Methanol II

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Investigations on CatalyzedSteam Gasification of Biomass

1.K. Mudge

S. L. Weber

D. H. Mitchell

L. J. Sealock, Jr.

R. J. Robertus

January1981

Prepared for theU.S. Department of Energyunder Contract DE-AC06-76RLO 1830

Pacific Northwest Laboratory

Operated for the U.S. Department of Energyby Battelle Memorial Institute % ,



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INVESTIGATIONS ON CATALYZEDSTEAM GASI F ICATION OF BIOMASS

L. K. MudgeS. L. WeberD. H. M i t c h e l lL. J. Sealock, Jr .R. J. Robertus

January 1981

Prepared forthe U.S. Department o f Energyunder Contract DE-AC06-76RLO-1830

Paci f ic Nor thwest LaboratoryRic hlan d, Washington 99352



SUMMARY

The P a c i f i c Northwest Lab ora tor y (PNL) i s conductin g an ongoing study,

sponsored by the Biomass Energy Systems D i v i s i on o f t he U.S. Department o f

Energy, on the ga si f i ca t i on o f biomass i n the presence o f catal ysts . The pur-

pose of the s tudy i s t o eva luate the technica l and economic fe a s i b i l i t y o f pro-

ducing spe c i f ic gas products v ia the c at a l y t ic gas i f ic a ton o f biomass. This

report presents the results of research conducted from December 1977 t o October

1980.

The study was comprised o f l ab or at or y studie s, process development, and

economic analyses. The la bo ra to ry stu di es were conducted t o develop op er at in g

condi t ions and cata lys t systzms for generat ing methane-rich gas, syn the sis

gases, hydrogen, and carbon monoxide; these st ud ies al so developed techni ques

f o r ca ta ly st recovery, regenerat ion, and rec ycl i ng. A process development unit

(PDU) was designed and cons tru ct ed t o eva lu at e 1aboratory systems at condi t ions

approximating commercial operat ions . The economic analyses, performed by Davy

McKee, Inc . f o r PNL, evaluated the fe a s i b i l i t y o f adapting the wood-to-methane

and wood-to-methanol processes to full - sc al e commercial operat ions. Pla nts

were designed i n t he economic analyses t o produce f uel -grade methanol from

wood and su bs t i tu te nat ur al gas (SNG) from wood v i a ca ta l y t i c g as if ic at io n wi th

steam.

Condit ions developed i n the l abo rato ry fo r generat ion o f a methane-rich

gas and a synthe sis gas f o r methanol pro duc tio n were confirmed i n t he f l u i d -

bed, PDU operat ions. No p ar t i c ul a r problems wi th scale-up were experienced.

Ca ta ly t i c steam g a si f i c a t i o n o f biomass has several advantages over con-

v en t i ona l gas i f i c a t i on .

No oxygen i s re qui red f o r ge nerat ion o f a methane-rich gas or a

methanol synth esi s gas, the ref ore , no oxygen p l a n t i s needed.

L i t t l e or no t ar i s produced req ui r i ng on ly s imple gas c lean ing and

waste-water treatment equipment.

No s h i f t reac tor i s requ i red f o r methanol synthes is.



Methanat ion requirements are low re su l t in g i n h igh conversion

e f f i c i e n c y .

a Yie lds and e f f i c ienc ies a re g rea te r .

The com ner c i a l ga si f i ca t i on p la nt s were designed t o process 2000 tons

(1800 t ) o f d r y wood per day and 10% of t h a t cap aci ty. For the large- sca le

methanol p lant , product ion i s 997 tons (900 t ) of methanol per day wi th a

hig her hea ti ng value (HHV) o f 9784 Bt u per pound. For the large- scale SNG3

pl an t, prod uct ion i s 21.6 m i l l i o n sc fd (610,000 nm /day) o f SNG w i t h a HHV3

o f 956 Bt u per sc f (35,600 kJ/nm ) . A l l y ie lds f rom the sma l l-sca le p lan ts

are 10% o f th e large- scale p lants. The p l an t design included a l l process and

support f a c i l i t i e s necessary t o convert wood. The p la n t lo ca t io n i s the

Pa c i f i c Nor thwest .

The cap i ta l costs fo r the la rge- scale wood-to-methanol and wood-to-methane

p l an ts are, on a September 1980 basis, $120,830,000 and $95,115,000, resp ec-

t i ve l y . P roduct i on cos ts tha t a l l ow fo r r e tu rn on cap i t a l were ca l cu la ted a t

va rious wood cos ts fo r bo th u t i l i t y and p r i va t e investo r f inanc ing . At a wood

cos t to the l a rge-sca le p lan t o f $20 per d ry ton ( $ 2 2 / t dry) , the methanol cost

was $0.55/gal ($0.15/L) and $0.69/gal ($0.18/L) f o r u ti i t y and p r i v a t e i nv es-

t o r f inanc ing , respect ive ly ; s i m i la r l y , methane costs were $ 6 . 5 0 / m i l l i o n Btu

and $ 8 . 1 0 / m i l l i o n Btu, r espec t i ve l y . S im i l a r cos ts fo r the smal l- sca le p lan ts

were 2-112 t o 3 t imes g rea te r than fo r the la rge- scale p lants. The co st ca lcu -

l a t e d b y t h e u t i l i t y f i n a n c i n g method i n c lu d es a r e t u r n on e q u i t y o f 15% and an

in te re st r a t e of 10% on debt. The pr iv at e inv est or f i nan ci ng method, which i s

100% equ i t y f i nanc ing , incorpora ted a discounted cash f low re t u rn on equ i t y o f

12%. Economics are presented as a fu nc t io n o f wood cos t and p la nt siz e.

The thermal e f f i c ie nc y o f the wood-to-methanol pl an t i s 53%. The thermal

e f f i c i e n c y o f t h e w o o d-to-methane p lan t , w i thou t tak ing cr ed i t fo r byproduct

char, i s 58%.

Continued research i n c a t a l y t i c g a s i f i c a t i o n o f b iomass w i 11 i nc lude labo-

ra t o ry stud ies t o examine more cat a ly sts , PDU s tu di es t o t e s t r e s u l t s o f t h e

c a t a l y s t development, and economic analyses t o develop a computer code th a t

w i l l determine t he ef fe ct s o f changes i n process y i e l ds and requirements on

overa 11 p l an t economi cs.



CONTENTS

SUMMARY .

INTRODUCTION.

CONCLUSIONS .

RECOMMENDATIONS

LABORATORY STUD1ES .

EXPERIMENTAL EQUIPMENT .

FEED MATERIALS .

EFFECTS OF ALKALI CARBONATES.

EFFECTS OF SECONDARY CATALYSTS

RECOVERY OF ALKALI CARBONATE CATALYSTS .

SYSTEMS FOR METHANE GENERATION

SYSTEMS FOR GENERATION OF METHANOL SYNTHESIS GAS

SYSTEMS FOR GENERATION OF AMMONIA SYNTHESIS GAS

SYSTEMS FOR GENERATION OF HYDROGEN .

SYSTEMS FOR GENERATION OF CARBON MONOXIDE.

EVALUATION OF ALTERNATIVE BIOMASS MATERIALS .

FUTURE LABORATORY STUDIES

PROCESS DEVELOPMENT UNIT STUDIES .

PDU DESCRIPTION

G a s i f i e r .

W o o d F e e d S y s t e m

.

G as S u p p l y S y s te m s .

G a s C l e a n u p S y s t e m .

I n s t r u m e n t a t i o n a nd D a t a A n a l y s i s .

iii



PDU OPERATION .

Test Procedures and Safety Considerations

Feedstock f o r th e PDU .

PDU Res ult s w i t h t he Agitated-Bed Gas i f i e r

PDU Re su lt s Wit h th e Fl uid-Bed Gas i f i e r .

Cal cul at io n o f Equ i l ib r iu m Gas Composi tions

FEASIBILITY STUDIES .

ECONOMICS OF CATALYTIC GASIFICATION

Wood-to-Methane Plant .

Wood-

to-

Methanol Plant.

COMPUTER MODELLING STUDIES .FUTURE MODELLING EFFORTS .

ACKNOWLEDGMENTS

REFERENCES



FIGURES

Sumnary o f S t u d y A c t i v i t i e s .

Batch-

Feed Reactor.

Continuous Wood-Feed Reactor .Volume o f Gas Pro duc ti on f o r Catal yzed and Noncatal yzed SteamGa si f i ca ti on o f Wood a t 5500C

Re1a t i v e Cata l ys t E f fect i veness a t 6500C

E f f ec t o f Temperature on Gas Prod uct ion With and WithoutPotassium Carbonate Catalyst .

Effect of Sodium Carbonate Concentration on Gas Yield at 6500C

Catalyzed Test Results on Remaining Carbon Versus Time .Uncatalyzed Test Results on Remaining Carbon Versus Time

React ion Constant Var ia t ion With 1/T for Catalyzed SteamGas i f ic at i on o f Wood .

Postu la ted Ef fect iveness Facto r

Hydrogen Chemisorption Apparatus .

Pore Di st r i bu t i on i n Fresh and Coked Cata l yst .Coked Ni-3266 Catalyst, 3340X .

Unused Ni-3266 Cata lysts, 10,021X .

Relative Concentrations of Elements on Fresh and ExposedCatalyst Surfaces Employed for Methanantion .Pore Di st r i bu t i on i n Cat a lys t Used fo r Synthesis Gas Generat ion

Regenerated Ni-3266 Catalyst Employed For Hydrocarbon SynthesisGas Production, 10,021X .

SEM Microgr aph o f Fresh Ni-3266, 3960X.

SEM Microgr aph o f Ni-3266 A f t er 28.6 h o f Operation, 13,200X .

SEM Microgr aph o f N i-3266 Af te r 61 h o f Operat ion, 13,200X .



SEM Microgr aph o f Carbon Lay ers on Ni-3266 After 72 ho f Operation, 132X .

SEM Micr ograph o f Coked N i -3266 After 88 h o f Operation, 132X

SEM Mi cr og raph o f Coked Ni-

3266 Aft er 117 h o f Operation, 132X

Relative Concentrations of Exposed Elements on Ni-3266Cata lyst Sur faces .

Schematic o f Or ig in al PCU .

Photograph of Or ig ina l PDU

Schematic of the St i rred-Bed Gasi f ie r

Bottom Head o f Sti rred-B ed G a s i f i e r .

View of Gasif ier Cross Section .Schematic o f Flui dized -B ed G a s i f i e r .

View o f Lock Hopper .

Gas Heater and Controls .

Gas Recycle System .

E l e c t r o s t a t i c P r e c i p i t a t o r

PDU Schematic As o f October 1980

.

Methane Yields from the Agitated-Bed Gasi f ie r as a Funct iono f Temperature and Cat al ys t .

Carbon Conversions f o r Methane Pr oductio n i n t heAgitated-Bed Gas i f i e r

Synthesis Gas Yields as a Function of Temperaturei n the Ag itated-Bed Gas i f i e r .

Carbon Conversions f o r Synthesis Gas Production i n th eAgitated-Bed Gas i f i e r

E f fec t o f S team-to-Wood Ratio on Gas Composition from theFluid-Bed Gas i f i e r .

Tempera tu re Pr o f i l es i n F luid-Bed Gas i f i e r .

v i ii



Comparisons of Catalyzed Versus Uncatalyzed Resultsf o r Methane Produ ction .

Comparisons of Catalyzed versus Uncatalyzed Resultsfor Synthesis Gas Production .

Eff ect s of Temperature on the Dry Gas Eq ui l i br iu m Compositiona t Atmospheric Pressure f o r Steam-to-Wood Ra ti o o f 0.33 .

Ef fe ct s o f Temperature on the Dry Gas Eq ui l i br iu m Composit iona t Atmospheric Pressure f o r Steam-to-Wood Rati o o f 1 .Potent i 1 Methane and Methanol Yields at Atmospheric Pressure

Effect of Temperature on Equi l ibr ium Standard Heats of React ion

Ef fe ct o f Pressure on Equ i l ib r i um Standard Heats o f React ion .

Effect of Pressure on Product ion of a Methane-Rich Gas .Ef f e c t o f Pressure on Produc tion o f a Methanol Synthe sis Gas .Wood-To-Methane Process Areas .

Ef fe ct o f Wood Pr ic es on Gas Cost f o r 2000 t o n l d a y Dry Wood .

Wood-To-Methanol Process Area .

Ef fe ct o f Wood Pr ic es on Methanol Cost f o r 2000 t o n l d a y Dry Wood



TABLES

1 Comparative Analysis of Wood and Bark . 13

2 Gas Yields at 6500C 19

3 Elemental Yields at 6500C 19

4 Gas Yield from Different Wood and Bark Materials at 6500C . 20

5 Reaction Rate Constants for Wood Char-Steam Reaction in thePresence of Catalysts 24

6 Activation Energies for the Wood-Char/Steam Reaction . 25

7 Yields from Batch Experiments of Wood-Steam Gasification

in the Presence of Catalysts.

28

8 Carbon Conversions and Energy Efficiencies in Batch Experiments . 28

9 Catalysts Tested for Methane Production . 3 1

10 Results of a Catalyst Lifetime Test for Methane Production . 32

11 Catalyst Performance Following Catalyst Regeneration . 33

12 Catalyst Characterization Tests Performed on a Ni-3266Catalyst Used For Methane Production 34

13 Lifetime and Regenerability of Catalysts Tested forMethane Generation

.

39

14 Catalysts Screened for Methanol Synthesi s Gas Production 40

15 Results of a Catalyst Lifetime Test for Hydrocarbon SynthesisGas Production . 41

16 Characteristics of a Ni-3266 Catalyst Employed for SynthesisGas Production . 41

17 Use of a Ni-Cu-Mo/A1203 Catalyst for Synthesis GasGeneration . 49

18 Lifetime and Regenerability of Catalysts Tested forHydrocarbon Synthesis Gas Production at 7500C 50

19 Catalyst Systems Tested for Ammonia Synthesis Gas Production . 51



Experi mental Result s o f Lodgepole Wood Impregnated w i t h17% K2CO3 i n the Presence o f Ai r , Steam, Gi rd le r 6-3and Si-A1 Cata lysts . 52

Experi mental Re su lt s o f Lodgepole Wood Impregnated w i t h17% K2CO3 i n the Presence o f A i r , Steam, and CommercialG i r d l e r 6-93 CO S h i f t Ca ta l ys t . 53

Catalysts Tested for Hydrogen Production Studies . 54

Ty pi ca l Res ult s f o r Lodgepole Wood Impregnated w i t h17 w t % K2C03 i n t he Presence o f Steam and Ca tal yst s 55

Exper imenta l Resul ts for CO Production from Lodgepole Wood . 56

Steam Ga si f i ca ti on o f Biomass f o r the Production ofHydrocarbon Synthesis Gas . 57

Compositions o f Feedstocks Used i n PDU Operat ions.

79

Ammonia Synthesis Gas Results from the Agitated-Bed Gasi f ie r . 86

Resu l ts o f PDU Tests Witn Nic kel Extrudate Cat aly st 89

Resu l ts o f PDU Tests with Nonporous Alumina Spheres 90

Resu l ts o f F lu id-Bed PDU Tests wi th Nicke l Cata lys t 93

Wood-to-Methane Plant Design Basis--2000 ton /day Dry Wood . 107

Ult im ate Analysis o f Feedstock . 109

Wood-To-Methane Capital Cost Summary--2000 t on /day Dry Wood . 112

Di rect Cost Sumnary for Wood-to-Methane Plant--2000 ton /dayDry Wood . 113

Total Capital Cost Requirement For Wood-to-Methane Plant- -2000 t on /day Dry Wood 114

Annual Direct Operating Costs For Wood-to-Methane Plants--2000 t on /day Dry Wood 115

Methane Cost - Ut i l i t y F inancing Method--2000 ton /day Dry Wood

.

118

Methane Cost - Equity Financing Method--2000 ton /day Dry Wood 119

Total Capital Requiremerit for 200- ton/day Wood-to-Methane Plant . 120

x i i



40 Annual Direct Operating Costs for 200-tonlday Wood-to-MethanePlant . 121

41 Gas Cost - Utility Financing Method--200 tonlday Dry Wood . 122

42 Gas Cost - Equity Financing Method--200 tonlday Dry Wood 123

43 Wood-to-Methanol Plant Design B1lsis--2000 onlday Dry Wood . 124

44 Wood-to-Methanol Capital Cost S~immary--2000 onlday Dry Wood . 128

45 Direct Cost Summary For Wood-to-Methanol Plant--2000 ton/dayDry Wood . 129

46 Total Capital Required For Wood-to-Methanol Plant--2000 ton/dayDry Wood . 130

47 Annual Direct Operating Costs For Wood-to-Methanol Plant--

2000 ton dry woodlday 131

48 Methanol Cost-Utility Financing Method--2000 tonlday Dry Wood . 133

49 Methanol Cost-Equ ty Financing Method--2000 tonlday Dry Wood . 134

50 Total Capital Requirement for 210-ton Dry WoodIDay Methanol Plant . 135

51 Annual Direct Operating Costs f ~ r200-ton Dry WoodlDayMethanol Plant . 136

52 Methanol Cost - Utility Financing Method--200 ton Dry WoodIDay . 138

53 Methanol Cost - Equity Financing Method--200 ton Dry WoodIDay . 139



GLOSSARY OF TERMS

acfm = cubic feet of vo lume per minute at actual condi t ions

atm, ATM = atmosphere or atmospheric pressure

BC = booster compressorBD = blow down

BFW = bo i le r feed water

Btu = Br i t i s h thermal un i t equal t o 1055 Joules

CS = carbon steel

DC = d i r e c t c o st

D/H = d i r e c t h i r e

DCF = discounted cash f low

DEA = diethanolamine

E,M, L = equipment, materials, and labor

F = degrees Fahrenheit

f t = f e e t

f t2

= area i n square fe et

f t3

= volume i n cubic f ee t

gpm = gal lons per minute

HO = home office

HHV = higher heat ing value

hp=

horsepowerHP = high pressure

HT = high temperature

I C I = Im per ial Chemicals, In cor pora ted

I D = inside d iameter

i n . = inches

K-0, KO = knock out

kW = k i l o w a t t

1b = pounds mass

LG = l eng th

LP = low pressure

M = thousand

MHR, MH = manhour



MHRS = manhours

MM = m i l l i o n

No., # = number

OC = overhead charges

OAH = o v e r a l l h e i g h t

, = diameter

PS = p r o f e s s i o n a l s e r v i c e s

P, AP = pressure, press ure drop

P/O = purchase order

p s i = pounds force per square inch

p s i g =ps i gauge, or pressure i n ps i minus 14.7 p s i

S/C = subcont rac t

S/T= s u b t o t a l

s c f = cub ic f ee t o f vo lume a t s t anda r d cond i t i ons o f 60°F and

1 atmosphere pressure

s c f h = cub ic fee t o f vo lume per hour a t s tandard cond i t ions o f

60°F and 1 atmosphere pressure

scfm = cub ic f ee t o f vo lume pe r m inu t e a t s t anda r d cond i t i on o f

60°F and 1 atmosphere pressure

SNG = s u b s t i t u t e n a t u r a l gas

SS = s t a i n l e s s s t e e l

T-T, T/T = t angent- t o- tangent measurement

TIC = t o t a l i n s t a l l e d c os t

TPD = tons per day

TPH, t / h = tons per hour

WH = waste heat

x v i



INTRODUCTION

The Pa ci f i c Northwest Laboratory (PNL) i s conduct ing st udi es on steam

ga si f i ca t i on of b iomass i n the presence o f catal ysts . These ,studies are spon-

sored by th e Biomass Energy Systems Di vi si on o f t he U.S. Department o f Energy

(DOE).

