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INSTITUTE OF TECHNOLOGY İZMİR

FIRE HEATER ENGINEERING ECONOMICS & DESIGN

Ş. Selcen KARAKOÇ

Instıtute of Technology İzmir

B.Dilhan CAM

Instıtute of Technology İzmir

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ABSTRACT

A fired heater is a direct-fired heat exchanger which uses the hot gases of combustion

to raise the temperature of a feed flowing through tubes. The process fluid is first heated in

the convection section preheat coil which is followed by further heating in the radiant

section. In both sections heat is transferred by both mechanisms of heat transfer, radiation

and convection. In this project a fired heater design, the natural gas and the 25% excess air

were reacted in order to obtain efficient combustion heat for the system. The heat duty was

found as 11.61 106kj/h . The temperature was 623 K and using 64 tubes.

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1. INTRODUCTION

1.1 Fired Heater

A heat exchanger is a device that is used for transferring thermal energy betweentwo

or more fluids, between a solid surface and a fluid, or between solid particulates anda fluid, at

diff erent temperatures and in thermal contact. Typical applications imply heating or cooling

of a fluid stream of concern and evaporation or condensation of single- or multicomponent

fluid streams.[1]

A fired heater is a direct-fired heat exchanger that uses the hot gases of combustion to

increase the temperature of a feed flowing through coils of tubes aligned throughout the

heater.Fired heaters are used throughout hydrocarbon and chemical processing industries

such as refineries, gas plants, petrochemicals, chemicals and synthetics, olefins, ammonia and

fertilizer plants. Most of the unit operations require one or more fired heaters as start-up

heater, fired reboiler, cracking furnace, process heater, process heater vaporizer, crude oil

heater or reformer furnace. The functions of fired heaters in chemical plants are many

ranging from simpleheating or providing sensible heat and raising the temperature of the

charge to heating andpartial evaporation of the charge, where equilibrium is established

between the unvaporisedliquid and the vapour. The mixture leaves the furnace in the form of

a partially evaporatedliquid in equilibrium.Fired heaters are usually classified as vertical

cylindrical or box-type heaters depending onthe geometrical configuration of the radiant

section.[2]

In the cylindrical-type furnace, the radiation section is in the shape of a cylinder with

a vertical axis, and the burners are located on the floor at the base of the cylinder.Heat

transfer to the tubes on the furnace walls is predominantly by radiation. The heatexchange

area covers the vertical walls and therefore exhibits circular symmetry withrespect to the

heating assembly. In the radiant section, the tubes may be in a circular patternaround the

walls of the fire box or they may be in a cross or octagonal design which willexpose them to

firing from both sides. The shield andconvection tubes are normally horizontal. Cylindrical

heaters with vertical tubes (Fig. 1) are commonly used in hot oil services andother processes

where the duties are usually small, but larger units, 100 million kJ/hr andhigher, are not

uncommon.Heat transfer in the shieldsectionwill be by bothradiation and convection. The

tubesizesusedwillnormally be between 75 and 150mm diameter.The tube size and number of

passesuseddepend on the application and the processfluidflow rate.[3][4]

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Cylindrical heaters are often preferred to box-type heaters. This is mainly due to the

moreuniform heating rate in cylindrical heaters and higher thermal efficiency.

Furthermore,cylindrical heaters require smaller foundations and construction areas and

theirconstruction cost is less.

Figure 1.1.1: Illustration of fired heater with side view of top section

Radiant Section: The radiant tubes, either horizontal or vertical, are located along the walls

in the radiant section of the heater and receive radiant heat directly from the burners or target

wall. The radiant zone with its refractory lining is the costliest part of the heater and most of

the heat is gained there. This is also called the firebox.

Convection Section: The feed charge enters the coil inlet in the convection section where it

is preheated before transferring to the radiant tubes. The convection section removes heat

from the flue gas to preheat the contents of the tubes and significantly reduces the

temperature of the flue gas exiting the stack. Too much heat picked up in the convection

section is a sign of too much draft. Tube temperature is taken in both convection and radiant

sections.

Shield Section: Just below the convection section is the shield (or shocktube) section,

containing rows of tubing which shield the convection tubes from the direct radiant heat.

Several important measurements are normally made just below the shield section.

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Stack and Breeching: The transition from the convection section to the stack is called the

breeching. By the time the flue gas exits the stack, most of the heat should be recovered and

the temperature is much less. From a measurement point of view, this location places fewer

demands on the analyzer but is much less desirable for the ability to control the process.

