.,KERNFORSCHUNGSANLAGE JULICHGESELLSCHAFT MIT BESCHRÄNKTER HAFTUNG

Institut für Reaktorentwicklung

The Pebble Bed High Temperature Reactoras a Source of Nuclear Process Heat

Volume T

Processes of Nuclear Process Heat

A Common Study byKernforschungsaniage Jülich GmbHand General Electric Company

by

R. Schulten, K. Kugeler, M. Kugeler, H. Nießen,H. Hohn, O. Wolke and J. H. Germer

Als Manuskript gedruckt

Berichte der Kernforschungsanlage Jülich - Nr. 1119Institut für Reaktorentwicklung Jül -1119 - RG

Dok. : Nuclear Process HeatPebble Bed High Temperature Reactor - Nuclear Process Heat

Im Tausch zu beziehen durch; ZENTRALBIBLIOTHEK der Kernforschungsanlage Jülich GmbH,Jülich, Bundesrepublik Deutschland

THE PEBBLE BED HIGH TEMPERATURE REACTOR

AS A SOURCE OF NUCLEAR PROCESS HEAT

Volume 7

Processes of Nuclear Process Heat

August 1974

R . Schulten, K . Kugeler, M. Kugeler, H. Nießen, H . Hohn * )

O . Woike, J .H . Germer **)

A Common Study by

Kernforschungsanlage Jülich GmbH and General Electric Company

Institute for Reactor Development, Kernforschungsanlage Jülich, Germany

* ~` ) General Electric Company, USA

Dr . Maya Röth-Kaurat deserves our warmest thanks for translatingthe report and for her part in assisting us with the preparation ofthe manuscript .

Kernforschungs-anlage Jülich

JÜL - 1119 - RG

Oktober 1974GmbH IRE

ABSTRACT

THE PEBBLE BED HIGH TEMPERATURE REACTOR

AS A SOURCE OF NUCLEAR PROCESS HEAT

Volume 7

PROCESSES OF NUCLEAR PROCESS HEAT

The important processes of nuclear process heat are given alongwith their flow diagrams . The following are especially described :- Hydrogen or synthesis gas production- Hydrogasification of coal- Direct reduction of iron ore- Hydrogenation of coal- Hydrocracking of heavy fuel oils- Steam gasification of coal-- Chemical heat pipe systems .The process conditions and typical consumption figures are given .

VOLUME 7

PROCESSES OF NUCLEAR PROCESS HEAT

7 .1 Hydrogen or H2 /CO - Production by Steam-Reforming of Methane

7 .2 Hydrogasification of Coal

7 .3 Direct Reduction of Iron Ores

7 .4 Production of Gasoline by Hydrogenationof Coal.

7 .5 Hydrocracking of Heavy Fuel Oils

7 .6 Chemical, Heat Pipe Systems

7 .7 Steam Gasification of Coal

7 .1 .

Hydrogen or H2/CO-Production by Steam Reforming of Methane

(ref .

1 )

The steam reforming of methane occurs in tubes, filled with catalysts,which are heated externally . The reaction temperature is about800 - 850 oC at a pressure of 30 to 40 bar . The steam/methaneratio varies between 2/1 and 5/1 depending on its application . Thefollowing reactions chiefly occur :

CH4

+ H2O -} CO

+

3H2-49

kcal/mole

CO

+ H20 ~ CO 2

+ H2+9kcal/mole

The composition of the gas mixture at the end of the reaction de-pends on the pressure, temperature, and the water/methane ratio .High temperatures, low pressures and a large water/methane ratio .influence the reaction of methane and water favorably (Equation 1) .Heating with nuclear heat from a helium gas coolant (at a pressureof 40 bar and a maximum temperature of 950 0C), results in the samehEat fluxes as in the case of conventional reformer furnaces whichare heated at peak temperatures of 1 .500 0 C . It is possible to re-duce this peak temperature because of the improved heat transferof high pressure helium . According to the overall diagram the uppertemperature range of the helium (750 - .950 oC) serves to performthe reforming reaction as well as to superheat the process gasesfrom 450 oC to 800 - 850 0 C . The lower temperature range of thehelium (250 - 750 0 C) generates steam for turbines and possiblycan be used to preheat the reactants from 20 oC to 450 oC . The pro-duct gas which is at a temperature of 800 to 850 oC has to be cooledto room temperature in the course of this process . The impuritiesalso must be removed from the gas during this process .

After the reaction, the following occurs :

a) Utilization of waste heat in a heat exchanger between the reac-tion temperature and approximately 400 0C ;

b) A high temperature shift to convert the CO in the product gasinto H2 by means of an iron oxide catalyst

c) A CO 2 scrubbing by means of either chemical scrubbers (hotpotash, Sulfinol, Purisol) or physical scrubbers (i .e . Rectisoi

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scrubber in which methanol is used as an adsorbant) ;d)

A low temperature

separation of the unreacted CH4 portion ;e) If necessary, methanation

to convert'the remaining tracesof CO and CO 2 into methane .

When using this above process for the synthesis of methanol, theshift reaction (b) is not utilized because a high percentage ofCO is required in the product gas . In the case of an ammonia planta secondary reformer is added after the steam reformertubes . Throughpartial oxidation with air and reaction of the remaining methane,a high quality gas is produced for which methane separation is notnecessary . An additional advantage is that the required N2/H 2 ratiocan be directly adjusted .

