lube of manufacturing brief report

8/18/2019 Lube of Manufacturing Brief Report

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TECHNO-ECONOMIC FEASIBILITY REPORT OF PRODUCTION OF

190,000 TPA LUBE OIL BASE STOCK AND WAXES VDU RESIDUE

Siddharth Saraswat


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Introduction to the format of the report

The complete project report consists of fourteen chapters, each chapter focusing on a specificdimension of the report.

i. Chapter 1 contains the summary of the project and the problem statement.

ii.

Chapter 2 is dedicated to the process selection and raw material specifications.

iii. Chapter 3 is fully devoted to material and energy flow information.

iv. Chapter 4 includes exhaustive process designing of all major equipment in the plant.

Mechanical designs of some of the equipment have also been presented.

v. Chapter 5 deals with material handling and storage facilities

vi. Chapter 6 consists of process instrumentation and control of the plant.

vii. Chapter 7 deals with the environmental protection and energy conservation in the

plant. Possible sources of pollutants of all kinds are investigated and their mitigate

measures suggested. Uses of alternative sources of energy are assessed.

viii. Chapter 8 gives in full detail of the utilities involved in the plant. Considerations are

taken for various types of utilities, their application range and their general facilities.

ix. Chapter 9 encompasses the organizational structure and manpower requirements of

the project.

x.

Chapter 10 covers the market prospects of the product with pivotal focus on analysis

of demand and supply, export potential and marketing network.

xi. Chapter 11 covers site selection. Alternative feasible sites are weighed and the best

site selected. It also covers the plant layout.

xii. Chapter 12 deals with assessment of the economic viability of the project. It shows in

detail the total project cost estimations, cash flow diagram, and break even analysis and

implementation schedule.


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xiii.

Chapter 13 constitutes of a reference list of books, journals, standard codes and

encyclopaedia which have been used for the project work.

Drawings of process flow sheet, material and energy information, process instrumentation

and control, mechanical designs and plant layout are included in the report.


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CHAPTER 1

SUMMARY OF THE PROJECT

REPORT


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Project at a Glance

1. Product Lube oil base stock

2. Capacity 190000 TPA

3. Location In or around Durgapur, West Bengal

4. Raw Materials Vacuum residue

5.

Important chemical processes involved

Propane Deasphalting

Furfural extraction

Solvent dewaxing

Hydrofinishing


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CHAPTER 2

PROJECT DESCRIPTION


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1)

Introduction

Lube oil Base Stocks

The base stocks used to formulate lubricants are normally of mineral (petroleum) orsynthetic origin, although vegetable oils may be used for specialized applications.

Synthetics can be made from petroleum or vegetable oil feedstocks and are "tailor

made" for the job they are expected to do. Lubricant base stocks influence additive

performance through two main functions: solubility and response. For example,

performance of surface active additives such as anti-wear (AV) or extreme pressure

(EP) depends largely on their ability to adsorb on the machine surface at the proper

time and place. Base stocks with poor solubility characteristics may allow these

additives to separate before they can fulfill their intended functions. Conversely, base

stocks with very high-solubility characteristics may keep the additives in solution, not

allowing them to adsorb.

DEASPHALTING UNIT

This Unit produces deasphalted oil (DAO) by removing asphalt from short Residue obtained

from vacuum Distilation Unit (VDU). Asphalt removal is effected by extraction with liquid

Propane. Deasphalted oil is sent to Furfural unit for aromatics removal.

Solvent deasphalting takes short residue feed from VDU and removes the paraffinic DAO

from the asphalt by counter current solvent extraction. In the extraction, the DAO component

in the feed is soluble in the liquid propane at operating temperatures of the extractor.

Whereas, the asphalt component remains undissolved and which settles down at bottom and

this property is used in extraction.

This propane-DAO mix leaving the top of extractor is sent to a ROSE(Residuum oil super

critical extraction) column. Here the feed is heated and maintained above the supercritical

temperature of propane(above 93degrees C).At this temperature, the solubility of DAO

propane is negligible and hence DAO separates out at bottom and sent out for further

recovery. The main advantage of ROSE process than the conventional process(which uses

distillation to separate propane from DAO mix) is the savings in energy.

Short residue feed is charged to the unit, by the offsite pump at a temperature of 1300C at the

unit battery limits.


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Vacuum residue enters the column frop top and propane solvent required for the extraction

enters the bottom of the Extraction Columns, thus providing counter current flow. The total

propane flow is such that 6 volumes (standard) of propane per volume (standard) of feedare

mixed in the Extraction Columns. The columns operate at a pressure of about 38Kg/cm2g at

the top.

Asphalt is insoluble in the solvent at the extraction conditions and therefore drops out of

solution and exit through the bottom of extraction column, taking some amount of propane

along with it.

DAO is soluble in propane at the extraction conditions. As the DAO/propane phase flows

upward in the Extraction Columns, it passes by a set of heating coils or heat exchangers

located in the top section of the extractors.

The extraction temperature effectively controls the DAO yield. Higher operating

temperaturesresult in less product extracted overhead. A lower operation temperature

produces more DAO, but of a poorer quality.

Product yield and quality is affected by

* Operating temperature

* solvent composition

solvent-to-oil ratio

* Pressure in the Extraction Column (to a lesser extent)

Since certain primary process parameters (i.e. total solvent-to-oil ratio, solvent composition,

and operating pressure) are fixed or set a relatively constant value, the Extraction Column

operating temperature is used as the primary performance control variable. The operating

temperature of Extraction Columns effectively controls DAO yield. Higher temperature will

result in lower yield at higher quality while lower

temperature will result in a higher yield but of a poor quality.

The Extraction Column overhead DAO/propane solution (i.e. rich solvent) is boosted to a

pressure above its critical value by the Extractor Overhead Booster


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Pump . Then it is heated above the critical temperature of the pure solvent by exchanging

heat with recovered solvent in the ROSE Exchanger and M.P. steam in the DAO Separator

Preheater (32-E-119). The DAO Mix then enters the DAO Separator.

Increasing the temperature of the pure solvent above its critical temperature takes advantage

of the solvent’s low-density properties in this region. As the temperature increases above the

critical point, the density of the solvent significantly decreases to values approaching that of

dense gases. At this increased temperature, the DAO is virtually insoluble in the solvent, and

a phase separation occurs. About 90+% of the extraction solvent is recovered by this

supercritical phase separation. The DAO Separator operates at a pressure of 43 bar and a

temperature of 1020C.

The DAO Separator operating conditions are set to achieve the required density difference

needed for good separation. Adjusting the discharge of the Booster

Pump controls the pressure. Adjusting the steam flow to the DAO Separator

Preheater controls the temperature.

Supercritical phase separation in the DAO separator and subsequent heat recovery in the

provides significant energy savings over conventional deasphalting process. The conventional

process requires substantial energy to vaporize and condense subcritical solvent in the solvent

recovery.

The DAO with some remaining propane enters the DAO Strippers on liquid level control

from the DAO Evaporation Columns where the pressure is reduced to 1.7-kg/cm2g. The

DAO is contacted with steam in the DAO Stripper to strip the remaining propane to low

levels in the product stream. DAO product is withdrawn from the bottom of each DAO

Stripper and pumped by the DAO Product Pump (32-P-4/104) through the DAO-Propane

Exchanger and the DAO Cooler to the battery limit.

FURFURAL EXTRACTION UNIT

The Furfural Extraction Unit (FEU) is designed to process the deasphalted oil to improve

viscosity index (VI) of lube oil.


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Some solvents dissolve hydrocarbons at all temperatures and form homogenous (single

phase) solutions; e.g. carbon tetrachloride, dichloromethane, ethyl ether etc. Some other

solvents dissolve selective components of hydrocarbons and leave others, thus forming a two

phase solution, up to a limited temperature (called the critical solution temperature- CST).

However, above the CST, two phases don’t exist and only one remains. These selective

solvents are of great importance in lube base stock manufacture.

Raw lube base stocks (vacuum distillate cuts) contain tars, asphaltenes, polycyclic aromatics

etc. which are undesirable in the finished lube oil base stocks. These undesirable components

behave unfavourably in respect of their viscocity characteristics at higher application

temperatures and also sometimes degrade to leave undesirable deposits during use. The

demulsibil ity of the base stock also gets affected unfavourably in the presence of the

undesirable components. Hence, there is need to eliminate them and this is achieved by

selective solvent refining.

Selective solvent refining involves mixing of the solvent with vacuum distillate cut in the

required proportion and allowing the mixture to settle and separate into two phases in an

extractor. The solvent, solvent/feed ratio and the extractor operating conditions are so

selected that the required degree of separation of undesirable components occurs and two

distinct phases of the desirable and undesirable hydrocarbons dissolved in solvent separate

out. These phases, called raffinate phase and extract phase respectively, are processed

separately to recover the solvent for re-use and leave the raffinate and extract for subsequent

processing/use.

The ability of a solvent to keep the hydrocarbon components in solution is called its solvent

power. For the given hydrocarbon feed and at fixed solvent/ feed ratio, solvents that can hold

more of selective hydrocarbons in solution can be termed as solvents with high solvent

power. High solvent power is desirable because this determines the solvent circulation rates

in the commercial plant and decides the equipment sizing and energy consumption for

pumping as well as for solvent recovery.

Selectivity of a solvent indicates the degree of preference with which a component or a group

of components is dissolved in it from a mixture. For the same solvent & feed quality at higher

solvent/feed ratios, the dissolving power increases and selectivity decreases. The result is a

raffinate of higher quality, but lower yield. Additives like water in phenol or furfural


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decrease dissolving power whereas addition of benzene or Toluene to furfural/phenol/ketones

increases the dissolving power.