Ga si fi ca ti on processes are commonly used t o con ver t carbonaceous matter,

i nc lu di ng biomass, i n t o gases and some con densib le 1i qu i ds ( t a r s ) . A1though

ther e are except ions wit h some re fr a ct o ry carbonaceous mate ria ls, the residues

f rom such processes comprise pr in c i pa l l y the i ne r t minera l const i tu ents o f the

feed mater i a1. These residues are essenti a1l y dev o id o f s i gn i f i c an t quan t i t i e s

of carbon or char. I n contra st, th e residues from many so-ca lled py ro lys i s

processes are c lassi f ied as chars because they contain a substant ia l f ract ionof the carbon f rom the or ig inal carbonaceous feed mater ia l .

Both gas if ic at io n and pyr ol ys is o f carbonaceous ma ter ia ls produce a mix-

tu re o f gaseous products as a re su l t o f the complex par al le l , compet i t ive, and

sequen tial chemical reac t io ns. The pr efe rr ed rea ct io ns, those th at produce the

most de si ra bl e products, are l i mi te d and slow a t conventional processing tem-

perat ures. However, th e ra te s can be enhanced by i nc lu si on of ce rt ai n cat a-

l ys ts . Fur thermore, a va r ie ty o f cata l ys ts are used in du st r ia l l y t o promote

certain react ions to emphasize format ion of desired products f rom other gasmix tur es. For example, methanation ca ta ly st s promote th e fo rma ti on o f methane

from hydrogen and carbon mor~oxide, s h i f t cata ly s ts promote the format ion o f

hydrogen from water, and carbon monoxide and ga s i f i c a t i o n ca ta ly st s promote th e

breakdown o f carbonaceous matt er t o gases.

The state- o f- t he- art procedure f o r convers ion o f a carbonaceous mate r i a l

to a synthesis gas involves several s teps:

1) ga si f i ca t i on by p a r t i a l combust ion wi th pure oxygen and steam t o form

a mi xt ur e o f gases (CO, Hz, C02, CH4, H2S),

2) s h i f t c on ve rs io n t o y i e l d a 2 : l H2:C0 r a t i o i n t he gas mix tu re,

3) acid gas removal,



4 ) r e f o rm ing of hydrocarbons,

5) mo is tu re removal, and

6 ) f ur th e r gas cleanup t o remove tra ces o f s ul f ur compounds th a t would

otherwise poison catalysts.

P a r t i a1 combustion o f the carbonaceous mat er i a1 wi th pure oxygen t o provid e

heat f o r t he endo thermic gas i f i ca t i on reac t ions i s cu r r en t l y p rac t i ced f o r

gasi f icat ion to produce synthesis gas mixtures. An expensive, energy-consuming

oxygen p lan t i s requ i red fo r t hese gas i f i c a t ion p rocesses t o avo id d i l u t i on o f

the product gases w i th nitroge n. Conversion of wood to hydrocarbon synthes is

gas i s s im pl i f ie d somewhat because of the n eg l i g i b l e s ul fu r content o f wood,

bu t t he conversion s t i l l i nvo lves the steps of gas i f i ca t ion , s h i f t convers ion,

reforming, C02 removal, and mo is tu re removal.

The ove ra l l ob j ec t i ve o f th e PNL s tud ie s i s t o evaluate the technica l and

economic f e a s i b i l i t y o f p roducing spec i f i c gas p roducts v i a t he ca ta ly t i c gas i-

f i ca t i on o f biomass. Spec i f i c products th at are be ing s tud ied inc lud e

(a) methane, (b) synthesis gases for product ion of ammonia, methanol, and

hydrocarbons, ( c ) hydrogen, and (d) carbon monoxide. I n th e stu die s, y i e l d s o f

high-value, gaseous products from a single- reac t ion stage are enhanced by the

proper choice o f reactants, operat ing condi t ions, and ca ta ly st combinations.

Gaseous products, successful ly produced by the var ious steam gasi f icat ion sys-

tems employed t o date , in cl ud e synth esi s gases f o r hydrocarbon, methanol, and

ammonia generation and a methane- rich gas.

Th is repo r t d et a i ls progress o f th e laborato ry s tudies , process develop-

ment u n i t (PDU) s tu di es, and economic anal yses conducted a t PNL from December

1977 t o October 1980. A c t i v i t i e s o f t h e p r o j e c t a re sumnar ized i n F ig ur e 1.

Appendices A through D present the economic analysis of wood-to-methane and

wood-to-methanol plants completed by Davy McKee, Inc., o f Cleveland, Ohio.

These conceptual p la nts use the c at al y t i c g as i f ic at io n processes developed a t

PNL.



CONCLUSIONS

The following conclusions are based on current results from PNL studies

on catalyzed steam gasification of wood:

a Steam gasficiation of wood in the presence of catalysts to produce a

methane-rich gas and a methanol synthesis gas is technically

feasible.

a A nickel-on-silica-alumina catalyst was effective both for generation

of a methane-rich gas and for generation of a methanol synthesis gas

at respective temperatures of 550' and 750'~.

a Catalyst systems need not include an alkali carbonate primary

catalyst.

a Presence of an active nickel catalyst essentially eliminates produc-

tion of condensable liquids during the gasification processes.

a Potential yields of methane and methanol are significantly greater

from the catalytic processes than from conventional processes that

employ oxygen-steam gasification of wood in a fixed bed.

a Cost of methane from wood by catalytic gasification in a 2000-ton

(1800-t) dry wood per day plant is competitive with projected futurecosts of substitute natural gas (SNG).

a Cost of methanol from wood by catalytic gasification in a 2000-ton

(1800-t) dry wood per day plant is competitive with the current price

of methanol.

a Major operating costs are for wood, labor, gasifier catalyst, and

taxes and insurance.

a Investment costs and high labor costs make production of methanol

and SNG in a 200-ton (180-t) dry wood per day plant economically

unattractive.

Further conclusions based on task studies are presented in the discussion.



RECOMMENDATIONS

Results from the studies on catalyzed steam gasification of wood show

considerable promise for commercial production of valuable chemical products.

Real ization of this promise depends on successful development of long-1 ife

catalyst systems and on avail abi 1ity of commerci a1 quantities of feedstock.

We, therefore, recommend continuation of studies on the development of design

parameters for a commercial facility. Such studies may logically include

pilot-plant operations for production of one or more products at a site with a

wood residue source of 10 to 50 wet tons (9 to 45 t) per day.

Near-term studies are required before pilot-plant operations can be defi-

ni tively considered. Laboratory-scale studies on evaluation of special ly pre-

pared catalyst compositions should continue. These studies will evaluate

yields from the new catalysts as well as catalyst lifetime and regenerability.

Thermal requirements for the processes should also be determined. Systems for

generation of an ammonia synthesis gas should be evaluated further.

We recommend PDU studies to evaluate yields at pressures up to 7 atm from

the catalyst systems developed in the laboratory. Information on catalyst

durability in fluid-bed operations should be developed.

Feasibility studies have indicated that processes for generation of SNG

and methanol via catalytic gasification of wood are economic at large scale.

Yields and process inputs have been improved since completion of these studies.

We, therefore, recommend the continued development of the computer code system

to evaluate the effects of these process changes on the overall economics.



LABORATORY STUDIES

The l abo ra to ry stud ies are designed t o develop a fundamental understanding

o f the ca ta ly t i c gas i f i ca t io n o f biomass, t o prov ide a sound bas is f o r design

and ope rat ion o f t he PDU, and t o pr ovi de support se rvi ces f o r PDU ope rat ion .Resul ts of these studies, presented i n the fol low ing sect ions, are intended t o

de f i ne f eas i b l e r eac tan t- catalyst combinat ions and operat ing condi t ions.

EXPERIMENTAL EQUIPMENT

Two types of experimental systems were employed i n th e l ab or at or y phase of

t he p ro jec t : 1 ) a ba tch- feed re ac to r shown i n Fi gu re 2, and 2) a continuous

wood-feed re ac to r, shown i n Fi gu re 3. The batch- feed r ea ct or was used t o com-

p l e te p re l im inary ca ta lys t screening s tudies, t o ob ta in k i ne t i c da ta on thesteam re act io n w it h wood char, and t o determine a l k a l i carbonate recovery

te st s. The system was designed f o r atmospheric pressu re ope rat ion and i s q u i t e

f l ex ib le w i t h respect t o co n t ro l o f exper imenta l va r iab les inc lud ing (1) sample

si ze s up t o 15 g, ( 2 ) steam fl ow rate s, (3 ) temperatures up t o 850°c, (4) aux-

i l i a r y gas addi t i on, and (5) recovery and qu an t i f i ca t i on of gases, o i / t a r s ,

and char produced during gasif icat ion.

The continuous wood-feed re ac to r was employed t o ob ta in da ta regard ing t he

generat ion of speci f ic products f rom wood, t o screen ca ta ly st systems, and t oinv est i gate cata l ys t l i fe t i me and regenerat ion. The reactor sys tem i s con-

stru cted pr im ar i l y of quartz, designed f o r atmospheric pressure operation, and

i s f l e x i b l e w i t h respec t t o con t ro l o f experimental va r iab les inc lud ing :

1) wood feed rates, 2) steam fl ow rate s, 3) temperatures up t o 850°c, and

4) au x i l i a r y gas addi t ion.

FEED MATERIALS

The wood samples used i n the lab or ato ry g as if ic at io n s tud ies i ncl ude sepa-

r a t e wood and bark samples o f western la rc h (tamarack), lodgepole pine, douglas

fir, and cottonwood. The samples were hand pi cke d from a small p r i v a t e ch ip -

pi ng op er at io n i n Usk, Washington. These samples were dr ie d and m i l l e d t o th e

desired p a r t i c l e s ize. Chemical and physica l char acte r iz at io n of these wood



WOOD AND I SS LOCK HOPPERI

GRADUATED

+COLD WATER OUT

LI QU ID CONDENSER

GAS

SAMPLE

BAG

FIGURE 2. Batch-Feed Reactor



T.C . 5 cm ABOVE

CATALYST BED

GRADUATED

5.7cm ID QUARTZ REACTOR

WIRE FINGER STIRRER

CATALYSTBED

WATER PUMP

rt -) WATER OUT

I /C- LIQUID CONDENSOR

n

3- WATER IN VENT

GA S

SAMPLE

COLLECTOR

WATER TEST

LIQUID COLD METER

COLLECTOR

FIGURE 3. Continuous Wood-Feed Reactor



samples are presented i n Table 1. Anal ysi s of a f i e l d cor n sample used as an

a l t e rna t i ve f eedstock i s a l so g iven i n Tab le 1. Lodgepole pine and douglas fir

were the pr imary feedstocks used i n the lab ora tory studies.

The potassium carbonate, sodium carbonate, and borax g a s i f i c a t i o n cat a-

l y s t s used i n the lab or ato ry st udi es were Baker analyzed reagents. Trona was

suppl ied by St au f fe r Chemical Company o f Wyoming, Green Ri ve r, Wyoming.

EFFECTS OF ALKALI CARBONATES

The c a t a l y t i c e f f e c t s o f d i f f e r e n t a l ka l i carbonates on the react ion o f

steam w i t h wood and wood char were determined i n th e lab or at or y. Resul ts o f

these determinat ions are presented i n the fol lo wi ng discussion.

I n v iew of the heterogeneous nature o f the wood-catalyst-steam interac-t i on s, we inv est ig ate d two methods of ca ta ly st mixing wi th wood--dry mixing and

so lu ti on impregnation. Sodium carbonate and lodgepole pi ne were sele cte d as

the reference materi 1s.

The above ca t a l ys t / wood mixtures and wood were gasif ied with steam at

5 5 0 ' ~ and 650'~. Results obtained at 5 5 0 ' ~ wi th a cata lys t concent ra t ion o f

3 x l om 3 g-mole sod iumlg wood are shown i n Fi gu re 4. Also shown f o r comparison

i s th e gas y i e l d f o r s imple pyro ly s i s o f wood under the same condi t ions. These

data show that impregnating the wood part icles with sodium carbonate catalysti s a be t te r contac t ing method than dry mix ing the equiva lent amount o f ca ta ly s t

w i t h the wood. Impregnation produces about 10%more gas than dr y mi xi ng w i t h

a s l i gh t increase i n ra te o f gas production. Both contac t ing methods s i gn i f i -

cant ly increase to ta l gas product ion (75% t o 90% increase) and r a te o f g a s i f i-cat ion. However, us ing a b oi l i n g solu t i on instead o f an ambient so lu t i on o f

cata lys t for impregnat ion had no addi t iona l e f fec t .

The ef fe ct iv ene ss of fou r di f f e re n t cat al yst s (potassium carbonate, sodium

carbonate, tron a and borax) i n concen trat io ns o f 3 x and 3 x g-mole

a l k a l i l g wood a t 550°C, 650°C, and 7 5 0 ~ ~was als o evaluated. These ca ta ly st s

were chosen from t he la rg e number o f po ss ib le cand idates because: 1) exper i-

mental in ve st ig at io ns w it h other mat er ia ls have shown these to be th e most

ac t i ve gas i f i ca t io n ca ta lys t s , and 2) of the large number of possible



gas i f i ca t io n ca ta lys t s , t hese are re l a t i ve l y i nexpensive . Ty p i c a l r e s u l t s

obtained at 6 5 0 ' ~ fo r a cata lys t concent ra t ion o f 3 x g e m o l e a l k a l i l g wood

and steam flow of 1.2 g l m i n are shown i n Fi gu re 5. These data show a di mi ni sh -

in g order o f e f fec t iveness for po ta ss i um carbonate, sodium carbonate, tr on a and

borax. The same r e l a t i v e ord er of effe cti ven ess was observed a t oth er tempera-

ture s and cat al yst concentrat ions. The ef f ec t of temperature wi th and wi thou t

potassium carbonate ca ta ly st on gas y i e l ds i s shown i n Figur e 6; the e f f e ct o f

sodium carbonate concen trat io n i s shown i n Fig ure 7.

The steam ga s i f i ca t i o n rea ct i ons produce a char res idu e and a gas mi xtu re

containing some condensable volati les . The y i e l d s of these produc ts (condens-

able vo l a t i l es n ot i nc luded) are shown i n Tables 2 and 3. These products are

apparent ly formed f rom py rol ys is o f wood t o y i e l d a mixture o f gas, condens-

able l iq u id s , and char. Fur ther ga s i f i ca t i on then takes p lace by in te ra c t io n

o f char and steam as shown i n Equat ion (1 ) and by i n t e r a c t i o n o f carbon monox-

id e and steam i n t he water-gas s h i f t re ac t i on shown i n Equation (2) :

A clo se examinat ion o f the data i n Tables 2 and 3 shows that the increased

y i e l d o f hydrogen and carbon d iox ide t ha t res u l ts f rom the use o f ca ta lys t s

cannot en t i r e l y be at t r ib ut ed t o Equations (1) and (2) . A l though the a lka l i

carbona tes a re e f f ec t i v e ca ta l ys t s f o r t hese reac t ions , t he increased y i e l d o f

carbon d iox ide and hydrogen i s a t t r i bu t ed t o the e f fec t tha t these a l k a l i car-

bonates have on the yi el ds from py ro lys is .

It has been shown th at a l ka l i ne ca tal yst s increase the y i e l d of char from

pyro lys is o f ce l lu los ic mater ia ls a t the expense o f condensable vo la t i les

(Mudge, Sealock and Weber 1979; Sealock, Weber and Mudge 1980; Mudge e t a l .

1980). The increased char y i e l d i l l u s t r a t e s t h e f a c t t h a t t h e c a t a l y s t s c o u l d

inc rea se the amount of hydrogen and carbon dio xi de wi tho ut a pr op or tio na te

reduct i on i n the f i n a l y i e l d o f carbon monoxide and char as expected f rom



CATALYST CO NCE NTRA TI 0 ~ = 3 x 1 0 - ~MOLEALKALI I g WOOD

-

T E M P E R A T U R E = ~ ~ O ~ C-

-

--

-

-

-

-

-

-

NO CATALYST

I

0 200 400 600 800 loo0 1200 1400 1600

TIME, s

FIGURE 5. Re1a t i ve Catalyst Ef fect iveness at 6500C

Equations (1) and (2 ) . The increased gas y i e l d i s a t t r i b u te d t o c racking ofta rs and other l i q u i d organics produced dur in g pyr oly s is.

Gas yi el d s from wood and bark o f f ou r di f f er en t species and from ce ll ul os e

wi t h and wit ho ut impregnated potassium carbonate ca ta ly st are given i n Table 4

f o r steam gas i f i c a t i on a t 650'~. Concentrat ion o f hydrogen i n the product gas



1 CATALYST CONCENTRATI O N = ~ X ~ O - ~MOLE K l g WOOD

TIME, s

FIGURE 6. Effect of Temperature on Gas Production With andWithout Potassium Carbonate Catalyst

was about 60 v o l% for catalyzed and about 15 v o l% fo r uncata lyzed tes ts . A l l

woods gave a sl i g h t l y high er hydrogen con cen trat ion and, i n most cases, a

s l i g h t l y lower carbon monoxide and methane con cen tra t io n i n the produ ct gas

than t h e i r resp ect i ve barks. Di f fere nces wi t h wood va r i et y were considered

i n s i g n i f i c a n t .



