Verslag behorende bij het fabrieksvoorontwerp: Waterstofproductie

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Contents lA. Summa:ry lEe Conclusions

2. Introduction

3.

4.

5.

6.

7.

starting points for the design Process Flow Diae;ram

Process conditions

Calculations of equipment reforrner

high temperat"ure shift converter

CO

2 removal system

metha..."lator

the overall energy balance

compressibility factors and fugacity coefficients pumps

Mass and heat balarrces References Appendix 1 pg 1 2 3 4 6

7

11

17

23

37

42 45 47 48

56

Computerprograms on reformer, hi~l temperature shift converter and methanator

Appendix 2

ComputerprogTams concerning calculations of enthalpy (subroutines)

and compressibility factors Appendix 3

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lA. Surnma!X,

A heat balanee , lnaterial balanee and some equipcnent elesign calculations for a hydrogen plant suitable for feeding a 100,000 tons/year cyclohexane

plant were perfo:>::mcd and are presented in this repo:d. Applying a benzene

hydrogenation route for the cyclohexane product ion anel assuming 330 operating

t

days per calender year the hydrogen plant's capacity was established at 450 :.

~

...,

~

kmoles H2/hr.

;c.

~) The plant' s feed is composed of two streams. The main feed stream is

r .

}.\...vf

Dte

purge gas from the cyclohexane plant having a typical composi tion of

v~

49

%

H

2 and 51

%

CR₄ (17) available at the asswned and the other is au LPG streaIll (liquid p:r-opal1e and

o

temperatuxe of 100 C, butane at 20 oe) with a

butane content of 0 to 100 mole percent to provide some operating flexihility. As the only purpose of the H

2 plant is feeding the cyclohexane pla..nt,

operation at 100

%

LPG- feed was not considcred.

A process scheme without J.OVl temperature shift conversion was adopted, (will resuH in a product purity of approximately 90

%

),

because hydrogen purity is not a critical variable in the cylohexane production as cyclohexane

plants are designed to ha..ndle hyclrogen feeels with hydrogen content as low

as 65

%.

The scheme employs a feed treating section, a reforming furnace

equipped with vertical, catalyst fillcd, tubes, a fixed bed high temperature

shift converter, a CO

2 removal system and a fixed bed methanator. The heat recovery equipment has been incorperated in the process scheme.

The most important operating variables at the reformer's outlet were

o

chosen as follows: Temperature and pressure of 770 C a..nd 25 ata

respecti vely and a steam to carbon ratio of 6 (molecules steam/ at om carbon). Such operating conditions will result in a hydrogen purity of 91

%

(dry basis).

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-IB. COITclusions

Hydrogen of sufficient purity for cyclobexane production may be

produced by the steam reforming of light hydrocarbons at moderate pressures

without the applicatioTIl of 101'1 temperature shift conversion. If a high

pressure cyclohexane process operating at

35

ata is chosen

(17)

and if

product puri ty has to be maintained at the same level, a choise must be

made between the installation of a booster compresser bet ween the two plants

0

;

the modification of the H

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3 -2: Introcluction

Hydrogen is one of the basic feed streams in the petrolerun and

~chernical industries. In the petroleum industry it is used mainly as feedstock in the various desulfurization operations encountered iru crude oil refineries. The source of hydrogen there is the H

2 rich off-gas leaving gasoline reformers. As sulfur specifications in fuels become strict er it is obvious that more hydrogen plants will be required in the future.

Hydrogen is further used in the product ion of ammonia, synthetic natural gas (coal gasfication) and the hydrogenation processes of the

~(;hemical industry.

steam reforming of light hydrocarbons ranging fror:J. natural gas to li@lt naphta has been applied successfully in industry for large scale

hydrogen production. In these processes hydrocarbons are mixed with

stearn at an appropriate ratio, preheated and passed through externally fired tubes packed with a nickel catalyst at temperatures of 650 to 900°C and pressures of 5 to 40 ata (3), (27). The hydrogen rich gas leavL~g the refonner is further processed for 00 conversion, CO

2 removal and methanation of the carbon oxides stillIeft. Suitable adjustment of the various variables will reffillt L~ a hydrogen product purity of 90

%

to

>

99

%.

The main reactions through which hydrogen is formed L~ the reformer are: CH 4 + H20~ 00 + 3H2 1 CH 4 + 2H20~ C02 + 4H2 2 00 ₊ HO~ C0 2 + H2 3 2 ~u.,r u\',a~.(1-~-~, ~ C0₂ + CH 4 ~ 2CO + 2H2 4

When the pl~~t's feed contains hydrocarbons heavier than methane i t has been reported that hydrocracking takes place and that even aftel.' very short contact times with the catalyst, methane is the only hydrocarbon which can be detected (12) and therefore one can proceed with calculations considering only methane.

In the shift converter reaction

3

occurs while in the methanator reactions 1 and 2 occur.

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-3. S~ting point s for the desigg

a Plant eapaeity was fixed on

450

kmoles/hr hydrogen which is suîfieient for the production of

100,000

tons/year of cyclohexane from benzene.

A plant sImt dovm period of

35

days/year VlaS assumed vlhieh resulted in

330

operating days per year.

b The plant feed consists of the cyclohexane plant purge gas composed of

49

%

H

₂

and

51

%

CH

4

at

100

°c and a mixture of propane / n-butane

t . .

0

100

ol b t t b ' t t t of

20

°C.

con alnlng - ;0 u ane a a n a verage am lcn - empera ure

For the CO

₂

rc~moval system a

30

%

by weight K

2C0

₃

solution is required.

---d The hy---drogen product has a composition of approximately

86

%

H2'

9

%

CH

4

and

5 %

water

(

91

%

H

2 on dry basis). '11he total aInOlmt of carbon oxides present is less than

10

ppm (volume).

e Waste streams:

f

g

(X)2

removal system of gas. Contains approximately

26

%

C0

2

a.nd

74

%

H2

0.

The maximum amount of C0

₂

vented is

4350

tons/year (butane feed).

A small K

₂

C0

3

-

XJI(X)3

solution stream (leakages etc.) from the C0

2

removal system.

A solid waste streaT!1 in the form of ZnS from the feed treating facilities.

The amount depends on the sulfur content of the LPG feed.

The hydrogen plant being a steam exporter will require facilities for

the supply of boiler feed water.

The physical constants of the various cornponents encounted in the process are as follows: T k

[K]

Pk [at a] M H

₂

_33.3

_12.8

_2.016

CH

4

190.7

45.8

16.04

C

3H

8

370

42

44.09

C

4

HIO

425.2

37~5

58.12

(X)

_133.0

_34.5

₂

₈

_.01

(X) 2

304.2

72.9

44.01

H

2

0

647.4

218.3

18.02

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-h Corrosion:

Norrnal carbon steels are not sufficiently resistant to attack by the

hydrogen rich, high temperature environment of the H

2 plant and

therefore special alloys will be required for the piping and equipment.

c-( - Mo steels were found to be suitable for this application. In

order to determine tJle concentration of

Cf'"

en Mo required, which is a

function of the hydroeen partial pressure and the temperature, the so

called N81son charts can be used. For more details a reference may be made to art icles discus:::ing this subject (29), (30).

COI-rosion in the K

2C0₃ .. KHC0₃ C02 removal system can be drastically

reduced by adding ~2

%

potassium di.chromate as an inhibitor (24). In

locations where high velocities occur (28) recommends the use of

stainless steel.

In andition precaution for the prevention of corrosion must be taken

wherever hot condensate saturated with C0

2 is formed. Such condensate

is kno,vn to be extremely corrosive for nonnal carbon steel so that the

use of stainless steel is recommended.

i The toxity limit of CO is 100 ppm (36).