BEST SOLVENT PROPERTIES

(i) High solvent power.

(ii) Low light-heavy selectivity and high group selectivity.

(iii) Boiling point much different from that of feed stocks processed.

(iv) Low melting point.

(v) Density of 1.0 to 1.2 gm/c.c

(vi) Low surface tension.

(vii) High thermal/chemical stability.

(viii) Low toxicity.

(ix) Low flamability.

(x) Low corrosivity.

(xi) Low latent heat & specific heat.

(xii) Low viscosity at working temperatures.

(xiii) High bio-degradability.

OPERATING VARIABLES

Solvent/feed ratio:

Higher solvent/feed ratio results in raffinate of better quality and lower yield This is because

of improved solvent power and reduced selectivity.

Top and bottom temperature:

Higher temperautres results in raffinate of higher quality and lower yield and conversely

lower temperatures result in lower quality and higher yield of raffinate. Top is maintained at

10°C less than CST. Bottom is maintained depending on pour point of feed stock and melting

point of solvent.

Water in furfural

Water contamination in furfural decreases the solvent power and has the same effect as

reduced solvent feed ratio.


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High temperautre operation with furfural

Operation of furfural at or above 232°C promotes decomposition of furfural and results in

polymer/coke deposition in furnaces/equipment. Operation at high temperature as above in

the presence of 02 further aggravates the situation of furfural decomposition. The acidity

caused by degraded products of furfural is neutralised by injection of a mild alkaline solution.

Excess alkalinity also promotes cracking and must be carefully avoided.

PROCESS

The feed oil is pumped to the extraction tower via a pump and furfural enters the column

from bottom, thus creating countercurrent flow along the extractor. Packed bed is used for

sufficient contacting.

The extract stream is preheated and is sent to a flash drum for furfural separation. This flash

drum operates at low pressure in the range of 150mmHg and temperature of 160-170 0C. the

bottoms are then sent to the stripper column where remaining furfural is stripped of the oil.

The final output that is oil free from aromatics is then sent to storage tanks.

Solvent Dewaxing Unit process the vacuum distillates obtained from VDU either directly

from vacuum Distillation Unit or after processing the same in Furfural Extraction unit. The

objective of this process is to remove paraffinic hydrocarbons to the extent that the lube base

stock produced will have enough low pour point so that is suitable for low temperature

applications.

Solvent dewaxing Is a complex process which includes extraction and crystallisation

followed by filtration. The solvent blend used in the process under the prescribed operating

conditions, extracts the useful part of the feed stock (i.e. the lube oil), crystallises and

precipitates the undesirable\ part (i.e. wax). The two phase mix thereby developed is filtered

and the useful filtrate solution is separated from wax solution. The solvent being used for this

process is a mixture of Methyl Ethyl Ketone (MEK) and Toluene in volumetrically equal

proportions.

Oil Solvent

A solvent is required to dissolve the desired components of the charge stock (e.g. Benzene,

Tolune). In our case it is Tolune.


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Anti Wax Solvent

Since the oil solvent (Toluene) has some dissolving power for waxes also, another solvent is

required which has the lowest wax solubility. This is called anti-wax solvent or wax rejector

(e.g. MEK,MIBK, Acetone). In our case this anit wax solvent Is MEK. MEK has poor oil

miscibility characteristics too. Toluene MEK blend shall be such to have little solvent effect

on wax and highest solvent effect on oil.

CRYSTALLISATION

Crystals are highly organised type of matter with it’s constituents arranged as space lattice

conforming to any of the standard crystallographic forms.

Each substance has definite crystalline form and chemically similar substances have the same

crystalline form. In the process of crystallisation, the solution when cooled to a temperature

of saturation or lower, the soild phase separates out in the form of crystals. Process of

crystallization takes place simultaneously with (i) Nucleation & ii) Growth.

Above supersaturation, the nucleation is spontaneous and then onwards, nucleation & growth

proceed together. Above saturation but below supersaturation, the growth occurs only when

nucleation is started by seeding with external agents. Nucleation would be

(i) Spontaneous (as in case of above super saturation)

(ii) due to presence of foreign substances

(iii) due to seeding of solution

(iv) due to attribution of existing crystals.

Crystallisation is a complex process involving both heat & mass transfer in a multiphase,

multicomponent system. In the dewaxing process ex- tractive crystallisatlon takes place.

Under suitable chilling condItions., the oil is extracted, the wax is precipitated to saturated /

super saturated condition in the wax phase and nucleation & growth of crystals proceed.

The hydrocarbon components (i.e. waxes) that precipitates may be normal

paraffins iso-paraffins, napthenes or traces of aromatics. The wax crystals that form may be

plate, needle, malcrystalline or microcrystalline.

The size and shape of wax crystals are affected by the following:


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(i) Nature of the feed stock

Depending on the origin of the crude, the distillate cut is likely to vary in respect of its

hycrocarbon composition.

(ii) Hydrocarbon composition:

Different types of hydrocarbons form different types of wax crystals. Normal paraffins form

plate type crystals, isoparaffin s & napthenes form needle and microcrystalline crystals. Plate

& needle types are easily filterable, whereas microcrystalline crystals pose some filtration

problems since it plugs the pores of the filter cloth.

(iii) Boiling range of the feed stock

As the boiling range of the distillate cut increases its viscosity increases and the hycrocarbon

composition changes progressively from normal paraffin to isoparaffin to napthene type. Also

the increased viscosity of the crystallisation media and to reduce the pressure losses in

chillers.

(iv) Rate of cooling/chilling

This is one of the critical parameters that control crystallisation and dictates filtration. Low

chilling rate promote large & uniform wax crystals resulting in improved filtration rate that in

return gives better yield of DWO.

Faster/shock chilling forms fine crystals that cause plugging of the filter cloth and also carry

more oil into crystal lattices. The chilling rate with different dilution modes/rates for different

feed stocks are normally decided based on data from the run of a pilot plant.

(v) Presence of Impurities/contaminatlon

This will effect the controlled process of nucleation and result in large range of crystal size

distribution which is not favourable for good filtration.

(vi) Rate of agitation

This keeps the crystals and solution in constant motion thus ensuring the steady growth of

crystals w.r.t. time in the crystallisation media whose concentration is also maintained as

uniform as possible due to the constant mixing. Absence of agitation would lead to deposition

of wax crystals (having low heat transfer coefficient) on the internal surface and affect the

performance & unit throughput.


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Non-uniform growth of crystals, thus increasing the range of crystal size distribution. This is

not favourable for good filtration.

Excessive agitation, however, results in breaking of already formed crystals which would

multiply the nuclei and result in large range crystal size distribution

The oil from storage tanks is mixed with equal amounts of toluene and MEK solvents at

ambient temptrature in a mixer. The mixing is exothermic and final output temperature is

around 400C. the oil solvent mix is then heated using saturated steam to about 950C in order

to dissolve any wax that is already coagulated in order to prevent uncontrolled crystallization

in chillers.

GOOD QUALITIES OF DEWAXING SOLVENT BLEND

(i) High solvent power for oil & low solvent power for wax.

(ii) Low viscosity at operating temperatures to reduce pressure drops

as well as helps in precipitation/crystallisation of wax.

(iii) Good filtrabIlity.

(iv) Low freezing point and temp. effect of dewaxing.

(v) High thermal/chemical stability.

(vi) Low corrosivIty.

(vii) Low explosivity.

(viii) Low toxicity.

(ix) High biodegradability.

(x) Boiling point much lower than that of the feed stocks.

(xi) Low latent heat/specific heat.

PROCESS

The solution is then cooled using cooling water to about 45 degree Celsius. The stream is

then divided into 6 parallel trains which then undergo crystallization in double pipe heat

exchangers.

The stream first exchanges heat with the cold filtrate stream and then flows inside two

scrapped surface heat exchangers where the loose heat to vaporizing ammonia flowing in

annular side.


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The process of crystallization takes place in these exchangers the stream is then sent to the

rotary vacuum filter. Suction pressure is maintained using inert gas due to which the solvent

is sucked in and the wax is deposited on the filter cloth as cake. Washing of wax cake is done

using solvent in order to remove the oil entrained in the wax.

The wax is removed by doctor’s knife and the solvent after passing through heat exchangers

is sent to the flash drum where .6kg/cm2 pressure is maintained and a temperature of 102 0C

is maintained.

Then the bottoms are sent to the stripper column where remaining solvent is removed by

stripping the feed stream by steam at 1.5 kg/cm2 pressure and 1020C.


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2a) Importance of the problem

The vacuum residua contain a considerable quantity of high viscosity components useful in

the manufacture of lubricationg oils and asphaltenes and resins which contribute an

undesirable dark color to the lube base stocks and form deposits of carbonaceous materials on

heating.

The global base oil market is going through a period of great changes and challenges. First,

the rationalization of Group I production is not anymore a possible scenario, but has become

reality. Secondly, the global capacity of Group II and Group III base oils has significantly

increased during the last years and new capacity has been announced to come on stream in

the coming years. In the unpredictability of global events, one thing appears quite likely and

that is that the future base oil market will be very different from the way it looks today. But

how is the base oil market changing? And why?