7 .2 Hydro-Gasification of Coal (see fig . 7 .2-1) (ref, 2)

The following basic equations describe the gasification of coal :

According to the final products desired in the gasification process- methane (i .e . synthetic natural gas), mixtures of carbon monoxideand hydrogen (e .g . as a reducing gas for the iron industry, asinitial gas for chemical synthesis), or pure hydrogen (e .g . for hy-drogenation reactions) - it is now possible to combine the individualsteps of the processes according to the equations . Therefore it isnot necessary that the reactions occur one after the other in sepa-rate systems . It is possible that some of these conversions occursimultaneously in one system . This will be especially the case whencoal consists not only of pure carbon, but also of oxygen and hydro-gen .

the

AH/kcal/molehydro-gasification C + 2H 2 = CH 4 ; -20 .6 (exothermic)steam reforming CH4 + H20(v) = CO :+ 3H2 ; + 49 .0 (endothermic)shift reaction CO + H20(v) = CO 2 + H2 ; - 9 .9 (exothermic)

a) The generation of methane (SHG)

Gasifier

2C+4H2 = 2CH4

; AH = -41 .2 kcal/mole Carbon 0Reaction temperature = 800 c

Steam reformer :

CH4+H20(v)

= CO+3H2 ; AH = 49 kcal/moleReaction temperature = 800 oC

Steam reformer

CO+H20(v) = CO 2+H2

; AH = -9 .9 kcal/male+ shift reactor :

Reaction temperature = 400 o C

Gross reactionfor

methane

.

2C +

2H20(v)

=

CH4

+ CO2

b)

The generation of synthetic gas

Gasifier

:

C+2H2 = CH4

; AH = -20 .6 kcal/mole Carbon

CH4+H 20(v)

= CO+3H2 ; AH = +49 .0 kcal/mole

C + H20 (v)

= C0+ H2

c)

The genera tion of hydrogen

C + 2H 20=00 2 + 2H2The hydro -gasification of coal has two advantages in comparisonto the gasification with steam .

a) The pre-heating temperature required for the gasification mediumis about 200 oC lower than that for the gasification with steam .

b) The required heat is not all transfered in the fluidizing bedof the gasification reactor, but can be transmitted to gas-phaseheat-exchangers outside the reactor

(CH4 conversion to H2 ) .This is because the hydrogasification is either an exothermicor, in

the presence of steam, an autothermal reaction . Theproblems of erosion or fouling of heat exchangers can be avoidedand potential corrosion problems will be minimized as a resultof low pre-heating temperatures,

Gasifier C+2H 2 = CH =4 ; 4H -20 .6 kcal/mole carbonSteam reformer : CH4 +H2 0 = CO + 3H2 ; AH = +49 .0 kcal/moleSteam reformer+ shift reactor : CO+H20 = CO 2+H2 ; pH = -9 .9 kcal/mole

Fig . 7 .2 - 1

Schematic of Hydrogasification of Coal

1 HTR 9 C0 2 f H2S - purification2 Circulator 10 reed water preheater3 Steam reformer 11 Hydro-gasifier4 Steam generator 12 Fluidized-bed drier5 Regenerator 13 Tube drier6 Shift-reactor 14 Back-pressure turbine

C02 purification 15 Feed water pump7 Recuperator8 Low temperature separator

- 6 -

7 .3 Direct Reduction of Iron Ores

(ref . 3)

Corresponding to the reactions

Fe 203

+

3H2 -}

2Fe

+

3H20

-

195

kcal. /kg

Fe

Fe203 +

3CO ~

2Fe

+ 3CO2 +

69 kcal/kg Fe

the generation of crude iron is also possible outside the blast fur-nace and without the use of coke . Numerous proposals have been

Typical examples of the 3 types of direct reduction mentioned hereare : (a) the H-iron process ; (b) the NU-iron process ; (c) theKorf/Midrex, Purofer, Hyl processes . The theoretical consumptionof gas in these processes is about 602 Nm 3 H2/t Fe . The actualconsumption is dependent upon the chemical utilization factor ofthe gas (in general 40 to 50 % per passing through the reduction fur--nace) .Cyclic processes achieve nearly theoretical, values . The pro-duction of the reducing gas itself as well as the gas processingand pre-heating of the raw materials require additional amounts o£energy . Thus the energy consumption of the known process is between3 and 5 x 10 6 kcal/t Fe .

Process Energy demandkcal/t Fe

Raw material Pressure TemPeeöC- PowerPowerttFe/d'm3_ ~~__~~_~=aa~r~nccam~gcmsex~esaraemßaus

H-iron 4 .8 x 10 6 Natural gas 36 540 25NU-iron 3 .5 x 10 6 Fuel oil 2 .5 700 15Korf 3 .2 x 10 6 Natural gas 1 .5 800Blastfurnace 4 x 10 6 Coke 1 1300 30 ., .50

made for this process of direct reduction . Basically there are thefollowing 3 groups of conditions :

a) Pure H2 T low, p highb) Pure H2 T medium, p mediumc) H2 /CO mixture, T high, p low

7 .3 .1

H-iron process

The H-iron process, for example, operates as follows :

Concentrated ore is reduced by pure H2 in a fluidizing reactor atabout 500 o C and 35 b . Hydrogen is pre-heated in a hot water/H2heat exchanger to 540 OC . The reactor itself operates with 3 super-

imposed fluidizing layers

(2 for the reaction,

1

for preheating incounter-flow) . Water is removed from the gas leaving the reactor .