The fla .... nmable ranges (% volume in air) of the hydrogen product and the

raw materials used for its production (in air at 1 ata) are as follows

4 - 74

%

5.3

- 14

%

2.2 - 9~5

%

.'j _{0 0 0} t""\

High Pressure

~

350

CyclOhexane Plant Purge Gas

~.

CV

10

I

[§]

Flue Ga. to Staek

e-

X

H 1 LPG VAPORIZER ~ 7 HIGH TEMP SHIFT CONVERTER H13 P2 LPG FEED PUMP H 8 HEATRECOVERY EXCHANGERS P14

R3 DESULFURIZER IH 9 REGENERATOR REBOILER T1~

H4 HEATRECOVERY COl LS

r

10 GAS COOLER H~6 R !l REFORMER V11 CONDENSATE DRUM

HO WASTE HEAT BOl LER T12 ABSORBER PH

"'" ,.-, ~ t'.

"'"'

~ ~ 20

~

(SB

~

H16 ...

~

₂₃₂ ....--'---... ~

LEAN SOLUTION COOLER

I

H18

LEAN SOLUilON PUMP V19 REGENERATOR H20 METHANATOR FEED/EFFLUENT

I

R21

EXCHANGER V22

WATER RECIRCULATION PUMP

Icw.

~

. . i./i\J 15 16 '.1" H9 OVERHEAD CONDENSER OVERHEAD DRUM START UP HEATER METHANATOR DRYER

PROCESS FLOW DIAGRAM

of HYDROGEN PLANT

~rCYCLOHEXANE PLANT

N. Yahya en A. RIJk~boer Juni 1976

o

Streamnr. [iJ Temp. In 'C 0~ In at~ CW.CooIln~ Wat.". BF.W .• Bolier ~ wat.,..

,-.. ~ ~ ~

~

I ~ I

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7

-5.

Process conditions

As was mentioned before a flow scheme without low temperature shift

conversion was chosen. The application of such scheme reduces significantly

the cost of catalyst required for the plant and results in a sjmpler process. In order to proceed wi th thc heat ~nd material balanee calculations some

process variables must be fixed. The three most important variables in

the production of hydrogen are the reformer out let pressure ~nd temperature

~nd -Lhe steam to carbon ratio of Jche gas mixture.

The reformer outlet pressD~e depends to a certain extent on the

operating pressure in the cyclohexane plant. The latter was 35 ata for a

process applying a single reactor (17) while IFP has developed a low

pressure catalyst reducing the operating pressure to 2.9 ata but requiring

two reactors (18). Thc best choise of a cyclohexffile process is not evident

and many factors such as operating costs, pl~nt investmen-t, catalyst cost,

royalties and cyclohexane losses (which increase at lower operating pressure

and at lower hydrogen p-Ll:r..'i ty (19)) must be considered first before a choise

can be made.

The influence of the reformer out let pressure and temperature on the

product purity is discussed in (27), anrl (3) mentions operating pressures

up to 40 ata. WDile a low operating pressure will result in large equipnent,

a high pressure plant (rv40 ata) will require a higher reformer outlet

temperature, to maintain a constant product ~lrity, and this in turn will

result in thicker reformer tubes. (Note that the reforming reactions being

endothermic and accomp~nied by an increase in volume, are favored by low

pressure and high temperature ). At this point it was decided to select a

moderate reformer outlet pressure of ~ta~

The reformer outlet temperature 'was limited to 770 oe to avoid high

-..;.--tube temperatures. The operating life of the tubes will increase and with

it the plant reliability while a reasonable product purity can be still achieved.

As far as the steam to carbon ratio is concerned the limiting factor

in the select ion of ~n appropriate value was not the danger of carbon

forma-tion in the reformer tubes but the fact that a sui table CO conversion

had to be achieved in a single high temperatUl'e shift bed in order to keep

the temperature rise in the methanator within reasonable limits~ A ratio

of 6 molecules steam per atom carbon resul ted that way which is the maximum.

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S -NJuilibrium constants:

The calculation procedure adopted here requires the equilibrium

constants of the followine; two reactions:

eH 4 + H

20::;::?': eo -I- 3H2

eo -I- H₂0 <~

co 2

+ H

2

Their equilibrimn constants can be wri tten as

i \

3

Peo'

%

2 and

Peo . 1\-1

2 2 Peo'~ 0 2

Tabulated values of Kpl and K

p2 versus temperature fron (4) were converted,

using a least squares method, to polynomials of the form

lnK

p a + biT + cjrr

2

+ •••

VIhere T is the absolute temperature [KJ. The followLng equations Vlere

established that way: lnK pl 31.1 - 2.S5·10 4 /T + 7'105/T2 315 ~ t:Ç 930 lnK p2 - 3~3S + 3.06·lÓ3/T + 7.4:10 5 /T2 lOS /T 3 205.{;, t ( 538 lnK p2 '" - 2.33 + 1.6'l02/T + 3040'l06/T2 9.2-'lOS/T 3 53S~ t~ 980

As ~Qll be seen later the various gas strearns in the plant can be considered

as being ideal and therefore no fugacity coefficients are req-uired~

Enthalpy correlations:

Heat exchangers' duties and reaction heat effects were both calculated

by means of enthalpy'correlations based on entha.lpy data given by (l)~

oe oe oe

(1) lists the enthalpies of H2' CH

4' e

3

HS' e4~0' eo, e02, H20 as a function

óf tempera:l;ure élJla_ pressure relative to their elements' enthalpies at 0 Kelvin

and 0 ata. Expressing enthWpies in sueh a manner is convenient because if

all components have the same zero state the reaction heat effects can be

calculated from the change of composition taking place during the reaction~

By observing the tabulated enthalpy values it was concluded that

pressure has a negligible effect on the enthalpy of the normal1y highly

superheated eomponents so that enthalpies were evaluated at a constant

pressure of 2005 a.ta (300 ps ia) except for e

3HS and e4H10 that occur as

saturated vapor and as pressure influence becomes more significant, then.

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-The enthalpy data vrere convereed, using a least squares rnethod, to polynornials of the form

2 .

h = a + bt + ct + •••

The following equations were established:

h = enthalpy

[~~u~

f = temperature [oF]

hydrogen:

100 ~f(. 600

The enthalpy of hydrogen at higher ternperatures and at a pressure of 0 ata

are given by (2). These data were t:reated in the same rnanner neglecting

the inîluence of pressure. The result was,

H = enthalpy [cal

J

_ Lgmole methane: T = temperature

[KJ

h

= -

1572~1

+ .5006'f + 2.292'10-4

'r

2 h

=

-

1531.0 + .3741' f + 3.G70'10-4 'f2 - 5.278'10-8'f3 propane: 500~T (1200

----

.

100~f~ 1000 1000 <, f"< 1600 h = -

761~7

+

.7

406

'î -

4.993'10-4 'f2 + 7.52'10-7 'f3 -

2~89'10-10'f4

180 (; î!(

600

n-butane: h

=

-

810.4 +

1~4242'f

- 2.1089'10- 3 'f2 + 2.3988'10-6· f 3 - 9.076 10-10 f4 270 «f<; 700 carbon monoxide: h

=

-

1635.1 + .2509'f + 3.28·10-6'f2 + 5:84'10-9 'f3 h = - 1626.7 + .2257'f + 3'10- 5 'f2 -

4~17'10-9'f3

carbon dioxide: h

=

-

3775.2 + .2270·f + 2.033·10- 5·f2 + 6072'10-9 'f

3

h = - 3787.1 + .2483'f + 1.738·10-5'f2 100 <f( 1000 1000<f<1600 100.( f~ 1000 1000 ~ f< 1600

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(' - 10 -water: superheated vapor: h - 5727.4 + 1.1637'f - 7.975 10-4 'f2 + 3.298'10-7'f3 h - 5351.3 + .0511'f + 3.467'10-4 'f2 -

7~81'10-8'f3

saturated vapor: h

= -

5560.4 + 1.5268·f - 6.1525·10-3· f2 + 1:33394'10-5' f3 -420~ f~ 1000 1000 ~ f~ 1600 - 1. 09621' 10-8, f4 100 <, f:( 700 saturated liquid: 6 -3 2 . 6 -6, 3 h = - 553.0 + .2831·f + 4.0154'10 ' f - 8.884 '10 ' f + + 7.0720'10-9'f4 100(f(700

The enthalpy of a gas mixture was assumed to be equal to the sum of the cornponents enthalpies at the mixture temperature multiplied by the component flow:

H

= "'h.·

_L ~ .

