrogen sulfide is an extremely poisonous gas. Hydrogen sulfide poisoning results

from breathing hydrogen sulfide gas (2$), even in very low concentration. Two forms of

poisoning occur - acute and subacute.

1. Acute Hydrogen Sulfide Poisoning

Breathing air or gas containing as little as 0.10% (40-60 grains of H2S per 100 standard

cubic feet) for ONE MINUTE can cause acute poisoning.

Much sour natural or refinery gas contains more than 0.10% (60 grains per 100 cubic

feet), so care must always be taken to avoid breathing such sour gas. The naphtha

hydrotreating recycle gas and high pressure stripper gas contain from 0.5 to 5% H2$,

while the low pressure stripper gases contain from 10 to 50% H2S,

These gases must NEVER be breathed. One full breath of high concentration hydrogen

sulfide gas will cause unconsciousness, and may cause death, particularly if the victim

falls and remains in the presence of such gas.

The operation of any unit processing gases containing H2S is perfectly safe, provided

ordinary precautions are taken and the poisonous nature of the gas is .recognized. No

work should be undertaken on the unit where there is danger of breathing H2S, and one

should never enter or remain in an area' containing it without wearing a suitable fresh air

mask.

2. Symptoms of Acute Hydrogen Sulfide Poisoning

Muscular spasms, irregular breathing, lowered pulse, odor to the breath, nausea. Loss of

consciousness and suspension of respiration quickly follow.



After apparent recovery, edema (dropsically swelling) of the air passages or lungs may

cause severe illness or death in 8 to 48 hours.

3. First Aid Treatment of Acute Poisoning

Remove the victim at once to fresh air. If breathing has not stopped, keep the victim in

fresh air and keep him quiet. If possible, put him to bed. Secure a physician and keep the

patient quiet and under close observation for about 48 hours for possible edema of the air

passages or lungs.

In cases where the victim has become unconscious and breathing has stopped, artificial

respiration must be started at once. If a Pulmotor or other mechanical equipment is

available, it may be used by a trained person; if not, artificial respiration by mouth-mouth

method must be started as soon as possible. Speed in beginning the artificial respiration is

essential. Do not give up. Men have been revived after more than four hours of artificial

respiration.

If other persons are present, send one of them for a physician. Others should rub the

patient's arms and legs and apply hot water bottles, blankets or other sources of warmth

to keep him warm.

After the patient is revived, he should be kept quiet and warm, and remain under

observation for 48 hours 'for the appearance of edema of the air passages or lungs.

4. Subacute Hydrogen Sulfide Poisoning

Breathing air or gas containing 0.01 to 0.6% H£S (6 to 40 grains per 100 cubic feet) for

an hour or more may cause subacute or chronic hydrogen sulfide poisoning.

5. Symptoms of Subacute Poisoning

Headache, inflammation of the eyes and throat, dizziness, indigestion, excessive saliva,

and weariness are all symptoms which follow continued exposure to H2$ in low

concentrations. Edema of the air passages and lungs may also occur.



6. Treatment of Subacute Poisoning

Keep the patient in the dark to reduce eyestrain and^ have a physician treat the inflamed

eyes and throat. Watch for possible edema.

Where subacute poisoning has been suspected, the atmosphere should be checked

repeatedly for the presence of H2S by such methods as testing by odor, with moist lead

acetate paper, and by Tutweiler determination to make sure that the condition does not

continue.

7. Prevention of Hydrogen Sulfide Poisoning

The best method for prevention of H2O poisoning 1s to stay out of areas known or

suspected to contain it. The sense of smell is not an infallible guide as to its presence, for

although the compound has a distinct and unpleasant odor (rotten eggs), it will frequently

paralyze the olfactory nerves to the extent that the victim does not realize that he is

breathing it. This is particularly true of higher concentrations of the gas.

Fresh air masks or gas masks suitable for use with hydrogen sulfide must be used in all

work where exposure to it is likely to occur. Such masks must be checked frequently to

make sure' that they are not exhausted. Whenever work is done on or in equip ment

containing appreciable concentrations of H2S, men must wear fresh air masks and should

work in pairs so that one may effect a rescue or call for help should-the other be

overcome.

As mentioned above, , the atmosphere in which men work may be checked from time to

time for small concentrations such as would cause subacute poisoning.