This gas is mixed with fresh gas and returned to the reactor .

Fig . 7 .3 .1 - IH- iron process for direct reduction

1 Compressor2 Recuperator3 Gas heating4 Reduction5 Off--heat utilization6 Water separation

This process can also be combined with a nuclear reactor as the heat

source for hydro-generation . Hydrogen production will then occur at

about 850 0 C, 30 b, with H2O/CH4 = 5/l . These conditions will steep

the remaining content of methane so low that no CH4 separation is

necessary . It is followed by only a normal shift-reaction and CO 2

wash . The pre-heating of H2 , according to fig . 7 .3 .1 - I, occurs in

an exit gas/H2 heat exchanger,followed by heating with fresh steam .

The problems o£ hydrogen diffusion into the helium cycle are there-

by avoided . Steam for the steam reforming process is supplied by

tapping the steam turbine .

Fig . 7 .3 .1 - 2 H-iron process with nuclear heat

The Korf-Midland-Ross process is briefly discussed here as an exampleof an autothermal process for direct reduction .

Product gas is producedgas and exit gas of thely led to the reduction

counter-flow andis purified andfeed material . Thewith flue gas fromfurnace is operated as aThis process is used for

in a steam reformer from both fresh naturalreduction furnace . Under heat this is directfurnace . The ore and the gas are preheated in-

reduced . The exit gas leases the reduction ; itpartially applied to heat the steam reformer and as

required combustion air is preheated recuperativelythe steam reformer .The lower part of the reduction

closed purification cycle to cool the sponge,commercial plants, e,g . in Hamburg/Germany,

1 Nuclear reactor 8 Direct reduction2 C.irculator 9 Recuperator3 Steam reformer 10 Water separation4 Steam generator 11 H2 heater5 Off-heat utilization 12 Turbine6 Shif t-reactor 13 Condenser7 CO 2 purification 14 Feed-water pump

7 .3 .2 Korf -Midland-Ross Process

and Georgetown/USA . The annual production amounts to 400,000tonsof sponge iron . The basic rata material is natural, gas .

Fig. 7 .3 .2-1 shows the flow schematic of the plant .

Fig . 7 .3 .2 - 1

Korf-Midland--Ross Process Schematic

t

~

shaker scraen~"~-"l"r~

canhuslion air

natural gas

undersixaJ gram

,_

This reduction technique in combination with a nuclear-heated steamreformer (see fig . 7 .3 .2 - 2) results in special specifications forthe reducing gas ;

temperature

about 800 oCpressure

about 3 bH2 /CO 'ti 1residual content of methane

< 2residual, content of 1120 + CO2< 5

The generation of such a quality of gas requires the operation of

a nuclear -heated steam reformer at a system pressure of about 5 b,which of course imposes an external pressure on thn reforming tubes .Since this gas conversion in the reduction furnace amounts to 50

per process, the exit gas has to be separated from H2O and CO 2 and

heated again in the cycle to 8000 C . This heat also can be taken from

the hot helium cycle . The diffusion of hydrogen for this pre-heatingprocess is inhibited by the CO-content of the reducing gas .

Steam reforming itself is operated at a H2O/CH4 ratioorder to achieve the desired high reduction potential .of this proportion is that it avoids the separation ofproduct conduit.

Fig . 7 .3 .2 - 2 ICorf-Midland-Ross Process with Nucl ear Heat

123456789

1011

HTRSteam reformerSteam generatorCirculatorReducing gas heaterReduction chamberGas coolerC02 and H2O removalTurbineCondenserFeed water pump

of 1 .5/1 inAn advantagesoot in the

H,a

HeavyOil

7 .4

Fig . 7 .4 - 1 Flow sheet of gasoli ne produc tion

drogenation of Coal (ref'. . 4)

18

In the future it might be possible to make use of the H-coal process,which represents a further development of the H-oil process used

for utilizing residuals of heavy oil . Gasoline could in this way be

generated from coal with the addition of hydrogen . For this

dried and ground coal (size of the kernel : about 0 .4 mm) is

with fuel cycle oil in the proportion 1 :1 and combined together with

the hydrogenating hydrogen at about 200 at . After pre-heating, it isintroduced into the hydrogenation reactor . This aggregation operates

with a catalyst which is kept in liquid suspension . Hydrogen, coaland mash oil flow upwards through this catalyst layer . A centraltube permits the recirculation of the raw materials . The gaseous,

liquid, and solid products are removed . Now it is possible to com-

bine this process with a nuclear reactor which delivers the re-

quired heat for hydrogen generation and the required electrical energy

for compressors and other aggregates .

process,mashed

1 Milling and drying 7 Distillation 14 Shift-reactionof coal a H2 compressor 15 Heat exchangers

2 Mashing 9 Low temperature 16 Steam reformer3 Mash compressor carbürization and preheating4 Preheating 10 Hydrocracking 17 C1 . . . C4 compression5 Hydrogenation reactor stage 18 Nuclear reactor6 Gas separation, 11 C1, . .C4/H2 19 Steam turbine and

desulfurization separation generator12 C02 purification13 Heat exchanger

For the integrated process :

At the same time coal is converted as described above . The liquid

hydrocarbons are hydrogenated in an additional hydrocracking stage .