1 1

where h

r

energy

l

~it mas~

~

_.

r

_b.m

mass it tim=1

l

In cases where gas streams at relatively 10w temperatures are involved (water exists only as liquid) the water vapor enthalpy was assumed equal to that of saturated steam enthalpy at the gas streamternperature.

As enthalpy calculations are of ten repeated two CPS subroutines

.tJ\

• p"<~ .. ~\ calculating a mixture' s enthalpy and performing all the necessary units

.\'J ,.~

i~r conversions were wri tten.These are "enthv" which calculates the enthalpy at both the higl"l and low temperature ranges and "enth" which calculates the enthalpy at the 10w temperature range. Using thè same correlations mentioned above two extra subroutines were used; "enthre" calculates the ~11t of the shift reaction and the subroutine "denth" calculates the first derivate of the enthalpy with respect to the temperature ~ dH/dt. The latter is required when the Nemon-Raphson convergence method is appiied during iterations. All four subroutines are presented in appendix 2.

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11 -6. Caleulations of eguipment Reformer

The cyelohexane plant pu.rge gas entering the H

2 rJlant at 100 °c is

mixed with the 136°C n-butane vapor from the LPG vaporizer to form a

combined hydrocarbon feed at 123 °C. This hydrocarbon stream is preheated

to 360 °c in the conveetion sedion of the reformer furnaee and treated for

sulfur removal as the reformer nickel eatalyst is poisoned by sulfur. The

organic sulflLr eompounds are convert,ed to H

2S anu a hydrocarbon on a

cobalt-molybdenum hydrosulurizatiorL catalyst bed (4) after which H

2S is

eliminated on a ZnO bed aceording to the reaetion:

ZnO + H

2S ---tl'-- ZnS + H20

The sulfur free gas is mixed wi th the required amount of superheated stearn

at 540°C to give the desired steam to carbon ratio of 6 and a combined

refonner feed temperature of 505 °c in accordance vvi th practical values

(10), (12).

It should be mentioned at this point that the exact reforrner feed

composition is not known as the hydrogen product purity has not been ;

I I "I

established yet. The procedure which was followed Vlas to assume the product \ \~ i

dry gas molar flow rate and i t ' s purity, to add the required amount of LPG

~

and to compute the reformer outlet eomposition. Then the whole plant is

calculated and a new value for the dry gas flow rate and purity is

established. '1'h1S procedure is repeated until both assumed and computed

values match.

(

To compute the reformer outlet composition a basis of 1000 kmoles/hr carbon

in the feed is ehosen. The reformer feed ean be expressed then as CH + rH 0

m 2

where m is it's hydrogen/earbon ratio and r it's steam/earbon ratioo As

was explained earlier itean be safely assumed that the only eomponents

present at the reformer outlet are H2' CH

4, CO, C02 and H20. If the outlet

temperature and pressure are fixed and equilibrium is assumed the outlet composition is fixed and five independent equations relating the five unknovm components' flowrates will be required for a solution. The overall carbon, hydrogen and oxygen balances provide three equations and any other two

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-As p1lblished equilibrium data vs. the ,temperature for the reactions CH

4 + H20 Ol! '" CO + 3H2

CO + H

20 ... C02 + H2

and

are available

(4),

their equilibrium constants expressions were chosen to

provide the two extra equations.

3

~mOletCO'~~ole,H2

2

PCO'PH P

K 2

pl

PCE . PH

°

~ mole,CH .~ .

(I~mole,i)2

4

mole,H20

4

2 PCO . PH ~ " ~ K 2 2 mole,C02 mole,H2 p2 PCO' PH 2û

o

mole,CO

'0

.mole,H20

Expressing all molar flow rates in terms of ~ _{mo e,} 1 CO and ~ _{mo e,} 1 C02 the following expressions (valid for the reformer outlet) can be derived:

~ _.mole,CH

4

m 2 - 2 + 3~ .mole,COmole,C0+

40

2 1 - ~ _{,ffiole, CO}

~

_L

~

_{. mo} 1 _e, . = m_{2 - 1} + r +

2~

1 CO +

2~

1 CO 1. ,mo e, . mo e, 2 H balance C ba12ulCe

°

bala.l1ce

where m is the ITJclrogen/ carbon ratio and r is the steaTJl/ carbon ratio.

Substitution in the expressions for Kpl and K

p2 and the'simultaneous solution of the resulting two equations in ~ 1 CO and ~ 1 CO will give the

. mo e, mo e, 2

desired outlet composition.

A computer program which calculates the LPG feed rate, solves the outlet cOIaposition and calculates the enthalpies of the in and outgoing streams was developed and is described in appendix 1. This program prints out values on the basis of 1000 kmoles/hr carbon in the feed. Values for the

actual plant capacity of 450 bnoles H

2/hr can be calculated by using a multiplication factor of

0.17.

The printout sheet is presented in the next page.

eguilibrium

Rather than assuming a complete cheIDlcalJ\the concept of "approach to equilibrium", expressed in degrees Celcius, is of ten used to describe the displacement from equilibrium. Au "approach" having a cerlain value means

"') .~ .... D ... .... 0 ... 0 ... .

,..,

. ... .

,..,

rcWf) "1 "ö(}";;11"r;·I(ïÖ;5;··~(i·~·~!l;7"1 ·ti;2"Ç"··· H[ f'T IHn"'fii' {') F()~"AF!'"I ... 8,'.S IS hr Ci'rh 0r'\ i·r, ir ~.~:g.~ .. ~.~./.!~T ... m0 1 !~ fri'C _{.. k.r1'<:'~r.~Ir.r. .... "..·.~ .. l .. ". ... J.r '" ('} 112 ·25C\.!;] r'I.(1o:~rt:; U)Q.ri n.~177]

r::i t, ... ···Z7ï'i .. ;·r .. (i·(';ï\''iir.? ? ····2tii':·iS···r.;·;.;;;ijïit··· ... .

:':3 c.r.r n.rrrr" tç·· .. · .. ···· .. ···· .. ·· ... · ... ·· .. · .... · .. ·· .. · .. ······· .. ·HT ... sëï· ... ··· .. 0:6in·_ë;··· .. · .. ··· ... . co (i,cr o.:,rrrr 2GJ.~C' (\ .r. '>:J r ~ ~··:··r(iii;i:'···· c 02 ...... n;tïr··· ·(i;(){h·,~!i ... 5~Ç~:~{ 1;2 Cl Gn0r.. rr 'Yn f" rr r.n ...•...•... ; ... ; ... : ... ; ...•...•... ,. ...•....•... TOTAL _{.................}G _.....711. _..........'1_....1 _... nrrrr I'Irrrr T "'l T !.:.~ ... ~.;::/.!.~.:. ... _ ... :. .J. ~:.:: .. 7 ..• '.: v • ~ ~ ;. .. ~~ r. .... .... . ,.rn!.HHp..Y ... ~.;.I~c:~lJ~E -èr.5JI'I? -? nr ,~:> ... , ... . l!O.f " bi'li1no::c lGrr.rr

'I ·::;i;i·i1n·c:ë:;···_···_···_···,.I·,:.:ï;·~:;:·.; .. p.j ....... , .. . rnr.,r,r

<> r>F.~.o i1 c:.h ... ~() COl' i 1 r~ri l:'.~ ... r()L. ti, f' r (' f (' r"' jr:~ .. r..~<:.c;:.r, Î'; n : 2. ? .??~.r" (' : ... n •

,...., . Er.. (.r.(~f9.r.~:i.r1[,l..:::. .~r..~ r, 7 5 i1 t 7 I, ? .. f .... <.<:'.~

.