Burner: Traditional premix burners on a process heater premix the fuel with the primary air

which is inspired to the burner by the fuel gas flow. The pressure of the fuel gas supply is

important since low gas pressure degrades performance. The primary air flow should be

maximized without lifting the flame off the burner. Most of the air (as primary air) is

delivered to the burner along with the fuel. Secondary air is introduced and adjusted with the

registers.

Figure 1.1.2:Premix burner with products of combustion

A correctly set burner, with good air-fuel mixing, produces the maximum flame

temperature in a compact conical flame. The flue gas contains a minimum of oxygen together

with levels of combustibles (CO and H2) in the 100 to 200 ppm range and a minimum of

NOx.

CH4 + 2O2 + 2.72N2 = CO2 + 2H2O + 2.72N2 + ppm CO + ppm H2 + ppm NOx

Too much or too little secondary air gives poor combustion. A minimum excess air

level is required for complete combustion but too much excess air reduces flame temperature

and drops efficiency.[2],[3]

Flame and Effective gas temperature

Flame and effectivegas temperatures are keyvariablesthatneed to be

accuratelydeterminedbeforeanalysis of the heat transfer in the radiantsection of

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firedheaters.Flametemperature is the temperatureprovided by the combustion of a fuel. This

temperaturedependsespicially on the calorificvalue of the fuel. A theoretical or ideal

flametemperaturemay be calculatedassumingcompletecombustion of the fuel and

perfectmixing. But evenwhencompletecombustion is assumed, the

actualflametemperaturewouldalways be lowerthan the theoreticaltemperature.

1.2 Flash DrumPart

In an equilibriumflash process, a feedstream is separatedintoliquid and vapor

streams at equilibrium. The composition of the streamswilldepend on the quantity

of the feedvaporized. The groupsincorporating the liquid and vaporflowrates and the

equilibriumconstantshave a general significance in separation process calculations

Figure1.1.3: Flash drumseperation process

In this project the vertical fired heater and flash drum design were done. For fired

heater the specifications were identified for each segments and design parameters were

calculated in terms of achieving desired process fluid outlet temperature.[3]After heated fluid

exit from the heat exchanger it is feeded to flash drum for separation and then the vapor part

was sent to distillation column.

Figure1.1.3: Layout of the system

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2.CALCULATION

Modeling of fired heaters is generally based on the two section, radiant and convection

sections. When the fired heater was designed, we calculated within each section and four heat

transfer elements. There are the process fluid, the flue gas, fluel gas and air. When fired

heater was designed some important key variables serve help us determination of heater

performance ;

- Heat exchange areas

- Heat transfer rates to the process fluid

- Process fluid and flue gas temperatures

- Flame and tube skin or tube wall temperatures

- Process fluid flow rate

- Fuel flow rate and combustion

- Process fluid pressure drop and the pressure profile in the heater and stack

2.1 Mass Balances

Figure 2.1: Boiling point diagram for methanol/Methyl oleate at 101325 Pa(chemcad result)

The diagram used for to determine the temperature of the entrance and exit liquid and

vapor stream. First of all the exit temperature was determined based on the chemcad result as

564.89 K and this is proved by checking the figure 2.1. Since the reboiler act like a tray

therefore the leaving liquid composition would be in equilibrium with the leaving vapor

composition. The methanol mole fraction in liquid was set as 0.003 and the vapor methanol

mole fraction in equilibrium with liquid phase was found as approximately 0.24.

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Figure 2.2: Equilibrium mole fractions for methanol/methyl oleate at 101325 Pa(chemcad

result)

In order to calculate the inlet liquid and outlet vapor, amount, temperature and

composition, the mass and energy balance were done around the reboiler.

The exit amount, temperature and composition for the reboiler was calculated for

previous project by using chemcad program.

V= The total amount at the exit gas phase from reboiler (kmol/h)

L= The total amount at the entrance liquid phase to reboiler (kmol/h)

W= The total amount at the exit liquid phase from reboiler (kmol/h)

xM= Fraction of methanol at the exit liquid phase from reboiler.

xB = Fraction of methyl oleate at the exit liquid phase from reboiler.

yM= Fraction of methanol at the exit vapor phase from reboiler.

yB= Fraction of methyl oleate at the exit vapor phase from reboiler.

xLM= Fraction of methanol at the entrance liquid phase to reboiler.

xLB = Fraction of methyl oleate at the entrance liquid phase to reboiler.