REMEMBER - JUST BECAUSE YOUR NOSE SAYS IT'S NOT THERE,

DOESN'T MEAN THAT IT'S NOT 1

8. Further Information

A more detailed information booklet, The Chemical Safety Data Shee t SD36, may be

obtained by writing to:



Manufacturing Chemists Association 1825 Connecticut Avenue, NW Washington, DC

20009

9.2 NICKEL CARBONYL FORMATION

Nickel carbonyl [Ni(CO)4] is known to be an extremely toxic gas. Its primary effect is to

cause lung damage with a lesser effect on the liver. The maximum average exposure to

nickel carbonyl recommended by NIOSH is a TLV of 0.001 ppm (1 ppb), and a

maximum spot exposure of 0.04 ppm (40 ppb).

In Naphtha Hydrotreating units, the potential for forming nickel carbonyl exists only with

catalysts containing nickel (S-6, S-7, S-15, S-16), and only during regeneration or during

the handling of unregenerated catalyst. Care must be used to ensure that the procedures

used will prevent the formation of nickel carbonyl. Data has been published showing the

equilibrium concentration of Ni (C0)4 versus temperature, pressure, 'and CO

concentration in a gas. The nickel carbonyl concentration drops rapidly with increasing

temperature and decreasing CO concentration. At 7 kg/cm2g (100 psig) with 0.5 mol-%

CO in the gas, the nickel carbonyl concentration is at the maximum recommended spot

level of 0.04 ppm at 149°C (300°F), and 0.001 ppm at 182°C (360°F).

The following practices should be followed to prevent the formation of nickel carbonyl:

1. Once a reactor containing a nickel catalyst has been exposed to oxidizing conditions

(regeneration), a measurable concentration of oxygen must be maintained until the com-

bustion of all carbon ceases and all CO2 has been purged from the system.

2. Once a reactor containing a nickel catalyst is in a reducing atmosphere and

regeneration is not desirable, maintain the system in a reducing or inert atmosphere until

all the catalyst has been cooled to at least 66°C (150°F). Unregenerated catalyst should be

unloaded with Ng purged before receiving used catalyst. Oxidation (burning) must be

avoided.

There are many published techniques for determining the concentration of nickel

carbonyl in air (such as a vessel to be entered for maintenance), and several direct reading



instruments are available commercially. For further information, see:

American Industrial Hygiene Assoc. Journal May - June, 1968 Jan. - Feb., 1965

9.3 SAFETY PRECAUTIONS FOR ENTERING A CONTAMINATED

ATMOSPHERE

Anyone entering a vessel which contains- an inert or contaminated atmosphere must

follow all prescribed standard safety precautions and regulations which apply. In

particular, when entering a reactor containing used catalyst, and which therefore can

contain some hydrocarbons and H2S along with possible pyrophoric iron sulfide deposits,

there are a number of additional precautions which apply and which should not be

overlooked. For this discussion, it is assumed that entry into a reactor containing used

catalyst under a nitrogen blanket is planned. In this case, the following precautions

should be included in the standard procedure:

1. The reactor should be isolated by positive action, such as blinding, to exclude all

sources of hydrocarbon, hydrogen, air, etc.

2. Just prior to entry, all purging of nitrogen through the catalyst bed should be

discontinued, and nitrogen purge lines should be inserted at points ABOVE the catalyst

bed.

This is to assure that there will be no forced flow of vapors passing upward through the

catalyst bed and into the working area.,

3. Install an air mover outside the reactor near the open man way nozzle to sweep away

the vapors leaving ,the reactor.

4. The man entering the reactor must be equipped with a fresh air mask in proper

working condition, with a proper air supply.

5. "There should be available and ready for immediate use and transfer to the man in the

reactor, a separate spare air supply which is independent of electrical power.

6. The man entering the reactor should wear a safety harness with a properly attached



safety line.

7. There should be a minimum of two backup men at the man way nozzle in continual

surveillance of the actions of the man in the reactor.

8. There should be a spare fresh air mask complete with its own separate air supply to

allow a second man to enter the reactor quickly in case of an emergency. Therefore, this

spare equipment must be compact enough to allow the second man to enter through the

man way while wearing the equipment.

9. It is recommended that any man working in a. reactor which is under a nitrogen

blanket not be permitted to descend through any appurtenance, such as a tray or quench

gas distributor. The reason for this precaution is that should the man develop some

difficulty while below a tray, for example, to the point where he could not function

properly or lost consciousness, it would be extremely difficult for the surveillance team

outside the reactor to pull the man up through the small tray man way by use of the safety

line.