.

..

.

r ... . .. n • •• D: ... .L~.I:if.P" J • , 5 (; ." t. . .}?.~ .. ~.0. ... ~ r r: r V TIOS _.......fT _.........(lUTLET _{............... C}f1 /C O,~r.lCl .... n •• • ••••••• ••• • ••• • ••••• •

r

r 4 ..-.h77S35... G, 5GS9757E-~~ ...-...~772J~2r1522n~E... , ... , ...-r~ ... . t"'I -., ,...,..., ,... "-Lv

1-I

c

c

c

( (,

c

I "

r

r

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~ 14

-that the equilibrium constant must be evaluated at a temperature lower than the actual operating temperature in the case of endothermic reaction and higher in the case of exothermic reaction the difference being equal to the

value of the approach.

( An approD,ch of 22.2 oe (40 oF) was assumed for the reforming reaction in agreement with published values (3), (14) while the shift reaction was

assumed to reach equilibrimn (zero approach).

)-

'\

I

The nickel oxide catalyst used in the reformer is paclced in centrifugally I

cast chrome-nickel tubes having an internal diameter of 5 to 20 cm (2 to 8 inches) and lenght of 3 to 12 meters (10 to 40 feet) (4). For this plill1t a tube inside diameter of 7.5 cm, to reduce the temperature difference across the tube' s wall (14), and a tube l enghr

of8

-

;-eter~-;;-re

seleded.

..r----..,....,...--..--Small diameter tubes are acceptable here as the plant capacity is relatively

low (16).

The conclusion of the literature SDTvey was that the heat transfer rate rather than the reaction kinetics determine the nU!llber of tubes and therefore

...ç.--the catalyst vollLrne. The most commonly used valuc for the heat flux is

2 - -:--2: --- . '- 2

54,000 kcal/m hr (v20,000 BrU/ft hr, N63 kW/m ) (11), (16) related to the

tubest _{internal surface. This value was also used here. The catalyst filled}

tubes are placed in the radiant section of the furnace and therefore the duty given by the compu:ter program is equal to the radiant duty of the furnace.

From the print out sheet,

radiant duty

=

49.55'10

6

.•

17

=

8.42°106 kcal/hr (= 9.63'103 kW) 8.42'10

6

2

54,000

= 156

m

tubes internal surface

,

internal surface/tube = r(' .075' 8 1.88 m 2

number of tubes required 83

The problem of carbon formation in the reformer tubes is discussed by

(13) and (31). (13) presents a plot of the equilibrium constant

Kn

vs~ the temperature for the carbon forming reaction

(X)2 +

e

-4 ~ 200 2 p,yOO yOO 2 where ~ 1 1

( I

c.

( ( C)

0

()

o

() ( \

<-( ( ( ( ( ( (

o

o

If - 15

-(

P'Yeo

_Yeo

2)

<

2 reformer outIet

KB (at the reformer outIet temperature)

then no carbon formation is expected. Applying this procedure using the reformer effluent composition shows that no carbon laydown is expected.

and

Y

_eo

2

~39

<

4

(31) presents ternary diagrams for the system H - e - 0 with carbon formation

isotherms which separate the carbon formation and carbon free zones. The reformer effluent having an elemental composition of 68.8

%

H, 26.8

%

0 and

4.4 %

e lies in the carbon free zone.

As both tests do not indicate carbon formation it can be concluded that

/ : -.. -...~--~

--operation under the conditions mainta:ined :in the reformer is safe.

The calculation of the furnace fired duty is performed in the following manner. First the ratio radiant duty/fired duty (so called radiant

efficiency) must be established. It'svalue depends on the geometry of the furnace and it's estimation involves elaborate heat transfer calculations.

(15) presents four furnace desi~1 studies with radient efficiencies ranging

fr om 38

%

to 48

%.

For the calculations here the average value of the four

(42 ra) was used. Thus,

fired duty 8.42'10

6

-.42 6 . 3 20.05'10 kcal/hr (= 23.32'10 kW)

The overall thermal efficiency of hydrogen plant furnaces is given by

(14) as 86

%

at a flue gas outlet temperature of 260 oe. Using this value,

e heat absorbed in the convection section ean be ealeulated:

total heat absorbed = 20.05.106'.86

=

17.24'106 kcal/hr (= 20.06'10 3 kW)

heat absorbed in the eonveetion seetion

=

(17.24 - 8.42)0106

16

-(

The overall heat balance for the furnace becomes then:

kcal/hr kilowatts

( radiant duty 8.42'10

6

9.80-103

steam superheating 3.31'106 3.85'103

hydrocarbon preheating .49.106 .57 '103

( 12.22'10

6

14.22.10 3

.: ,,'1.'

available for recovery 5.02'106 5.84.10 3 /'" 'I,v- 'J~

total heat absorbed 17.24.106 20.06.103

( losses 2.81'10

6

3.26'103

total heat fired 20.05'106 23.32.103

(

(

c

o

c,

r

17

-The removal of 00 from the reformer effluent gas must be accomplished

in order to increase the hydrogen product purity and to limit the temperature

rise in the methanator where carbon oxides are converted to methane. Compared

with the direct removal of CO, it 's conversion to C0

2 followed by C02 removal

is advantageous as hydrogen is produced simultaneously. The reaction is,

co

+

_H20

4 . . CO

2

+

H2

CO shift converters have been applied successfully in industry in the

temperature range

350

to

450

°c

(high temperature shift) and

200

to

300

°c

(low temperature shift)

(4).

Low temperature conversion is applied when low

00 concentrations are desired as the shift reaction b'eing exothermic is

favored by low temperatu~es.

The reaction takes place in a fixed bed reactor with an iron-c~omium

oxides catalyst for high temperature conversion and with a copper-zinc catalyst for low temperature conversion.

As mentioned before the process' variables were chosen in such a way

that the desired à.egree of CO conversion can be affected in one high temperature converter. With a chosen reactor inlet temperature of

350

°c

the reforrr,er effluent must be first cooled in a waste heat boiler producing

steam at the SaIne timeo

The kinetics of the shift reaction has been studied by many researchers, for example,

(20), (21), (22)

and

(23).

Of the four refe~ences only the kinetic relations presented by

(20)

and

(21)

were used in the calculations.

(20)

proposes a kinetic expression having the form

r-

_{\ (}

n

'

('-Where C. J. k K p =k'(C 'C -00. ~~

concentration of component i

[~~leJ

rate consta.'1t

b

m 3

ruoJ

kmole

( ( ( ( ( ( (

o

o

o

18

-The shift catalyst activity undergoes a rapid clecline during the first

days of operation af ter which the decline becomes JflOre graduale It is

therefore clea.r that a "design" rat.e constant rat.her than a rate constant

based on 12.boratory experiments must be used in order to determine the

required catalyst volume.

Once a consistent set of units is chosen the folloy,ing expressions for

the rate constant at atmospheric pressure are given by (20):

rate constant from

(16.88 - 4~93)

Gn3

jkrnole

h~

laboratory experiments k e (16.00 - ~899)

Ga

3

jkrnole

rnd

'1_' "desié'Jl" rate constant k == e

where T is the absolute temperature

[KJ

As the reaction rate increascs wi th inc:ceasing operating pressure a mul-Liplier

taking this effect into account must be used. For an operating pressure of

24.5 ata its value is

4

(11), (20).

The last factor \'/hich must be considcred y..,-hen cornputing the catalyst

vohlJne is the conversion to be aehieved in the converter. Again, the approach

to equilibrium concept is used. The approach recorrunended by (20) is 27 ~ 8.