The gaseous hydrocarbons (C 1 . . .C 4), resulting from the hydrogena-

tion process, are fed into the steam reformer with saturated steam

after desul£urization and pre-heating . Under addition, of nuclear

reactor heat, the process produces hydrogen . After the usual re-

fabrication of product gas (cooling, conversion, CO 2 purification,H2/CH4 low-temperature separation), about 80 % of the hydrogen is

led to the hydrogenation reactor,aud the remaining 20 % to thehydrocracking stage .

A brief estimation of the conversions gives the following result :The hydrogenation stage (H coal reactor + hydrocracking plant) hasthe following balance (calculated in analogy to the Bergius--Pierprocess) :

1 .7 t of soft coal + 2600 I3m3H2 4- 1 t of gasoline + 0 .49 t ofC 1 -C4 + losses

(coke,

gas,

etc .)

The amount of 490 kg C 1 . . .C 4 is used in the steam reformer for ge-neration of required hydrogenating hydrogen, and decomposed to-gether with saturated steam .

490 kg of

C 1'*'C4 + 4500 Meal of

reactor heat -} 2600 Nm3H2

1 .7 t C + 4500 Meal reactor heat + 1 t gasoline .

In addition to heat, electrical energy must be available to performthe hydrogenation process, namely for compression of hydrogen,grinding of coal, moulding of pasta etc . Per ton of gasoline pro-duced, this will require 600 kith of electricity . This energy canbe produced in a ,steam turbine plant with 1300 Meal of reactor heat,assuming an efficiency of 40 % .

The following table considers that amount of coil that a nuclearreactor of 3000 MWth could process within 8000 h/a .

Conversion

1 .7 t C + 5800 Mcal RR -* 1 t gasolineHeat capacity of the

nuclear reactor

2.06 x 10 10Mcal/A 'Generation of gasoline

3.56 x 10 6 t/aConsumption of coal

6.05 x 10 6 t/a

The technical parameters of a nuclear-heated steam reformer, suchas pressure, temperature,

gas analysis,

etc.,

are given in Vol .

3 .

7 .5

Hydrocra cking of Heavy Fuel Oils (ref . 5)

The catalytic cracking of hydrocarbbns under a high partial pressure

of hydrogen is called "Hydrocracking" . The different porperties of

raw materials and of the inventory materials (i .e . the content of

asphalt, salts, vanadium, and nickel) result in differences in the

process conditions (p, T), catalyst inventory, reaction mechanism,

reaction heat, consumption of hydrogen, and qualities of the pro-

duct . One must therefore distinguish between hydrocracking of

distillation residuals and hydrocracking of distillates . The pro-

cess of hydrocracking of residuals (as opposed to hydrocracking of

distillates) does not appear to be adequately developed, because

the heterogeneous components of residuals presently cause a de-

activation of the catalyst . Thus the catalyst cannot fulfil the

task of supporting the cracking reaction and accelerating the

hydrogenation .

Application of Distillate

The process conditions are :

p

100 - 140 atT =

320 - 400 o CThe consumption, of hydrogen depends upon the rate of conversionof distillates into light fractions of hydrocarbons and amountsto

250 - 450 Nm 3 H 2 /t inventory .

Hydrocracking of heavy fuel oils is applied in, one or two-stageprocesses to produce a mixture of components for gasoline or middledistillates . In the case of a two-stage process, the first stageserves to remove nitrogen, and sulfur from the inventory materialswhile the real conversion of the inventory material into gasolineoccurs in the second stage . The development of new Pd--containingcatalysts with molecular sieves, as opposed to conventional ca-talysts

(Ni/Mo/W sulfides on Al 203 /Si0 carriers),

results in adecreased sensitivity of the catalyst to deactivation. by nitrogencompounds . By introducing a hydro-refinery catalyst of the Ni/Wsulphide-type in molsieve catalysts, it is possible to perform aone-stage process with an unrefined initial material .

The following table shows the operation data of a one-stage processplant for generation of mixed components for carburetor fuels ofheavy straightrun distillates and recycle oils of catalytic crackingplants . A. plant of 1,5 x 10 6 t/a operating according to the uni-cracking-THC process is described and has a total throughput off2 .5 x 10 6 t/a .

H2- consumption : 272 Nm3/m 3 feed

Feed 0ut2 utDensity (t/M3 0 .879 Cl-.-C3 (Nm3/m3 feed) 17 .0Boiling : i - C 4 (Vol %) 12 .9Starting point ( 0 C) 165 n - C (Vol %) 7 .310 Vol % converted 263

4C 5 - 82 o C (Vol %) 33 .750 Vol % 331 82 - 205 oC (Vol %) 69 .490 Vol % 403 Total C +

4(Vol %) 122,3End point (0C) 449

Sulfur (% by weight) 0 .66Nitrogen (Ppm) 230

Product Parameters

C5 - 82 o C

82 - 250 o C1 03Density(at

.Octane

numbeAfterROZ +

After MOZ +

1) BTA = Tetraethyl-lead

OPERATION DATA OF A ONE-STAGE UNTCRACKING-IHC PLANT

As the table above shows, hydrocracked gasoline can be used as

a mixed component for carburetor fuels . This fact can be explained

by the dominating formation of isoparaffins and hydrogenation

of aromatics, and by the ring openings of the condensed aromatic

systems . For example :

2 H., H2-CNfGH-CH5

CH,

_

H: CH4H-CH.,

Nallbthaline trt_t ;aliüC i-Butyl- Benzene i-ButaneBenzene

For the fabrication of -heavy fuel oils into light hydrocarbons,

further processes have been developed which operate according to

the fixed-bed process, similar to those of the Unicracking-IHC

process of Union Oil Co . of California, and Esso Research and

Engineering Co .