10. As an added precaution, it is suggested that the man in the reactor have available to

him in the reactor, an emergency self-contained air supply and appropriate associated

equipment. Preferably, the emergency air supply could be connected to the fresh air mask

he is wearing. Such "reserve air supply" systems are available commercially.

Chapter #10 Environmental Impact


CCHHAAPPTTEERR 1100

EENNVVIIRROONNMMEENNTTAALL IIMMPPAACCTT

Petroleum refining is one of the largest industries in the United States and a vital part of

the national economy. However, potential environmental hazards associated with

refineries have caused increased concern for communities in close proximity to them.

This update provides a general overview of the processes involved and some of the

potential environmental hazards associated with petroleum refineries.

10.1 DEFINITION OF A PETROLEUM REFINERY

Petroleum refineries separate crude oil into a wide array of petroleum products through a

series of physical and chemical separation techniques. These techniques include

fractionation, cracking, hydrotreating, combination/blending processes, and

manufacturing and transport. The refining industry supplies several widely used everyday

products including petroleum gas, kerosene, diesel fuel, motor oil, asphalt, and waxes.

10.2 BACKGROUND

The United States is one of largest producers and consumers of crude oil in the world.

Based on data from the U.S. Department of Energy (1998), in 1995 the United States was

responsible for about 23% of the worlds‘ refinery production. With a record high of 324

refineries in the early 80‘s, the U.S. was able to produce about 18.6 million barrels per

day. However, because of changes in oil prices, a shift to alternate fuel use and an

increasing focus on conservation, by 1985 the industry lost several primarily small,

inefficient refineries that could not continue to compete. Over the last decade, the number

of refineries has continued to shrink from about 194 to the current 155. This decrease has

been due in part to increasing requirements placed on the facilities for producing cleaner

fuels along with a number of mandated federal and state clean air and water regulations.



10.3 PROCESSES INVOLVED IN REFINING CRUDE OIL

The process of oil refining involves a series of steps that includes separation and blending

of petroleum products. The five major processes are briefly described below:

1 Separation processes: These processes involve separating the different fractions/

hydrocarbon compounds that make up crude oil based on their boiling point differences.

Crude oil generally is composed of the entire range of components that make up gasoline,

diesel, oils and waxes. Separation is commonly achieved by using atmospheric and

vacuum distillation. Additional processing of these fractions is usually needed to produce

final products to be sold within the market.

2 Conversion processes: Cracking, reforming, coking, and visbreaking are conversion

processes used to break down large longer chain molecules into smaller ones by heating

or using catalysts. These processes allow refineries to break down the heavier oil

fractions into other light fractions to increase the fraction of higher demand components

such as gasoline, diesel fuels or whatever may be more useful at the time.

3 Treating: Petroleum-treating processes are used to separate the undesirab le

components and impurities such as sulfur, nitrogen and heavy metals from the products.

This involves processes such as hydrotreating, deasphalting, acid gas removal, desalting,

hydrodesulphurization, and sweetening.

4 Blending/combination processes: Refineries use blending/combination processes to

create mixtures with the various petroleum fractions to produce a desired final product.

An example of this step would be to combine different mixtures of hydrocarbon chains to

produce lubricating oils, asphalt, or gasoline with different octane ratings.

5 Auxiliary processes: Refineries also have other processes and units that are vital to

operations by providing power, waste treatment and other utility services. Products from

these facilities are usually recycled and used in other processes within the refinery and

are also important in regards to minimizing water and air pollution. A few of these units

are boilers, wastewater treatment, and cooling towers.



10.4 ENVIRONMENTAL HAZARDS OF PETROLEUM REFINERIES

Refineries are generally considered a major source of pollutants in areas where they are

located and are regulated by a number of environmental laws related to air, land and

water. Some of the regulations that affect the refining industry include the Clean Air Act,

the Clean Water Act, the Safe Drinking Water Act, CERCLA (i.e. Superfund:

Comprehensive Environmental Response, Compensation, and Liability Act), Emergency

Planning and Community Right-to-Know (EPCRA), OSHA (Occupational Safety &

Health Administration), TSCA (Toxic Substances Control Act), Oil Pollution Act and

Spill Prevention Control and Countermeasure Plans. Here is a breakdown of the air,

water, and soil hazards posed by refineries:

1 Air pollution hazards: Petroleum refineries are a major source of hazardous and toxic

air pollutants such as BTEX compounds (benzene, toluene, ethyl benzene, and xylene).