°c

(50 oF) and this value was used tr ... roughout all shift converter calculations.

where

(21) p:coposes a kinetic expression having the form dCO

- - ==

k-C

CO - CO )

dt e

CO

,

any magnitude representing the amount of CO present (pa:t."'tial

pressure, concentration, mole fraction, flow etc.)

CO ::: the amount of CO at eQuilibrium (using consistent units)

e

k rate constant

~Vse~

-.'

The expression for the rate constant given by (21) -is: ..,

(25~33

( ( ( ( ( ( (

o

o

19

-The displacement from equilibrium CO .• CO is calculated as follows:

e

If CO, H₂0, CO₂,,H₂ denote the respective amounts of CO, H

20" CO2 and H2

present in the system at equilibrium,

K (T) := p CO 'H 2,e 2,e CO 'H 20 e e (C0 2 + x}(H2 +

x)

(CO x)'CH 20 - x)

where T is the absolute temperature of the system and x is equal to CO - CO •

e Rearranging,

~p(T)

- lJ'X2 - [Kp(T}(CO + _H20) + CO₂ + H2J·x + K/T)'CO'H₂₀ - _{CO 2'R2} :=

°

setting b and

the applicable solution for x is

x :=

- b -Vb

2 - 4c

2

r

Again, the rate constant must be multiplied by 4 in oruer to talce into

, acco~nt the influence of pressure.

As the shift converter operates adiabaticaly thc temperature rise

accompanying the reaction must be ta1cen into account. This is normally done

by deviding the reactor to small steps and ass1J11ling each step to be an

ideally mixed reactor.

Two computer programs were used. for the calculations. One is the

"FIXJ3ED" program which was made available by H. Hoekstra and the other is a

self developed CPS program ("shift ,,) which was writtern in order to check

the resul ts calculated by "FIXJ3ED" ~ . The print out sheets of both programs

are attached and i t can be -seen that the catalyst volumes predicted by the

two programs are in good agreement. The "shift" program predicts a somewhat

lower catalyst volume

~reason

being thé magnitude of the spacetime per

step selected. Reducing this value (increasing the number of steps) will

reduce the difference. The programs are further described and explained in

( ( ( ( ( ( C

o

o

N/l.AV, H "HfI) STR~\ IN STRfJ. UIl FIXBED l FLOW P ~lüL/S t. TA 23("/.34 24.5C 24.50 G ~ \-JE, ~i S r ,7 CO iN ~ R. ~ 1 E l; [ ;:(:: 11<1

REAC10R21~~NSIES: VJLU~E=

20

-Ref. (20), design. T H Vj H20 C KC/,LlS 3~)O.CLt -791-54.9 C.01530 0.561900 376.69 -79f:.5f_,.9 {).O05164 0.535534 KS: 11.2~G2 K~= 11.100& 2: .• CnO 1-\3. LENGTE= 4.071.3 ~l ~-

----PJQOSllElT=O.5000 VERtiU JFTlJJ=: 2.4000 SEC

tg.".", t:Lj:,! p

f~ :~rl::: 1 CJ l·lOLlS A Tl.

STRI1 lr\J 23,;r.3l , ?l, • 50

STfLV; UH 2::'07.34 2.' •• ~) 0

G~WE~STE CJ~V~~Si~ b~~Ej~T

REAC1'O~JI~~NS1~S: VGl.U~IE=

Ref. (20), lab. exp. T H co n2v C KCALIS 3~.0 .C4 -79Li54·.9 O.C3153C ('.561900 -Y(6.76 -7935l_l.9 0.Oi)!,C,93 _0.535'.63 t~:;= 1l.4é6<, K,'.= 11.17/(, 9":~J LE~Gl[= 3.r~19 ~

PORJSITEIT;O.5S00 ViRt;llJr-TI JO=c 1. OOCl) SEC

N".~~: FLOW p

HNHi: ID M(JUS A TA

Slil.~: IN 2307.34 é~4. 50

STK~I UIT 2307.34 24.5C

GEWENSTE CON~ERSIE bEREIKT

REAC T OiZO 11-1 E:. S lES: ~'\J LUI'\E=

Ref. (2l) , labo expo

T H CO H20 C KCAL/S 3~O.(itl _{-19854.9 0.031530 0.561900} 316 ~91 -79854.9 O~C04953 0.535323 KS= 11.8194 Ka= 11.1511 6.4363 M3, LE~GTE= 2.6329 H

POROSITEIT=0.5000 VER8LlJFTIJO= 0.6500 SEC

CDi t;Z CH'"t 0.06'.450 0.317110 0.024410 0.090816 0.344":16 0. 02f_,4 iO CLl2 1i2 ~H.:.,. 0.0(.4',-50 0.31771 0 0.0/."'110 0.090ÛG7 C.34f_{dl,.1 O.CL'1}/_.1') 2. r:; ~il ~! C02 H2 CH4 0.O641,'50 0.311110 C •. 0:<:4',10 0.091027 0.344287 0.024410 DIAMETER= 1.7641 H

' ) "-'J' ...... ···,···0 ···'··0 "·,,,,···· ,.,.,"). ,,,.,,,,,,,.,, .. ,, .... ,, .. ,,.,, .. ,,.1*',.,,,,,.,,.,, ..... "." ... :.". ".,." . ..-,." .•... , .. ,,,.,, , .. " "" .• .1""'\" " "."" " " " .... " ..... " " " .. ".,...., .. ".; .. " .. " .... " " ".-'"""" ... /"'"'r" """ .. """ .... " r

...•... -...

.• 1 ••••• . ... , .... _ ... _ ...•...•... )(C'Ci

... .,!5..c ~.2Ji .. ,.5 .. oi •• 7..r,~}.9 ... ~.?~.?0..f .. , .. ?;.j .. ~ .. 0J .. ~.~.:. .. I. .. ~}:.~ .. ! .. ~.t.I .. :~ . .c.r.?~

..

~

..

~

.

.I.J ..•.. ~.I. .. , ... ~.I. .. • .. 0.? ............. .

.... "." .. "." ... "" ... " .. "" .. " .. _ .. "-" .... "" .. " ... """"."""""."".",, .. """"""""iïr t T ·i~ïr"'iiT"r;';·i",iï' 11 ~ t ~ "r r .. """ ... " .. ,,, " .. "

.. ""._ ... "~.l:5J,).L ... ln...r.Q,,):r:~0J."',,!':I.~!r ... ,,L"-,':,>~.r. .... 1,.T' .. .r.r.: . .f..':'.r..~:.r: r f. .. r:'.~'.~.' .. ~ .. ' ... .

. " ... " ... 1._ .. _ ... \0.':0 .. 1,.r..';:.I.~:.t ... ~:.0..L~.,.f.r.~.::; ......... "" ... ::.r:~.~,~ .. :.:.~.l.1.~T" ..... ~.~.~ .. ~ .. ~ ..... (~ :~ ~ ...

.

'1 ~ 7. ~ 7, " (\ ? rl~, , 7 7 J ... ~ ~ i " " " r I, I, " r,

... ·r·tl

ç-_

..

..

" .. _

...

.

_

...

2'"

:<

"ï

:,

'

"

...

ri:

'Ä~ I: ·4·'· ... · ... .' ','r,? : '"ir.'" ... " ·:·~:!i;·i;r

.. ·

... ·

....

..

... · ..

·

... ·

...

·

....

· ...

..

...

.

... -

...

. _

..

..

... ..

(:.) ? r J • i '" r . C! j r'"t, !~ \ • ,.. .., ('1 t ,., 0 5-' ~

·····, .. ·· ... · __ ·_····c ()·2· .. ···_··_ .. ~-_·_··· .. ··_···_·_···,"s"?,'~' "f .~. f··· .. · ... ~ ... ~. 'Ö'5 'ij"4' :; ... _ .• _ ... "'7' ~ 'ï~"~ "~".;' "~"""''''r';'''~'' "';'ij' 6 ·t:·!'i ...... , ... _ ... : ... ..