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Material balance for methanol;

L=V+W (1)

L* xLM= V* yM+ W* xM(2)

W= 37.962 kmol/h

xM= 0.003

yM= 0.24

Energy balance for methanol;

For the energy balance the sensible heat of vaporization and latent heat of the

methanol and biodiesel were calculated.

hL= Enthalpy of liquid entering the reboiler (j/kmol)

hW=Enthalpy of liquid leaving the reboiler (j/kmol)

HV= Enthalpy of the vapor leaving the reboiler (j/kmol)

CpyM= Vapor specific heat capacity of methanol (j/kmol K)

CpyB= Vapor specific heat capacity of methyl oleate (j/kmol K)

CpM= Liquid specific heat capacity of methanol (j/kmol K)

CpB= Liquid specific heat capacity of methyl oleate (j/kmol K)

λM= Latent heat of methanol (j/kmol)

λB= Latent heat of methyl oleate (j/kmol)

qR=Reboiler duty (j/h)

hW= xM(CpM (T-Tref))+ (1-xM ) (CpB (T-Tref)) (3)

Tref was set by considering the low boiling point component. The lowest boiling

point compound is methanol as 337.6 K.

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HV= yM(λM + CpyM (T-Tref))+ (1-xM ) (λB + CpyB (T-Tref)) (4)

At 564.89 K

CpyM= 64501.38 j/kmolK

CpB= 848026.59 j/kmolK

λM= 3.802* 102 j/kmol

λB= 7.18 * 107 j/kmol

Since at 564.89 K methyl oleate was about to changing phase. Therefore while the

leaving vapor enthalpy was calculated the sensible heat of vaporization of methyl oleate was

neglected just latent heat of vaporization of methyl oleate was considered.

In order to calculate leaving liquid enthalpy , just the methyl oleate sensible liquid

heat was considered because at that temperature , the amount of liquid methanol could be

neglected.

hW= (1-0.003) (848026.59 j/kmolK (564.89 K-337.6K))

hW= 190597303.4 j/kmol

HV= 0.24(3.802* 102 j/kmol + 64501.38 j/kmolK (564.89 K-337.6K))+ (1-0.24)

*(7.18 * 107 j/kmol)

HV= 71823997.17k/kmol

qR= 11612.969*106 j/h

L*hL+ qR= W*hW+ V*HV (5)

If the equation (1) , (2) and (5) were calculated simultaneously and by usind the trial

and error method by xLM assumed as 0.17;

L = 120 kmol/h

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While the xLMwas 0.17 by using fig 2.1 temperature was found as 365 K and hLwas

found as 1509532.8 j/kmol . In order to satisfy this result enthalpy of liquid entering the

reboiler is calculated by using equation (3) . Again the liquid enthalpy of methanol was

negligible because of the amount of methanol in liquid phase is too small at that temperature.

At 365 K

CpB=662563.7045 j/kmolK

hL= (1-0.17) (662563.7045 j/kmolK (365 K-337.6K))

hL= 15078391.55 j/kmol

When this two results were compared , they were really closed to each other then the

fallowing results can be considered as correct results.

L= 120 kmol/h

V=82.038 kmol/h

XLM=0.17

2.2. The Energy Balance

We consider at the above parameter but stack was negligible. Energy balance was calculated

in the heater.

Qin =Qout (6)

Qin = Qrls + Qair + Qfuel + Qflu+Qfluid (7)

Qout = Qu + Qlosses (8)

There are source of heat input; the combustion heat fuel, Qrls and the sensible heat of

the combustion air, Qair, natural gas ,Qfueland natural gas when applicable fluid,Qfluid ,heat

flue gas from combustion Qflu. This heat input equal in the radiant section is absorbed by in

the radiant QR , shield Qshld section and heat loss Qloss. Heat loss is by radiation through the

furnace walls. The heat of radiant, QR includes radiation, Qr and convection, Qc.

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Firstly Qin calculated, we set 25% air combustion with fuel gas. The composition of

fuel gas were obtained from the literature and it is stayed in Table 2.1 after that by using

these value the specific heat and ratio of fuel gas and air calculated as 12.9 by using

software program [5] . Different flow rate of air and fuel gas flow rate were tried to

determine how much air and fuel gas need to heat process fluid. Methane and air combustion

equation determine how much flue gas occurs from combustion.