For example :

Isomax process

of Chevron Research Corp . and

of Universal-Oil Products Co .

H .G . process

of Gulf Research and

Development Co . and

of Houdry Process and Chemi-

cals Co .

5 .6C)in t/m0 .663 0 .77

r

3 ml BTA 1) 99 .5 84 .93 m1 BTA 1 ) 102 .1 82 .8

Application of Residual Oils

Hydrocracking of heavy residuals upto a distillation end point

of mar . 600 o C requires more well-defined operating conditions

than are used for hydrocracking of distillates . The operating

pressure is between 140 and 210 at, depending on the distillation

end point of the feed material. . The consumption of hydrogen for

hydrocracking heavy residuals is dependent upon the conversion rate

as in the case of distillates . A typical range of H2 consumption

is between 165 and 350 Nm3Jt of feed material .

The available hydrocracking products differ from products of

primary crude oil fabrication of corresponding distillation plantsmainly by the higher degree of refinement . This fact can be ex-plained by a dominating hydroalkylation of hydrocarbons withstraight--chain fractions .

Two kinds of processes have been developed for hydrocracking ofheavy residuals which are characterized by a different arrangementof catalysts ;

(a) fixed-bed process(b) fluidized-bed process

The fluidized-bed process is favoured for large installationsbecause here, as opposed to the fixed-bed process, the catalystactivity is maintained by removing and replacing some parts ofthe catalyst . Moreover, a nearly isothermal cracking is possiblewith a fluidized-bed which is difficult in case of fixed-bed cata-lyst because of the exothermic reaction . The latest development3of the fluidized-bed process prefer smaller dimensions of thecatalyst kernel (< Smm), which needs no cycle pumps for keeningthe catalyst in suspension . The development of this process bearsmore and more similarity to the coal hydrogenation process,

FeedResidual oil (atmospheric bottom)Dens.ity(at15 .6 O C) 0 .968 t/m3

Table 7 .5 - l Use of Products of a H-Oil Plant for Conversionof Fuel Oil-S into Middle Distillates (Kerosene)

H2-consumption : 310 Nm3 /t Feed

The following processes have been tested in pilot plants :

RCD-Isomax

based on the fixed-bed process . It is used for

desulfurization of heavy fuel oil .

H-oil process

of Hydrocarbon Research Inc . (HRI) . It is based on

the fluidized-bed process . The technical plants

have a capacity of up to 1 .4 x 10 6t/a, with reac

tor diameters of 4 m .

Table 7 .5 - 1 shows the utilization

of such a plant which con-

verts the atmospheric residual of a Kuwait oil and returns the

resulting heavy oil gas to the H-oil reactor .

Heavy fuel oä.1 can be converted into gasoline fractions by two

process stages . During the first stage, heavy fuel oil is con-

verted into middle distillates in a H-oil plant, and in a se-

cond stage these are converted into gasoline . A limitation with

regard to the concentration of sulfur in heavy fuel oils does not

Sulphur

Product

4 .06

Wt . %

Wt .7-

Sulphur in Wt .--% 'Density in-t/m3

H2S, NH3 4 .1 - -C 1 . . .C 3 7 .0 -- -

C 4 1 .9 -- -C5 0- 82 C 7 .2 - 0 .66982 0- 177 C 12 .6 0 .01 0 .75177 0- 343 C 59 .3 0 .05 0 .84Residue > 566 oC 10 .4 3 .0 1 .12

102 .5

Fig . 7 .5 - 2

Process Schematic for Generation of Hydrogen from

Heavy Fuel Oil with Nuclear Process Heat

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exist . The light hydrocarbons obtained by hydrocracking are forthe most part free from sulfur, which permits their direct intro-duction into the steam reformer catalyst . As opposed to the straightrun gasoline, the cost for a desulfurization stage are not necessary .

- 20 -

7 .6

Chemical Heat Pipe Systems

(ref . 6)

Th e Principle of the Chemical Heat Pipe

Nuclear power plants could be sited far away from densely populatedareas if .nuclear heat could be delivered by a system which trans-ports chemical heat via cold gases . The most favourable energycarriers are the gases CO and H2 which are produced at high tem-perature by steam reforming of methane with saturated steam . Fig .7 .6 - 1 shows the flow scheme of the process .

Fig . 7 .6 - I Flow Scheme of the Chemical Heat Pipe

CH, r H?0 -s- C0+3Ni -50kcollmk1 district heal,electricity,woste heat

of nuclear heat, according to the following :

CH4

+ H2O -} CO +

3H2-

50 kcal/mole .

Methane and steam are converted catalytically, with the addition

Without the addition of a gas processing plant, the cooled productgas is compressed to a pressure of ti 64 b (typical. for natural gaspipelines), and delivered through a cold pipe to a customer, who isfar away from the reactor plant . This customer utilizes the re-verse reaction in which the gas is again converted catalyticallyinto methane and steam, Here the heat added by the nuclear reactorfor the steam reforming process is recovered . According to today's

1 Nuclear reactor 4 Blower 7 Methanation2 Steam reformer 5 Use of waste heat 8 Heat exchanger3 Pre-heater 6 H , CO compressor 9 CH4 compressor

technology it is possible to achieve gas temperatures of 450 oCby this methanization process . This amount of heat is sufficientto produce ,hot, water (130 0C/10 b) or steam (e .g . 350 . . .400 0 C/100 b) .