'12 () , 11 r r 7 • :; ~ r . 5 61 "() ,,:. 'I ,~ • r, ~ r .5) :' S 5

... _ ... _ ... _ ... , ... , .. , ... , ... , ... , ... ' ... , ... _ ... -. __ ... _ ... I.~l . .T..!.:J. ... _._ ... _ .. ? .. 7: .. ~.r.,

.

.

:

.

J J .(\"r0n _.........~~"r_...._.....~l _...........1._..."(\""r _{..................} , ... ' ... ,' ... , ... , ... ..

... _ ... c ... T.::-:.II.L .. ~::.I.!.\.r ... _ .... _ ... _ ... l.).~ .. ~.?~ .. , .. ~ ... _ .... .J..~},:}.I: ..• ~ ... .

. T ["'H' d (' fT r :3 :-n • (\ . 3'1 (; .r

... _ . . . ... . . . " ••••• _1 ... - .... _ ... , ... , ... _ ... _ •••.•• __ ... _ ... _ ... _ ... , ... . N

"

... r..q.~.~.~ ... !~ .. t!1 ... "',, ... " ... " ... " .... ~.~.! ... ~.:! ... " ... " ... "., ... , ... , . ... P7fY.I .. ~.Y .. .JJ:.,Ç.~Jj.l::r...:::.?r?tLC? ... ,... -? ~ 7 :, ~ ?

... Î..t:_Ht .. .Ys..:;: ... ~F.:.','l.: ... _ ... y.Q1Yr.:~ ... ~ .. ~.,.(f.f... .. r.I.~.~! .. , ... ~~ .. 1 ,-r-r t " r ?'?) n"

.: ... _._._ ... __ .. _ .... '-.. _.{.ç:Q2.t~C)2J.J .. (.r;.0 .. l..~.!..!.:.~.<:') ... r..~.~ . .!..r.: ... : ... ) . .1: .. ! .. ~ .. C.~ ... _.'.'..t. ... r. .. r: .. '~.r..t,.~.r. ... r:.~., .. t..l~.: .. t. ... , ... , ... ..

..

..

.

..

_

... _

... _

....

.r

..

Q ... :: ... J) ...• ..1.?.f. ... ~ .. t .... j .. ; .. ~ .. L~:· ... ~.: ~: .. ~ .... ( ... ~ ... ~:.r.: .... ~, ... ):.!.1 ... ~ .. :. .. ~ .. ~~ ...... J' .. !... ~ '1'" • r (' r I'r 'r ... _ ... t.0..tQ .. 1... .. s..~r,:. r.;; .... J .... ,J.H ... .. .. ... , ... :' ... , ... _ ... .. ... ... , ... , .. " ... , ... , ... .

,

-( ( ( ( ( ( (

c

\

c

o

Cl

22

-The fo1lowing catalyst volUlues were ca1cu1ated based on bed porosity

of ~5

ca1culated cata1yst computer program rat e const ant reference

vo1Ulne

[m~

used

23.8 FIXBED design

23.4 shift design

9.9

FIXBED laboratory expo

6~4 FIXBED laboratory expo

A cata1yst volUlne

of

23.8 m3 was selected in order to account for the

dec1ining cata1yst activity.

(20)

(20)

(20)

(

r.

r

(

r

o

C 23 -CO 2 removal system

Following the reliability and operational simplicity criteria guidiP.€

the design of the hydrogen plant a sirnple, and if possible low cost, CO

2

removal process should be selected.· Reviewing the various treating schemes

available the practical alternatives are reduced to two, namely, absorption

by an 1~ (mono-ethanol amine) solution and absorption by a hot K

2C03 solution. The absorption of CO

2 by a hot potassium carbonate solution is

advan~ageous as both absorption and regeneration are conducteel at approximately

the same temperature so that no heating is required for the rich solution

leaving the absorber (stream no. 23). This fact results in a lower steaIn

consumption in the regenerator and in the elimination of the lean solution

-rich solution (st~ceam nol s 27 anel 23 respectively) heat exchanger which is

typical for amine treating plants. Another reason for the lower steam

consumption in K

2C0₃ treating plants is the lower heat of reaction (37

%

of

the reaction heat in amine solutions (24)).

Generally spea..1_{dng, treating with a potassiuIIl carbonate solution is more}

economic than NIEUI. treating fOT 00

2 partial pressures greater than

L 4

atrn

(/\/20 psia) anel vice versa for 00

2 partial pressures lower than 1.4 atm (32).

The conclusion of this short discussion i~ that hot potassium ëärbonate

treating offers significant savings when bulk CO

2 removal is elesirecl. For the hydrogen plant under discussion, reducing the CO

2 content of the gas

leaving the absorber to .6

%

(volume) Vlas f01mel sufficient for the purposes

aimeel at here. This corresponds with a CO

2 partial pressure of .14 atm at an

absorber outlet pressure of 23:5 atm. At the inlet to the absorber the

00

2 partial pressure is

4.3

atm. Rather than using two types of solutions,

it was decided to use a split stream hot potassium carbonate treating scheme

as described by

(7),

(25) and (28',).

The reactions taking place during CO

2 absorption are:

CO₂ (~) + R₂₀ (sol) • . po H_2C03 (sol)

ISC03 (sol)

+

_H2C0₃ (sol) ~ _2KFIC03 (sol)

The following operating conditions were assumed for the absorber:

entering cold lean solution temperature (stream no. 21)

₌

88

temperature of gas leaving the absorber (stream no. 20)

₌

88

temperature of gas entering the absorber (stream no. 18) ::: 115

temperature of rich solution leaving the absorber (str~no. 23)

₌

115

cold lean solution stream as wt

%

of total lean solmion

°c

°c

°c

°c

( ( ( ( ( ( ( ( l

o

24

-In addition it was assumed that the solution strength is 30

%

by weight K

2C0₃ in water. According to (7) this is the maximum recommended strenght

with which complete conversion to KHC0

3 can be achieved without the danger •

of bicarbonate percipitation.

Equilibrium composition, vapor pressure, specific gravity and other

thermodynaTJlic and physical properties data jn the form of curves presented

by (26) and based on the work of Benson and Field (24), (25) Vlere used in

the calculations. In cases where gas streams saturated with water are

encountered, the value of the water partial pressure was determined from

steanl tables (9).

*

The calculation of the gas composi tion at the inlet to the absorber

Ct

=

115 °C, p

=

23.8 ata):

From the shift converter printout the following molar flow rates in kmoles/hr

are established: H 2 485.84 CH 4 34.47 CO 7.31 C0 2 128.22 655.84

vapor pressure of water at 115

°c

absorber inlet pressure

₌

23.8 ata

••• YH

°

=

1.67/23.8

=

.0702

2 . . .

water vapor flow rate at absorber inlet =

= 655~84·YH 0/(1 - ·YH 0)

=

49.52 kmoles/hr

2 . 2 . .

M

=

average molecular weight

=

'M.·

y. = 11. 73

~ L 1 1

*

The calculation of the gas composition at the absorber outlet

(t 88

°c,

p

=

23.5 ata):

A regenerated lean solution loading of 3 SCF C0

2/gallon solution which

has an equilibrium C0

2 partial pressure of .068 atm at 88

°c

was selected.

The CO

2 partial pressure at the absorber outlet is .14 atm which means that

( ( ( ( ( (

o

o

n

25 -A loading of 3 SCF C0

2/gallon means that one gallon (3~785 lit) of the

solution selected (30 %wt K;C0

3 unconverted) at 60 OF contains 3 standard

cubic feet of C0

2 (3 cubic feet at 60 OF and 1 ata).