-

-- 3 x l o 4 MOLE OF Nalg WOOD

INa2C03)

-

- 3 x MOLE OF Nalg WOOD(Na2C03)

LODGEPOLE PINE

1 1200 400 600 800 lo00 1200 1400 1600 1800 2000 2200 2400

TIME, s

FIGURE 7. Effect of Sodium Carbonate Concentration on Gas Yie ld a t 6500C

Experiments were conducted t o examine the k inet i cs of c ata lyzed biomass

gasif ic ation. In it ia l l y, the ra te correlat ions were expected to be patterned

after those reported by Gardner, Sarnuels and W i lks (1979) and Johnson (1974).

This could not be accomplished, as our data based on gasification products

were not comparable to t he thermogravimetric da ta based on t he t o t a l weight

lo ss used in th es e ref erences. However, the dat a could be used t o generate

ra te co r rela t ionsi n

a manner similar to that reported by Lefrancois, Barc layand Skaperdas (1967) for their bench-scale studies on coal and coal char

gas i f ica t ion .



TABLE 2. Gas Y ie lds a t 6500C

Volume o f Gas, m l l g wood ( bH CO C H C H CH CO T o ta l-2- -2 -2-4 -2-6 4 -

Potassium Carbonate 895 578 3 7 57 145 1685Sodium Carbonate 581 407 4 8 59 116 1175

Trona 424 350 5 10 73 111 973

None 94 94 6 4 47 235 480

(a) Cata lys t concent ra t ion = 3 x 10-3 g - m ol e a l k a l i l g wood.(b) Steam flow = 1.2 glmin, 25 min for react ion wi th 10 g wood.

TABLE 3. Element al Yi el ds a t 6500C

' ( a )Cata l ys t

W t i n Gas Product s( W t i n Wood W t C i n Res idue

( W t i n Wood

Potassium Carbonate 130 188 75 22

Sodium Carbonate 97 138 58

Trona 79 132 56

None 32 69 38 31

(a ) Cata lys t concent ra t ion = 3 x 1 0-3 g -m o le a l k a l i l g wood.

(b) Steam flow - 1.2 glmin, 25 m in fo r react ion w i th 10 g wood.

When wood i s con tacte d wi t h steam a t ele vat ed temperatures, a ser ie s o f

complex phy si cal changes and chemical rea ct io ns are in i t i a t e d . These phenom-

ena inc lude 1 ) r ap id vo la t i 1i z a t i on (py ro ly si s) t o y i e l d gases, condensable

l i qu id s, and a char residue and 2) steam rea ct io n wi th th e py ro ly si s products.

Steam react ion w i th the char res idue occurs a t a s low ra te re la t ive to

pyro l ys i s .

The model used t o analyze la bo ra to ry data assumes that , as discussed

above, th e ove ra l l changes occur i n two consecutive stages: 1 ) p y r o l y s i s o f

the wood and 2) steam ga si f ic at io n o f the char residue. For t h i s analysi s the

reac t io ns i n these stages are assumed t o occur c onsecut ive ly and t o be

independent.



TABLE 4. Gas Yield from Di ff er en t Wood and Bark Ma te ri al s a t 6500C

Mater i a1

Douglas f i r bark

Lodgepole bark

Tamarack bark

Cottonwood bark

Lodgepole wood

Tamarack wood

Douglas f i r wood

Cottonwood wood

Cel lu lo se

Volume ga s , ( a ) L /g biomass

3 X gemole K2g3/g biomass No ca ta ly st

2.2 0.72.1 0.7

2.0 0.6

2.0 1.0

2.0 0.6

2.0 0.5

1.9 0.5

1.7 0.7

1.5 0.6

(a) Steam flow = 1.2 g/min, experiment 25 min in duration.

Kinet ic dat a were obtained a t fi v e concent ratio ns of impregnated po tass i um

carbonate: 0, 4, 8, 12, and 17 wtX . Temperatures studies included 550°c,

650°c, 750°c, and 850 '~ . The total carbon gasified at each sampling was

assumed to be the carbon appearing in the product gas and was determined from

gas volume and gas composition dat a. The percent carbon was determined by di f -

ference and plotted versus time on semi-log paper. Examples are shown in Fig-

ures 8 and 9 f o r catalyzed and uncatalyzed t e s t s , res pec tiv ely . The re su lt s

i n Figures 8 and 9 can be interpreted by the two reaction stages previously

def i ned.

During th e f i r s t 160 s , the l oss es correspond to p yrol ysis of the wood.

The base carbon remaining i n the residue then reacts with the steam in the

gasif ication stage.

I t appears from the r es ul ts in Figure 8 that more base carbon remains in

the residue after the pyrolysis s tage at 6 5 0 ' ~ than after the pyrolysis s tage

a t 550°c, and i t appears from the re su lt s in Figure 9 that more base carbon



40 II I I 1 I

100 200 300 400 500 600

TIME, s

FIGURE 9. Uncatalyzed Test Results on Remaining Carbon Versus Time

remains a f te r the pyro lys is s tage a t 7 5 0 ' ~ than at 650'~. Since condensable

vo la t i le s were no t inc luded i n the de te rmina t ion o f remain ing base carbon, t h i smay or may not be the actual case.

The da ta p l o t t ed i n F igures 8 and 9 dep i c t reasonably l in ea r re la t io nsh ips

fo r the ga si f i c at io n stage. Assuming pseudo - f i rs t- o rde r k i n e t i c s , t h e f o l l o w-

i ng express ion f o r the gas i f i c a t i on s tage reac t i on ra t e cons tan t i s ob ta ined :

where C1 i s the carbon remaining a t the t ime tl and C2 i s the amount o f carbon

remain ing a t the t ime tp .



The activation energy Ea i s o btai ne d f ro m a p l o t o f i n k vs 1000/T (see

Figure 10) as in dic at ed by the Arrhenius equat ion:

Gas prod ucti on rates, r a t e constants, and ac ti va ti on energies were determined

from these studies using the techniques discussed. Carbon-steam reac t i on ra te

cons tan ts fo r the var ious a l ka l i carbonate ca t a l ys t concent ra tions are g iven i n

Table 5. Steam flow rates of 0.4 g/min and 1.2 g/min were employed i n the se

experiments. An increase i n steam f l ow ra t e s i gn i f i ca nt l y increased the value

of the carbon-steam reac t io n ra te constant . Th is i s an ind ica t i on tha t the

FIGURE 10. React ion Constant Var iat ion With 1/T fo r Cata lyzedSteam Gasif ic a tion o f Wood

- 2 5

-3.0

Y -352?-

-4. 0

-45

0 WOOD IMPREGNATED WITH 3 x 1 0 - ~

MOLE Klg WOOD

-

-

-

- I I I

1 1.1 12



TABLE 5. Reaction Rate Constants for Wood Char-SteamReact ion i n t he Presence o f Cata lys t s

ReactionTemperature,

ocSteam Rate,

g/mi n

0.4

0.4

0.4

0.4

0.4

1.2

1.2

1.2

0.4

0.4

0.4

0.4

0.4

1.2

1.2

1.2

0.4

0.4

0.4

0.4

0.4

1.2

1.2

0.4

0.4

0.4

0.4

0.4

Rate Constant,s-1

2.6 x

2.5 x

2.5 x 10'~

2.2 low5

2.6 x

1.4 x 10'~

1.1

1.0

1.3 x 10'~

1.2

8.6 x 10'~

8.0 x 10'~

4.0 x

3.4

9.3

9.9

3.1 x 10'~

3.1 x

3.1 x 10'~

2.5 x 10'~

8.2 x 10'~

9.9

2.0



reaction ra te i s influenced by the mass tra ns fe r or diffus iona l properti es in

the char. The act iva tio n energies calcu lated a t the lower steam r a t e ar e

presented in Table 6. A discussion of the postulated mechanism follows.

Experiments completed early in the program with a continuous wood-feedreactor allowed visual observation of the reaction process. The presence of

th e potassium carbonate cata ly st was observed to induce swelling of the wood

pa rt ic le when exposed to reac tion temperatures. Diffusion is the cont ro ll ing

mechanism for alkali carbonate catalyzed steam gasification of wood char as

indicated by 1) the fact that the reaction kinetics are dependent upon the

steam flow rate and 2 ) the fact that particle swelling is induced by the alkali

carbonate.

I t i s postulated that the swelling of the wood pa rt ic le increases the porediameters thus making the char surface more accessib le. In cont ras tin g a small

pore without the alkali carbonate to a large pore with the alkali carbonate,

the reaction rate in the smaller pore will be more diffusion limited at the

higher temperatures. Fi rs t order reaction ra te equations th at include dif fu-

sional 1imitations conclude that the calculated reaction rate constant, k , i s

proportional to the square root of the tr ue reac tion r a t e constant k t , or1 2kakt . Therefore, the true activation energy,EaTRUEi s equal to two times

the observed or calculated activation energy, E a OB S , or EaTRuE-- 2EaOBS The

effectiveness factors associated with diffusion limitations is represented

graphically in Figure 11 and discussed below.

TABLE 6. Activation Energies for the Wood-Char/Steam Reaction(a1

( a ) Steam flow r at e i s 0.4 g/min.



mL = L \ l k l ~

FIGURE 11. Postu lated E f f ec t iveness Fac tor

Sin ce t h e r e a c t i o n r a te i s d i r e c t l y p r o p o rt i o n a l t o t h e r e a c ta n t concen-

t r a t i o n i n a f i r s t o rde r reac t ion, the average concen t ra t i on near the pa r t i c le

surface, A, compared t o t he bul k stream concentra tion, A, i s a measure o f how

much the rea c t i on ra te i s a f fec ted by d i f fus ion . The r a t i o A /A i s de f ined as

the e f fec t i veness fa c t o r E shown i n F igure 11. .Therefore, th e observed reac-

t i on r a te cons tan t k i s equa l t o the t r ue reac t ion r a t e cons tant , kt, m u l t i -

p l i e d b y t h e e f f e c t i v e n es s f a c to r , E. Figure 11 i l l u s t ra te s how E va r ies

w i t h i n t h e po re w i t h a p l o t o f E vs mL, or E vs L m where:

m = W Dk = observed react ion rate constant

D = d i f f u s i v i t y c o ef f ic i en t

L = l eng th o f po re

Id ea l l y E would equal mL, which equals 1. Th is wou ld i nd ica te tha t the l im i ta -

t i o n b y p o re d i f f u s i o n i s e s s e n t i a l l y absent.



Major conc lus ions obta ined from the k inet ic s tudies are as fo l lows:

a The e f fe c t o f a l k a l i carbonate on the char-steam r e a c t i o n r a te i s

neg l i g ib l e at temperatures be1ow 550'~.

a Char-s team reac t ion ra te cons tan ts inc rease w i th inc reased a lka l icarbonate concentration at temperatures from 6 5 0 ' ~ t o 850'~.

a The c at al ys t appears t o enhance t he char-steam reac t i on ra te by

increa s ing carbon sur face area av ai l ab le f o r react i on.

a The carbon-steam rea ct io n r a t e appears t o become more d i f fu s i on al

l i mi t ed as a l k a l i carbonate concentrat ion i s decreased.

EFFECTS OF SECONDARY CATALYSTS

Data obtained i n s tudies wi th a l k a l i carbonates show th at a t 5 5 0 ' ~ potas-

s ium carbonate had l i t t l e e f f ec t on the carbon-steam react ion but had an ef fect

on the i n i t i a l wood pyro lys is reac t ion and y i e l ds of py ro lys is p roduc ts.

Experiments were designed t h a t considered the f a c t t h a t le ss carbon, a gre ate r

qu an t i ty o f l i qu id s, and a smal ler qua nt i t y of gases are produced dur i ng gas i-

f i ca t i on i n the absence o f the a l k a l i carbona te a t 550'~. Theore t i ca l l y an

increase i n gas y i e l d could be accomplished i f l iqu ids produced could be hydro -

cracked to t he desired gases i n the presence of steam and ca ta ly st s. Data

obta ined from studies completed i n th e batch reacto r demonstrated th a t a t tem-

peratures ranging from 5 5 0 ' ~ t o 7 5 0 ' ~ a 1:l s i l i ca- a lumina (S i-A l ) t o NiO o r N i

ca ta ly st combinat ion s i g ni f i c an t l y increased th e gas volume, energy ef f i c ie nc y,

and carbon conversion to the gaseous phase over the yields with the alkal i car-

bonate. Table 7 presents yields comparing batch experiments with potassium

carbonate and those with the Ni /S i -A1 ca ta ly st mixture. Carbon conversions and

energy e f f i c i en c i es are shown i n Tab le 8.

The Ni and NiO ca ta l ys ts u t i l i ze d i n these s tud ies a re p roven and we l l

known th roughout indus t ry fo r the i r hydrogenat ion capab i l i t ies . The strength

o f t h e s i l i c a- a l um i na c a t a l y s t l i e s i n i t s a c i d i c c ra ck in g f u nc t io n. This

ca t a l ys t inc ludes a h igh concent ra t i on o f Brons ted and Lewis a c id centers .

Lewis a cid s i t e s promote cra cki ng by removal o f hyd rid e ion s from the molecules

o f the reactants , thus g i v i ng a carbonium io n in termediate. Bronsted ac i d



TABLE 7. Y ie l ds fro m Batch Exper iment f Wood-Steam Gasif icat ioni n the Presence of Cata lys ts a ?

Yields, w t%

Exper iments Gases L i qu ids Char Water

- - -( a) A l l t e s t s u t i 1ized a 10-g charge of lodgepole wood

wi th a s t ream rate o f 0 .4 g lm in .

TABLE 8. Carbon Conversions and Energy E ff i ci e n c i e s i n Batch ~ x ~ e r i m e n t sa )

% CarbonExper iments Converted t o a Gas Btu Gas/Btu Wood x 100

At 5 5 0 ' ~

17% K2C03 25 16

N i 1404:Si-A1 (1:l) 48 46

( a) A l l t e s ts u t i l i z e d a 10-g charge of lodgepole wood with a streamr a t e o f 0.4 g lm in .



s i te s y i e ld a pro ton t o the reactant mo lecu les, i e. , o le f ins and aromatics,

r e s u l t in g i n an int em ed i ate carbonium i o n for matio n (Andrew, Ma rt in and Noher

1971). With the lar ge qua nt i t y of l i qu id s produced from the wood durin g

py ro ly si s and the knowledge th a t an aromatic r i n g must f i r s t be hydrogenated

before i t can be cracked, a hydrocracking rea cti on should be id ea l f o r t h i s

process. Combining these two ca ta ly st s i n t o a si ng le system gives both

requ ired c a t a ly t i c functions, hydrogenation and cracking, o r hydrocracki ng.

The presence o f steam gives a dd it io na l reforming of th e organics. The data

and results presented here are reproducible and support the fact that at the

550 '~- to -750 '~ temperature range, a Ni o r Ni0:Si-A1 ca ta ly st system i s a more

ef fect ive ca ta lys t system for gas i f i ca t ion o f wood than 17 w t% impregnated

potassium carbonate.

RECOVERY OF ALKALI CARBONATE CATALYSTS

Recovery and recycling of a1k a l i carbonate i s requ ired f o r those processes

th at use a pr imary ga s i f i c a t io n ca ta lys t . As presented i n la te r sect ions, the

alkal i-carb onat e pri mary c at al ys t i s needed f o r ge neration o f an ammonia syn-

thesis gas and a hydrogen- rich gas. The ob ject ive o f the a lka l i carbonate

recovery tests was to determine primary catalyst loss per unit of biomass pro-

cessed. Residues from g as if ic at io n experiments were used t o conduct th e te st s.

The residues were selected from batch experiments conducted with a potassiumcarbonate cata lyst at 7 5 0 ' ~ with a steam rate of 0.4 g lm in which gave an aver-

age o f 92% carbon conversion t o gas.

The recovery experiment consisted of washing the residue with water or a

water-solvent mixture. The wash was f i l t e r e d and the f i l t r a t e and residue

were retained for potassium analysis. The potassium i n t he r esid ue and f i l -

t r a t e was analyzed wi t h an in du ct io n couple plasma ICP spectrometer ins trume nt.

The parameters of the studies included:

Leaching Solutions:

(a) water,

(b) water lmethanol ( 1 I ) ,

(c) acid ic (pH 3),

(d) water /ace tone (1 : l )



a So l u t i on /Res i due Weight Ra ti o: 10 and 100

Leaching Solution Temperature: 9 5 ' ~ (water only) and 2 5 ' ~

a Leaching Time: 2 h (water only) and 1 h.

Results of these tests can be summarized as follows:

a 98% potassium recovery i s achieved wi th a s o l u t i o n / r e s i d u e weight

r a t i o of 100 fo r 1) a water wash, 2) an ac id wash, or 3 ) a water/

methanol wash at 25'~.

a 95% potassium recovery i s obtained wi th a 10 /1 water wash at 25 '~ .

a A leach ing t ime of 1 h y ields the same resul ts as a 2-h leaching

time.

a 2 5 ' ~ i s a bet ter leach ing temperature than 95'~.

X-ray d i f f r ac t i on s t ud i es were performed t o i d en t i f y t he t y pe o f a l k a l i

compound remaining i n th e char re sidue . Resu l t s ind ica te t ha t K2C03*H20 and

K2C03 are the onl y a l k a l i compounds i n the residue.

With 95% recovery, th e co st o f potassium carbonate and sodium carbonate

catalysts at September 1980 prices would be $2.60/dry to n o f wood and $0.40/dry

ton o f wood, resp ect i vel y. No los s i n ef fect i veness o f the recovered ca ta ly st

was observed.

SYSTEMS FOR METHANE GENERATION

The obj ec t i ve of these studie s was t o f i n d ca ta ly st combinat ions and oper-

a t ing cond i t i ons f o r genera tion o f a methane- rich gas. Favored systems devel-

oped i n the labora tory d id not inc lude the use of a pr imary a l k a l i carbonate

c a t a l y s t .