CH4 + 2O2 +2.72N2--- > CO2 + 2H2O+2.72N2

Air% 0,25 Fuel gas in %

CO2 0,08234 CH4 0,8043

N2 0,7179 C2H6 0,0902

O2 0,0382 C3H8 0,0454

SO2 0,00188 N-C4H10 0,0032

H2O 0,15966 I-C4H10 0,002

N-C5H12 0,0002

I-C5H10 0,0004

N2 0,01735

CO2 0,0352

H2S 0,0009

Table 2.2.1 The composition of natural gas and 25% excess air

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For the combustion process, flame temperature had to be determined. Flame

temperature is the temperature attained by the combustion of a fuel. According to methane

flame temperature is define theoretical flame temperature 2000K but actual flame

temperature, 1855. A theoretical or ideal flame temperature may be calculated assuming

complete combustion of the fuel and perfect mixing. But even when complete combustion is

assumed, the actual flame temperature would always be lower than the theoretical

temperature.[5]

Qcombustion=ΣWi*ʃ Cp * dt (9)

Qcombustion = Heat of combustion of fuel based on the gross calorific value.

Wi = Mass of a flue gas component. The five components considered are CO2, N2, O2, SO2

and water vapour.

Cpi = Molar heat of a flue gas component

t1 and t2 = Initial and final temperatures.

Qcombustion=711.42kmol/hr 27.95 kg/kmol 1.40 kj/kg K (1855K-298K)

=43.6 10 6kj/hr

Qlosses=5% Qcombustion [3] (10)

mfuel kmol /hr mair kmol /hr mflue kmol/hr

30,91003274 400 711,4235166

46,3650491 600 1067,135275

54,09255729 700 1244,991154

61,82006547 800 1422,847033

69,54757365 900 1600,702912

77,27508184 1000 1778,558791

85,00259002 1100 1956,414671

92,73009821 1200 2134,27055

100,4576064 1300 2312,126429

108,1851146 1400 2489,982308

115,9126228 1500 2667,838187

Table 2.2.2 Flow rate of natural gas, air and flue gas from combustion

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Qlosses=2.18 106kj/hr

Inlet temperature of process fluid fed at 365 K and boiling point of process fluid is 565 K so

that out fluid temperature must be reach 623 K. The outlet fluid temperature is higher than

boiling point of the process fluid temperature because of there can be some heat losses along

the transportation between the fire heater to flash distillation unit. The reference temperature

Tris set 288.5 K. the specific heat of process fluid was taken by ChemCad

Qfluid= Qoutfluid-Qinfluid (11)

Qin,fluid= N Cpfluid,in (Tin-Tr) (12)

Qin,fluid=120 kmol/hr 577.1 kj/kmol K (365K-288.5K)=5.3 106

kj/hr

Qout,fluid= N Cpfluid,out (Tout-Tr) (13)

Qout,fluid=120 kmol/hr 853.8 kj/kmol K (565K-288.5K) =34.2 106kj/hr

Qfluid =34.2 106kj/hr-5.3 10

6 kj/hr=28.9*10

6kj/hr

Qair= Wair*Cpair*(Tin,air –Tr) (14)

Qair=400 kmol/hr*28.97 kg/kmol*(298-288.5)*1.005 kj/kg K=11.06*104kj/hr

Qfuel=Wfuel*Cpfuel*(Tin,fluel-Tr) (15)

Qfuel=30.9 kmol/hr*19.9 kg/kmol*2.34 kj/kgK*(298-288.5)= 13.73*103kj/hr

Qflue=711.42kmol/hr*1.4 kj/kmolK*(1278K-288.5K)= 9.91*105kj/hr

Qin = Qrls + Qair + Qfuel - Qflu+Qfluid (16)

Qin =43.6 10 6kj/hr-9.91*10

5kj/hr+13.73*10

3kj/hr+11.06*10

4kj/hr-28.9*10

6kj/hr

=11.6*106kj/hr

Heat transfer is affected by radiation in theory by the Stefan-Boltzman law for Qr=σT4

[6]

In design of the fired heater, the mathematical solution of radiative heat transfer is

more complicated however as it involves the calculation of heat exchange factors as a

function of the furnace geometry and the calculation of the absorptivity and emissivity of the

combustion gases. There are also other factors to be considered such as the emissivities of the

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surface and the effects of re-radiation of the tubes. The construction materials were selected

from Table 2.2.3. Temperature of flue gas is 1005 K so 60 Cr and 40 Ni stainless steel was

chosen.