The water vapor pressure . above such a solution at 88

°c

is .51 atm •

• 51/23.5 .02l7

(specified)

Assuming no change in flow for H2' CH 4 and CO,

Y_H + Y_CH + Y

_co

= 1 - .0060 - .0217

=

~9723 2 4 . H₂ + CH 4 + CO =.: 527.62 lanoles/hr C0₂

=

527.62'.0060/:9723 3.26 kmoles/hr H 20 = 527.62·~02l7/.9723 = 11.78 lanoles/hr

The outlet composition in Y~oles/hr becomes:

H 2 485.84 CH 4 34.47

co

7.31 H 20 11.78 542.66 ,.

~The calculation of the lean sollition rate required (stream no. 25):

1 SCF CO

2 = 28032.10- 3/.08206'288.9

=

1~195·10-3

kmole CO 2

total CO

2 absorbed = 128.22 - 3.26

=

124.96 kmoles/hr

The equilibrium C0

2 loading of a 30

%

K2C0₃ solution having a C02 partial

pressure of 4.3 atm (the CO

2 partial pressure at the absorber inlet) is

7.6 SCF/gallon at 115 °C.

minimum solution rate 124.96/lc195·10-3 (7.6 - 3)

=

22.7-103 gallons/hr (60 OF, unconverted)

Use 1.25 x minimum flow rate

lean solution rate = 28.4'103 gallons/hr

~

6

0 _.

Specific gravity of unconverted 30 ~o K

2C03 solution at 0 F =

1.294

lean solution mass flow rate =

( ( ( ( ( ( (

c

'

o

- 26

-Ass1llning that the amount of water absorbed has a negligible effect on the rich solution volume;

rich solution loading at the absorber outlet

=

3

+

124.96/1.195'10-

3

'28.4'10

3

=

6.68

SCF/gallon

The ca1.--rying capacity is equal to

3.68

SCF/gallon (6~68

-

3) which, according

tö

(25),

is still acceptable.

*

Heat balance on the absorber:

'As the temperature of the rich solution leaving the absorber was fixed, the purpose of the heat bala'1ce is to establish the tempera:ture of the entering lean solution. In order to simplify the calculations the following

assumptions were made.

a The specific heat (weight basis) of the lean solution (3 SCF CO/gallon) is assumed to be equal to the specific heat of unconverted

30

%

solution. The value is

.72

kcal/kg °C.

b The liquid phase enthalpy is assumed to be equal to the sum of the component's enthalpies.

c The enthalpy of CO 2 in the liquid phase is assumed to be equal to the

C0

₂

gas enthalpy at the sol ut ion temperature plus the heat of reaction

(-7722

kcal/kmole CO

2

=

-13900

BrU/lbmole

C0

2

).

d The enthalpy of water in the liquid phase is aSSQmed to be equal to the steam enthalpy at the solution temperature reduced by the latent heat of water.

e The datum \emperature (enthalpy = 0) for the lean K

2

C0

₃

solution was

chosen as the unknown temperature of the lean solution leaving the regenerator (stream no.

27).

With the aid of the above ass~mptions the enthalpy of the rich solution is calculated as follows: Heo

= -

2067

kcal/kg 2

~20

=

3000

kcal/kg at

115

oe at

,

115

oe

latent heat of water

=

529.3

kcal/kg at

115

°c

C0

₂

absorbed

124.96

kmoles/hr

( ( ( ( ( ( ( (

o

o

() 27

-rich solution enthalpy "" 14.36-104 '.72 (115 - t) +

+ 124.96-(- 2067'44.01 -'7722) +

+ 37.74°18.02'(- 3000 - 529.3)

28.42'10 5 - 1.0339·10 5·t kcal/hr

where t is the lean solution inlet temperature.

cold lean solution enthalpy = 14.36'104 '.33':72 (88 - t) =

=

30~03'105

- .3412·105

·t

kcal/hr

entering gas enthalpy (stream no. 18) - 136.71-10 5 kcal/hr

outgoing gas enthalpy (stream no. 20) 3.70'105 kcal/hr

The overall heat balance can be written now as follows:

- 136.71'10 5 + 30.03'105 - .3412·105·t

=

= -

3.70-10 5 -28.42'10 5 -

1~0339-105,t

,', t = 74.56'105/.6927'105 = 107.6

°c

rich solution enthalpy = - 28:42'10 5 - 1.0339'105'107.6 =

- 139.71'105

kcai/h~

co1d 1ean solution entha1py

=

30.03 105 -

~3412·105'107.6=

6.70'10 5

kca1/h~

lean solution cooler duty (H 13) - 6.70'10 5 kca1/hT (= 779 kW)

*

The ca1c171ation of the regenerator reboiler duty:

Under the operating conditions of .6

%

C0

2 in the treated gas and C02 partia1

pressure at the absorber inlet of 4.3 atm the regeneration efficiency (the

number of standard cubic feet CO

2 removed by 1 lb of condensing steam) is

equal to 7.5 SCF COj1b steam (25). Assuming a steam latent heat of 520 }\:Qal/kg the reboi1er duty becomes,

/

-3

6

/

reboiler duty

=

124.96'520 1.195'10 '7:5-2.205 32.89'10' kcal hr

( ( ( ( ( ( (

o

- 28

-The calculation of the regenerator overhead vapor temperature:

The regenerator pressure apd the lean solution composition will fix the

lean solution temperatureo A lean solution t emperature of

108

0e will result

in an operating pressure of about

1.6

ata at the bottom of the column and, allowing

.3

atm for preS&'1ITe d:cop across the packing, the pressure at the top of the column becomes

1.3

ata.

Assuming that the CO

2

cont ent of the recycled wat er (stream no.

29)

is

negligible, the amount of

(X)2

leaving the regenerator in the overhead vapor

is equal to

124.96

kmoles/hr. Furthermore if the

(X)2

removal system is

assumed to be in balance with respect to water (i.e. no make-up water is required) the amount of water l eaving the system in the vent gas (stream no.

36)

must be equal to the amount of water absorbed in the absorber n8lnely

37.74

kmoles/hr. Assuming the pressure in the drum

(V 19)

i s l~l ata the

partial pressure of water i s,

PH 0

=

1.1

0

37.74/(37.74

+ 124~96)

=

~26 atm

2 .

o

which corresponds with a condenser outlet temperature of

65.7

e. Assuming the overhead vapor is saturated wi th water the overhead

temperature ca.n. be calculated from a heat bal ance on the regenerator by solving the equation:

where H is the enthalpy in kcal/hr and the subscripts indicate stream numbers. A trial and error procedure is applied. A value for the overhead temperatu~e

is assumed then corrected until the heat balance equation is fullfilled. The overhead vapaD temperature which was faund this way is

97.7

oe.

a

PH 0

at

97.7

e =

.92

atm

2

YH 0

= ~92/1.3

= .7077

2

water in overhead vapor =

124.96·YH 0/(1 - YH 0)

.

2

·

2

.

~20

(g)

(66

oe) = -

3019

kcal/kg

~20

(1)

(66

oe)

3580

kcal/kg

H(X)

(66 oe) = 2

- 2075

kcal/kg

~20

(g)

(98

oe) =

- 3007

kcal/kg

HH

2 0 (g)

(98

Oe)

2068

kcal/kg

302.55

kmoles/hr

( ( ( ( ( ( (

c

'

( l

o

o

H. ln 29 -Hout ::= H27 + H28 =

o

+ 302.55'(- 3007)-18.02 + 124.96'(- 2068)'44.01

6

/

::= - 27.767'10 kcal hr

The calculation of the condens er (H 18) duty:

condens er duty H36 + H29 - _H28 37.74'(- 3019)'18,02 + 124.96-(- 2075)'44.01 - 13.465'106 kcal/hr condenser duty = (- 13.465 - 170083 + 27.766)'10 6 ::= - 2.782'106 kcal/hr (= - 3236 kW)

?' The determination of the columns dimensions:

The diameter of any packed column used for gas-liquid contacting depends on the gas and liquid mass flow rates, the operating preSS~Te and temperature, the type and size of the packing used and to alesser extent on the physical

properties of the gas and liquid which are being handled. In practice collli~

diameters are based on superficial mass velocities (gas mass flow rate per

unit empty cohurm cross section [kgJhr

m~)

which are approximately (for Raschig rings) 60

%

to 80

%

of the superficial gas mass velocity that will cause flooding (33). The flooding phenomenom itself is described in (8). For column diameter calculations the plot of

vs.

under flooding conditions (8) may be used.