Numerous secondary ca t al ys t systems f o r g enera tion o f a methane- rich gas

a t 5 5 0 ' ~ were tes ted w it h the equipment shown i n Fig ure 3. F av o r ab le ac t i v i t y

f o r methane prod uct i on was observed wi th some cat al ys ts tes ted i n these stu di es

(see Table 9). Ca ta lys t s t ha t a re e f f ec t i v e f o r genera tion o f a methane- rich

gas were determined t o have the fol lo wi ng proper t ies: a ni ck el con centrat ion

of 25 w t% or gr eater, an aci di c support, a BET surface area o f grea ter than2 2

100 m /g, and a n i c ke l surfa ce area of 30 m /g or greater .



TABLE 9. Catalysts Tested for Methane Production

Harshaw Ni-0101

Harshaw ~ i - 1 4 0 4 ( ~ )

Harshaw ~ i - 3 2 6 6 ( ~ )

Harshaw N i 3 2 1 0 ( ~ )Harshaw ~i5 1 4 2 ( ~ )Harshaw Ni-Cu-Mo

Harshaw N i Mo

Harshaw Ni-W

Hars haw Co-Mo

Grace Si-A1

Union Carbide Xeolte

G i r d l e r ~ i - ~ - 5 6 ( ~ )G i r d l e r ~ i - ~ - 8 7 ( ~ )

N i Ni/A1203 catalyst prepared at Brigham Young Un iv er si ty a )

Strem Pd

Strem Ru

Strem R ~ I ( ~ )

(a) Active for methane generation.

Af te r sel ect i on o f sui tab le ca ta ly st systems, at te nt io n was focused on

catalyst l i fet ime and regenerat ion. These studies coincide with the develop-

ment o f a s u i t a b l e c a t a l y s t f o r o p e ra ti on i n t h e PDU.

Lifetime of a Harshaw Ni-3266 catalyst was determined for generat ion of a

methane-r ich gas. Table 10 presents te s t re su l t s and condi t ions p r i o r t o cata-

l y s t deact ivat ion. A wood-feed t o ca tal ys t weight r a t i o of 10 was obtained

before no t iceab le loss o f ca ta lys t ac t i v i t y . Loss o f ac t i v i t y appears t o be

caused by carbon de po si t i on on th e ca ta ly st . The dea cti vat ed ca ta ly st was

regenerated with steam at 6 5 0 ' ~ t o remove carbon from the c at al ys t surface;



TABLE 10. Result s of a Cata lys t Lifetime Test f or Methane Production

Catalyst

Reactor Temperature

Catalyst Bed Temperature

Wood Feed Rate

Steam Rate

Carbon Conversion to a Gas

Wood-Feed/Catalyst Weight Ratio

(BTU Gas/BTU Wood) x 1003 a )scf CH4/ton Dry Wood ( nm / t j

Gas Composition

(a) Includes methanation of H z and CO.

29.5 vol%

34.3 ~ 0 1 %

Trace vol%

25.4 vol%

10.8 vol%

this was followed by reduction with hydrogen at a space velocity of 95h - I

and a temperature of 450'~. The regenerated catalyst was then subjected to

th e t e s t condit ions employed before deac tiv ati on. The regenerated ca ta ly st

showed more renewed a c t i vi t y and longer l i f e than the fr es h ca ta ly st . A wood-

feed t o cat al ys t r a t i o of 52 was obtained before a second deacti vat ion

occurred. Result s following regene ration are presented in Table 11. The major

difference between gasification results shown in Tables 10 and 11 i s a decrease

in carbon conversion with the regenerated catalyst, a possible result of some

loss in cata lyst act ivit y. No furt her regeneration of the cata lys t was

attempted.

Catalyst characterization studies were completed on catalysts exposed to

gasification conditions for extended periods and on regenerated catalysts.

Ca ta lyst ch ar ac te ri za ti on equipment items included a hydrogen chemi sorp tion



TABLE 11. Catalyst Performance Fol lowing Catalyst Regeneration

Catalyst

Reactor Temperature


Wood Feed Rate

Steam Rate

Carbon Conversion to Gas

Wood-FeedICatalyst Weight Ratio

(BTU GasIBTU Wood) x 100

3 (a)scf CH4/ton Dry Wood (nm It) ,

Gas Composition

33.6 vol%33.6 vol%

22.5 volX

10.31 vol%

Trace

(a) Includes methanation of Hz and CO.

apparatus, a BET surface area analyzer, a porosimeter, and a scanning electron

microscope (SEM) working in conjunction with an electron microprobe (EM). A

schematic of the hydrogen chemisorption apparatus is shown in Figure 12. Rela-

tive activities of the nickel catalysts are determined by measuring chemisorp-

tion of hydrogen on nickel which occurs as follows: Hz + 2Ni(s) + 2NiH.

Table 12 presents results of surface area studies on a Harshaw Ni-3266 catalyst

employed for methane production. A significant decrease in the BET and nickel

surface areas occurs upon deactivation (coking).

The pore distribution in a fresh and coked (deactivated) Harshaw Ni-3266

.

catalyst is shown in Figure 13. The catalyst was coked in operations to gener-.

ate a methane-rich gas. Coking of the catalyst reduces the concentration of

smaller diameter micropores, which explains the drastic loss of BET and nickel- surface areas as noted in Table 12.



PURGE

BUBE""

ROTAMETER

4THREE WAY I

VALVE 1'1

SAMPLE CEL

.GAS STORAGE

BULB

MANOMETER

ru:RESERVOIR

FIGURE 12. Hydrogen Chemisorpt ion Apparatus

TABLE 12. Ca ta ly st Ch ar ac te ri za ti on Tes ts Performed on a Ni-3266 CatalystUsed For Methane Production

Unused Ni-3266 Coked Ni-32662

BET Surface Area, m /g 145 47-452

Nickel Surface Area, m /g 56 7

The SEM was used i n con ju nc ti on w i t h t he EM t o study coke deposit i on on a

NI-3266 ca ta ly st . Fig ure 14 presen ts a micrograph o f a coked Ni-3266 cata lyst

employed for generation of a methane-rich gas.



P= ABSOLUTE PRESSURE, PSI

D = PORE DIAMETER, MICR ONS

100 10 1 o 0.10 0.010

FIGURE 13. Pore Distribution in Fresh and Coked Catalyst

0.201

0.18-

0.16

0.14-

g 0.12-

0.10-

K

5 0.08-2

0.06-

0.04

0.02

0.00

The micrograph shows that the majority of the catalyst surface is coveredwith carbon. This can be contrasted to a micrograph of a fresh Ni-3266 cata-

lyst shown in Figure 15. The contrast in surface morphology explains why the

coked catalyst was inactive. The EM supplies information regarding the rela-

tive concentration of surface nickel sites available for reaction. Figure 16

illustrates a comparison of surface nickel concentrations of coked and fresh

catalyst. This figure indicates that the catalyst surface is being covered

with carbon prohibiting contact of reactants with the active sites on the cata-

lyst. A carbon-hydrogen-nitrogen (CHN) analysis indicated that the carbon con-

tent of the coked catalyst was 15 wt%.

1 1 1 1 1 1 1 I I 1 1 1 1 1 I I ' 1'""' ' ' I"'

/.-- /- COKED CATALYST /

FRESH CATALYST /

/

I

-

-

I 1 I l l l l l l I 1

Studies on catalyst lifetime and regeneration are in progress to determine

catalyst requirements. Catalyst systems that show favorable activity for gen-

eration of a methane-rich gas are also being studied. The weight'ratio of wood

1 10 100 1,000 10,000



FIGURE 14. Coked Ni-3266 Catalyst, 3340X



FIGURE 15 . Unused N i -3266 Ca t a l y s t s , 10,021X



FIGURE 16. Relative Concentrations of Elements on Fresh and Exposed CatalystSurfaces Employed f o r Methanantion

-

-UNUSED Ni - 3266 Ni (Ka)

- -COKED N i

- 3266

I

Ni (KP)

Fe (Ka)



processed t o ca ta l ys t charge before deact iv at ion i s g iven i n Table 13 f o r the

f res h and regenerated cata lys t s tes ted t o date. A ca ta ly s t w i t h an i n h i b i to r

for coke deposi t ion should give extended l i fe. Such catalyst systems are being

developed.

SYSTEMS FOR GENERATION OF METHANOL SYNTHESIS GAS

The obj ec t i ve o f these s tud ies was t o develop ca ta l ys t systems and oper-

at ing condi t ions for the direct generat ion of a hydrogen and carbon monoxide

gas mixt ur e th at could be used f o r methanol or hydrocarbon synthesis. Treat-

ment o f th e gas mixtu re f o r carbon dio xid e removal was intended t o be the o nl y

necessary opera t ion t o make th e mixtu re s ui ta bl e f o r methanol synthesis.

A1i s t o f ca ta l ys ts tes ted f o r product ion o f a methanol , or hydrocarbon,

synthesis gas i s presented i n Table 14. Favorab le a c t i v i t y fo r synthesis gas

gener atio n was observed i n some ca ta ly st s as noted. Tests o f combinations o f

these cata lys ts wi th a pr imary (a lka l i carbonate) gas i f ica t ion cata lys t showed

th at gas y i e l ds were s imi la r w i t h or w i thout the pr imary cata ly s t . Therefore,

the pr imary gas i f ica t ion cata lys t was e l im inated f rom fur ther s tudy.

Pre l im inary tes ts per formed wi th the Ni-3266 f o r synthesis gas product i on

looked promis ing. Typica l re su l t s are presented i n Table 15. A t 7 5 0 ' ~ a wood-

f eed t o c a t a l y s t charge r a t i o o f74

was obta ined pr i o r t o cata lys t deact iva-

t i o n . Po ten t i a l methanol y i e l d i s shown as a weight f ra ct io n o f d ry wood.

TABLE 13. Li fe t i me and Reg ene rab il i ty o f Catal ysts Tested f o rMethane Generation

Wood-Feed t o Wood-Feed ToFresh Cat aly st Cat aly st Charge Ra ti o

Cat aly st Charge Ra ti o Fol low ing Regenerat ion

Harshaw Ni-3266 10 52

Laboratory-Prepared Ni /S i -A1 13 6

Laboratory-Prepared Ni A1 203 5 Not Determined

Harshaw Ni-3266 28 Not Determined

Ni-3266:Si-A1 ( 3 : l ) 15 30



TABLE 14. Ca ta ly st s Screened f o r Methanol Synthesi s Gas Pr od uc ti on

Harshaw ~ i - 1 4 0 4 ' ~ )

Harshaw Ni-1404:Grace Si-A1 (Var ious ra t ios) ( a )

Harshaw Ni-Cu-Mo ( a )

Harshaw ~i3 ~ 6 6 ( ~ )

Harshaw N i -3266:Grace Si-A1 (Var ious ra t ios) ( a )

G i r d l e r Ni-G-56:Grace S i A1 (Various r a t i o s ) ( a )

G i r d l e r C u-Zn

Gi rdl e r Fe203

G i r d l e r N i ~ - 1 3 - 3 ( ~ )

Grace Si-A1

Strem Mordenite Molecular SieveStrem Type Y Molecular Sieve

Harshaw Co-Mo: M ole cul ar Siev e

10% N i impregnated on Si-A1

( a )30% N i impregnated on high surface area A1203

15% N i impregnated on low surface area Si-A1

Ni-Ni A1203 ca ta ly st prepared a t BYU

Prepared Ni-Cu-Mo ( a )

Grace Ni

(a) Favorab le ac t iv i t y shown

Attempts t o regenerate the cat a l ys t w i t h steam gas i f ic a t io n o f the deposi ted

carbon and subsequent hydrogen re du ct io n were unsuc cessf ul. Ca ta ly st charac-

t e r i z a t i o n stu di es were subsequently performed. Table 16 pre sen ts BET and

hydrogen chemisorpt ion re su lt s for the Ni-3266 catalyst before and af ter

tes t ing for synthes is gas product ion.

The pore dist r ibut ion determined for f resh and regenerated Ni-3266 cata-

l y s t used f o r synthesis gas generat ion i s shown i n F igure 17. The concentra-

t i on o f sma ll d iamete r m ic ropores i s obv ious ly d r as t i ca l l y decreased i n t he

regenerated ca ta lys t . Th is decrease re su l t s i n loss o f over a l l and ac t i ve su r-

fa ce area as shown i n Table 16.



TABLE 15. Resul ts o f a Cata lys t L i fe t ime Test for HydrocarbonSynthesis Gas Production

Cata lys t

Reactor Temperature


Wood Feed Rate

Steam Feed Rate

Carbon Conversion t o a gas

Wood-FeedICatalyst Weight Ratio

H2/C0 Mole Ratio

(BTU Gas/BTU Wood) x 100

Potential Weight Fraction Wood as Methanol

Gas Composition

53.0 v o l%

16.9 v o l%

Trace v o l%

Trace v o l%

3.1 v o l%

27.0 v o l%

TABLE 16. Charac te r i s t i cs o f a N i-3266 Cat al yst Employed f o rSynthesis Gas Production

Unused Regenerated2

BET Surface Area, m /g 145 1.42

Nickel Surface Area, m / g 56 0.4

The i na b i l i t y t o r es t o r e c a t a l y s t a c t i v i t y was undoub tedl y t he r es u l t o f

this permanent loss of ac t i ve surface s i t es. SEM studies show simi lar resul ts.

Figure 18 shows a micrograph of a regenerated Ni-3266 cata l ys t . Contrast ing

t h i s w i th the micrograph o f an unused cata ly s t presented ea r l ie r i n F igure 15

reveals a decrease i n sur face concent ra t ion o f ac t ive s i te s on regenerated



D = PORE DIAMETER, MICRONS

0.161---- COKED Ni-3266 CATALYST

I

P=ABSOLUTE PRESSURE, PSI

0.14-

.

0.12-

0.10-

PNak 0.08-2

0.06

0.04

0.02

FIGURE 17. Pore Distribution in Catalyst Used for Synthesis Gas Generation

FRESH Ni -326 6 CATALYST

-

0

/0

- H~--e*)- -------

e4

& 9 ~ ) ~O . O O ~ :

.

I " ' I I I I I 1 1 1 1 I I I I r 1 1 1 1 1 I r I I I 1 1 1 1 1

1 10 100 1,000 10,000



FIGURE 18. Regenerated N i -3266 Catalyst Employed For HydrocarbonSynthesis Gas Production, 10,021X

43



ca t a l ys t . Th i s i nd i ca t e s t ha t the r egene ra t ed ca t a l y s t has been s in t e r e d and

s u pp o rt s t h e d a t a p r es e n te d i n T ab l e 16 and t h e p o r e d i s t r i b u t i o n

determ inat ions .

A SEM was used i n con ju nc t i on w i t h an EM t o examine the su rfa ce changes ofa Ni-3266 ca t a l y s t w i t h exposur e t o cond i t i on s f o r syn t hes i s gas gene r a ti on .

Ca ta ly st samples were removed f rom the rea cto r a t v ar i ous exposure t imes i n a

t e s t f o r g e ne r at i on o f s y nt h es i s g as. The micrographs presented i n Figure s 19

through 24 show the gradua l bu i ldu p o f char ( carbon) on the n i ck e l ca ta l ys t

w i th inc reased exposure t ime. The po r e spaces a re f i l l e d i n i t i a l l y , see F i g-

ure 21, wh ich decreases the sur face area ava i la b l e f o r reac t i on . Carbon con-

t i n u es t o d ep o si t on t h e c a t a l y s t u n t i l lay ers coat th e sur face , see F ig ure 22.

F igur es 23 and 24 show th a t the sur face i s comple te ly coated wi t h la yer s o f

carbon a f te r exposure t imes o f 88 and 117 h, re sp ec t i ve ly .

FIGURE 19. SEM Mi cr og ra ph o f Fr es h Ni-3266, 3960X



FIGURE 20. SEM Micro graph o f Ni-3266 After 28.6 h of Operation, 13,200X

FIGURE 21. SEM Microg raph of Ni-3266 Af te r 61 h of Operat ion, 13,200X

45



FIGURE 22. SEM Mic rograph o f Carbon Lay ers on Ni-3266A f t er 72 h o f Operation, 132X

FIGURE 23. SEM Mi cr og ra ph o f Coked Ni-3266 Af ter88 h o f Operat ion, 132X



FIGURE 24. SEM Mi cr og ra ph o f Coked Ni-3266 A ft er 117 h of Operat ion, 132X

Elec t ron m ic roprobe scans o f the exposed ca ta lys t sur faces ar e shown i n

F i g u r e 25. The r e l a t i ve concen t r a t i ons o f n i cke l , s i l i co n and aluminum

decrease wi th ca ta ly s t exposure t ime. The gradua l decrease i n the n i ck e l peak,w i t h ex po su re t i me , f u r t h e r i n d i c a t e s t h a t t h e c a t a l y s t s u rf a ce i s b e in g c oa te d

wi th carbon.

As a r e s u l t o f t h e i n a b i l i t y t o r e ge ne ra te t h e Ni-32 66 c a t a l y s t , a t t e n t i o n

was sh i f t e d t o eva lua t i on of a Ni-Cu-Mo a l l o y ca ta l ys t on an alumina suppor t

f o r sy n th e s is gas g en e ra t io n . P r e l i m i n a r y s tu d i e s w i t h t h i s a l l o y c a t a l y s t f o r

syn t hes i s gas gene r a t i on has i n d i c a t ed ex ce l l en t r es i s t a nce t o car bon depos i-

t i on . Work on methanat ion re ac t i on s wi t h a Ni-Cu a l l o y c a t a l y s t by A r ak i and

Ponec (1976) demonst rated th a t th e presence o f copper d i l u t e s th e number o f

n i c k e l s i t e s , t h u s i n h i b i t i n g c ar bo n f o rm a ti o n . T he r ef or e, an a l l o y c a t a l y s t

f o r syn thes is gas produc t i on may have th e l i f e t i m e needed f o r economic

f e a s i b i l i t y .



Experiments to evaluate a Harshaw Ni-Cu-Mo a l lo y ca t a ly s t on a lumina

showed a s ign i f i c an t i ncrease i n ca ta l ys t l i f e t i me . Data i n Table 17 presen t

average res u l t s ob tained w i t h the a l lo y ca t a ly s t . Ear ly i n the exper iment,

ca t a l y s t ac t i v i t y decreased s l i g h t l y . No fu r the r change i n ac t i v i t y occu rredfor the next 175 h. A l though s t i l l ac t ive , the ca t a ly s t was then subjected t o

regenerat io n wi th steam a t 8 0 0 ' ~ f o l l owed by hydrogen r educ t i o n a t 450'~.