Qout = Qu + Qlosses (17)

Qu=Qradiation+Qconvecrion+Qshld (18)

The tube wall temperature is calculated by the average tube wall temperature [3]

Tw=100 0.5

(19)

Tw=100 0.5

= 347 K

Qradiation=σ α Ac F (Tg4-Tw

4) (20)

α: the relative absorption effectiveness factor, 0.835

F: radiation exchange factor, 0.97 [5]

σ: Stefan-Boltzman constant, 2.041*10-7

( kj/hr.m2K

4) [5]

Ac: Cold plane area of the tube bank (m2) [5]

Ac=Ntube*tube space*length (m2)

Ntube= number of tube in radiation section

Qradiaition=2.041*10-7

*0.835*0.97(52*2.2*0.1)*(12784-347

4)=50.9*10

5kj/hr

Table 2.2.3 According to flame gas, tube material [7]

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Qconvection= hconvection*At*(Tg-Tw)

hconvection: film convective heat tranfer coefficient, 30.66k (j/hr.m2.

C) [5]

At=Area of the tubes in bank (m2)

At=п*Dtube*L*Ntube

Qconvection=30.66*(3.14*0.99*2.2)*52*(1005-74)

=10.15*105kj/hr for radiation section

Qconvection=30.66*(3.14*0.99*0.37)*12*(1005-74)

=39.4*103

kj/hr for convective section

Qshld= σ α Ac F (Tg4-Tw

4)shld (21)

Ac= (Ntube-1)*tube space*length (m2) (22)

Qshld=2.041*10-7

*0.835*0.97((52-1)*2.2*0.1)*(12784-347

4)

=49.9*105kj/hr for radiation section

Qshld=2.041*10-7

*0.835*0.97((12-1)*0.37*0.1)*(12784-347

4)

=18.1*104kj/hr for convective section

Qout = Qu + Qlosses

Qout=50.9*105+18.1*10

4+49.9*10

5+39.4*10

3+10.15*10

5+2.18 10

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Qout=11.603*106kj/hr

Qout was calculated for each different diameter. According to the general energy

balance Qin is equal Qout. Qoutwhich was calculated for 0.099m diameter is equal Qin,

11.68*106kj/hr.

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mfuel kmol /hr mair kmol /hr Q in(kj/hr) Di(m) Space(m) Qout(kj/hr)

30,91003274 400 11681791,88 0,025400051 0,031750064 3482683,905

46,3650491 600 35160980,6 0,050800102 0,055880112 6194455,848

54,09255729 700 45850574,97 0,088900178 0,093980188 10454841,75

61,82006547 800 56540169,33 0,099060198 0,101600203 11681794,01

69,54757365 900 67229763,69 0,127000254 0,132080264 14715227,66

77,27508184 1000 77919358,06 0,152400305 0,157480315 17555484,93

85,00259002 1100 88608952,42 0,177800356 0,182880366 20395742,2

92,73009821 1200 99298546,79 0,203200406 0,208280417 23235999,47

100,4576064 1300 109988141,1 0,228600457 0,233680467 26076256,74

108,1851146 1400 120677735,5 0,254000508 0,259080518 28916514,01

115,9126228 1500 131367329,9 0,279400559 0,284480569 31756771,28

The heat exchanger lenght,diameter,tube diameter and other parameters obey and

satify the standarts. The heat exchanger lenght and diameter were set as 2.7m and 1m. They

were set by looking at the standarts. For different tube diameters the heat duties were

calculated after that when the chemcad heat duty result fit with the calculated result ,the tube

diameter was found. The results were tabulated in Table 3.2

The thermal efficiency of the fired heater can then be written as:

N=(Qu/Qin)*100

N=(9.4*106/11.6*10

6)*100=80.6 %

2.3 Tube inside parameter calculations

2.3.1 Density of the Mixture

The diameter of the tube was calculated in the previous calculations as 0.09m. The

tube lengths were calculated as 0.37m for the convectional part and 2.2 m for the radiation

part.

Table 2.2.4 Calculation parameter

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Liquid density of liquid leaving the reboiler at 564.89 K ;

ρM= Pure methanol density kg/h m3

ρB= Pure methyl oleate density kg/h m3

ρmix= xM * ρM +xB * ρB

ρM= 788.39 kg/h m3

ρB= 656.94 kg/h m3

ρmix= 0.03*788.39 kg/h m3+ 0.997*656.94 kg/h m

3

ρmix= 678.6209 kg/h m3

Molecular weight of the methanol= 32.4 kg/kmol

Molecular weight of the methyl oleate = 296 kg/kmol

Relative mass of the entry liquid= xLM* MWM + xLB* MWB

Relative mass of the entry liquid= 0.17*32.4 kg/kmol + 0.83*296 kg/kmol

Relative mass of the entry liquid=251.1268 kg/kmol

Relative mass of the exit vapor= yM* MWM + yB* MWB

Relative mass of the exit vapor = 0.24*32.4 kg/kmol + 0.76*296 kg/kmol

Relative mass of the exit vapor =232.736 kg/kmol

Relative mass of the exit liquid = yM* MWM + yB* MWB

Relative mass of the exit liquid = 0.003*32.4 kg/kmol + 0.996*296 kg/kmol

Relative mass of the exit liquid =296.084 kg/kmol

The volume of the vapor is calculated from fallowing equation;