G' ::= gas mass velocity

~b/hr ft~

L' ::= liquid mass velocity

~b/hr ft~

_,....

Cf ::= packing characterizationfactor

~

i

::= liquid viscosi ty [centipois~

g~

::= conversion constant ::= 4018 '10

8

~bm

ft/lb

f

hr~

__

( ( ( ( ( ( (

o

o

o

30

-The packing characterization factor Cf increases rapidly with decreasing

packing size. The result is that column diameters become larger when the

packing size is reduced. Using the same packing as used by (25) (!"

porcelain Raschig rings) in the CO

2 removal system will result in wider

and shorter columns which are unfavorable at high pressure operations

because of the high material costs. It was decided to use 1!" porce1ain

Raschig rings as packing for both the absorber and the regenerator. The

value of Cf for such packing is 95 (8).

The absorber diameter is calculated as follows:

Bottorn part conditions (highest gas and liquid flows) will determine

the diameter and therefore wil1 be used in the calculation.

~ - mass rate of gas en.tering absorber

=

8272 kg/hr

f'm,g

-d mass rate of rich solution leaving absorber

=

)'lrn 1

, = 14.36'104 +

124.96·44~01

+ 37.74'18.02 = 14.98'104 kg/hr

. ,

assuming ideal gas behaviour, .

D

=

M 'pjRr

=

11.73'23.8/.0821'388.2 8.76 g/lit .Ig av ~1

=

1~28 kg/lit

14.98

0

104V

.

8.76 1 8272 1280

entering the curve (8) at this va1ue,

A'

₁ .58 CP .0135

1.5

(G,)2

:0135'4.18'108'8.76'10-3'l~282'6Q.373/Cf'o58·2'62.37

3.513·108/Cf

-cr

,;,

V

3.513'108/951 = 1923 1b/ft2 hr (at flooding)

operating at 70

%

f100ding gives,

G~11 = ~7'1923

=

1346 lb/ft2 hr

=

6564 kg/m2 hr

. 2

column inside cross section =~. /6564 = 1026 m

m,g

.'. inside diameter = 1.27 m

( (

r

( ( ( ( C,

o

o

31

-It should be mentioned here that the gas flow rate in the absorber

decreases significantly as a re sult of the CO

2 absorption. If the volume reduction is large enough it may be economic to reduce the diameter of the column's top section. This possibility was not investigated further.

The column diameter which results if

!"

packing is used is also calculated for the sake of comparison;

Cf = 640

!"

porcelain Raschig rings

G~l

= 1346\195/640 1 = 518.6 lb/ft2 hr 2529 kg/m2 hr

L;

= 24380\195/640 1 = 9393 lb/ft2 hr

=

45800

kg

/

~2

hr

i~side diam

e

te~

= 1.27'(640/95)·25= 2.05 m

When calculating the regenerator diameter it can not be decided

beforeh&~d which conditions are determining namely those at the top or at

the bottom of the column. To solve the difficulty the column diameter was calculated twice using both thc top and bottom conditions.

Bottom section:

The flowrate of the vapor leaving thc reboiler must be first established. This can be done by using the following data which were already calculated.

reboiler duty = 32.89-10 5 kcal/hr pressure at bottom = 1.6 ata lean solution loading = 3 SCF C0

2/ga1lon

The stream leaving the reboiler (stream no. 30) is a mixture of vapor and liquid in equilibrium. The separation of both liquid and vapor takes place in the column so that the vapor entering the packing is in equilibrium with the lean solution leaving the column (stream no. 27). From charts (26), the partial pressure of 00

2 above the 1ean solution at 108

°c

is equa1 to .122 atm.

•• yoo

2

.122/1.6

=

.0763

It is assumed that all the reboiler heat input is utilized for KHC0 3 dissociation and water evaporation. As will be seen later the amount of water vaporized in the reboiler is small (~4

%)

relative to the 1ean

solution flowrate so that the assumption of constant temperature evaporation. is acceptable. In addition, this assumption will result in a slightly

( (

r

( ( ( (

o

o

- 32

-higher vapor flowrate than the actual and wil1 give a column diameter slightly larger than the actual diameter required.

heat of reaction for CO

2

=

-

7722 kcal/kmole CO2 latent heat of water qt 108

°c

= 534 kcal/kg H

20

A heat balance on the reboiler gives then,

32.89'105

=

~

1 < (YFI 0'18.02'534 + Y

_co

'7722)

. mo e,g 2 2 .

where,

~ _{mo e,g} 1 = mo1ar f10wrate _. of the vapor leaving the reboi1er

with Y

_co

=

.0763 2 and _{YH 0}

_{= 1 -}

_Y

₌

_.9237 2 . CO2 . ~ 1 = 347 kmoles/hr .mo e,g . ~ = ~ + ~ = 5776 + 1165 = 6941 kg/hr .m,g m,H 20 .m,C02

average molecular weight = M = ~ /~ 1 = 20.0

av . m,g .. mo e,g

~ 1

=

total mass flowrate of liquid to the reboiler

=

m, =

~

+ 14.36'104 = 15.05'104 kg/hr m,g D = M

'p/m

= 20'1.6/.0821'381 1~02 g/lit )g av .PI: = 1.245 kgJ1it

~\/~' ~ ~m,l~~

I

~ 15.05'l04~.02

I

~

.62 G'

fî

~m,g

!i

6941 1245

entering the çurve (8) at this value,

g~

<?g'fi

PI

.036

Ai

= .61 CP 2 6 8 -3 2 6 3/ 6· 2. 6 7 (G')

=

.03 4.18 10 -1.02 10 '1.245' 2.37 Cf" 1 2.3 1.022<108/C f . G'

V

1.022<108/95 I = 1037 lb/ft2 hr (at flooding) operating at 70

%

f100ding gi ves,

( ( ( ( ( ( ( (

o

o

() 33

-column inside cross section

inside diameter

=

1~58 m 15750 1b/ft2 hr Top section: 2 L96 m 76800 kg/m2 hr

In order to calculate the vapor flow leaving the packing and the liquid

flow entering the packing, stream number 26 must be calcu1ated. Being confronted with two phase flow this means that the flows and compositions

of both vapor and liquid phases must be estimated. The two equations which

must be satisfied are:

6

H23

=

_H26

=

-13.971"10 kcal/hr (= - 16248 kW)

~m,23

=

~m,26

The trial and error procedure which was applied together with the calculations

is presented in appendix 3" The re sul t of the calcu1ation was the fo1lowing: flash temperature

=

102.8

°c

vapor phase: 53.2 kmo1es/hr 1iquid phase:

~ =

14.59.104 kg/hr m

A material balanee around the bottom part of the regenerator gi,ves,

~ _{mo1e,g -}- ~ _mOle,g,C0 + ~

2 mo1e,g,H20 Assuming ~ _mo 1 1 is pure water,

e, ~ = 71.76 kmoles/hr mole,g'(X)2 ~ _mole, 1 - (87.91 - 37.74) ~3770 ~6230

~ ~

~

Hg mole,l mole,g

~-

---

~----~~~ -l-~

~

-r-~

~ ~

:

I Hf,l

~mole,l

- 50.17

~. 1 1 can be determined from a heat balance on the bottom part of the mo e,

regenerator. Such a balanee gives,