A following stage of development should involve a gas temperatureincrease to 600 0 C, by which one can also operate a steam turbine(530 oC/180 b) . This allows the transmission of energy through acombination of heat and electricity, accompanied by all the associatedthermodynamic advantages . As will be explained in a followingchapter, this achievement will depend upon the use of a suitablecatalyst . The produced methane is returned again to the steamreformer via the long-distance pipeline . After ä correspondingpre-heating it is utilized again together with water as feed ma-terial for the steam reformer . The process proposed here can,according to the chosen circuit, transmit 50 to 75 % of the reac-tor power into energy at a long distance from the reactor . Theremaining heat is transformed into either electrical energy andheat (hot water at 145 oC/15 b), or only electrical energy . Apartfrom small leakage losses, the system does not consume raw ma-terial . It can also supply other customers with heating gas and,

the possibility of energy accumulation occurs because o£ the ex-tensive pipe system .

Steam Reformer

Helium from the HTR in the temperature range between 950 and600 oC is used tq perform the steam reforming reaction and to

superheat feed material from 450 oC to the final temperature of825 oC (see figure) . After heat transfer from the inner pigtails

to the feed gas the product gas leaves the steam reformer at 600 0C .

Helium enthalpy between 600 and 350 oC is used for steam generation

(535 OC, 195 b) . This steam is led to a back-pressure turbine

in which electrical energy is generated at a very high efficien-

cy

(85 %) . 75 % of the steam is removed in an intermediate

- 2 2 -

stage at 335 o C and 47 b . This amount serves as an additive of sa-turated steam to methane in the steam reformer, after the mixture

has passed a pre-heating stage where it was heated to 450 o C andthe gas cooled to 484 0C . After the remaining 25 % of the turbinesteam has been expanded to 6 b, it is, as the figure shows, eitherconsumed to generate local heat or led to a condensingsystem to generate electricity .

30LW1~

Fig . 7 .6 -- 2 Steam Reformer Flow Schematic :Flexible System for BothShort- and Long-D istance E nergy -

P

450°C535°C

~Plgrail5

1747bh

335°C

159°C

05 b 6b

728MW SpOCB heat535MW

1070 Vh591 Vh

175OMW

600°C

1250MW

149'C

40°C

40

5qunnlyliV Co

ZBVol -%H 2 443Vol-%C02 S,2Vo1=/.CH4 11,6Vdr/.H 20 3L1 VbF%.

950°C. 40b 6001C235MW

M"C

2300 1/h 314°C220 b

170°Cslacehed333MW

39MWelMethan!1, 07 106Hrr/h

56̀ T

10°C

Fresh water639äh

z0-C, 20b

y:.VCO

IWO=/.H2 642VM/.C02 7,6V6W.CH4 16,9Sb1~/

6 3Flow : 3,00-10 Mm/h IWal

40°C

--r Waste heat

~83MW

25°C

The residual enthalpy of product gas below 484 o C serves to pre-heat feed methane, to pre-heat feed water and to generate heatin form of hot water for local room heating, etc . The residua].enthalpy of product gas below 80 o C cannot be used and representswaste heat of this combined process . The portion of 190 MW is very

Long-distance energy : 1772 MWShort--distance energy : 868 MWElectrical power 255 MWWaste heat 190 MW

low .

23 -

The . process

itself is performed at

40 b .

The gas mixture

(CO,

H2 ,CO 2 , CH4), after having the water removed, is compressed to 64 barand delivered to the consumer (methanation) . Methanereturning from the consumer is compressed from 20 to 45 bar andthen returned into the steam reforming process .The process scheme suggested here uses about 60 % of the nuclearreactor pourer for the chemical heat pipe . This percentage couldbe somewhat increased by variation of helium temperatures . 29of the nuclear reactor power is released in the form of hot wateror steam for local district heating, etc . The electrical power of255 MW is required for the most part to cover the requirements forgas compressors, helium blowers and feed-water pumps . In principle,it is possible also to convert electrical, energy into short-distanceenergy (see fig . 7 .6 - 2) .

Small variations of the upper and lower helium temperatures in theprocess present a completely different proportion between theenergy released locally and that transmitted by the chemical heatpipe . Also the variation of helium outlet-temperatures from thesteam reformer can deliver a proportion in favour of the chemicalheat pipe system . An increase of helium inlet-temperatures in thereactor to 450 oC (as is proposed for helium turbine plants) wouldallow one to apply about 75 % of the reactor power to the chemical.heat pipe .

40M

- 24 --

Fig . 7 .6 - 3 Steam Reformer Flow SchematicVariant with Release System for both Electrical-and Long-d istance Energy (Chemical Heat pipe

3 "10gVm 31hVolCO 11,3C0 2 7,6CH,,16,91~64,2

36.1f rn3!Vol

CO 7,8C0 2 5,2CH 4 11,6H2 44,3

H20 31.1

Methanation

The principle question of how to recover nuclear heat used forthe gas-mixture in the steam reforming process is answered infig . 7 .6 - 4 and 7 .6 - 5 . The reaction rate, which is defined asthe number of moles of CH4 produced compared to those of C of

theavailable product gas, is platted as a function of reaction pressureand reaction temperature . It should be recognized that at a reac-tion temperature of 450 oC and a pressure of 40 b almost 95of the feed-material is converted . Although this incomplete con-version does not involve any energy losses the transportationcosts are raised by 5 % because of the added volume flow ofgas .