Vv=zc*V*T*

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zc= compressibility factor

P=101325 Pa (operating pressure)

T= 564.89 K

R= 8314.34 Pa K/ kmol

zc= 0.9848

Vv= 0.9848*82.038 kmol/h * 564.89 K/ *8314.34 PaKm3/kmol /101325 Pa

Vv= 3745.545 m3

The volume of liquid= W* Relative mass of the exit liquid /ρL

VL= 37.962 kmol/h*296.084 kg/kmol/678.6209 kg/h m3

VL= 16.56 m3

The exit density of the reboiler was calculated as fallows ;

ρexit=W* Relative mass of the entry liquid / Total volume

ρexit

ρexit= 8.01 kg/m3

2.3.2 Friction Loss

Cross sectional area and total cross section area of the tubes the tube numbers were

found in the previous section as 12 for convection part and 52 as radiation part. All the

calculations were done for the optimum values as diameter, lengths of the tubes and the

number of the tubes.

Table 2.3.2.1 Specifications of the tubes

Diameter Length Area Total area

0.09m 0.37 m 0.0785m2

0.0942m2

0.09m 2.2 m 0.0785m2

0.4082m2

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Mass flux and homogenous velocity were calculated as fallows;

G=L* Relative mass of the entry liquid/total area of the tubes

ν= G/ ρexit

Table 2.3.2.2. Mass flux and homogenous velocity of fluid

Total area Mass flux Homogenous

velocity m/h

Homogenous

Velocity m/s

0.0942 319906.7516kg/m2h 39937.45261 m/h 11.09m/s

0.4082 73824.635kg/m2h 9216.335 m/h 2.6m/s

Calculation of reynold and fanning factor;

Viscosity of the fluid = 4.356 m Pa/h

Re=

Fig 2.3.2.1. Fanning factor(taken from Geankoplis C.)

ε was found from Geankoplis for commercial steel

Table 2.3.2.3. Reynold number and fanning factors

Homogenous

velocity

Reynold

number

ε ε/D Fanning factor

39937.45m/h 7344 0.000046 0.00046 0.01

9216.33 m/h 1694 0.000046 0.00046 0.015

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At tube exit, pressure drop per unit length;

ΔPf= 4*ƒ*ΔL/D*ν2/2

Table 2.3.2.4. Pressure drop per unit length

ΔL ƒ ΔPf

0.37m 0.01 72.95 Pa

2.2m 0.015 34.65 Pa

At tube entry, liquid only, pressure drop per unit length;

Velocity= G/ ρmix

Velocity for 0.37m length tube = 471.40m/h

Velocity for 2.2 m length tube = 108.78m/h

Table 2.3.2.5. At tube entry, liquid only, pressure drop per unit length;

Velocity Reynold ƒ Pf Average Pf Pressure drop over the

tubes

471.40m/h 7344 0.01 4.65Pa 38.08Pa 14.35 Pa

108.78m/h 1694 0.015 0.37Pa 17.5Pa 38.5Pa

Average Pf was calculated as ;

(At tube exit, pressure drop per unit length+ At tube entry, liquid only, pressure drop

per unit length)/2

Pressure drop over the tubes was calculated as;

Average pressure drop * Length of the tubes

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Static pressure of the tubes;

νi= 1/ρL

νo=1/ρexit

Ps=

g= 9.8 m/s2

Table 2.3.2.6. Total pressure drop over the tubes and P head

Ps Total pressure drop over the tubes P head

130.48 Pa 144.84Pa 2460.68Pa

775.83Pa 814.36Pa 14631.07 Pa

Total pressure drop over the tubes was calculated ;

(Pressure drop over the tubes+ Static pressure of the tubes)/2

Phead = ρL*L*g

As it can be seen from the table available head is larger than the total pressure drop

over the tubes. It shows that this pressure drop is available for this project.