The changes in the base oil market can be understood more easily if we start from a basic

consideration. In the past the base oil industry was strictly linked to the lubricant industry,

and base oil supply has been driven, both volumetrically and technically, by the lubricant

demand. However, if we look at the trends in the base oil industry, it clearly appears that the

link between base oil supply and the lubricant industry has become weaker and is going to belargely overruled in the coming years. This means that if we want to understand the changes

in the global base oil market, we have to look outside the lubricant industry and evaluate the

impact and the consequences of “external” factors.

The first, and probably most important, factor to be considered is the impact of mandatory

clean fuel investments. The growth in demand of diesel fuel, combined with the increasingly

stringent regulations on sulphur content will have a strong impact on the global refining

capacity, with significant new investments on new refinery hydrocracking capacity. The

implications of the mandatory clean fuel investments on the base oil market are numerous,

but probably the most relevant way in which clean fuel related investments are impacting and

altering the global base oil landscape is linked to the synergy between hydrocracking

investments and Group II and Group III base oil production. In fact, most of the new

hydrocrackers produce potential feedstock for Group II and Group III base oil production and

a significant portion of new Group II and III capacity is indeed coat-tailed on clean fuel

investments.


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Another factor that is also “favouring” Group II and III production is crude selection. Due to

availability and price issues, several refiners have moved away from light and sweet crudes

(e.g. crudes with high API gravity and low sulphur content) towards heavy sour crudes (e.g.

crudes with low API gravity and high sulphur content). Unfortunately, this shift poses serious

problems to base oil production. In fact, light crudes are good “lube” crudes, or in other

words, crudes that give high yields of base oil feedstock, while heavy crudes are “non-lube”

crudes, that is crudes with low base oil yields.

The shift from a “lube” to a “non-lube” crude is a problem primarily for Group I base oil

production. The reason for this is that Group I base oils are produced mainly by separation

processes, which means that the “lube” molecules must be present in the distillate. Instead,

Group II and III base oils are produced by conversion processes, which means that new

(lube) molecules can be formed and the chemical composition of the final product can be

influenced.

Finally, the last external factor driving future base oil supply is commonly referred to as ‘the

technology paradox”, which is that the highest quality base oil has the lowest cost of

production. Group II and III base oil plants produce high quality base oils, higher base oil

yields and higher value products and by-products than Group I plants and have lower capital

and operating costs.

To summarize, Group I plants are more sensitive to crude selection, have bad economics and

do not present any synergy with a fuel strategy. On top of this, if we look at the demand,

which is historically driven by the automotive industry, we observe an increasing use of

Group II and III base oil and a decreasing use of Group I base oils.

For this reason, as the announced oncoming Group II and III capacity will most likely lead to

a large oversupply, several market analysts agree that Group I refineries will close to

compensate for new G II and G III capacity. The refineries that are more likely to close are

higher costs and small scale operations. Another factor that will be determining is whether

the operations are strategic to the overall business. Also, refineries with excess fuel hydro-

cracking capacity are more likely to close down the base oil line. The time frame of the

closures will be mainly decided by the time of the coming on stream of the new capacity.

However, as we are already seeing, the recession will have an impact and accelerate closures,

as a result of lower demand and lower prices leading to shrinking margins and promotingrationalization of operations.


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2bi/ii) Available processes for the production of the product

Techno-economic appraisal of alternative schemes

Petroleum lubricating oils are made from the higher boiling portion of the crude oil that remain

after removal of the lighter fractions. Crude oils contain varying amounts of compounds of

sulfur, nitrogen and oxygen, metals such as vanadium and nickel, water and salts. All of these

materials can cause problems in refining or subsequent product applications. The manufacture

of the lube base stocks from crude oil involves a series of subtractive processes to remove these

undesirable components, resulting in a base oil that meets performance requirements. The

manufacture of the lube base oils involves following processes.

Vacuum Distilation Process

Vacuum distillation process separates the atmospheric residue mixture into a series of

fractions representing different molecular weight ranges or viscosity ranges from 90-100

neutral to the 500 neutrals.(the neutral number is the SUS viscosity at 100℉) The residuecontains the heavier base oils such as the bright stocks. (150 to 250 SUS at 210℉) The latteris separated from asphaltenes and resins prior to introduction into the extraction process.

Solvent Extraction


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Extraction process involves removal of impurities in the base oils like aromatics, polars,

sulfur and nitrogen compounds. Especially, aromatics make poor quality base oils because

they are among the most reactive components in the natural lube boiling range. Oxidation of

aromatics can start a chain reaction that can dramatically shorten the useful life of a base oil.

Conventionally, solvent(furfural) extraction was adopted as the purification process, in which

aromatics are removed by feeding the raw lube distillate (vacuum gas oil) into a solvent

extractor where it is countercurrently contacted with a solvent. The resulting product is

usually referred to as raffinate. Hydrocracking is a more recent form of purification process.

It is done by adding hydrogen to the base oil feed at higher temperatures and pressures. Feed

molecules are reshaped and often cracked open into smaller molecules. A great majority of

sulfur, nitrogen and aromatics are removed. This massive reforming process produces

molecules that have improved viscometrics and thermal and oxidative stability than product

from solvent extraction process.

Solvent Dewaxing

The next step in the lube base oil manufacture is the dewaxing process. Solvent dewaxing

process utilizes dewaxing solvents like MEK(methyl-ethyl-ketone), which is one of the most

popular ones, to be mixed with the waxy oil. The mixture is then cooled to a temperature 10

to 20 below the desired pour point. The wax crystals are then removed from the oil by

filtration. More desirable alternatives to the solvent dewaxing are i) catalytic dewaxing and ii)

wax hydroisomerization. Catalytic dewaxing removes long n-paraffins and waxy side chains

from other molecules by catalytical cracking them into smaller molecules. The wax

hydroisomerization process, more advanced form of the catalytic dewaxing process,

isomerizes n-paraffins and other molecules with waxy side chains into branched molecules

with very desirable quality as lube base oils rather than cracking them away.

Hydrofinishing

The final process in the manufacturing of lube base oil is hydrofinishing to improve color and

thermal/oxidative stability of base oil. In hydrofinishing process, hydrogen is added to base

oil at an elevated temperature in the presence of catalyst. By reaction of hydrogen with some

remained sulfur and/or nitrogen containing molecules, these sulfur/nitrogen containing

compounds are decomposed into smaller molecules to improve product color and stabilities.


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2B iii) Selection of technologies/schemes:

Choice of solvent

The essential characteristics of a good solvent are its selectivity and its solvent power, or

solvent capacity.

Selectivity corresponds to the solvent’s affinity for one substance rather than another, so that

it will be able to extract this compound preferentially from the hydrocarbon mixture making

up the feed of the extraction unit.

The solvent power, or capacity, is expressed by the amount of feed oil that can be dissolved

per unit of volume or weight of solvent.

A good extraction solvent for aromatics must therefore have high selectivity for aromatics

molecules and good solvent power in order to perform the extraction with a small volume of

solvent.

Besides these two characteristics, the following points also enter into consideration in the

choice of the solvent:

High extraction temperature for good mass transfer;

Easy recovery, if possible simply by flash;

Low vapour pressure to make high pressure equipments unnecessary;

High specific gravity for rapid separation of the oil and solvent phases;

No emulsion for rapid separation of the oil and solvent phases;

Stability, i.e., no thermal or chemical degradation;

Adaptability to a wide range of feeds;

Availability at a reasonable cost;

Non-corrosive toward conventional construction metals;

Non-toxic for the environment and on-the-job safety.

The most often used solvents are:

Furfural

Phenol

N-methyl-2-pyrrolidone, or NMP.


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Environmental and toxicity criteria have taken on an increasing importance in recent years, and

this explains why the phenol process, which was very widely used in 1937, has remained static.

NMP, by itself, is not corrosive to carbon steel. However, because of NMP’s high dielectric

constant, other corrosive compounds will readily ionize in NMP and become very aggressive.

The NMP condensing circuit may be at risk to accelerated corrosion from accumulated

corrosive elements or

corrosion/erosion from high velocities.

It is clear that furfural is applicable for wide range of raw materials, it has a low potential to

form emulsion, it separates easy from oils, and it gives higher raffinate yield in some cases. It

is also better in circulation of the solvent and it requires lower corrosion protection. The most

important reasons for the use of furfural are its low toxicity, low price, availability, better

selectivity and excellent extraction is there is relatively low solvent/oil ratio. Therefore, furfural

is the most often used solvent for extraction of paraffin and naphthene distillates nowadays.

Dewaxing Solvent Selection

As solvent is added to the waxy raffinate the oil is diluted and the

viscosity of the oil solvent mixture decreases allowing filtration to take place

more easily. The polarity of the oil-solvent mixture increases, decreasing the

solubility of the wax and promoting the formation of more compact wax

crystals. But as solvent is added the resulting filtrate becomes more dilute,

loading up filtrate pumps and solvent recovery facilities.

Properties to Consider in Selecting a Dewaxing Solvent:

1. Solubility


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2. Selectivity

3. Solvent boiling point lower than the boiling point of the oil

4. Low heat capacity

5. Heat of vaporization

6. Low viscosity

7. Non-Toxic

8. Non-corrosive

9. Low freezing point

10. Inexpensive

11. Readily Available

Ketone units typically use a dual solvent system consisting of MEK and

either MIBK or Toluene. The MEK acts as an antisolvent to reject wax

molecules from solution. This reduces refrigeration requirements but

excessive MEK may cause oil phase separation. The second solvent keeps the

oil in solution but also dissolves some wax. MIBK and toluene act as

prosolvents.