C a ta l ys t a c t i v i t y was r es tor e d t o t h a t o f t h e o r i g i n a l ca ta l ys t by t h i s pr oce-

dure. The experiment was cont inu ed f o r an add i t iona l 210 h w i th on ly the

i n i t i a l s l i g h t d e ac t i va t io n . R ege ne ra ti on agai n r e s tor e d t h e s l i g h t a c t i v i t y

l o s t after which the experiment was continued another 24 h before the catalyst

was removed f o r ch ara ct er i za ti on studi es.

TABLE 17. Use o f a Ni-Cu-Mo/A1203 C a ta l ys t f o rSynthesis Gas Generation

Reactor Temperature

Catalyst Temperature

Wood Feed Rate

Steam Rate

Wood Feed ICa t a l y s t Weight Ratio 188 .O

Carbon Conversion t o a Gas 85.0%(BTU Gas/BTU Wood) x 100 94.6

H2/C0 Rat io

Gas Composition

L i f e t i me and reg ene rab i l i t y of ca t a ly s ts t ested t o da te a re p resented i n

Table 18. Conclusions der ived from the cata lyst systems studied for hydrocar-

bon synthesis gas production are given below:



Properties of the most effective catalyst tested for the production

of a methane-rich gas include a nickel content of 30 wt% or greater,2

an acidic support, a BET surface area of 100 m /g or greater, and

2a nickel surface area of 30 m /g or greater.

The nickel catalyst lifetime employed for methane production is

short. However, the catalyst is regenerable with steam and subse-

quent hydrogen reduction. Further tests need to be completed to

determine the nickel catalyst activity after repeated regenerations.

A Ni-Cu-Mo alloy catalyst looks promising for hydrocarbon synthesis

gas production, exhibiting good catalyst lifetime as well as

regenerability.

TABLE 18. Lifetime and Regenerability of Catalysts Tested forHydrocarbon Synthesis Gas Production at 7500C

Wood-Feed Wood-Feed to Catalystto Catalyst Charge Ratio

Catalyst Charge Ratio Fol lowing Regeneration

Harshaw Ni-Cu-Mo 188 (a)

Harshaw Ni-3266 74 4.4

Grace Si-A1 loo(b1

Harshaw Ni-1404

Girdler C-13-3

5 2 5

17 Not Determined

Prepared Ni -Cu-Mo 30 1st:23 2nd:10

15% Ni On Low Surface

Area Support 17 18

30% Ni On High Surface

Area Support 50 1st:16

Grace Ni Spheres 70 Not Determined

(a)' No deactivation - was subjected to regeneration scheme twice.(b) No deactivation at 850%.



SYSTEMS FOR GENERATION OF AMMONIA SYNTHESIS GAS

The objective of these studies was to generate a gas mixture with the

proper hydrogen-to-nitrogen ratio for ammonia synthesis. The gas mixture is

produced by reacting air, steam and wood in the presence of catalysts. A

summary of catalyst systems tested is given in Table 19.

TABLE 19. Catalyst Systems Tested for AmmoniaSynthesis Gas Production

Girdler 6-3 Fe-Cr: Strem Si-A1 (3:l)

Girdler 6-93 Co-Mo

Girdler 6-3 Fe-CrGirdler G-9 Cu-Mn

Harshaw HT-100 Ni-Mo

Girdler 6-101 V205: Strem Si-A1 (3:l)

Girdler 6-66 A Cu-Zn

Harshaw 0301 Fe203

Effective catalyst systems for ammonia synthesis gas production were

found to be combinations of primary and secondary catalysts. The primary cata-

lyst was potassium carbonte (17 wt%) impregnated in the wood, and the secondary

catalyst was a Girdler 6-93 cobalt molybdate: si 1ica-alumina (Si-A1 ) system at

a weight ratio of 3:l respectively, or a Girdler 6-3 chromium promoted iron

oxide and silica-alumina at a weight ratio of 3:l. Typical results using the

Girdler 6-3 catalyst with the silica-alumina are given in Table 20.

Various operating parameters, temperature, steam space velocity, and

wood-feed rate, were studied for production of ammonia synthesis gas. Results

obtained using the 6-93 Co-Mo CO shift catalyst are presented in Table 21.

Examination of the CO concentrations presented in Table 21 shows that

increasing the steam rate decreases the CO concentration via the water gas

shift reaction (CO + H20 + C02 + Hz) and that a gas consisting almost entirely

of Hz, N2 and C02 is produced. Carbon monoxide yields as low as 2% were



TABLE 20. Experimental Results of Lodgepole Wood Impregnated with 17%K2CO3 in the Presence of Air, Steam, Girdler 6-3 andSi-A1 Catalysts

Secondary CatalystReactor Temperature

Catalytic Bed Temperature

Air Rate

Wood Feed Rate

Carbon Conversion to a Gas

H2/N2 Ratio

Cold Gas Efficiency

Gas Composition

obtained by increasing steam rates. The increased steam rate magnifies the

water gas shift reaction thus producing more Hz and C02, and less CO. Typical

results obtained by increasing steam rates while at the same time maintaining

the H2:N2 ratio at 3:l are included in Table 21.

Major conclusions of the completed ammonia synthesis gas studies are sum-

marized below:

a Girdler 6-3 chromium-promoted iron oxide or a Girdler 6-93 cobalt-

molybdate CO shift catalyst produced an optimum H2:N2 ratio of 3:l.

Of the catalysts tested, these catalysts produced the highest overall

conversions (65%) to the gaseous phase while maintaining the desired

product yields.



TABLE 21. Exper imen tal Resu lt s o f Lodgepole Wood Impregnated w i t h17% K2CO3 i n the Presence o f A i r , Steam, and CommercialG i r d l e r 6-93 CO S h i f t C a t a l y s t

Steam Rate

0.64 g/min 1.0 g/min 1.2 g/min 1.6 g/min

Run number

Secondary Catalyst

Reactor Temperature, OC

Catalyst Bed Temperature,

Oc

A i r Rate L/min

Wood Feed Rate, g/min

H2/N2 Rat ioCold Gas Efficiency, %

Gas Composition, v o l %

2

N2co

A maximum steam r a t e o f 4 t imes the wood feed rate produced CO y i e l d s

as low as 2% by volume fo r bot h ca t a l y s t systems desc rib ed above.

a An a i r i n l e t r at e o f approximately 40% by weight o f the wood feed

r a t e produced the desire d H2:N2 r a t i o f o r bo th ca ta l ys t systems.

Steam g a s i f i c a t i o n r a t e s a t 5 5 0 ' ~ are too low t o maintain a desired

H2:N2 r a t i o o f 3 : l at a s ign i f icant gas product ion ra te . Constant

wood feed ra tes a re c r i t i c a l i n ma in ta in ing a 3: l H2:N2 product

r a t i o .

An incre ase i n temperature enhances the o ve ra ll conve rsion t o gaseous

products.



SYSTEMS FOR GENERATION OF HYDROGEN

The objective of these studies was to select catalysts and operating con-

ditions for the direct production of hydrogen. A total of 22 experimental runs

were completed. The catalysts tested are presented in Table 22. Various oper-

ating parameters and catalyst systems were investigated. Twelve specific cata-

lyst systems were evaluated, and three temperature ranges (550°c, 650°c, and

750'~) were studied.

Two catalyst systems were found to be effective for hydrogen production:

1) Girdler Fe-Cr CO shift catalyst and a Harshaw Ni-1404 catalyst in a weight

ratio of 3:1, respectively, and 2) Girdler 6-93 Co-Mo CO shift catalyst and

Ni-1404 with the same 3:l ratio. In all cases the wood feed was impregnated

with 17 wt% potassium carbonate. Table 23 presents typical results employing

preferred catalysts at each temperature range. Data presented in Table 23 show

that 64 vol% hydrogen is obtained with a carbon conversion to gas of 78%.

TABLE 22. Catalysts Tested for Hydrogen Production Studies

Girdler 6-3 Fe-Cr

Gi rdl er 6-93 Co-Mo

Girdler 6-64 Fe203

Girdler 6-64 Fe203:Girdler 6-72 D ZnO (1:l)

Strem Si-A1

Girdler 6- 3 Fe-Cr:Grace Si-A1 (1:1), (2:1), (3:l)

Gi rdl er 6-66 Cu-Zn

Girdler 6-3 Fe-Cr:Harshaw Ni-1404 (3:l)

Girdler 6-93 Co-Mo:Grace Si-A1 (2:l)

Girdler 6-3 Fe-Cr:Grace Si-A1 :Harshaw Nil404 (4:2:1)

Girdler 6-93 Co-Mo:Harshaw Ni-1404 (1:l)



TABLE 23. Typ ica l Re su lt s f o r Lodgepole Wood Impregnated w i t h17 w t % K2CO3 i n th e Presence o f Steam and Cat aly sts

Ca ta l ys t

Steam Rate, g/min

Wood Feed Rate, g/min

Cold Gas Eff iciency, %

Carbon Conversion

t o Gases, %

Gas Composition, v o l%

Conclusions from experimental results for the hydrogen production case

are:

a 7 5 0 ' ~ i s the pref er re d operat ing temperature of the thr ee tempera-

t u res i nves t iga ted.

Hydrogen production i s poor at 550'~.

a Approximately 80% o f th e carbon i n t he wood was converted t o gas i n

t e s t s a t 750'~.

a The two cata lyst systems previously descr ibed are su i tab le for hydro -

gen production.

a High steam-to-wood feed rates (order of 4 : l ) a re requ i red t o ob ta in

the des i red CO s h i f t reac t io n needed f o r hydrogen production .



SYSTEMS FOR GENERATION OF CARBON MONOXIDE

The objective of the CO studies is to optimize the selective production

of CO by employing steam, oxygen, and C02 as reactants to gasify wood in the

presence of catalysts. Maximum results for CO production obtained up to this

point are presented in Table 24. Results to date appear unfavorable. There-

fore, the direct production of CO from biomass is not recommended.

EVALUATION OF ALTERNATIVE BIOMASS MATERIALS

Laboratory studies to investigate the production of hydrocarbon synthesis

gas (syngas) from alternative biomass materials were initiated this year. A

corn feedstock was examined to determine the applicability of this type of

feedstock. A proximate and ultimate analysis of the corn is presented in

Table 1. Hydrocarbon synthesis gas studies at 750'~ indicate that product

TABLE 24. Experimental Results for CO production from Lodgepole Wood

Primary Catalyst

Secondary Catalyst

17% K2C03

Grace Si:Al

Reactor Temperature 850'~

Catalyst Bed Temperature 725'~ to 770'~

Wood Feed Rate 0.6 g/min

Steam Rate 0.1 g/min

O2 Rate 0.08 g/min

C02 Rate 1.5 g/min

Cold Gas Efficiency 127%

Carbon Conversion to Gas 93%

Gas Composition

2 15.7 vol%

C0243.8 vol%

'zH4 0.6 vol%

'zH6 0.1 vol%

CH43.1 vol%

CO 36.7 vol%



yie lds and conversion ef f ic iencies for corn and wood are simi lar . A comparison

o f steam ga s i f i ca t i on o f corn and wood i s dep ic ted i n Tab le 25. A c a t a l y s t

system cons is t ing o f Ni -1404 lS i -A1 a t a 1:l weight r a t i o was found t o be su i t -

ab le f o r steam gas i f i c a t i on o f corn. A NiOISi -A1 ca ta lys t system or a1:3 we ight r a t i o o f the Ni-1404/Si-A1 ca ta ly st system proved to be unsat i sfac -

t o r y fo r co rn feeds tock . These catalyst systems had worked equal ly wel l with

wood as the feedstock. The po ten t ia l methano l y ie ld ind ica tes tha t about 5 ga l

o f methanol can be der i ved from one bushel o f corn. Fur the r st udie s on a l t e r -

native feedstocks are planned.

TABLE 25. Steam Ga si fi ca ti on o f Biomass f o r the Pro duc tio n

o f Hydrocarbon Sy nthes is Gas

Feed Corn Lodgepole Wood

Ca ta ly st System Ni-1404:Si-A1 N i-1404: S i- A1(1: 1) (1:3)

Reactor and Catalyst BedTemperature, OC 750

Feed Rate, g lm i n 0.4 0.4

Steam Rate, g/min 0.2 0.2

Carbon Convers ion t o a Gas, % 0.83 0.82

Cold Gas Eff iciency, % 94 98

Poten t ia l Weigh t Fract ionYi el d o f Methanol 0.63

Gas Composition, v o l%

"2 55.7 57.8

CO 30.3 27.7

co2 12.7 12.7

CH4 1.4 1.8



FUTURE LABORATORY STUDIES

Catalyst development, lifetime, regenerability, and characterization

studies will continue for both methane and hydrocarbon synthesis gas produc-

tion. Nickel alloys on various substrates will be obtained in the near future

from The Davison Chemical Division of W.R. Grace & Co. These will be tested

for both methane and hydrocarbon synthesis gas production. As a result of

these tests, catalysts may be prepared in the laboratory for testing.

Investigations to produce an amonia synthesis gas will resume this year.

Operating conditions at 750'~ wi 1 1 be optimized. Catalyst 1ifetime and regen-

erability will be studied for the favored catalyst system.

Alternative feedstocks, such as corn, alfalfa, and wheat straw, will beinvestigated regarding the feasi bi 1ity of producing specific gases.

Thermal studies to determine the heat requirements (or release) in biomass

(wood, 1ignin, cellulose) pyrolysis were initiated in FY-1980. These studies

will continue and will include alkali carbonate catalyzed biomass. Methods

that will be used to study pyrolysis and gasification of biomass include dif-

ferent a1 scanning calorimetry (DSC), thermal gravimetr ic analysis (TGA), and

thermal mechanical analysis (TMA). Support studies for the PDU will also

cont nue.



PROCESS DEVELOPMENT UNIT STUDIES

The main objective of the PDU studies was to evaluate catalysts and oper-

ating conditions developed in the laboratory on a scale approximating commer-

cial operations to identify heat and mass transfer problems associated with

scale-up. Operations were structured to allow determination of material and

energy distribution in the system.

Design and procurement of equipment began in January 1978. The initial

design included a stirred-bed gasifier. By March 1980 the PDU was modified by

converting the stirred bed to a fluidized bed. The PDU designs and operations

are presented in the following discussion.

PDU DESCRIPTION

Design of the PDU with the stirred-bed gasifier was initiated before

design criteria were fully established in the laboratory. Therefore, the PDU

was designed for general purpose operation to develop the following

information:

a process yields, conversion efficiency, thermal efficiency, throughput

a required or desirable process features (catalyst recycle, feed mate-

ri a1 preparation, product gas treatment)

a equipment requirements (materials hand1ing, instrumentation, con-

struction materials, unit operations)

utility requirements

a emission control requirements.

A flow schematic of the original PDU with the stirred-bed gasifier is shown in

Figure 26. A photograph of the completed PDU is shown in Figure 27.

Gasifier

The stirred-bed gasifier is illustrated in Figure 28. The gasifier is a

250-psig pressure vessel. The shell is 2 ft (60 cm) sch 80 carbon steel pipe

10 ft (3 m) long. The bottom head of the gasifier contains the gas inlet



FIGURE 26. Schematic o f Original PDU



FIGURE 27. Photograph of Original PDU



l I 4 i N . TUBE BUNDLE

FOR S A M P L l NG AND

TEMPERATURE PROBES

CASTABLE REFRACTORY

TUBE L I N E Z S - 3 I N . I

FEED PORT

Dl STR lB UTOR PLATE

SOLIDS SAMPLER HOUSING

CERA MIC TUBE L I N E R S -

I D O P EN I NG 4 I N .

FIGURE 28. Schematic o f the Stirred-Bed Gasifier



nozzle and gas distributor plate. A drain line in the bottom head allows an

operator to remove materials that fall through the distributor plate. Three

small lines permit sampling of solids from three different levels in the gasi-

fier. Figure 29 is a photograph of the bottom head section with the solid sam-

pling and removal lines.

Three nozzles above the distributor plate on the gasifier allow addition

of biomass at different levels in the four foot bed. We have, however, only

used the lowest feed port. An internal auger about 8 in. (20 cm) in diameter

was used to stir the 4-ft-deep bed of catalyst and char.

Energy was supplied to the gasifier by heaters surrounding an 11-in.

(28-cm) diameter, 310 stainless steel 1 iner. Heater leads leave the gasifier

through glands in three flanged ports. In the early experiments two types of

heaters were tested. The original heaters were nichrome resistance heaters

embedded in ceramic. Several different configurations were tried, however,

these heaters could not withstand the reducing environment. They were rep1aced

with incoloy-sheathed heaters wrapped around the stainless steel liner. These

heaters lasted we1 in the gasifier, however, their heat transfer characteris-

tics did not allow operation at bed temperatures above 550'~.

The reactor has an internal refractory lining 14 in. (36 cm) I.D. by

16 in. (41 cm) O.D. The remainder of the shell insulation is provided by

~aowool8. Figure 30 shows the refractory and ~ a o w o o l ~in the reactor as well

as the three electrical ports at 120' spacing and three biomass feed ports.

The stirred-bed gasifier was equipped with three instrument bundles. A

bundle consisted of seven stainless steel tubes welded in a close-packed con-

figuration. Each tube extended to different levels in the gasifier. One bun-

dle was used to measure pressures at seven points in the bed. Bundle two had

seven thermocouples at different levels. The third bundle was used for sam-

pling the gas composition at different levels.

The exterior of the reactor was painted with temperature sensitive paint

which changes colors as the temperature exceeds certain values. This is useful

8Registered Trademark of Babcock and Wilcox.



FIGURE 29. Bottom Head o f Stirred-Bed Ga s i f i e r

64



in identifying hot gas circulation between the shell and refractory. Seven

thermocouples were glued to the shel exterior to monitor shel temperatures.

From October 1979 to March 1980 the PDU was modified extensively. The

most significant change was the modification of the gasifier vessel to operate

in a fluid-bed mode.

During the agitated-bed testing it became apparent that the agitated sys-

tem was difficult to model, and very difficult to scale up. Also, and perhaps

most importantly, the agitator was presumed to be the cause of severe catalyst

attrition. Finally the system did not have good heat transfer characteristics

which prohibited operation at temperatures above 600'~ without the use of oxy-

gen. For these reasons it was decided to convert the PDU to allow fluidized-

bed operation.