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2.4 Flash Distillation calculation

Calculate maximum permissible vapor velocity, uperm (m/s)

Uperm=Kdrum

ρl: liquid density, 649.06 kg/m3

ρv: vapor density, 2.64 kg/m3

Kdrum: K factor (m/sec)

Figure 2.4.1 Flash distillation

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K Uperm=K*((ρl-ρv)/ρv)^0,5

0,05 0,78239386

0,06 0,938872632

0,07 1,095351403

0,08 1,251830175

0,09 1,408308947

0,1 1,564787719

Calculate cross-sectional area, Ac (m)

Ac=

V: vapor flow rate, 0.0225 kmol/h

MW: molecular weight vapor, 246.46 kg/kmol

Table 2.4.1. Separator types and K factors[8]

Table 2.4.2.Upermfor different Kdrum

25

Ac=V*MW/Uperm*ρv

2,793655152

2,32804596

1,995467965

1,74603447

1,55203064

1,396827576

Calculate drum diameter, D, and height, h

Table 2.4.3. Cross sectional area

Figure 2.4.2: Flash distillation minimum size

26

The height of the vessel is composed of a number of terms. Droplet settling length is

the length from the center line of the inlet nozzle to the bottom of the mist eliminator. There

is height equal0.75 D [8]. Then height from the bottom of the inlet nozzle to liquid surface is

required to prevent nozzle flooding. The feed in the flash is minimum 0.5D. Another

important term liquid height is minimum 0.25 D.

D( diameter)

feed

bottom(0.75*D) liquidhight(0.25*D)

height

above(D+0.5*D) total height

1,886477239 1,414857929 0,47161931 2,829715858 3,301335167

1,72211023 1,291582672 0,430527557 2,583165344 3,013692902

1,594364265 1,195773198 0,398591066 2,391546397 2,790137463

1,491391207 1,118543405 0,372847802 2,23708681 2,609934612

1,406097114 1,054572836 0,351524279 2,109145672 2,46066995

1,333940848 1,000455636 0,333485212 2,000911272 2,334396484

All of values provide the desired vapor amount and by looking at result the minimum

value height and diameter were selected for flash drum design because it gives the minimum

fix cost.

Table 2.4.4. Process of flash distillation size

parameter

27

2.5.Cost Analysis

For combustion process the natural gas and %25 excess air were reacted in vertical fired

heater. According to excel calculation the required amounts were found as 30.9 kmol/h for

natural gas and 400 kmol/h excess air

The plant operation time was set as 8500h and in this time period natural gas consumption

rate was calculated as follows.

Utility cost

Density of natural gas: 110kg/m3

Moleculer weight=19.99kg/kmol

30.9kmol /h*19.99kg/kmol*8000 = 4941528kg/year

4941528kg/110kg/m3=44922m

3/h

Naturel gas cost(tl/m3)=0.89tl [9]

44922m3*0.89tl /m

3= 39981tl=50609$/year

Fixed costs

The equipments are the vertical fired heater and drum.

For the vertical fired heater the cost was calculated by looking at the heat duty.By using

equipment cost program was found as 531000$ (2007) [10]

531000*590/500=626580$

For the drum cost analysis;

Cp=20000

Bare model cost (CBM)=Cp*(B1+B2*Fp*Fm)

Fp=1

Fm=1.5

B1=1.74

28

B2=1.55

Cbm=20000*(1.74+1.55*1.5*1)=81300$

The total fixed costs=626580$+81300$=707880$

Cost of manufacturing(COM)=0.280 Fixed Costs+ 1.23*(Cost of raw material+cost of

utilities)

COM=0.280*707880+1.23*50609$=260456 $

29

3. RESULTS AND DISCUSSIONS

Table 3.1. Process data sheet for box-type fired heater

Unit process conditions

Process fluid Methanol+methyloleate

Fluid flow rate (kmol/h) 120

Inlet temperature 365

outlet temperature 623

ınlet pressure(atm) 1

outlet pressure(atm) 0,9733

efficiency 0,80

Fuel characteristics

Type of fuel Natural gas

Nett calorific value (kJ/kmol) 927844.41

Molar heat (kJ/kmol.K) 2.34

Temperature (◦C) 25

Flow of fuel (kmol/h) 30.9

Molecular weight (kg/kmol) 19.99

Composition (% mol)

CH4 (80.43), C2H6 (9.02), C3H8

(4.54),

iso-C4H10 (0.20), n-C4H10 (0.32),

iso-

C5H12 (0.04), n-C5H12 (0.02), CO2

(3.52),

H2S (0.09), N2 (1.735).