Typical Dewaxing Solvent Properties

Solvent Wax

Solubility

g/100 ml

Viscosity

@ 0°C, cSt

BP, °C Latent Heat

of

Vaporization,

cal/g

Specific

Heat,

cal/g-oC

MEK 0.25 0.40 80 106 0.55

MIBK 0.90 0.61 116 87 0.46

Toluene 13.0 0.61 111 99 0.41

MEK/MIBK refrigeration requirements are lower than MEK/Toluene

because the Pour-Filter spread is smaller due to the lower wax solubility. The

Pour-Filter spread is the difference between the Dewaxed Oil pour point and

the filtration temperature required to meet the Dewaxed Oil pour point

specification. Wax has a higher solubility in Toluene than MIBK and


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MEK/Toluene systems will require a lower filtration temperature to achieve the same pour

point. MEK/MIBK solvent mixture viscosity is lower than MEK/Toluene. Filtration rates are

higher for MEK/MIBK. Toluene costs less than MIBK.

2B iv) Raw Materials

Vacuum Residue

Composition: cut data

Temperature range % distilled

390-400 .0049

400-420 .0080

420-440 .0112

440-460 .0315

460-480 .0509

480-500 .0951

500-520 .1249

520-540 .1245

540-560 .1109

560-600 .0922

600-650 .0743

650-700 .1095

700+ .162

API gravity = 14.841

Specific gravity = .966

Propane

1. Formula: C3H8

2. Boiling point: -42 °C

3.

Density: 493.00 kg/m³

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4.

Molar mass: 44.1 g/mol

5. IUPAC ID: Propane

6.

Melting point: -188 °C

7. Classification: Alkane

Furfural

Furfural is an organic compound derived from a variety of agricultural byproducts, including

corncobs, oat, wheat bran, and sawdust. The name furfural comes from the Latin word furfur,

meaning bran, referring to its usual source. It is a colorless oily liquid with the odor of

almonds, but upon exposure to air samples quickly become yellow.

Boiling point : 161.7 0 C

Molar mass : 99.06g/mole

Density : 1.16g/cm2

Furfural structure.

Hygrogen gas is used for hydrofinishing o lube oil base stock.

Methyl Ethyl Ketone (MEK):

Butanone, also known as methyl ethyl ketone, is an organic compound with the formula

CH₃CCH₂CH₃. This colorless liquid ketone has a sharp, sweet odor reminiscent of butterscotch and acetone.

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1.

Formula: C4H8O

2. Boiling point: 79.64 °C

3.

Density: 805.00 kg/m³

4. IUPAC ID: Butan-2-one

5. Molar mass: 72.11 g/mol

6.

Melting point: -86 °C

7.

Soluble in: Water

Toluene:

Toluene, formerly known as toluol, is a colorless, water-insoluble liquid with the smell

associated with paint thinners. It is a mono-substituted benzene derivative, consisting of a

CH₃ group attached to a phenyl group.

1. Density: 866.90 kg/m³

2.

Boiling point: 110.6 °C

3. Molar mass: 92.14 g/mol

Methyl ethyl ketone and toluene are used as a solvent in 1:1 ratio for the dewaxing process.

Availability: Indigenous/Imported

Light Arabian Crude is imported from Saudi Arabia through pipelines

https://www.google.co.in/search?es_sm=93&q=butanone+formula&stick=H4sIAAAAAAAAAGOovnz8BQMDgxkHnxCnfq6-gaFJboqBlnp2spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2KVll-UW5qTGKSnvOh-wFqj0JjrKpZLrW5PdxDOBQBZJ2wDUwAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CIoBEOgTKAAwDghttps://www.google.co.in/search?es_sm=93&q=butanone+boiling+point&stick=H4sIAAAAAAAAAGOovnz8BQMDgw0HnxCnfq6-gaFJboqBlm52spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2KVlJ-Zk5mXrlCQn5lXoh6zc9KCiKiTc0-cvzp7yezEn3YLDAB3mX6_WQAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CI0BEOgTKAAwDwhttps://www.google.co.in/search?es_sm=93&q=butanone+density&stick=H4sIAAAAAAAAAGOovnz8BQMDgxkHnxCnfq6-gaFJboqBlnp2spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2KVkppXnFlSeeZHxLI1X5-mXPjSxaibX5fFe052LwAgTmEYUwAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CJABEOgTKAAwEAhttps://www.google.co.in/search?es_sm=93&q=butanone+iupac+id&stick=H4sIAAAAAAAAAGOovnz8BQMDgzkHnxCnfq6-gaFJboqBlkZ2spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2KVWVqQmKyQmfJA-k_VkpJa7bY6z7qg5Te9-VJvRAMAVMGM2VQAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CJMBEOgTKAAwEQhttps://www.google.co.in/search?es_sm=93&q=butanone+molar+mass&stick=H4sIAAAAAAAAAGOovnz8BQMDgyUHnxCnfq6-gaFJboqBllZ2spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2KVm5-TWKSQm1hcrJmwrHhWdkbyTq9WW85oGfdTgYf2AADepRLGVgAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CJYBEOgTKAAwEghttps://www.google.co.in/search?es_sm=93&q=butanone+melting+point&stick=H4sIAAAAAAAAAGOovnz8BQMDgw0HnxCnfq6-gaFJboqBlm52spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2KVm5pTkpmXrlCQn5lXwsxTlsPBGsdj23NT596ycj-jRAZhAJ4nMsFZAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CJkBEOgTKAAwEwhttps://www.google.co.in/search?es_sm=93&q=butanone+soluble+in&stick=H4sIAAAAAAAAAGOovnz8BQMDgyUHnxCnfq6-gaFJboqBllZ2spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2JVnJ9TmpSTqpCZV3O5-6Xz6vT_a3sTWZ4H1Hy8oTG3CADSOWxhVgAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CJwBEOgTKAAwFAhttps://www.google.co.in/search?es_sm=93&q=h2o+water&stick=H4sIAAAAAAAAAGOovnz8BQMDgzMHnxCnfq6-gaFJboqBEgeIaWFskaallZ1spZ-ckZqbWVxSVAlhJSfmxCfn5xbkl-alWBXn55Qm5aQqZOY5qJlvm5L2u-Ws1pPZ_M63nPSdJS0At38jjmAAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CJ0BEJsTKAEwFAhttps://www.google.co.in/search?espv=2&biw=1517&bih=714&q=toluene+density&stick=H4sIAAAAAAAAAGOovnz8BQMDgykHnxCHfq6-gXm2WbyWenaylX5yRmpuZnFJUSWElZyYE5-cn1uQX5qXYpWSmlecWVLp8W1ftsf7yb93MoaaZFhLywcX35gLAOlaNZFSAAAA&sa=X&ei=hQQmVZWsMtKjugSu3YCYDg&ved=0CI8BEOgTKAAwFQhttps://www.google.co.in/search?espv=2&biw=1517&bih=714&q=toluene+boiling+point&stick=H4sIAAAAAAAAAGOovnz8BQMDgzUHnxCHfq6-gXm2WbyWbnaylX5yRmpuZnFJUSWElZyYE5-cn1uQX5qXYpWUn5mTmZeuUJCfmVdiZsb275pK3gnOjvnFAj78LkkpKiUA5m-ENFgAAAA&sa=X&ei=hQQmVZWsMtKjugSu3YCYDg&ved=0CJIBEOgTKAAwFghttps://www.google.co.in/search?espv=2&biw=1517&bih=714&q=toluene+molar+mass&stick=H4sIAAAAAAAAAGOovnz8BQMDgwUHnxCHfq6-gXm2WbyWVnaylX5yRmpuZnFJUSWElZyYE5-cn1uQX5qXYpWbn5NYpJCbWFz830nVeofExidO9-0k1m09I-j6QHEWAK7XDJxVAAAA&sa=X&ei=hQQmVZWsMtKjugSu3YCYDg&ved=0CJUBEOgTKAAwFwhttps://www.google.co.in/search?espv=2&biw=1517&bih=714&q=toluene+molar+mass&stick=H4sIAAAAAAAAAGOovnz8BQMDgwUHnxCHfq6-gXm2WbyWVnaylX5yRmpuZnFJUSWElZyYE5-cn1uQX5qXYpWbn5NYpJCbWFz830nVeofExidO9-0k1m09I-j6QHEWAK7XDJxVAAAA&sa=X&ei=hQQmVZWsMtKjugSu3YCYDg&ved=0CJUBEOgTKAAwFwhttps://www.google.co.in/search?espv=2&biw=1517&bih=714&q=toluene+boiling+point&stick=H4sIAAAAAAAAAGOovnz8BQMDgzUHnxCHfq6-gXm2WbyWbnaylX5yRmpuZnFJUSWElZyYE5-cn1uQX5qXYpWUn5mTmZeuUJCfmVdiZsb275pK3gnOjvnFAj78LkkpKiUA5m-ENFgAAAA&sa=X&ei=hQQmVZWsMtKjugSu3YCYDg&ved=0CJIBEOgTKAAwFghttps://www.google.co.in/search?espv=2&biw=1517&bih=714&q=toluene+density&stick=H4sIAAAAAAAAAGOovnz8BQMDgykHnxCHfq6-gXm2WbyWenaylX5yRmpuZnFJUSWElZyYE5-cn1uQX5qXYpWSmlecWVLp8W1ftsf7yb93MoaaZFhLywcX35gLAOlaNZFSAAAA&sa=X&ei=hQQmVZWsMtKjugSu3YCYDg&ved=0CI8BEOgTKAAwFQhttps://www.google.co.in/search?es_sm=93&q=h2o+water&stick=H4sIAAAAAAAAAGOovnz8BQMDgzMHnxCnfq6-gaFJboqBEgeIaWFskaallZ1spZ-ckZqbWVxSVAlhJSfmxCfn5xbkl-alWBXn55Qm5aQqZOY5qJlvm5L2u-Ws1pPZ_M63nPSdJS0At38jjmAAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CJ0BEJsTKAEwFAhttps://www.google.co.in/search?es_sm=93&q=butanone+soluble+in&stick=H4sIAAAAAAAAAGOovnz8BQMDgyUHnxCnfq6-gaFJboqBllZ2spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2JVnJ9TmpSTqpCZV3O5-6Xz6vT_a3sTWZ4H1Hy8oTG3CADSOWxhVgAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CJwBEOgTKAAwFAhttps://www.google.co.in/search?es_sm=93&q=butanone+melting+point&stick=H4sIAAAAAAAAAGOovnz8BQMDgw0HnxCnfq6-gaFJboqBlm52spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2KVm5pTkpmXrlCQn5lXwsxTlsPBGsdj23NT596ycj-jRAZhAJ4nMsFZAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CJkBEOgTKAAwEwhttps://www.google.co.in/search?es_sm=93&q=butanone+molar+mass&stick=H4sIAAAAAAAAAGOovnz8BQMDgyUHnxCnfq6-gaFJboqBllZ2spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2KVm5-TWKSQm1hcrJmwrHhWdkbyTq9WW85oGfdTgYf2AADepRLGVgAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CJYBEOgTKAAwEghttps://www.google.co.in/search?es_sm=93&q=butanone+iupac+id&stick=H4sIAAAAAAAAAGOovnz8BQMDgzkHnxCnfq6-gaFJboqBlkZ2spV-ckZqbmZxSVElhJWcmBOfnJ9bkF-al2KVWVqQmKyQmfJA-k_VkpJa7bY6z7qg5Te9-VJvRAMAVMGM2VQAAAA&sa=X&ei=hQQmVZuhMs_luQTeiYDoAg&ved=0CJMBEOgTKAAwEQhttps://www.google.co.in/search?es_sm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Prevailing Prices