They are also a major source of criteria air pollutants: particulate matter (PM), nitrogen

oxides (NOx), carbon monoxide (CO), hydrogen sulfide (H2S), and sulfur dioxide (SO2).

Refineries also release less toxic hydrocarbons such as natural gas (methane) and other

light volatile fuels and oils. Some of the chemicals released are known or suspected

cancer-causing agents, responsible for developmental and reproductive problems. They

may also aggravate certain respiratory conditions such as childhood asthma. Along with

the possible health effects from exposure to these chemicals, these chemicals may cause

worry and fear among residents of surrounding communities. Air emissions can come

from a number of sources within a petroleum refinery including: equipment leaks (from

valves or other devices); high-temperature combustion processes in the actual burning of

fuels for electricity generation; the heating of steam and process fluids; and the transfer of

products. Many thousands of pounds of these pollutants are typically emitted into the

environment over the course of a year through normal emissions, fugitive releases,

accidental releases, or plant upsets. The combination of volatile hydrocarbons and oxides

of nitrogen also contribute to ozone formation, one of the most important air pollution

problems in the United States.

2 Water pollution hazards: Refineries are also potential major contributors to ground



water and surface water contamination. Some refineries use deep- injection wells to

dispose of wastewater generated inside the plants, and some of these wastes end up in

aquifers and groundwater. These wastes are then regulated under the Safe Drinking

Water Act (SDWA). Wastewater in refineries may be highly contaminated given the

number of sources it can come into contact with during the refinery process (such as

equipment leaks and spills and the desalting of crude oil). This contaminated water may

be process wastewaters from desalting, water from cooling towers, storm water,

distillation, or cracking. It may contain oil residuals and many other hazardous wastes.

This water is recycled through many stages during the refining process and goes through

several treatment processes, including a wastewater treatment plant, before being released

into surface waters. The wastes discharged into surface waters are subject to state

discharge regulations and are regulated under the Clean Water Act (CWA). These

discharge guidelines limit the amounts of sulfides, ammonia, suspended solids and other

compounds that may be present in the wastewater. Although these guidelines are in place,

sometimes significant contamination from past discharges may remain in surface water

bodies.

3 Soil pollution hazards: Contamination of soils from the refining processes is generally

a less significant problem when compared to contamination of air and water. Past

production practices may have led to spills on the refinery property that now need to be

cleaned up. Natural bacteria that may use the petroleum products as food are often

effective at cleaning up petroleum spills and leaks compared to many other pollutants.

Many residuals are produced during the refining processes, and some of them are

recycled through other stages in the process. Other residuals are collected and disposed of

in landfills, or they may be recovered by other facilities. Soil contamination including

some hazardous wastes, spent catalysts or coke dust, tank bottoms, and sludges from the

treatment processes can occur from leaks as well as accidents or spills on or off site

during the transport process.



MARKET AND ENVIRONMENTAL FORCES CHANGING THE FACE OF THE

PETROLEUM INDUSTRY

The U.S. petroleum refining industry has come under considerable strain because of

several important factors and changes in the industry. Over the years, there has been an

increased demand for petroleum products and a decrease in U.S. production; however,

there has been no new major refinery construction in the United States in the last 25

years. This lack of infrastructure growth has caused a tremendous strain on the industry

in meeting existing demand, and the U.S. has had to increase the amounts of imports to

meet these needs.

The Clean Air Act and stringent state regulations have also caused the industry to incur

extremely high costs for environmental compliance. These costs are accrued because

refineries must produce reformulated, cleaner-burning gasoline, which require companies

to replace or modify existing equipment with devices for controlling emissions. These

costs of compliance are having a detrimental effect on refineries trying to expand and to

keep pace with the country‘s increasing demand.

The cost of meeting environmental regulations has led many petroleum companies to join

with the federal and state governments in reducing the amounts of hazardous air

pollutants being released. Consent decrees between the petroleum industry and EPA have

been made to reduce air emissions by refineries. One particular agreement was made

between the state of Delaware, Louisiana and the Northwest Air Pollution Authority to

reduce air emissions of nitrogen oxide and sulfur dioxide from nine refineries by more

than 60,000 tons per year (EPA, 2001). The settlements a re an effort to reduce the

amounts of illegal releases of harmful air pollutants from these refineries by installing up-

to-date pollution control devices and reducing emissions from leaking valves, flares and

process units within the refinery. This type of collaboration between refineries and the

state and federal governments provides a cooperative effort towards addressing

environmental concerns within the industry.

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