The modifications involved removal of the agitator and the agitator

support-drive system, installation of a new gas inlet section, a new distribu-

tor plate, a new stainless steel liner and gasifier heaters.

With the auger removed from the bed, the effective cross-sectional area of2

the reactor was 0.66 ft2(610 cm2) instead of 0.35 ft (330 cm

2) with the

auger. The larger area meant that the gas flows would need to double to main-

tain the same linear velocity. For this reason a new liner for the bed was

2 2constructed for an inside effective area of 0.31 ft (290 cm ) . The new liner,

also made of 310 stainless steel, had an I.D. of 7.75 in. (20 cm).

A new gasifier bottom head and gas distributor were procured to accommo-

date the smaller diameter bed. This assembly was intentionally designed to

decrease the heat loss in the gas inlet section. Only one bed-draining device

is used in the new bottom head assembly. The distributor plate was reposi-

tioned from 14 in. (36 cm) to 2 in. (5 cm) below the wood-feed port. This

allows the feed to be reacted with the hottest gas in the system.

Ceramic fiber heaters, with nichrome elements were clad to the new bed

liner. These heaters failed after only several days use. They were replaced

by sheathed heaters which are still in service.



A problem remained in getting enough heat into the gasifier bed. This was

solved by installing six cartridge heaters directly into the bed of char and

catalyst. These heaters have a total capacity of 30 kW and have performed very

well to date.

The top head of the gasifier had to be replaced for the installation of

the cartridge heaters. The modified head does not utilize the instrument bun-

dles described earlier. Instead, individual tubes for thermocouples and pres-

sure sensors extend into different levels of the gasifier from the top head.

Figure 31 shows a schematic of the fluidized-bed gasifier as it now

exists.

Wood Feed System

Wood is introduced to the gasifier using a lockhopper and an auger feeder.3 3

The lockhopper is a 4-ft (0.11-m ) chamber with pneumatically operated ball

valves on the top and bottom. Figure 32 shows the lower portion of the lock-

hopper. A level indicator in the lockhopper signals when wood loading is

required. Operators raise weighed batches of wood to the hopper via a

hydraulic-electric crane. The hopper valves are closed and the hopper is

purged with Cop. After purging, the top hopper valve is opened, and the wood

is added to the hopper. The top valve is closed. Air is purged from the hop-

per, and the bottom hopper valve is then opened. The wood drops to the auger

feeder. Interlock mechanisms prevent operators from opening both hopper valves

at the same time. Upon loss of power or supply air, the bottom valve fails to

the closed position, and the top valve fails open to ensure safety.

The screw feeder consists of two metering screws approximately 2 in.

( 5 cm) in diameter. These operate with a variable-speed hydraulic drive at

about 20 to 40 rpm. They push the wood onto a third screw oriented 90' from

the metering screws. The third screw (injector screw) is also about 2 in.

( 5 cm) in diameter. It operates at about 150 to 200 rpm with a variable speedhydraulic motor. At this speed, wood is rapidly moved into the gasifier. The

injector screw also is connected to a hydraulic ram that gives the screw 6 in.

(15 cm) of travel. This feature is helpful in breaking jams in the injection



15-cn

TUBE

FEED

I PIPE WITH CERAMIC

LINERS - 8-

cm

PORT

FEED PORT

METAL LINER

20 cm ID x 137 cm

LONG

GAS INLET NOZZLE

PRODUCT GAS

OUTLET

KAOWOOL BLANKET

8 cm THICK

CASTABLE REFRACTORY

3 cm THICK

6 EACH SUBMERGEDCARTRIDGE HEATERS

EXTERNAL BED

HEATERS

120-cm BED CATALYST

AND CHAR

DISTRIBUTOR PLATE

BED SAMPLING DEVICE

FIGURE 31. Schematic of Fluidized-Bed Gasifier



FIGURE 32. View of Lock Hopper

6 9



section. Controls for the feeder include two tachometers, control valves, and

an electric interlock that prohibits operation of the metering screws without

the injector screw in operation.

Gas Supply Systems

Reactant gases include steam, oxygen, air, and carbon dioxide. Steam is

supplied with a 150-lb/h (68-kg/h) 100-psi (690-kPa) generator. The saturated

steam is heated to 110'~ with trace heaters prior to the steam control valve

and orifice meter. A pneumatic integral orifice meter and controller are used

for steam control. Oxygen can be fed to the gasifier if desired. Since oxygen

requirements are low, the oxygen is supplied with 380 1b (170 kg) liquid oxygen

storage cylinders. Oxygen flow is measured by a rotameter. Low pressure air

is available as a feed gas. Air is metered using a pitot tube sensor. Carbondioxide is available from a 6-ton (5-t) storage tank and is metered using a

pitot tube apparatus. Carbon dioxide is sometimes used for purge gas on the

plant. Nitrogen is available from standard gas cylinders. It is used as a

quench gas in case of automatic shutdown. It is also used for purge gas. All

feed gas temperatures are recorded on a datalogger. Pressures and flow rates

are recorded in data books by operators. All feed gas lines contain check

valves to eliminate backflow in the lines.

The feed gases mix in a header system and then travel through a 1-in.(2.5-cm) ball valve. This valve is interlocked with the datalogger. If gasi-

fier temperatures exceed a preset point (usually 800°c), the valve will auto-

matically close, thereby preventing runaway temperatures in the reactor. When

the valve closes, another valve on the nitrogen quench system opens, and the

gases in the gasifier are flushed out of the system. The feed gases then pass

through the gas preheater. This unit uses electric heaters to bring the gases

to about 700'~. The controls for this heater and a portion of the heater are

shown in Figure 33.

Provisions were made for recycling a portion of the product gas through

the gas heater to the gasifier. This would be useful for boosting velocities

in the gasifier. The blower used for recycle gas was a rotory lobe-type with

a variable speed drive. A packed-bed contactor was placed upstream of the



recyc le b lower to coo l the recyc le gas by d i rect con tact w i th water i f neces-

sary. Figure 34 shows the contactor and recycle blower. Recycle gas flow

rates were measured by d i f ferent ia l pressure produced wi th a p i t o t tube. The

automatic shutdown system described previously a1so stops the recycle b lower .

Hot feed gases e x i t the gas heater through a 3- in. (7.5-cm) s ta in less

s t e e l l i n e l e ad in g t o t h e g a s i f i e r i n l e t n oz zl e. Th is l ine (about 20 f t ( 6 m)

long ) was we ll i nsu lat ed, however, heat loss was s t i l l h igh. Heaters were

added t o po r t i ons o f the feed l i ne t o coun teract the hea t l oss . Feed gases

now reach the gasif ier at temperatures between 7 5 0 ' ~ and 800 '~ .

Gas Cleanup System

Gas cleanin g i s requ ired to separate and recover the products f o r analy-

s i s . Th is in vo lves separa t ing the so l i ds and l i qu id s f rom the gas and coo l i ng

the gas.

The pr imary means o f so l i ds separat ion i s by a dry cyclone operat ing about

10 f t (3 m) downstream o f t he ga s i f i e r. The cyclone us ual ly operates i n the

temperature range of 3 0 0 ' ~ t o 5 0 0 ~ ~depending on gasi f ier condi t ions. I n s i d e

d iameter o f the cyc lone i s 6 i n . (15 cm). The cyclone i s ins ula ted t o prevent

condensat ion on i t s wal l s. Sol ids were dra ined from the cyclone in t o a drum

which was weighed and sampled after a day's test. I t was de si ra bl e t o sample

more f r equ en t l y to ob ta in co l l ec t i on ra tes a t s teady s ta te . A lock hopper sys-

tem was b u i l t f o r the cyc lone. It consi sted of a 2- in. (5-cm) p ipe w i th ba l l

valves at each end. This cyclone hopper i s drained at 1/2-h i n t e r va l s .

Downstream of the cyclone (see Figure 26) , a wet venturi scrubber was used

for gas coo l ing and f i n a l par t i cu la te remova l. The scrubber contacted water

w i t h t h e h o t gas i n a va r i a b l e t h r o a t ve n tu r i . Water cou ld be c i r cu la te d i n

the system or passed through the system on a once-through basis. C i r c u l a t i o n

was d i f f i c u l t because f lo w meters and water l in es p lugged wi th char and ta r .

Most o f the t ime the water to the scrubber was not reci rc u l at ed. Yie lds of

char and ta r were determined by c ol l e c t i ng t i med l i q u i d samples and perform ing

so l i ds ana lys is and measur ing to t a l o rgan ic carbon i n the l i q u i d phase.

The cl ea n gas was t hen measured w i t h a p i t o t tube and subsequently with

an o r i f i c e me te r. The p i t o t tube was unre l i ab l e w i th i t s opening tend ing t o



plug. The o r i f i c e me te r has been a r e l i a b l e gas me te r ing dev i ce f o r t h i s

a p p l i c a t i o n . Af ter sampl ing and meter ing, the gas i s burned i n a gas burner.

Poor pe rformance o f t he ve n tu r i sc rubber re su l t ed i n i t s r ep lacement w i t h

an e l ec t r os ta t i c p re c i p i t a to r (ESP) shown i n F i gu re 35. The p r e c i p i t a t o r i s a

5- i n . (13-cm) p ipe w i th a 1- in. (2.5-cm) e lec t ro de down the cent er o f the p ipe.

I on i z i ng d i sks a re l oca ted on t he l owe r po r t i o n of t he e l ec t rode . Gas enters

a t t h e b o t t o m o f t h e p r e c i p i t a t o r a n d e x i s t s a t t h e t o p . The ESP i s desi gned

f o r 1 00 p s i g ( 69 0 kPa) opera t ion . The un i t can ope ra te w i t h e i t he r a d ry wa l l

o r a wet ted wa l l . C u r r en t l y , w at er i s c i r c u l a t e d th ro ug h t h e p r e c i p i t a t o r t o

coo l th e gas and remove th e pa r t i c u l a t es co l le c t ed on the wa l l . A schematic of

th e PDU co nf ig ur at io n as o f October 1980 i s shown i n Fi gur e 36.

Inst rumentat ion and Data Analysis

Ana ly t i c a l equ ipment f o r gas ana ly s is cons is ts o f con t inuous ana lyzers

f o r i ns tantaneous composi t i ons and gas chromatographs f o r pe r i od i c de ta i l ed

ana l ys i s .

Four cont in uous u n i t s are used. Two Beckman in f rared analyzers measure

carbon monoxide and carbon di ox id e l ev el s. Oxygen i s measured w i t h a Beckman

pola rogr aphi c sensor. Hydrogen i s measured by thermal con du c t i v i ty i n a Mine

Saf ety Appl iances analyzer. Cont inuous reading s are recorded eve ry 15 min by

an operator.

Detai led gas analyses are processed every 20 min wi th a Car le thermal con-

d u c t i v i t y gas chromograph. The chromatograph i s combined w i t h a Spectr ophysi cs

SP 4000 micropr ocess or fo r automatic int eg ra t i on . The gases measured are CO,

C02, Hp, N2, O2 and C1 t o C4 pa r a f f i ns and o l e f i n s . Occas i ona l ly a Hew l e tt

Packard f lame phot ometr ic chromatograph i s used f o r hydrogen s u l f i d e measure-

ment i n the par t s per m i l l i on range .

Samples from the PDU are drawn to the analyt ica l equipment wi th a vacuumpump. The cont inu ous samples are drawn from th e re ac to r e x i t gas l i n e . The

chromatograph samples are nor mal l y taken a f t e r the gas i s cleaned and cooled,

however, th es e samples can be drawn fr om 10 o t he r l o c a t i o ns i n t h e p l a n t u s i ng

a 12- posi t ion va lve on the sample l ines.



Water content of the gas f rom the gas i f i er i s an important fa ct or i n the

mass balances. The water con ten t was o r i g i n a l l y measured w i t h a dewpoint

hygrometer. However, t h i s u n i t was un rel ia bl e, and poor se rvi ce from the fac-

to r y res ul t ed i n i t s discont inued use. Water content i s now determined severaltimes du ri ng a t e s t by drawing a gas sample, condensing the water, and mea-

suring the water collected and the dry gas volume.

Temperatures are scanned co nti nuou sl y and recorded a t 20-min intervals on

a 60-channel, pr i n t i n g datalogger. The datalogger can als o auto mati cal ly

shutdown gas feed t o th e PDU shoul d a temperatur e excursion occur. The tem-

peratu res monitored on th e datal ogger in cl ud e each feed gas, the heated gas,

7 po in t s i n the gas i f i e r , 7 re act or surface temperatures, o f f gas temperatures

and co ol in g water temperatures. Alarms on th e datal ogger al lo w opera tors t osee where problems may be developing and t o ta ke a pp ro pr ia te a cti ons. Tempera-

tu re s from the datal ogger tapes are keypunched, and temperature p l o t s are gen-

erated by a computer.

Pressures, d i f ferent i a1 pressures, co nt ro l l er set t ings, and meter readings

are recorded at 112-h in te rv al s i n a data book by an operator . This data book

also contains records o f wood add i t i on and cyclone co l l ec t i on rates. Af t er a

te st , these data along w i t h gas compositions and recorded temperatures are

processed on a minicomputer to gain a prel i min ary est imate o f the t e st re sul ts.The results of analyses subsequently performed on the l iquids and solids are

la te r entered in to the computer f o r a f i n a l t es t analys is. Data are s tored on

magnetic disks. Or ig ina l raw data are a lso kept for fu ture re ference.

PDU OPERATION

The purpose of the PDU op er at ions was t o t e s t c a ta l y s t systems developed

i n t he l abo r a to r y and t o ev alua te t he i r p r a c t i c a l i t y f o r l a r ge-scale operat ion.

PDU op er at ions were completed wi t h the s ti rr ed - bed gas i f ier and the f lu id-bed

gas i f ie r ; resu l t s a re p resented i n t he fo l low in g d iscussion .

Successful operat ion of the f lu id - b e d g a s i f i e r i l l u s t r a t e d t he t e ch ni ca l

f e as i b i l i t y o f t he c a t a l y t i c processes. Res ul ts w i t h the s ti r red- b ed g a s i f i e r



were o f l i t t l e value s ince operat ions wi thout oxygen were l im i t ed t o tempera-

tures under 6 0 0 ' ~ and since scale-up o f the ga si f ie r appears t o be imprac t ic al .

Test Procedures and Safety ConsiderationsDocuments were prepared t o assure s a fe ty o f t he PDU op er at io n bef or e

st ar tu p was per mit ted . The f i r s t document, "Operat ional Safety Analysis Review

for Cata ly t ic B iomass Gas i f ica t ion PDU," contained an analys is of po te nt ia l

problems and hazards o f each pi ec e o f equipment i n th e PDU and descr ibe d th e

safety features to be incorporated in to the or ig ina l p lant des ign. The second

document, "Safe Operating Procedures," 1is ted the s teps fo r p lan t s ta r tup ,

shutdown, normal operation, and emergency conditions.

Normal te s t procedures begin wi t h heat ing the p l an t t o th e des i red tem-perature, which takes 3 t o 4 h. When a t th e de si re d temperature, wood i s

in t roduced, and the p lan t i s operated at s teady condi t ions f o r 3 t o 5 h. Wood

feed i s then stopped, and the p la nt i s e i t her p a r t i a l l y or t o t a l l y shutdown

depending on whether or not i t w i l l be operated the fol lowing day.

Feedstock f o r the PDU

Wood feedstock f o r th e PDU was l i mi te d i n s ize by the screw feeder, which

could de li ve r chips wi th a maximum dimension of about 112 i n . (13 mm). O r i g i-

n a l l y , wood was purchased th a t had been s p ec i a l l y processed i n a hammer m i l l .

Later we found a supply of head r i g sawdust (maple and al de r) t h a t was su i ta b l e

f o r use i n th e PDU.

Another mat eri al t ha t was used f o r several te st s was fo r es t residue. This

was c ol le ct ed from th e f o re st i n the same manner wood residue s are c ol le ct ed

commercially. The wood has a different composition, containing much more ash

from the s oi l s than the normal pl an t feedstock. The composit ions o f the two

feeds are given i n Table 26.

Wood mo is tu re cont en t markedly i nf lu en ce s PDU performance. Wet wood tends

t o br idg e i n the lock hopper and gre at ly increases the heat load t ha t must be

supp l ied t o t he gas i f i e r . For these reasons, th e wood feedst ock was us ua l l y

d r i e d t o 1 t o 5 w t % moistu re b efo re use i n t he PDU.



TABLE 26. Compos itions of Feedstocks Used i n PDU Oper at ions

Feedstock Headr i g Sawdust Fo re st Residue

Weight %, dry bas is :Carbon 46 46

Hydrogen 6 6

Oxygen 48 39

ash 0.1 9

Heat o f combust ion , dry basis,

B t u / l b ( k J / k g ) 8300 (19,300) 8700 (20,200)

PDU Results wi th the Agi tated-Bed Gas i f ie r

The i n i t i a l operat ion o f the PDU wi th the agi tated-bed ga si f i er began i n

the f a l l o f 1978 and con tinued u n t i l t he f a l l o f 1979. Mod i f i ca t ions t o

improve PDU o p e r a b i l i t y were in i t i a te d a t the end o f t h i s t es t pe riod. Opera-

t i o n o f th e PDU wit h the agi tated-bed ga s i f i e r was successfu l i n obta in ing data

on generat ion of several specif ic gas products--a methane- ri ch gas, a methanol

sy nt he si s gas, and an ammonia sy nt he si s gas. Resul ts are presented i n the f o l -

lowing discussion.

Problems were encountered i n opera t ion o f the agitated-bed gas i f i e r

inc lud i ng fa i l u r e o f t he e l ec t r i c a l hea te rs , un re l i ab le wood feed ing , and

plu ggi ng o f the vent ur i scrubber. The prima ry problem was in s t a l l a t i o n and

operat ion o f the in ter na l e le c t r i ca l heaters . Several designs were t r i ed , but

none proved t o be complete ly suc cessfu l. The maximum temperature obtained was

600°c, which i s 1 5 0 ' ~ below the desired upper temperature for operat ion.

The wood feeding problems were el imi nate d by replacement of th e i n i t i a l

single-screw design. The or i g i n a l feeder used a si ng le screw f o r metering and

s low ly in je c t i ng wood to t he gas i f i e r . We specu lated tha t t a r s depos i ted i n

the screw channel, causing severe bi nd in g o f th e screw. The system currently

i n use separates the meter ing and i n j ec t i on func t ions . Two screws meter wood

from the lock hopper onto a th i rd screw which rapidly in jects wood into the



bed o f ca ta l y s t . The i n j ec t io n screw i s connected t o a hyd rau l ic ram capable

o f a 6- in. (15-cm) t ra ver se f o r removal o f any p lu g th a t might form.