Vacuum residue’s prevailing price is Rs. 19/kg.

Testing Procedures for raw materials

Vacuum residue testing procedure includes ASTM-189.

Summary of test method

A weighed quantity of sample is placed in a crucible and subjected to destructive distillation.

The residue undergoes cracking and coking reactions during a fixed period of severe heating.

At the end of the specified heating period, the test crucible containing the carbonaceous residueis cooled in a desiccator and weighed. The residue remaining is calculated as a percentage of

the original sample, and reported as Conradson carbon residue .


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CHAPTER 3

MATERIAL AND ENERGY

BALANCE


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3a) Material Balance

Deasphalting Un it

Feed to deasphalting unit : Vacuum residue

Flow rate of vacuum residue = 1710000 TPA

Solvent (i.e Propane) flow rate = 10260000 TPA

Product Deasphalted oil produced = 487350 TPA

Propane present in the deasphalted oil stream = 8878461.473 TPA

Fraction of oil in top stream = .052035

Fraction of propane in top stream = .947965

Asphalts obtained as bottoms = 1381538.527 TPA

Propane present in the bottom stream = 1068104.713 TPA

Material balance equation:

Mass In = mass out

Mass In = propane as solvent + vacuum residue

= 10260000 + 1710000 TPA

= 11970000 TPA

Mass out = 487350 + 8878461.473+1381538.527+1068104.713 TPA

= 11970000 TPA

So, mass in = mass out

FLASH DRUM :

flash inlet stream flow rate = 1170725.85 kg/hr

= 9365811.473 TPA

Top (propane vapor) stream flow rate = 8556914.582 TPA

Bottom flow rate = 808896.8914 TPA

= 101112.3 kg/hr


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Mass in = flash inlet stream

Mass out = top stream + bottom stream

= 808896.8914 + 8556914.582

= 9365811.473 TPA

Mass in = mass out

EVAPORATOR:

Feed to evaporator = 101112.3 kg/hr

Propane vapors removed = 32154.84 kg/hr

Bottoms (i.e) concentrated stream of oil = 68957.46

Total mass out = top +bottoms stream

= 32154.84 + 68957.46

= 101112.3 kg/hr

Stripper column of deasphalting unit:

Pressure = 1.7 kg/cm2

Temperature of feed = 1250C

Temperature of steam = 1500C

Temperature pf top tray = 119.9 0C

Temperature of bottom tray = 119.7 0C

Feed to the flash column = 68957.46 kg/hr

Steam used for stripping = 427.5 kg/hr

Vapor generated from top = 8430.87 kg/hr

Vapor composition:

Water = .1096 wt fraction

Propane = .8904 wt fraction

Bottoms flow rate = 60954.4415 kg/hr

Water wt. fraction = .04%


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Oil wt. fraction = .9996

Mass in = 68957.46 + 427.5 kg/hr

= 69384.96 kg/hr

Mass out = 8430.87 +60954.4415

= 69385.3115

Mass in ≈ mass out ( error due to rounding off of values)

Final output : deasphalted oil with very little amount of moisture present = 60954.4415 kg/hr

= 487635.532 TPA

Fur fur al Extraction materi al balance

Feed entering the furfural extractor unit = Deasphalted oils

Feed entering the furfural extractor = 64162.57 kg/hr

Solvent Flow rate = 76995.084 kg/hr

Aromatics fraction in feed (w/w%) = 0.558

Raffinate Flow rate = 43502.2 kg/hr

= 348017.6 TPA

Furfural in raffinate = 15399 kg/ hr

Aromatics fraction in raffinate (w/w%) = 0.0837

Extract Flow rate = 20660.37 kg/hr

Furfural in extract = 61596.0672 kg/hr

Aromatics fraction in extract (w/w%) = 0.9163

Raffi nate F lash column

flash inlet stream flow rate = 43502.2 kg/hr

= 348017.6 TPA


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32

Top (furfural vapor) stream flow rate = 5220.264 kg/hr

Bottom flow rate = 38281.93 kg/hr

Mass in = flash inlet stream

Mass out = top stream + bottom stream

= 5220.264+38281.93

=348017.6 TPA

Mass in = mass out.

Raffi nate Stri pping column

Pressure = 200 mmHg

Temperature of feed = 1700C

Temperature of steam = 1500C

Temperature pf top tray = 160 0C

Temperature of bottom tray = 160 0C

Feed to the column = 38281.93 kg/hr

Steam used for stripping = 48.02 kg/hr

Vapor generated from top = 1914.096 kg/hr

Vapor composition:

Water = .03 wt fraction

Propane = .97 wt fraction

Bottoms flow rate = 36367.8 kg/hr


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33

Dewaxing unit material balance :

Feed entering the dewaxing unit = 34059.6375 kg/hr

Mixer:

Mixture of Toluene and ketone at ambient temperature in 1:1 ratio are mixed with the waxy oil feed

stream in a mixer.

Solvent to feed ratio = 2.2

Feed stream = 34059.6375 kg/hr

Toluene flow rate = MEK flow rate = 37465.60125 kg/hr

Flowrate of outlet stream = 108990.84 kg/hr

Mass in = mass out = 108990.84 kg/hr

1st Steam heat exchanger of dewaxing circuit:

Inlet = 108990.84 kg/hr

Outlet = 108990.84 kg/hr

2nd heat exchanger where heat is exchanged with cooling water

Inlet = outlet = 108990.84 kg/hr

Flowrate of cooling water = 164215.4325 kg/hr

Chiller section of dewaxing circuit:

Now the stream is divided in 6 exchangers uniformly and the streams flow through a network of

double pipe heat exchangers where it first exchanges heat with the filtrate and liquid ammonia.

Flow rate of each stream = 108990.84/6

=18165.14 kg/hr of waxy oil

Flow rate of filtrate used for cooling = 17586.02 kg/hr

Mass in = mass out= 18165.14 kg/hr

No further mixing is done in heat exchangers ahead and flow rate of the stream is maintained at

17300.07 till the stream enters rotary drum filter.

ROTARY DRUM FILTER

Stream entering the rotary drum = 18165.14 kg/hr

Toluene and MEK mixture is added at the rate of = 798 kg/hr

Filter cake (i.e wax with oil and solvent entrained in it ) obtained = 1376.265 kg/hr

Flow rate of filtrate recovered = 17586.02 kg/hr


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34

15% waxes are removed and it carries over with it 30% oil and about 30% solvent.

i.e the wax obtained from the filter containes 810.9675 kg/hr of wax 243.751 kg/hr of oil and 256.5

kg/hr of solvent.

Mass in = 18165.14 + 798 kg/hr

= 18963.14 kg/hr

Mass out = 17586.02 + 1376.265

=18963.14 kg/hr

Mass in = mass out

(slight difference in values is a result of rounding of)

Flash column of dewaxing unit:

Feed is filtrate recovered from rotary drum vacuum filter.