Long distance energy : 1772 MWShort distance energy : --Electric Power : 384 MWWaste Heat : 944 MW

Fig . 7 .6 - 4Reaction Rate of Methanation

0600 6Ö0

- 2 5 -

800 1D00-~ TI°C

Fig . 7 .6 - 6

Energy Conversion of .Methanation

0, 2001m

400 600 800 1000

if the methanation temperature is increased to 600 °C, the conver-

sion is reduced to about 90 % . However, it is possible toperform

another methanation at lower temperatures so that a conversion

of 95 % could occur in any case . Fig . 7 .6 - 6 shows the ratioof

energy withdrawal in methanation to the energy demandfor steam

reforming as a function of pressure and temperature . Thispropor-

tion, however, must not be interpreted as the efficiencyrate ; it

is only a measure of additional demands of the gastransportation

system compared to that required with completeconversion .

While methanation catalysts far a reaction temperatureof 450 . . .

500 ° C should be acknowledged as well-'tested, one mustobtain

further data for a temperature of 600 °C . Such amethanation tem-

perature has not been of interest until now . Forexample, the

American gasification of coal, which requires thisprocess stage,

only aims at the elimination of residual contentsof CO and H2 ,

and not the exploitation of energy at highestpossible tempera-

tures . High methanation temperatures are likelyto cause mechani -

cal instabilities of the carrier material andrecrystallisatian

effects . High temperature$, however, do not requirea high activi -

26 -

ty of the catalyst . It is therefore important to test the available

catalysts and to determine their life-time .

,A basic circuit of methanation is given in fig . 7 .6 --6 . The incom-

ing cold gas is pre--heated to 70 0C, enters the first reactor and

leaves at 450 OC, then heat is removed in form of hot water or

steam . The product gas of the first stage is mixed with fresh gasand led to the second stage where it is heated to 450 oC as a

result of reaction energy . After another removal of heat andaddition of fresh gas, it is heated in a third methanation reactorto 350 0C . A portion of the product gas of the third reactor isreturned to the first reactor, and the remainder is recooled to40 0C and

fed

into

the methane pipe .

-2 7

Fig . 7 .6 - 6

Methanation Flow Schematic

250°C

45860M3h

120200Nm3h

45.5b321°C

700kW

42b

350°C

450°C

46930N ~l

45000

1258D1`-m3

FA&r428BCC60°C

FrRA41°C!50°C

310°C

FA&r4

4566

130°C

82700kglh

70°C

91°C

20°C

45b

1300C, 8b

300700kglh

70°C840000kg/h

105370Nm31h

Feed gas pipe

39363 Nm3h

MethanePipe

21750kglh

7 .7

Steam Gasification of Coal (ref . 7)

The following reactions are basic to the gasification of coal :

-- 28 -

C

+

2H 20

=

CO 2

+

2H2

By suitable combinations of these reactions, it is possible toproduce methane, CO + H2 , or hydrogen .

a) Production, of Methane (SNG)1 . -Gasifier

2 C + 2 H 2 O = 2CO + 2H2 ; AN = +56 .8 kcal/2 moles C(endothermic)

Reaction temperature = 900 o C

2 . Shift reactor :

CO + H 2 O =CO2

+ H2pH = -9 .9 kcal/male

(exothermic)Reaction temperature = 400 o C

3 . Methanatlon :

CO + 3H2 = CH + H 2 O 4H = -49 .0 kcal/mole(exothermic)

Reaction temperature = 400 o C

Overall Reaction :2C + 2x2 0 = CH4 + co t

b) Production of CO + H2 :C + H 2 0 = CO + H 2 OH = 28 .4 kcal/mole

c) Production of hydrogen :C + H 20 = CO + H2 AH - 28 .4 kcal/moleCO + H 2 0 W Co 2 + li2 4H = -9 .9 kcal/mole

Gasification reaction :

C + H2 O = CO + H 2 DH = +28 .4 kcal./mole

Shift reaction :

CO + H 2 0 = CO 2 + H 2 AH = -9 .9 kcal/mole

Methane s ynthesis_ :

CO + 3H2 = CH4 + H 2 O AH = -49 .0 kcal/mole

The actual heat consumption is much greaterwould be calculated from the stoichiometricduction of methane by steam gasification,ofheat is 4500 - 6500 kcal per standard cubic145 kcal/mole CH0'

Fig . 7 .7 - 1 Steam Gasification of Coal (schematic)

Ti, 950'C

(1)High temperature reactor

2 Intermediate heat exchang.3 Helium circulator

@Coal gasifierQSteam generator

coasteam (coal drying, gas purlikallW

Helium circulatorAsh + tar removal

Waste heat util

®C021 H2S_removali0 Methonation

than that whichdata . For the pro-coal, the requiredmeter

of

CH4

(100

-

Steam turbineCondenserFeed water pump

The figure shows a schematic of the coal gasificationprocess

coupled to a high temperature reactor . Heat is transferedinto

a secondary circuit (He, H2O or C0 2 ) through a heatexchanger .

The temperatures given in the figure are typical. forthe gasifi -

cation of bituminous coal .

Considerable development work will . be required tocouple the steam

gasification process to a high temperature nuclearreactor .