The ven tu r i scrubber gave problems wi th t a r r y so l id s p lug ging the water

c i r cu la t io n l i ne s . Th is p rob lem was reduced by once- through scrubbing wi th

water ins tead o f water rec yc l in g .

Fig ur es 37 and 38 summarize t he e f f e c t s o f temperature and ca t al ys t com-

b ina t ion s on p roduc t ion o f a methane- r i ch gas i n the ag i tated- b ed g a s i f i e r .

I n i t i a l l y we hoped the p r imary ga s i f i ca t i on ca ta l ys t , po tassium carbonate ,

would promote the steam-cha r reac t i on t he reby i nc reas ing y i e l ds o f CO and Hz

and producing a hi gh methane gas i n t he presence o f a secondary methanation

ca ta ly s t . This d i d not happen, however. At low temperatures, th e g as i f i ca t i on

ca ta lys t re ta rded bo th methane produc t ion and carbon convers ion re la t i ve to

y i e l ds w i t h seconda ry ca ta l y s t s on l y . For purposes o f comparing ot her

processes, F igu re 37 g ives the methane y i e ld s i n t o t a l s poss ib le- -assum ing a l l

CO and Hz was converted. I n t he 525OC t o 5 7 5 ' ~ temperature range, re s i du al Hz

and CO were both 10% t o 15% i n the PDU product gas. Another methanation step

would be needed t o co nver t these res id ua ls t o methane.

The shaded areas i n Fi gur es 37 and 38 re pre sen t maximum s c a t t e r o f t he

data. The so l i d l i n e s show tren ds fo r the oper at io ns where few problems

occurred. The t o t a l pos s ib le methane prod uct i on was d e f i n i t e l y lowered by

potass ium carbonate at low temperatures but po ints were too scat tered to see a

r e a l di f f er en ce between con cent rat ions o f 5% and 10% by weight potassium

carbonate.

We specu la te t h a t one e f f ec t o f t he p rima ry ca t a l y s t was t o a t t ack t he

n icke l methanat ion ca ta lys t and deac t i va te i t. At t he same time, th e potassiu m

carbonate l o s t some o f i t s e f fec t i vene ss . Several observa t ions suppor t t h i s

hypothes is bu t a re no t necessar i l y conc lus ive . F i r s t , t h e c a t a l y s t p e l l e t s

were d iscolored af ter these these tests but not when the potass ium carbonate

was absent. Hydrocarbon synthes is gas t es ts wi th out the n i ck el c at a l ys t

(potassium carbonate + SiA l on l y ) show a bene f i c i a l e f f ec t o f t he po tass i um

carbonate was absent. Hydrocarbon synt hesi s gas t es t s wi th ou t th e ni ck el

ca ta lys t (po tass ium carbonate + SiA l on l y ) showed a ben e f i c i a l e f f ec t o f t he



AVERAGE BED TEMPERATURE, OC

*ASSUMES ALL CO + Hz

ARE CONVERTED TO METHANE

S iA l + N i ONLY-

-

- PDU 172 K 2 C 0 3 + S i A l + N i

K2COJ + S iA l + N i

I I I

FIGURE 37. Methane Yields from the Agitated-Bed Gasi f ier as aFunction of Temperature and Catalyst



A LAB STUDY S i A l + N i

AVERA GE BED TEM PERATU RE, OC

FIGURE 38. Carbon Conversions for Methane Productionin the Agitated-Bed Gasifier

potassium carbonate i n terms of increasing gas yields. Our ammonia synthesis

gas tests clearly showed the effectiveness of the alkali in promoting the

water-gas-shift reaction (H20 + CO + H z + Cop).

Laboratory studies showed that the alkali carbonate primary catalyst

increases the yield of char, increases gas yield, and reduces condensable



l i q u i d y ie l d a t temperatures be low 6 5 0 ' ~ wi t h no secondary ca ta l ys t. The

secondary s i l i c a a lumina and n i cke l ca ta l ys ts a lso e f fe c t iv e l y c rack and re form

l iquids f rom pyrolys is of wood thereby increasing gas y ields over those

obtained wit h the pr imary ca tal yst . F igure 38 shows the same ef fe ct s of these

ca tal yst s. The constant carbon conversion wi th temperature i n the presence of

secondary cata lys t ind icates that on ly pyro lys is o f the wood i s a f f ec ted and i s

as complete at 4 5 0 ' ~ as at 600'~. The water-gas react ion (C + H20 + CO + Hz)

becomes s i gni ficant at about 650'~.

Tests on methanol, o r hydrocarbon, syn the sis gas pro duc tio n i n the

agitated-bed gasi f ier concentrated on generat ion of a 2:l H2:C0 m o la r r a t i o

i n th e gas stream. Fig ur es 39 and 40 summarize f i nd ings fo r t hese tes t s . We

found th at the ga s i f ic a t i on ca ta l ys t was ef fe c t iv e both i n increas ing gas

yields and carbon conversions. The ben ef i t s o f us ing the a l ka l i were most

apparent at higher ( - ~ 7 0 0 ~ ~ )temperatures. The primar y ca ta ly st (a1k a l i car-

bonate) e f f e c t i v e l y promoted the water-gas react ion (C + H20 + CO + Hz) a t

temperatures above 650'~.

The preferred temperature for generat ion of a 2H2:1C0 mole ra t i o gas i s

850'~. The heate rs were inadequate f o r mai nta in in g temperatures above 6 0 0 ' ~

i n the PDU, and pure oxygen was needed t o o bta in even 700'~. Yields are

markedly improved at the higher temperatures and without addit ion of pure oxy-

gen. A reac tor conf igurat ion for s team g a s i f i c a t i o n of the wood at 8 5 0 ' ~ was

needed. Conceptual ly, tube bundles i n a f l u i d i ze d bed o f ca ta lys t could be

used to provide the required endothermic heat of react ion.

From an ope rat i ona l standpoint, th e ga si f i ca t i on ca ta ly st caused problems.

When lo ca l h ot spots developed, the carbonate melted fu si ng to get her the sec-

ondary ca ta ly st and any char i n th e bed. Thi s problem occurred sever al t imes

and prompted a search for other catalyst (non melt ing) combinat ions that would

al so enhance gas y i e l ds and carbon conversion. The ni ck el and sil ic a- a lumina

combinat ion looked very promising i n the laborato ry. The low gas yi e l d s i n

the PDU were a l i t t l e puzz l ing. We no ti c ed a f t e r t hese t es t s t ha t l a r ge

amounts of the secondary catalysts were entrained from the gasif ier. ( I t i s

no t easy t o de tec t t h i s du r ing a t es t . ) We now speculate the re was i n s u f f i -

ci en t cat al yst i n the bed to promote ga si f i ca t i on o f the wood and char. Also,



- 0 4% K2C03

A S i A l ONLY

STEAM ONLY

-

-

0

- S i A l + Ni

- AYAI 1 I

500 600 700


FIGURE 39. Synthes is Gas Yi eld s as a Funct io n o f Temperaturei n t h e A gi ta te d-B e d G a s i f i e r

we have since determined t h a t the presence o f oxygen des troy s th e nic k e l cata -

l y s t . L a t e r t e s t s w i t h t h e f l u i d - b e d g a s i f i e r ( pr es en te d i n a l a t e r s e c t io n )

showed exce l len t resu l ts w i th a n icke l-on-alumina secondary cata lyst .

Compar ing the carbon convers ion p l o t f o r syn thes is gas p roduct ion w i t h th e

s im i l ar p l o t f o r methane produc t ion shows an apparent co ntr ad i ct ion . We see

h ig h carbon convers ions a t l ow temperatu res f o r a l ka l i - c a ta ly z ed t e s t s i n F i g-

u re 40 and low conversions i n F igu re 38. However, no methane t e s t ev er used

a i r o r oxygen t o mai nta in bed temperatures, whereas syn thes is gas t e s t s used

some oxygen. At t imes, th e re ac to r heaters were no t working so even f o r low



10%K2C 0 3 + SiA l

- 0 4% K2C03

n S iA l ONLY, PDU

-A S i A l ONLY, LAB STUDY

S i A l + N i + K2CO3, LAB STUDY

S i A l + Ni, LAB STUDY

500 600 700 . 800

.


FIGURE 40. Carbon Conversions f o r syn t hes i s Gas Productioni n the Agi ta ted-Bed Gas i f ie r

temper ature cases oxygen was added. For a p a r t i c u l a r synt hes is gas t e s t show-

i n g 80% carbon conversion (with O2 add i t i on ) t he f i gu r e drops t o 65% when the

carbon burned i s deducted.

A second order d i l u t i o n ef fe ct was also p ossibl e i n the methane product i on

runs. Because produc t gases were rec ycl ed, on ly 75% (or less) o f the feed gas

to the gasi f ier was a reactant . By contra st, 100% o f the feed gas t o the gasi -f i e r d ur ing synthesis gas product ion was a reactant .

The amnonia synthesis gas studies were conducted to determine the effects

o f temperature and ga si f i ca t i o n ca ta ly st on produ ct gas f lo ws and composit ions,

as we ll as wood conversions. A l l te s ts used a si li ca - a lumina c racking ca ta ly s t

plus an iron-chromium sh i f t cata lys t . Prior laboratory studies recommended a

4 t o 1 steam/wood weigh t ra t io t o m in im ize CO i n the produc t gases. The PDU



s tud i es used ra t i o s c l ose t o t h i s . We recogn i ze t h a t t he h i gh steam f l ows

r e pr e se n t a s i g n i f i c a n t e ne rg y p e n a l t y f o r any commercial opera t ion .

I n order t o maint ain d esi red bed temperatures, pure oxygen was added t o

the a i r s t ream to burn a por t ion o f the wood feed . From an operat ional stand-

point our procedure was 1) e s t a b l i s h t h e d e s ir e d ope ra t i ng tempera tu re us ing

p r e c a l c u l a t e d o x y g e n l a i r r a t i o s , 2 ) note hydrogen concentrat ions and product

f l ows at steady state, and 3) read just oxygen and a i r f l ows to g ive the des i red

3:l H2 /N2 m ol e ra t i o .

A d e t a i l i s impo rtan t here when examining the gas composi t i ons shown i n

Table 27. Ni t rogen composi t i ons depend so le ly on the a i r f l ow ra tes. Wi th

a d d i t i o n o f p u r e O2 and ai r , the two sources of oxygen can be trade d t o get

whatever n i t rog en concen t r a t ion i s des i red . Only opera to r i nexper ience and

TABLE 27. Ammonia Synthes is Gas Re su lt s fr om th e Ag it ated-B ed G a s i f i e r

Bed Temperature, OC 627 654 732 0' 735

W t % K2C03 ( a 0 10 0 10

kg s teamlkg dry wood 5.2 4.0 4.7 3.8

kg a i r / k g dry wood

kg 02/kg dry wood

Dry Gas Composition, v o l%

Carbon Conversion to a Gas, % 94 95 91 933m d r y gas/kg dry wood 1.37 1.81 1.37 1.75

Cold Gas Eff iciency, % 4 6 63 53 57

(a) Dry mixed with wood



imprec ise f l ow con t ro l p reven ted ou r ge t t i ng exac t l y t he 3 : l H2:N2 r a t i o i n t h e

product gas. The important numbers are residual CO and gas yields as tempera-

tures change and/or g a s i f i c a t i o n c a t a l y s t i s added. Observat ions f rom these

t e s t s a r e :

T he gas i f i ca t i on ca ta l ys t ( 10 w t % d r y m ix K2C03) m a r k e d l y s h i f t s t h e

CO t o Hz a t b o t h 6 5 0 ' ~ and 750'~.

A d d i t i o n o f t h e g a s i f i c a t i o n c a t a l y s t i nc re as ed gas y i e l d s n o t i c e a b l y.

A l though i t i s no t shown i n T ab le 27, t he wa te r con ten t o f t h e e x i t gases

i s o f i nt e re s t . I n a l l cases i t exceeded 70 mole% o f th e prod uct stream. Th is

suggests a huge excess of steam was present and supports our ear l ier conten-

t i o n th at more work i s needed a t lower steam f l ow ra tes . The e f f e c t o f the

decreased steam flow on CO concen t ra t i on w i l l be o f p r i me i n t e r e s t .

PDU Resul ts With the F luid-B ed G a s i f i e r

The PDU w i t h t h e f l u i d - b e d g a s i f i e r was op er at ed f o r a s e r i e s o f t e s t s i n

the temperature range of 5 0 0 ' ~ t o 7 5 0 ' ~ wi th severa l secondary cata lys ts and

w i t hou t secondary ca ta l ys t f o r va r ious steam ra tes . Most t es t s i nc luded a t

l e a s t 4 h o f opera t ion a t s t eady s t a t e .

The i n i t i a l t e s t s er i es used a Harshaw N i-3 26 6 c a t a l y s t i n t h e fo r m o f a

1116-in. (1.6-mn) ext rudate . These te s t s were a l l a t a temperature o f about5 0 0 ' ~ f o r genera t ion o f a me thane- r ich gas. Because o f the cata lys t s ize ,

sub s ta n t i a l amounts o f r ecyc le gas had t o be used t o f l u i d i z e t he c a t a l ys t bed.

The hig h gas v e l o c i t y and the low temperature i n the bed were proba bly respon-

s i b l e f o r t he low carbon conv ersi on compared t o subsequent t e s t s a t lower

l i ne ar v e l oc i t ie s and h igh er temperatures . F igur e 41 shows the e f fe c t o f vary-

i n g t h e steam-to-wood r a t i o on t he p roduct gas compos i ti on f o r t hese t es t s .

Table 28 summarizes the data f rom these test wi th the Harshaw Ni -3266 ca ta l ys t .

Severalo f t he t es t s used f o re s t r es idue o f t he compos i ti on shown i n

Table 26. The fo re st r es id ue had a hig h ash conte nt compared to th e ty p i c a l

h e a d r i g sawdust t h a t was used f o r most t es ts . No d i s c e rn a b l e d i f f e r e n c e i n t h e

per formance of the PDU wi th the fores t res idue feedstock was noted.



CATALYST Ni 3266 FLU ID BED = 5 0 0 ~ ~

0 I I 1 I I I I I I I0 0.1 0.2 0.3 0 .4 0.5 0.6 0.7 0.8 0.9 1 . O

WE1GHT RA TI O OF STEAM TO WOOD

FIGURE 41. E f f e c t o f Steam-to-Wood Rat io on Gas Composi t ionf r o m t h e F l u i d-B e d G a s i f i e r

A t t r i t i o n p ro ble ms w i t h t h e e x t r u d a t e c a t a l y s t p ro mp te d t h e p ur ch as e o f

s p h er ic a l m a t e r i a l f o r use i n t h e f l u i d b ed. The m a t e r i a l was s i z e d t o a l l o w

f l u i d bed o p e r a t io n a t a lo we r v e l o c i t y t ha n r e q u i re d f o r t h e e x tr u d at e . I n e r t

spheres o f nonporous alu min a (U.S. Screen -20 + 40 mesh) and c a ta l y s t spheres

(U.S. Screen -16 + 30 m esh) o f n i c k e l on a ce ra m ic s u b s t ra t e we re o b t a i n e d f o r

t e s t i n g t o compare r e s u l t s w i t h and w i t h o u t a c a t a l y s t . T e st s w i t h t h e i n e r t

spheres were completed at 550°, 600°, and 750 '~ . R e s u l t s o f t h e s e t e s t s

a re shown i n Tab le 29.

O p e r at io n w i t h s p h e r i c a l p a r t i c l e s and h e a t e r s s ubmerged i n t h e b ed

i mp ro ve d t h e g a s i f i e r pe r fo r ma n ce i n s e v e r a l ways. H ea t t r a n s f e r was g r e a t l y

improved a1 l o w i n g o p e r a t i on a t 750'~. Tempera tu re con t ro l was marked l y

i mp ro ve d. The b ed became i s o t h e rm a l w i t h f i v e th e rm oc o u p le s i n t h e be d d i f -

f e r i n g l e ss t h an 5 '~ . W i th t h e a g i t a t e d b ed and w i t h t h e f l u i d bed o f l a r g e r

c a t a l y s t s i z e , t h e t e m p e ra t u re s i n t h e bed v a r i e d as much as 1 0 0 ~ ~ .F i g u r e 4 2



TABLE 28. Results of PDU Tests With Nickel Extrudate Catalyst

Test No. 1 2 3 4 5 6 7

Temperature, OC 550 470 495 495 450 505 510

Feedstock Headr i g Head r i g Head r i g Head r i g Forest Forest Headr i gSawdust Sawdust Sawdust Sawdust Residue Residue Sawdust

Weight Ratio

Steam-to-Wood .93 .29 .34 .54 .40 -0- -0 -


241 26 28 34 31 18 15

CH4 15 2222 22 23 27 30

CO 4 8 8 6 33 7 11

C02 4043 41 38 42 48 43

20.0 0.2 0.2 0.0 0.0 0.0 0.2

C30.0 0.3 0.2 0.2 0.2 0.0 0.1

Cold Gas

Ef f i c iency , % 62 36 42 66 43 48 43

Carbon Conversion, %

gas 67 44 47 64 49 52 481i q u i d - - 1 1 1 1 1 2

s o l i d 4 44 51 33 23 30 17

P o t e n t i a1 CH4

sc f / t on 9500 5400 6400 9000 6700 5900 6000

( nm3/t (300) (170) (ZO O ) (280) (210) (180) (190)

shows a comparison of th e bed temperature p r o f i l e s f o r a case wi th 1) t he

extruded catalyst without submerged heaters and 2) t he spher i ca l ca ta lys t w i t h

submerged heaters.



TABLE 29. Re su lt s o f PDU Tes ts w i t h Nonporous Al umina Spheres

Test No.