Feed flow rate =6 x 16748.595 x 1.05 kg/hr

= 105516.1485 kg/hr

Vapour generated = 70182.5552

Bottoms obtained =35333.58 kg/hr

Stripper column :

feed is bottoms from flash drum of dewaxing circuit

feed flow rate = 35333.58 kg/hr

top flow rate = 7334.28234 kg/hr

top composition :

MEK = .1589

Toluene = .1393

Water = .7018

bottom flow rate =27999.29 kg/hr

bottoms composition:

.0024 mole fraction toluene

.0001 mole fraction MEK

.0103 water is present


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35

Steam is used for stripping:

Steam flow rate = 1496.25 kg/hr

Top stream temperature : 118.6 0 C

Bottom stream temperature = 114 0C

Total mass in = 1496.25 + 35333.58

= 36829.83 kg/hr

Total mass out = 7334.28234+ 27999.29

= 36829.83 kg/hr

Mass in = mass out

Bottoms which is dewaxed oil is the feed to the hydrofinishing unit

Flow rate = 27999.29 kg/hr


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36

CHAPTER 4

PROCESS DESIGN


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37

4) List and no. of equipments

EQUIPMENT NO. EQUIPMENT NAME

E 01 HEAT EXCHANGER

E 02 HEAT EXCHANGER

E 03 HEAT EXCHANGER

E 04 HEAT EXCHANGER

E 05 HEAT EXCHANGER

E 06 HEAT EXCHANGER

E 07 HEAT EXCHANGER

E 08 HEAT EXCHANGER

E 09 HEAT EXCHANGER

E 10 HEAT EXCHANGER

E 11 HEAT EXCHANGER

E 12 HEAT EXCHANGER

E 13 HEAT EXCHANGER

E 14 HEAT EXCHANGER

E 15 HEAT EXCHANGER

E 16 HEAT EXCHANGER

E 17 HEAT EXCHANGER

E 18 HEAT EXCHANGER

E 19 HEAT EXCHANGERE 20 HEAT EXCHANGER

E 21 HEAT EXCHANGER

P 01 PUMP

P 02 PUMP

P 03 PUMP

P 04 PUMP

C 01 EXTRACTOR

C 02 EXTRACTOR

F 01 FLASH TANK

F 02 FLASH TANKF 03 FLASH TANK

F 04 FLASH TANK

EV 01 EVAPORATOR

S 01 STRIPPER

S 02 STRIPPER

S 03 STRIPPER

S 04 STRIPPER

ST 01 STORAGE TANK

R 01 REACTOR

SP 01 SPLITTERFU 01 FURNACE


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38

M 01 MIXER

SE 01 SEPARATOR

4a) PROCESS DESIGN

Extractor C01

Taking Vacuum Residue as dispersed phase and Propane as continuous phase

Flow Rate of dispersed Phase = 4114.16 ft3/h

Flow Rate of continuos Phase = 49370 ft3/h

Density of continuous phase = 439.855 kg/m3

= 27.45575 lb/ft3

Density of dispersed phase = 965.917 kg/m3

= 60.29254 lb/ft3

Viscosity of continuous phase = 0.0716 cp

= 0.173208 lbm/ft.h

Interfacial tension = 34 dyne/cm

= 971448 lb/ft

Using ceramic Intallox saddles packing 50mm

Superficial area of packing,a = 42.8 ft2/ft3

Void Fraction = 0.78

Density Difference dp = 32.83679004 lb/ft3

Flooding Correlation

( ). ∗ ( ∆) ∗ .


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39

= 100.41

Using Crawford and C.R. Wilke, Flooding Correlation for packed extraction towers

1000 = . . ∗ µ

VC/VD= C/D = 6

Vd = 20.4 ft/hr

Vc = 122.448 ft/hr

Considering 50% flooding Vd+Vc lie in the packed towers

Cross-sectional Area = D/Vc

= 33.59924 ft2

D = 6.542292 ft

= 1.994 m

Height of the column

Droplet Velocity

= ∆∗

Ρd>ρc, Cp= 0.8

= 8.33 mm

Superficial Velocity

=


40/142

40

= 122.53 ft/h

Mean column to surface diameter of droplets at zero flow rates

= 0.92 ∗ ∆∗ = 2.3 mm

Pratt empirical relationship for kc and kd

= 0.6−. = 1.15−.

Kc = 0.1228 lb.mol/hr.ft2

Kd = 0.235 lb.mol/hr.ft2

Partition Coefficient m =

= 2.18

1 = 1 Kc = 0.045 lb.mols./hr

Packed Height H

= ∗∗∗∆ W = lb. Mols asphalt extracted = 10815.6 lb. mol

S = Cross-sectional area of tower = 33.575 ft2

Δcm = log mean driving force, lb.mol.units

= 0.14815

H = 28.28 ft

= 8.6 m


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41

Total height includes 2ft above and below for inlet and outlet

= 32.28 ft

= 9.84 m

Pressure Drop

Sauter Diameter Dp

= 1.55 ∗ ∆∗.

= 1.62135 ft

∆ = 1.6 Kpa

F lash Tank F01

Feed flow rate = 1232343.6 kg/hr

Liquid flow rate = 106434 kg/hr

Vapor generated = 1125910 kg/hr

Properties and equilibrium data was generated from aspen plus.

Vapour density = 126 kg/m3

Vapour phase is dense propane in gaseous phase just above its critical temperature and

pressure\.

Flv = ((liquid flow rate)/(vapour flow rate))(√ ÷ )

= 106434 ÷ 1125910 ∗ √ 451/126


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42

= .18

K = ^Flv ClnFlv DlnFlv Flv A = -1.887747

B = -.81458

C = -.187074

D = -.014523

E = -.001.15

K = .3789

K ≅ .38 = √ /

= 451 kg/m3 = 126 kg/m3

Permitted velocity = V perm = .61 m/s

Diameter of flash drum = √ 4 ÷ V ∗ ∗ Diameter = 2.469 meter ≅ 2.5 meterHeight of column

Height above feed line = 36” + nozzle diameter/2

Flow velocity max= 100/√ Flow velocity min = 60/√ Flow velocity max = 8.57 m/s

Flow velocity min = 5.19

Lets take flow velocity = 8m/s


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43

Nozzle size = 61cm

Height above feed line = 36” + .305

= .9144 + .305

= 1.2194 m

Heignt above feed line = 12” + .305

= .6098 m

Height of liquid pool = V/(base area)

Taking 8 min residence time

Height of liquid pool = 106434*60*8*4/(451*3600*3.14*2.5^2)

= 6.41m

Total height = H = Height above feed line + Heignt above feed line +Height of liquid pool

= 1.2194+.6098+6.41

= 8.2392 =8.24 m

H/D = 8.24/2.5 = 3.29

This value lies between 3 and 5 and so it is acceptable.

Str ipper column S01

Bottoms mass fraction

Temprature =119.70c

Flow rate = 64162.57 kg/hr

Water = .0004

Density of liquid at bottom = 883.6 kg/m3

Density of vapor(steam) at bottom = .9951 kg/m3


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44

Low pressure superheated steam at 1500c use d for stripping

Vapor leaving the column = 8874.6 kg/hr

Water = .1096

Propane = .8904

Above values in mole fraction

Temprature = 119.90c

Density = 2.041 kg/m3

Number of trays as found from aspen hysys simulation in order to minimize solvent in the bottoms with minimum utility composition = 10

For any no of trays less than 10 any feed inlet temperature 1250C the simulation does not

converge

liquid feed entering the system = 72586.8 kg/hr

top section;

Flv top = (64125/8874.6049) ..

= .346

Flv bottom = (64162/450)(.9950/883.6.5)^(.5)

=4.78

taking plate spacing = .6m

K1(top) = .068

K1(bottom) = .01


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45

flooding velocity top = k((pl – pv)/pv)^.5 m/s

= 1.4161 m/s

flooding velocity bottom = .297 m/s

assuming 85% of flooding velocity

for base uv = 1.2036 m/s

for top uv = .25245 m/s

Area of column

Top

volumetric flow rate = mass flow rate/ density

= 1.207 m3/s

net area = 1.207/1.2036

= 1.002 m2

Bottom

volumetric flow rate = .1256 m3/s

net area = .497 m2

take downcomer area as 12% of total

so now base area = .497/.88 =.56477 m2

top area = 1/.88 = 1.136m2

Diameter of column:

top diameter = (1.136x4/3.14)^(.5)

= 1.20 m

bottom diameter as we can see from area will be less than top diameter so we choose top dia

as column diameter.


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46

Provisional Plate design

plate spacing = .6m

column diameter = 1.20 m

column area = 1.136 m2

downcomer area = .13632 m2

net area = Ac - Ad

= .99968 m2

Active area = .86336 m2

hole area = 6% Aa = Ah (Taking for first trial)

= .0502 m2

weir length = lw = .62Dc

(From Figure 11.31, get lw/DC with help of Ad/Ac)

lw = .62Dc = .62x1.20

= .744 m

assuming weir height = 50 mm

hole diameter = 5mm

plate thickness = 5mm

Check for weeping

maximum liquid flow rate Lw = 64125kg/hr

= 17.81 kg/s

min liquid flow rate, 70% turn down = 12.468 kg/s

maximum depth of the crest of liq over weir

how = (Lw/( lw * ρl ))2/3


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47

= 89.9 mm

Minimum how = 71.2 mm

at

hw + how =50+71.2 mm

= 121.2

(From graph (11.30) in Richardson & Coulson, vol-6)

K 2 = 33

The minimum design vapour velocity, uh is given by:-

Where,

uh = minimum vapour velocity through the holes (based on the hole area), m/s,

dh = hole diameter, mm,

K 2 = a constant, dependent on the depth of clear liquid on the plate, obtained

uh = (33- 0.90 *(25.4-5))/ (2.041)0.5

= 10.24 m/s

Actual minimum velocity = minimum vapour rate / Ah

= 0.7*1.20/ 0.0502

=16.73 m/s

This will cause no weeping.