(a) Transfer of heat from the nuclearreactor on a commercial

scale .

(b) What is the required reactortemperature?

(c)

Design of

the gasif ier .

Several heat transport concepts have been proposed, utilizing theflow of solid materials like cake or sand, or an excess of steamas the heat carrier . With solid materials, very large heat ex-changers and complicated solid transport systems are required .This seems to be too complicated and too expensive .An attractive method would be to heat a fluidized bed of coal andsteam by means of tubes containing flowing hot gas . The advantagesof this method are :

(a) highest overall heat transfer coefficient, and therefore aminimum of required area in the hot gas high alloy steel, tubes .

(b) highest conversion rate per unit volume .(c) maximum gasifier output for a given nuclear reactor power and

outlet temperature .

Because of maintenance requirements, two heat transfer circuitsare needed :

(a) the primary helium circuit (40 bar) for cooling the nuclearreactor,

(b) a secondary circuit (42 - 45 bar) with He, C0 2 , or steam .For the design of the gasifier one must have further information onthe kinetics of the coal -steam reactions, as well, as the heat trans-fer properties of the tubes in the fluidized bed . The reaction ki-netics have been studied on a laboratory scale

Fig . 7 .7 - 2

Reaction Rates for Lignite, Hard Coal, and HardCoal with Catalyst

-

30

--

For lignite the reaction begins at a temperature of about 600 oCand reaches a conversion rate of 5 % carbon per minute at 690 OC .The reaction of hard coal begins at about 700 oC and reachesa conversion rate of 5 % carbon per minute at 800 0C . The same reac-tion rate is achieved at about 750 o C when catalysts are used .

The throughput of coal is also limited by the heat transfer tatefrom the heated tubes .In fig . 7 .7-3, steady-state operating pointsare represented by the intersections between the curves for heattransfer and for required heat .

Calculations are made for hard coal (lignite) in fluidized bedswith a density of 0 .2 t/m3 (0 .1 t/m3 ) and a heat requirementof 1 .4 Gcal/t (1 .1 Gcal/t) . Also it is assumed that the temperature ofthe helium leaving the gasifier is 50 oC greater than the gasifi-cation temperature in the fluidized bed . Also, the outlet tempera-ture of the nuclear reactor must be about 50 o C higher than thegasifier inlet temperature, because of the intermediate heat ex-

changer .

Fig . 7 .7 - 3 : Steady- State Operation between transferred andrequired heat

I) dry2) with catalyst

The following operating points are derived from the curves of thefigure above :

Production Data of a Gasifier

One can also gasify hard coal with a nuclear reactor heliumoutlet temperature of 1000 0C. This, however, reduces the produc-tion by 25 % . With a single nuclear reactor, one would have seve-ral gasifiers, with the numbers depending upon the size of theunits and the output of the reactor .

In designing a large gasifier (ca . 50 t/hr), one mus.t consider thevelocities in the fluidized bed as well as the practical limita-tions of overall size . Fig . 7 .7' - 4 shows a horizontal cylindricalgasifier with a length of 25 m and a diameter of 5 .5m . This is thelargest size that can be transported . The unit should be installedwith a downward slope toward the ash exit in order to facilitatecontinous ash removal .

Nuclear ReactorHelium Outlet-Temperature°0

HardFlow

(t/hr)

CoalNet heat valueGas production

(Gcal/hr)

Flow

(t/hr)

LigniteNet heat valueGas

(Gcal/hxjton

1100 6075 2)

4806002 }

115 700

1000 4560 2)

3604802 )

90 550

900 45 2) 3602) 70 420

ash

Fig. 7 .7-k : Horizontal Cylindrical Gasifier

l:Iti

25.000 ----,~--

raw

33 -

TPT ' II'I Vafi

WW.,.X~~rri; yr~r~

walat

char

REFERENCES

- 34 -

re£ . 1

Eickhoff, li .G :Technologische und wirtschaftliche Möglichkeiten,die sich durch den Einsatz des Hochtemperaturreak-tors für die künftige Mineralölversorgung der BRDergebenJÜL-1017-RG, November 1973

ref . 2

Hüttner, R . ; Teggers, H. :Braunkohle, Wärme und Energie,1971, Heft 4, S . 133

ref . 3

Maschlanka, W. ; Kehl, P . ; Knapp, H . ; Voss, H . :Ilse of the Midland-Ross Direct Reduction in anElectric Steel Hill of Small to Medium CapacityExemplified by Hamburger Stahlwerke GmbH .International Symposium on Direct Reduction ofIron Ore, Bucharest, Sept .1972

ref .

4

Kugeler,

11 .Energieprognose für die Bundesrepublik Deutschlandunter Berücksichtigung des Einsatzes von Kernwärmezur Vergasung fossiler RohstoffeJÜL-745-RG, März 1971

ref . 5

see ref . 1)

ref . b

Bohn, Th . ; Dietrich, G . ; Kugeler, K. ; Kugeler, M. ;NieSen, H.F . ; Schlenker, H .V . :Nuclear District Heating and Nuclear Chemical HeatPipe Energy (in German)JUL-1477-RG, June 1974

ref . 7

Van Heek, K.H . ; Jüntgen, H . ; Peters, W . :Journal of the Institute of Fuel, July 1973 .Van Heek, K.H . ; .TÜntgen, H . ; Klein, G . :Gordon Research Conference on Coal Science,July 2-6, 1973