Air characteristics

Molar heat (kJ/kmol.K) 1.005

Flow of air (kmol/h) 400

Air temperature (◦C) 25

Percentage of excess air 25%

30

The chemcad results gives the heat duty as a 11200000kj/h, the heater length was fixed as 2.5

m and with different tube diameter the heat duty was calculated. Radiation heat duty was

added to convection heat duty. In Table 3.2 QR+Qshield* parts shows the our term by term

calculation. However QR+Qshield part was calculated by making over all energy balances

around the heater and then these terms were taken from the equality . The very close values

were obtained for tube diameter 0.09m. The optimization was done by considering the heat

duty since costs were not change in this process .There was an desired heat duty and we had

to provide this value by using combustion heat duty.

Table 3.2. The optimization of the process with various tube diameter

The number of total tubes was 64, pressure drop was also found as 0.027atm.it is very small

value which means there wont be any problem in distillation column because columnt is

working at 1atm.The QR+Qshield values are so close and also these values provides the

chemcad reboiler heat duty which was 11200000kj/h.

Tube

diameter(m)

Tube

space(m)

# of

Tubes

pressure

drop(atm) lenght(m) QR+Qshield* QR+Qshield

0,025400051 0,031750064 112 0,0014 2,5 3488712,289 11651594,8

0,050800102 0,055880112 95 0,0082 2,5 6205937,611 37012276,5

0,088900178 0,093980188 76 0,0134 2,5 10474647,34 47156549,1

0,099060198 0,104140208 64 0,0267 2,5 11612969,93 57300821,8

0,127000254 0,132080264 57 0,056 2,5 14743357,06 67445094,5

0,152400305 0,157480315 39 0,067 2,5 17589163,54 77589367,2

0,177800356 0,182880366 25 0,07453 2,5 20434970,03 87733639,8

0,203200406 0,208280417 19 0,08965 2,5 23280776,51 97877912,5

0,228600457 0,233680467 14 0,0965 2,5 26126582,99 1080221853.

31

Table 3.3.Final design parameters

Fired heater

Inside diameter of the tube, (m) 0,09

Tube spacing(m) 0.104

Tube Length L (m) 2.2

Number of tubes Nt 64

Heater dimeter(m) 1

heaterlenght(m) 2.5

pressure drop(atm) 0.026

construction material stainless steel

Flash Drum

height(m) 2.33

diameter(m) 1.33

.

32

4.CONCLUSION

A fired heater is a direct-fired heat exchanger which uses the hot gases of combustion

to raise the temperature of a feed flowing through tubes. The process fluid is first heated in

the convection section preheat coil which is followed by further heating in the radiant section.

In both sections heat is transferred by both mechanisms of heat transfer, radiation and

convection, where radiation is the dominant mode of heat transfer at the high temperatures in

the radiant section and convection predominates in the convection section.

In this project a fired heater and the flash drum design were done. For fired heater

design, the natural gas and the 25%excess air were reacted in order to obtain efficient

combustion heat for the system. The heat duty was found as 11612969,93 kj/h. The heat

which was calculated for the providing enough heat in order to raise the outlet process fluid

temperature. The temperature was found as 623K. The tube diameter is 0.09m, number of

tubes for radiant part is 52, for convective part is 12 were found. The heat exchanger diameter

is 1m and height is for radiant part 2.2m, for convective part is 0.5m.The pressure drop

during the tube was calculated as 0.027 atm. After fired heater design the flash drum design

was done. The height of drum was obtained as 2.33 m, the diameter of the drum was

calculated as 1.33m. The total cost was calculated as 260456 $ in terms of operating costs and

fixed costs.

33

5.REFERENCES

[1]Online Resources, web link, [Available 2012]

http://media.wiley.com/product_data/excerpt/10/04713217/0471321710.pdf

[2]WildyF.,Fired Heater Optimization, AMATEK process Instruments,p:1-6

[3] Al H.,IbrahimH.,Fired process heaters,Al-Baath University, p:327-329

[4]Sınnott R.,TowlerG.,Chemical engineering design,Elsevier, p:934

[5] Online Resources, web link, [Available 2012], http://heaterdesign.com/design0.htm

[6]IncroperaF.,DewittD.,Bergman T., LavineA.,Fundamental of heat and mass

transfer,Wiley,sixth edition.

[7]OnlineResources,weblink,[Available2012]

http://kolmetz.com/pdf/ess/PROJECT_STANDARDS_AND_SPECIFICATIONS_fired_heat

ers_Rev01.pdf

[8]BubbicoR.,Gas liquid separators,“Sapienza” University of Rome

[9] ] faturalandırmaveenerjipiyasasıdüzenlemekurumu

[10]OnlineResources,weblink,[Available2012]

http://www.matche.com/EquipCost/Crystallizer.htm

34

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