Plate pressur e drop

Maximum vapour velocity = 1.2078/.086336 = 24.059 m/s


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48

Percent perforated area = 0.06 (Ah/Aa, approximately)

Plate thickness/hole diameter = 1


Orifice coefficient, Co = 0.81

Dry plate drop (hd):-

hd = 51 * (24.059/0.81)2 * (2.041/882.4) mm liquid

= 122.88 mm liquid

Residual head (hr ):-

hr = 12.5 * 103/ ρl

= 12.5 *1000/882.4

= 14.16mm liquid

Total pressure drop ht = hd + (hw + how) + hr

= 122.88 +121 +14.16

= 258.04 liquid

Downcomer liquid back-up

Downcomer pressure loss

Take hap = hw – 5 = 45mm

Area under apron = .86336* 45 * 10-3 = 0.03885m2

This is less than Ad

hdc = 166 * {17.8125/ (882.4*0.03885)}2 = 44.81 mm

Back-up in downcomer = (121 + 44.81 +258.04) = 423.85 mm = 0.424m


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49

0.424< (plate spacing + weir height)/2

Acceptable

Checking residence time

tr = (0.424 * 0.13632 * 882.4)/17.81

= 2.87 s

Perforated Area

Angle subtended = 99

Angle subtended by the edge of the plate = 180-99 = 81

Mean length, unperforated edge strips = (1.2 – 0.05) * 3.14 * 81/180 = 1.624950 m

Area of unperforated edge strips = 50 * 10-3 * 1.20 = 0.0812475 m2

Mean length of calming zone, approx. = weir length + width of unperforated strip

= .744 + 0.05

= .80m

Area of calming zones = 2(.80 * 0.05) = 0.08 m2

Total area of perforations, A p = .744-0.0812-0.08 = .5826m2

Ah/Ap = 0.0502/.5826 = .086

Lp/dh = 3 acceptable

Number of holes

Area of one hole = 1.9625 * 10 -5 m2

No. of holes = 3473 holes

Heat Exchanger E-102


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50

DAO mix with steam

Liquid tube side

Temperature goes from 870-1020 C

Flow rate= 1232343 kg/h

Latent heat of steam λ=2114 kj/kg

Steam in shell side

Flow rate = −.x. Flow rate steam= 45795.8 kg/hr= 12.72 kg/sec.

Assuming U= 800 W/m2 oC

ΔTln = −

=55 oC

Area=

Q

UΔT

=611.138 m2

Select tube of ID 25 mm and OD 30 mm

Area of one tube =0.459596 m2

No. of tubes =Tota area

Area of oe tube = .. =1332.

Sq. pitch=37.5

D b =30x(.)^(

.)

=1567

No. of tubes in centre row Nr =D b/Pt = 1567/37.5


51/142


52/142

52

=1.2721 m/sec.

hi=4200(1.35+ 0.02t)Ui0.8/di0.2 =0.023xRe0.8xPr 0.33

Re= =

...^−

=1.855145x105

Jh=3.5x10-3

Pr =

=.^−.

.^− =5.582

Nu=jhRe(Pr)0.33

hD/K=(2.5x10-3)(1.855145x105)(5.582)0.33

hi =1864.3 W/m2 oC

Overall coefficient

= + + + + x10-3 ln U=536 W/m2 oC

Now using 2 heatexchanger

Initial U=530 W/m2 oC

Q=1232343x5.07x1.033x(102 – 87)/2

A=461.23 m2

Tubes ID=25mm , OD= 30 mm

No. of tubes =461.23/0.459596 =1004

Using one shell two tube pass

D b=30 (1004/0.156)(1/2.291) =1379 mm


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53

Tubes in centre row =37

√=45795.8/(3600x4.88x1004)

=1.296x10-3 kg/s

hc= (0.95)x(0.686)[−.x..x. ]x1/3(25)^-1/6

=1247 W/m2 oC

close enough to 1500 W/m2 oC so no correction is required.

Tube side coefficient:-

Tube area= = 3.140.025/4x1004/2=0..24625 m2

ρliq=350 kg/m3

tube velocity=1232343.6/(3600x350x0.24629x2)

=1.988565 m/sec.

acceptable

hi=4200(1.35+ 0.02t)Ui0.8/di0.2 =0.023xRe0.8xPr 0.33

Re=

=..

.^−

=2.8956x105

Jh=3.5x10-3

Pr =

=.^−.

.^− =5.582

Nu=jhRe(Pr)0.33

hD/K=(2.5x10-3)(2.8656)(5.582)0.33


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54

hi =4074 W/m2 oC

overall coefficient

=

+

+

+

+

x10

-3

ln

U=628.46 W/m2 oC

Error=18%

Acceptable design

Shell ID= 1379 +80 =1460 mm

Cross flow area =− Dsd b=10.2x2.1316 = 0.42632 m2

Mass flow rate= 14.91 kg/m2

de=29.69 mm

μ=1.39x10-2cp

Re= 31,772

Jf =.04

vs= 5.94 m/s

ΔP=2.33 kPa

Negligible pressure drop

Tube side

μ=6.04x10-3

cp

Re= 288750

Jf =0.0022

ΔPt=8190.20 N/m2=8.19 kPa

acceptable


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55

Extractor E-201

Taking Deasphalted oil as dispersed phase and Furfural as continuous phase

Flow Rate of dispersed Phase = 1940 ft3/h

Flow Rate of continuos Phase = 2328 ft3/h

Density of continuous phase = 1068.3 kg/m3

= 66.68329 lb/ft3

Density of dispersed phase = 529.882 kg/m3

= 33.07 lb/ft3

Viscosity of continuous phase= 1.032 cp

= 2.4965 lbm/ft. h

Interfacial tension = 42.2 dyne/cm

= 1205738 lb/ft

Using Intallox saddles packing 50mm

Superficial area of packing,a = 42.8 ft2/ft3

Void Fraction = 0.78

Density Difference dp = 33.6118 lb/ft3

Flooding Correlation

( ). ∗ ( ∆) ∗

.

= 214.44

Using Crawford and C.R. Wilke, Flooding Correlation for packed extraction towers

700 = . . ∗ µ


56/142


57/142

57

= 0.6−. = 1.15−.

Kc = 0.1042 lb.mol/hr.ft2

Kd = 0.199 lb.mol/hr.ft2

Partition Coefficient m =

= 0.608

1

=

1

Kc = 0.092 lb.mols./hr

Packed Height H

= ∗∗∗∆ W = lb. Mols asphalt extracted = 1118 lb. mol

S = Cross-sectional area of tower = 33.575 ft2

Δcm = log mean driving force, lb.mol.units

= 0.24625

H = 26.2 ft

= 7.9 m

Total height includes 2ft above and below for inlet and outlet

= 30.2 ft

= 9.20 m

Pressure Drop


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58

Sauter Diameter Dp

= 1.55 ∗ ∆∗.

= 1.83155 ft

∆ = 2.3 Kpa

Double-pipe heat-exchanger oil mix exchanging heat with cold fi ltr ate E16

Filtrate :

viscosity = 1.13*10^(-3) kg/m-s

filtrate flow rate = 14022.34 kg/hr

C p = 1.5126 KJ/kg

k = .15689

density = 859 kg/m3

waxy-oil mix:

flow rate = 18210.83 kg./hr

C p = 1.63 kJ/kg

k = .15119

density = 889.391 kg/m3

on heat balance we get temperature of outlet stream to be 320C

properties at mean temprature are mentioned above

we use 20 ft long 4x3 inch pipes


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59

LMTD = 18.95 K

a check from table 6.2 kern shows inner area to be higher so hot fluid in inner pipe

annular area = 4.14insq

= .002015 sqmeter

Ga = 14022/area

= 7011000 kg/hr-m2

De= (D2^2 - D1^2)/D1

= (4.5^2 - 3.5^2)/3.5

= .0579 m

viscosity = 1.13x10^(-3) kg/m-s

Rea = GaxDe/viscosity

=359236

Jh = .0022

K = .15689

Pr= 10.89

ha = jhRePr^(.33)k/D

= 4709

Tube side

area = 7.38 sqinch

= .00474m2

Gt = 3841772 kg/hm2

inner dia = 3.068 inch


60/142

60

= .0777m

viscosity = 2.48x10^3 kg/m-s

Ret = DGt/viscosity

= 120365

jh = .0029

K= .15119

Pr = 26

ht = 2.9x10^(-3)x120365x(20)^(.33)x.15119/.0777

= 2012.139

hio = hi*ID/OD

= 1763.68

clean overall coefficient

Uc = hiho/( hi+ho)

= 1283.17

taking fouling factor for waxy oil feed to be 2000

U dirty = 781.6

Q =UAdT

A = 282768/(781*18.95)

= 19.106 m2

for a 3inch pipe outer surface area per foot = .917ft2/ft

= .278 m

length of pipe required = 19.106/.278 m


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61

= 68.72 m

no of 20 feet pipes needed = 68.72/6.08

= 11.3037

= 12 pipes

i.e 6 hairpins in series

now total area = 12x6.08x.278

= 20.282 m2

Ud = Q/Adt

= 735.684 W/m2- 0C

R d = (Uc - Ud)/UcUd

= (1283-735.68)/1283*735.68

= .0005798

Pressure Drop

annulus

De' = D2 - D1

= 4.5-3.5 inch = .0253 m

Re' = 156971

jf = .0021

dP = 8jf (L'/di)Ut2/2= 8913 KPa

Tube side

Jf = .0028


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62

dP = 8jf (L'/di)Ut2/2= 12562.009

acceptable pressure drop values

Rotary drum filter D01

Volume of filtrate ,V = 4.082 m3

Mass of solids per volume of filtrate ,C s = 42.9281 kg/m3

Viscocity of filtrate , = 0.003916 kg/msfraction of the rotational time , f = 1/3

Time for the drum to complete one full cycle, tc = 785 s

Cake resistance , α = 1.6 *1010 m/kg

Pressure drop , ∆ = 147.099 KPa Porosity, ∈ = 0.6For area of rotary drum,

V =

2 −∆

Area = 91.06 m2

Therefore, diameter of rotary drum = 4.83 m

Shell and Tube heat exchanger E14

Waxy oil solvent mix cooled with cooling water

Flow rate of oil = 109264 kg/hr

Heated from 45 to 95 0c

K= .126261 w/m2-k


C p= 2.018 kj/kg


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63

Viscosity = .7*10^(-3) kg/m-s

Water being heated from 25 to 40 degree celsius

Water flow rate as calculated by energy balance = 164627 kg/hr

Lmtd = ((95-40)-(45-25))/ln((95-40)/(45-25))

= 43.16

R =(95-45)/(40-25)

= 3.33

S = (40-25)/(95-25)

= .214

f t = .95

Lmtd = 41

Waxy oil solvent mixture flows inside the tubes

Initially assuming u = 350w/m2

Area = q/udt

= 200.07 m2

Tube size:

Outer diameter : 19.05 mm

Length 5m

Inner diameter = 14.83inch

Surface area = .29936 m2

No of tubes = 668.33

= 672 tubes

Assuming 1 shell 4 tube passes


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64

Tubes per pass = 168

Area of crosss-section of one tube =.000173 m2

Total area =.029 m2

Velocity = v*density/(3600*area)

= 1.1m/sec

Re= 889.37*1.1*.01483/(.7*10^(-3))

= 20726.13259

Pr= 11.188

Jh = 3.9*10*(-3)

Jf = .004

Hd/k =jhrepr*(.33)

Ht d/k= 179.33

Ht = 1526.86

Dp = n p(8jf (l/di)+2.5)ut^(2)/2= 28.589 kpa

= 4.1 psi

Shell side (bells method)

D b = d0(n/.175)^(1/2.285)

= 705.553

= 706mm

Clearance = 64mm

Ds = 770mm

Baffle spacing = ds/2


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65

As = (pt - d0) dsl b/pt

= .05929 m2

Gs = ws/as

= 771.288

V = .8275 m/s

De = 1.27(pt^2 - .785d0^2)/d0

= .01881 m

Re =gsde/viscosity

= 19520

Jf = .046

Dp = 62571.8 kpa

Jh = .0045

Pr = 4.99

Nu = jhr e pr ̂ (.33)

Hd0/k = 150.85

Re = 19722.19

Hoc= 4949

H b = d b/2 -ds(.5 - bc)

= 706/2-770*(.25)

= 160.5

Ncv = (706-321)/20.7168 = 18.58

Fn = 1.03 tube row correction factor

R w = 2nw/nt


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Nw = nt x r a'

H b/d b = 160.5/706 = .227

from figure 12.41 coulson and richardson vol.6

R a' = .16

r w = 2 x .16

= .32

Fw = 1.06 window correction factor

Bypass correction factor

F b = exp(-a(a b/as)(1-(2ns/ncv)^(1/3)))

A b = l b(ds - d b)

= .024640

As = .05929

A b/as = .41558

From fig 12.34 f b = .57

Now we set ns/ncv = 1/5

F b = .863

Leakage correction factor

Fl = 1-atb + 2asb)/al atb = ct3.14d0(nt - nw)/2

Asb = csds(2*3.14 - )/2Ct = .8

Nw = 107.52

Atb = .0135 m2


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Cs = 4.8mm

Asb = 4.8 x 770(2 x 3.14-2.1)/2

= .00772 m2

Al = atb + asb)/2= .021224

Al/as = .3579

= .26

Fl = .64547

Hs = f lf bf wf hhoc

= 2303.00

Dps = dpif b'f l'

Dpi = 8jf us^(2)ncv/ 2Jf = .05

Dpi = 2313

F b' = .64 solved using a= 4

Fl' = .509

Dpc = 753

Window zone pressure drop:

uz = (uwus)^(2)

Uw = ws/aw = .0491/aw

Aw = 3.14(.77)x.77x .16/4 - 107.52 x 3.14 x .01905^(2)/4

= .5934


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Nwv = h b/ pt'

160.5/(.78 x 1.25 x 19.05)

7.723

Dpw = f l' x (2 + .6nwv)density(velocity)^2)/2

= 1089

End zone pressure drop:

Dpe = dpi(nwv+ncv)/ncv x f b

= 2094

Dps = 2dpe + dpc(n b-1) + n b x dpw

N b = 12

So substituting values we get

Dp = 25539

It is below 10 psi so is acceptable

Fouling factor

For shell side = .00033

Tube side = .00053 reference coulson and richardson vol.6

K= 50 w/m2

1/u = 1/2303 +(19.05/14.83)x1/1526 + .00033 +(19.05/14.83) x .0005 +

(.01905/(2*50))ln(19.05/14.83)

u = 367 w/m2

Value is very close to our assumed value so no further iterations required.


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Stripper column S04

Bottoms mass fraction

Toluene = .0024

Mek = .0001

Temprature =1140c

Flow rate = 34582 kg/hr

Water = .0103


Low pressure superheated steam at 1500c use d for stripping

Vapor composition

flow rate = 7352.6 kg/hr

Mek = .1589

Toluene = .1393

Water = .7018

Temprature = 118.60c


Number of trays as found from aspen hysys simulation in order to minimize solvent in the

bottoms with minimum utility composition = 10

liquid feed entering the system = 40434.68

top section;

Flv top = (40435.68/7352.6) ..


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70

= .2426

Flv bottom = (34582/1500)(1.037/887.5)^(.5)

=.788

taking plate spacing = .6m

K1(top) = .075

K1(bottom)= .045

flooding velocity top = 1.69 m/s

flooding velocity bottom = 1.31539 m/s

assuming 85% of flooding velocity

for base uv = 1.4365 m/s

for top uv = 1.25 m/s

Area of column

Top

volumetric flow rate = mass flow rate/ density

= 1.18194 m3/s

net area = 1.181.94/1.4365

= .82279 m2

Bottom

volumetric flow rate = .3956

net area = .3539 m2

take downcomer area as 12% of total

so now base area = .3539/.88 =.40 m2

top area = .82279/.88 = .93498 m2


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Diameter of column:

top diameter = (.93498x4/3.14)^(.5)

= 1.09 m = 1.1m

bottom diameter as we can see from area will be less than top diameter so we choose top dia

as column diameter.

Provisional PLate design:

plate spacing = .6m

column diameter = 1.1m

column area = .935m2

downcomer area = .1122 m2

net area = Ac - Ad

= .8228 m2

Active area = .7106 m2

hole area = 10% Aa = Ah (Taking for first trial)

= .07106 m2

weir length = lw = .76Dc

(From Figure 11.31, get lw/DC with help of Ad/Ac)

lw = .76Dc = .76x.935

= .7106 m

assuming weir height = 50 mm

hole diameter = 5mm

plate thickness = 5mm


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Check for weeping

maximum liquid flow rate Lw = 40434.68 kg/hr

= 11.23 kg/s

min liquid flow rate, 70% turn down = 7.68 kg/s

maximum depth of the crest of liq over weir

how = (Lw/( lw * ρl ))2/3

Minimum how = 53.22mm

at

hw + how =50+53.2 mm

= 103.22 mm


K 2 = 31

The minimum design vapour velocity, uh is given by:-

Where,

uh = minimum vapour velocity through the holes (based on the hole area), m/s,

dh = hole diameter, mm,

K 2 = a constant, dependent on the depth of clear liquid on the plate, obtained

uh = (31 - 0.90 *(25.4-5))/ (1.728)0.5

= 9.61 m/s

Actual minimum velocity = minimum vapour rate / Ah

= 0.7*1.181/ 0.071


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=11.64m/s

This will cause no weeping.

Plate pressure drop:

Maximum vapour velocity = 16.63 m/s

Percent perforated area = .10(Ah/Aa, approximately)

Plate thickness/hole diameter = 1


Orifice coefficient, Co = 0.84

Dry plate drop (hd):-

hd = 51 * (16.63/0.84)2 * (1.728/887.5) mm liquid

= 39mm liquid

Residual head (hr ):-

hr = 12.5 * 103/ ρl

= 12.5 *1000/887.5

= 14.08 mm liquid

Total pressure drop ht = hd + (hw + how) + hr

= 39 + 103 +14.08

= 156 mm liquid

Downcomer liquid back-up

Downcomer pressure loss

Take hap = hw – 10 = 40mm


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Area under apron = .7106 * 40 * 10 -3 = 0.0284m2

This is less than Ad

hdc = 166 * {11.23/ (887.5*0.0284)}2 = 33mm

Back-up in downcomer = (50 + 53+33 +156) = 292 mm = 0.292m

0.292< (plate spacing + weir height)/2

Acceptable

Checking residence time

tr = (0.292 * 0.1122 * 887.5)/11.23

= 2.6m/sec

Perforated Area

Lw/Dc = 0.76

Angle subtended = 99

Angle subtended by the edge of the plate = 180-99 = 81

Mean length, unperforated edge strips = (1.1 – 0.05) * 3.14 * 81/180 = 1.4836 m

Area of unperforated edge strips = 50 * 10

-3

* 1.4836 = 0.07428 m

lube of manufacturing brief report

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