Temperature, OC


CH4co

C02

C2,C3yC4

Weight Rat io Steam:Wood

s c f p r o d u c t / l b wood( nm3/kg )


gas

1i q u i d

char

Cold Gas Eff iciency, %

P ot en t i a1 CH4

s c f / t o n

(nm3 / t

Using spher i ca l med ia i n the bed a l lowed opera t ion w i th ou t r ecy c l ing p ro -

duc t gas. Steam was the onl y re ac ta nt gas. As a r e s u l t of the consequent low

ve l o c i t i e s i n t h e g a s i f i e r , t h e ch ar re si d en ce t i me i nc re ased . Gas y i e l d s

increased dra mat ic a l l y . The char decreased i n s iz e and a f i n e char dust was

blown fr om th e g a s i f i e r t o th e gas cleanup system. The system was incapab leo f adequa te ly c lean ing th i s gas a t low ve loci t ies and the downstream p ip ing

p lugged f requen t l y .



0

0 33 66 99 132 165 198 231 264 297 330

T l M E ( m i n ) AFTER START OF WOOD FEED

PROFILE WITH EXTRU DATE CATALYST

AT SCREW FEEDER

1 ft ABOVE FEEDER

A 2 A ABOVE FEEDER

0 3 ft ABOVE FEEDER

v 4 ft ABOVE FEEDER

0

0 22 44 66 88 110 132 154 176 198 220

T I M E ( m i n ) AFTER START OF WOOD FEED

0 AT SCREW FEEDER

0 1 ft ABOVE FEEDER

* 2 ft ABOVE FEEDER

o 3 A ABOVE FEEDER

v 4 A ABOVE FEEDER

FIGURE 42. Tempera ture Pro f i l es i n Flu id-B ed G a s i f i e r



A spherical catalyst was obtained from W. R. Grace & Co., The Davison

Chemical Division of Baltimore, Maryland. The catalyst composition is proprie-

tary but has the following general properties:

main constituents - alumina, nickel

U.S. Screen Size - -16 +30 mesh

particle density - 92 1b/ft3(1480 kg/m

3)

bulk density - 55 1b/ft3(880 kg/m

3)

Results of tests with the spherical catalyst before and after regeneration are

presented in Table 30.

The catalyst was regenerated with steam at 750'~ followed by hydrogen

reduction at 450'~ after ~ e s t4. Regeneration was effective as can be deter-

mined from the data in Table 30 (Test 5). However, the catalyst appears to

deactivate rapidly after regeneration (Tests 6 and 7).

Graphical comparisons of catalyzed versus uncatalyzed results are shown in

Figure 43 for methane production and Figure 44 for synthesis gas production.

The advantages of catalyzed gasification are obvious.

Calculation of Equilibrium Gas Compositions

A computer program that calculates chemical equilibrium between solid car-

bon and gases was developed in 1980. The program was needed to determine thedeviations from equilibrium in both the laboratory-scale gasifier and the PDU

gasifier. The program was useful for estimating gasifier performance at 10 atm

pressure for the economic evaluation of methane and methanol production.

Finally, the program was used to determine the effects of varying the steam-

to-wood ratio at 7 atm to help provide design data for future pressurized

operation.

The equilibrium program describes a batch system. Wood and steam are put

into the system and equilibrium compositions at a specified temperature andpressure are calculated. It can be easily modified for other feedstocks and

other reacting gases.



TABLE 30. Resul ts o f F l u i d Bed PDU T es ts w i t h N i c k e l C a t a l y s t

Test No.

Temperature, OC


H2

CH4CO

C02

2

C3

Weight Rat io Steam:Wood 1.0 0.60 0.63 0.76 0.75 0.6 0.7

Product Gas Yield

s c f / 1b wood 21.7 19.4 21.9 18.6 27.2 18.5 16.1

( nm3/kg) (1.35) (1.21) (1.37) (1.16) (1.70) (1.15) (1.01)


gas 78 7 4 70 7 3 8 2 66 7 2

l i q u i d 1 1 1 3 1 1 2

s o l i d 26 15 18 23 22 21 23

Potent ia l Methane Y ie ld

s c f / t o n 11,000 10,900 9,800 9,200 11,500 8,500 8,900

( nm3/t (340) (340) (310) (290) (360) (260) (280)

P o t e n t i a1 Methanol Yieldw i tho u t re form ing ,w t / w t dry wood 0.34 0.29 0.44 0.35 0.58 0.37 0.28

Cold Gas Ef f i c i en cy 79 77 75 71 88 65 6 7



UNC ATALYZED AT 750'~

WE lGHT TO TO TO CHq CO Hz CO2 WE1 GHT RATIO,RATIO GAS SOLIDS LIQUID METHANOL TOSTEAM

CARBON CONVERS lONSMOLE FRACTIONS IN GAS WOOD

TOWOOD

FIGURE 44. Comparisons of Catalyzed versus Uncatalyzed Results for Synthesis Gas Production



The assumptions used i n the program are:

no l i q u i d phase presen t

the only gas species considered are H2, H20, CH4, CO, and C02

any s o l i d i s 100% carbon and has an a c t i v i t y o f 1a l l hydrogen and oxygen i n t he wood are con verted t o gas.

The method of sol ut io n i s an i t e r a t i v e Newton-Raphson solution of seven

equa tions and seven unknowns. The equa tions inc lu de an element balance f o r

carbon, hydrogen, and oxygen, (t hr ee equ ati ons ), th re e equ i 1ibrium equations,

and one equation expressing Dal ton 's law o f p a r t i a l pressures. The eq ui l i br iu m

equat ions are for the fol lowing heterogeneous react ions:

Equi l ibr ium constants for these react ions were obtained for a temperature

range of 7 0 0 ' ~ t o 1 2 0 0 ~ ~f rom Gumz (1950) and f i t t e d t o equat ions o f the form2

K = exp (A + B/T + C/T ), where K i s t he equ i l i b r ium cons tant , T i s absolu te

temperature, and A, B, and C are constants.

Figures 45 and 46 show the effects of temperatures on the dry gas equil ib -

r ium composit ion at atmospheric pressure for steam-to-wood ra t ios o f

0.33 and 1, resp ect ivel y. These r at i o s were suggested f rom ea r l y lab ora tory

experiments f o r methane produ ct ion and synthesis gas product ion, re sp ect iv el y.

Note the increase i n the carbon conversion t o gas wi th i ncr eas ing temperature.

Resu lt s from one PDU t e s t and la bo ra to ry - s c ale r eac t or t es t s a r e p l o t t ed w i t h

the 1 atm eq ui l i br iu m cal cu la t i ons f o r comparison. Po te nt ia l methane and metha-

nol y ie ld s are p l o t te d i n F igure 47. The e f f e c t o f temperature on equ i l i b r i um

standard heats of re ac t io n i s shown i n Fig ure 48. Figure 49 shows the ef f e c t

of pressure on the standard heats of react ion.

For the economic scale-up of b iomass gasi f icat ion for methane or synthesis

gas production, i t may be de sir abl e t o operate a t e levated pressures t o reduce

th e downstream compression cos ts. The eq u i l i b r i um program has been used t o

es timate t he e f f ec t s o f opera t ing a t 10 atm fo r use i n t he economic f e as i b l i t y

studies.



I ' CARBON

0. 8LABORATORY SCALE

/CONVERS I 0 1

O REACTOR DATA

1.0

0.9

PDU DATA

CARBON iCONVERSION r

-P = l a t m

/

- STEAM :WOOD = 0.333 :DRY BAS lS

CONVERS I0 V f -- -- 2

0

TEMPERATU RE (O C 1FIGURE 45. Effects of Temperature on the Dry Gas Equilibrium Composition

at Atmospheric Pressure for Steam -to-Wood Ratio of 0.33



TEMPERATURE (OC

FIGURE 46. Eff ect s o f Temperature on t he Dry Gas Eq ui l i br iu m Compositiona t Atmospheric Pressure f o r Steam-to-Wood Ratio of 1



400 500 600 700 800 900

TEMPERATURE, OC

FIGURE 47. P o t e n t i a1 Methane and Methanol Yieldsat Atmospheric Pressure

Figu re 50 shows th e eq ui l i br iu m dr y gas composit ion of a simulated SNG

produc t ion case. Methane produc t ion i s gre at ly favored by increased pressure.

Carbon conversion t o gas i s s l i g h t l y af fec ted.

F igure 51 shows the pressure effect of a simulated synthesis gas case. By

equi 1ibr ium calc ula t io ns, carbon convers ion i s 100% a t 750'~. Again, methane

product ion increases wi th pressure. Methane i s not desi rabl e above cer ta in

leve ls o f methanol synthes is. The f i n a l dec is ion o f the des i red gas i f i ca t io n

pressure w i l l have t o weigh the increased methane con ten t o f th e gas ag ain st

the reduced compression costs.

One other advantage o f pressure operat ion i s t ha t the o ve ral l reac t io n

becomes le ss endothermic and i n some cases becomes exothermic . Th is f a c t can

gr ea t l y reduce the complexi ty o f the ga si f i er . Less heat t ra nsf er area w i l l

be needed at higher pressures.



I I 1 I I I I450 500 550 600 650 700 750 800

TEMPERATURE OC

FIGURE 48. Effect of Temperature on Equi l ibriumStandard Heats o f Reaction



FIGURE 49. Ef fe c t o f P ressure on Equ i l i b r i umStandard Heats o f Reac tion



0.6 1 CARBON

/-H 4- C02

S T = 55O0c'0.3z STEAM :WOOD= 0.3333:1

o DRY B A S 1 S

0 10 20 30 40

PRESSURE (dm)

FIGURE 50. Ef fe ct o f Pressure on Product ion o f a Methane-Rich Gas



T = 750'~

STEAM: WOOD = 1:lAL L CA RBO N CONVERTED

DR Y B A S I S

PRESSURE (atrn)

FIGURE 51. E f f e c t o f Pressure on Produc t ion o f a Methanol Synthes is Gas



FEASIBILITY STUDIES

Information on process economics is needed to determine which options

offer significant cost advantages and should, therefore, be further investi-

gated. For this study, Davy McKee of Cleveland, Ohio evaluated the economics

of using catalytic steam gasification of wood in two commercial operations.

Two process plants were designed for the study: 1) for the conversion of wood

to methane and 2) for the conversion of wood to methanol. Results of the study

are sumnarized in the following discussion and detailed in Appendices A through

D of this report.

Since the catalytic processes are still in the development stage, process

conditions and outputs are constantly changing. Already, yields from catalytic

gasification are improved over those transmitted to Davy McKee for use in their

evaluations. Because these changes can significantly affect process economics,

a computer code is being developed to evaluate the effect of process modifica-

tions on overall economic feasibility.

ECONOMICS OF CATALYTIC GASIFICATION

Two different plants and plant capacities were considered for application

of catalytic gasificaton of wood: a wood-to-methane plant and a wood-to-

methanol plant, each at 2000 and 200 tons (1800 and 180 t) per day dry wood.

Design bases and descriptions of these plants are presented in the following

discussion. Capital and operating costs estimated by Davy McKee are summarized

for the plants. Product selling prices are presented for utility and private

investor financing. Details are presented in Appendices A through D of this

report.

Wood-to-Methane Plant

The design basis for the wood-to-methane plant was developed in the lab-oratory and PDU studies and adjusted to operation at 10 atm pressure. The main

adjustment was an increase in the methane concentration in the gases from the

gasifier. Overall gas yield from the gasifier used in the design is less than

more recent PDU yields as presented in earlier discussion.



The design basis for the plant i s summarized i n Table 31. A block diagram

for t he wood-to-methane p la nt i s shown i n F igure 52. Wood t o th e pl an t i s con-

sider ed t o be 60% fo re st residue, 20% fir, and 20% al de r wi th the composit ion

shown i n Table 32. The design basis f o r the smal l p la nt i s s imply scaled t othe capaci ty of 200 t o n s l d a y dry wood (180 t l d a y ) .

Char and catalyst (nickel on si l ica - alumina) are co l lec ted f rom the gas

by the gas cleaning system and combined with the overf low from the gasif ier.

The ra t i o o f char t o c a ta l ys t by weigh t i s 50: l . H a l f t he c a t a l y s t m a t e r i al

i s recovered i n the raw gas from the gasif ier and half comes with the overf low

f rom the gas i f i e r . The en t ra ined par t i c u la te has th e fo l l ow i ng s i ze

d i s t r i b u t i o n :

Size, p Weight %

+ I 49 10-149 + l o 5 5-105 +74 5-74 +53 5-53 +10 55

-10 20

F ines i n t he char- cata lys t mix ture are pr imar i ly char . Very l i t t l e cata-

l y s t i s i n the f i ne s th at are removed by screening. The char and cata ly s t i nthe coarse fr a c t i o n can then be separated by a magnetic r o l l separator. Cata-

l y s t recove ry i s 95% fo r the wood-to-methane plant.

The ca ta ly st i s regenerated by passing steam at 6 5 0 ' ~ over the cata lys t

f o r 20 h. The t o t a l s team requirement i s 20 t imes the weight o f cata ly st . The

ca ta ly st i s reduced by product gas f rom the ga si f i er . The consumption o f

hydrogen t o reduce the c at al ys t i s 6.4 X 1 b - m o l e l l b ( g - m o l e lg ) o f c a t a-

l y s t . A l l other operat ions i n th e schematics shown i n F igure 51 are based on

commercial ly avai lable technology.A

b r i e f d e s cr i pt i on o f t he p l a n ts i s p re-

sented i n the f o l lo win g d iscussion.

The wood-to-methane p l an t s were designed t o process 2000 t o n l d a y

(1800 t l d a y ) o f dr y wood and 10% o f t h i s capaci ty. P roduc t ion f o r t he la rge



TABLE 31. Wood-to-Methane Plant Design Basis--2000 t on /day Dry Wood

Wood-to-Methane

Capacity, Dry Tons Woodlday ( t /day) 2000 (1800)

Locat ionProduct Spec i f i ca t ions

Pac i f i c Nor thwestSNG:

HHV - 960 B t u / s c f3- (36 mJ/nm )

0.1 ~ 0 1 %o

H2S - 0.25 g/100 s c f

- (0.57 g/lOO nm3)

Water - 7 l b /106 s c f6 3- (110 kg/ lO nm )

Gas i f i e r Cond i t i ons

Temper a ture, OC 550

Pressure, atm 10

Steam Rate, w t / w t Dry Wood 0.33

Char production, w t / w t M A F ( ~ )Wood 0.24

Char Heating Value, B t u / l b ( k J / k g ) 13,500 (31,400)

Gas Production, w t / w t MAF wood 1.09

Cold Gas Eff iciency, % 65


38.0

11.2

CH423.6

C02 19.5CO 7.7

(a ) MAF i s mois tur e ash free .



WOOD DELIVERED

TO PLANT

FRESH

FLUE GAS CATALYST A

RECYCLECATALYST

0 .PRIMARY

METH AN ATION

WOOD RECEIVIN G WOOD CATA LYTI C-ANDSTORAGE DRYING GASIFICATION ?

I tAI R

EXCESS

CHAR CHAR CATALYSTRECOVERY ANDREGENERATION

FLUE GAS,-COMPRESSION

S H I F T

CONVERSION4

w

-

-STEAM

b 3

AI R b. WASTEWATER FINAL

• BLOWDOWN -. TREATMENT c METHANATI ON

AC lD GAS

REMOVAL

1 1 1RAW WATER RAW WATER

ASH

TREATMENT TO HEAT RECOVERY.TREATED SLUDGE

WATER

L I M E

I

1PRODUCT

GAS DRYINGSYSTEMS - AND FINAL

1COMPRESSION

COOLING PRODUCT GAS

TOWERILOSSES

FIGURE 52. Wood-To-Methane Process Areas



TABLE 32. Ult imate Analysis of Feedstock

W t % Dry

Carbon

Hydrogen

Nit rogen

Oxygen

Ash

w t % Moisture (as received)

Bt u / d r y 1b

F i rWeightedAverage

plant is 21.6 m i 11i on sc fd (610,000 nm3 /day) o f SNG w i t h a HHV of 956 Bt u / sc f

(35,600 k ~ / n r n ~ ) .A1 1 process and support faci 1it ies necessary to conver t

wood t o SNG are inc luded i n t he cost estimates.

The thermal ef f ic ie nc y o f the plants, as def ined by Equat ion ( 8 ) , i s

58.3%. When the heat ing value of th e excess char i s inc l uded i n the output ,

the thermal e f f i c i en cy i s 62.6%.

SNG, HHV% = looWood, HHV + E l e c t r i c i t y + Diesel Fuel

Wood storage, handling, and dry ing are major cap i ta l cos t out lays for the

p lan ts . A b r i e f desc r ip t ion o f t he systems f o r t he la rge p lan t f o l l ows .

The bulk o f the feedstock i s receiv ed alrea dy chipped. The chips are

de l i ve red by t ruck - t ra i ler r igs that are weighed on one of the two t ruck scales

when entering and leaving the unloading area. S ix un loading s ta t ions are

i n st a l l ed t o permit a maximum unloading r at e o f 1200 t o n / h (1090 t / h ) , whichrepresents 48 t r u c k s l h . The capac ity t o handle t h i s number o f tru ck s ensures

cont inuous unloading part icular ly when t rucks are making del iver ies f rom sev-

er a l logg ing s i t es and may a r r iv e i n groups. Each truck-unloading s ta t ion con-

s i s t s o f a hy dr au li c t ru ck dumper, a tr uc k dump hopper, and a cha in feeder.



Each chain feeder transports the chips onto a t r i p p e r / s t a c k e r be l t c on-

veyor or onto a reclaiming belt conveyor. The t r i p p e r / s t a c k e r conveyor deliv -

ers the chips to one of the two pr imary storage pi les v ia a double wing

stacker. Each p rimary p i l e i s l i m i t ed t o approximate ly 40f t

(12 m) hig h, assome bark and fines are supplied along with the chips. The stacker builds a

25-day ca pa ci ty storage; any enlargement of th e storage i s done by two b u l l -

dozers spreading the pi le s. Each o f the two primar y storage p i l e s i s 2800 ft

(850 m) lo ng and 80 f t (24 m) wide f o r t he 25-day capaci ty, wi th poss ibl e

enlargement to 240 f t (73 m) wi dt h f o r a 125-day capacity.

The reclai ming o f chips f rom the pr imary storage i s done wi th 14 chain

rec la imers (8 f t (2.4 m) wide) , seven f o r each p i l e . Each recl ai me r has a

capac i ty o f 700 t o n / h (630 t / h ) which represents the to ta l requi red rec la iming

rate based on 8 h/day, 5 days/week. The bull doz ers are used t o push chip s

towards reclaimers when needed. Two re cl ai min g be l t conveyors, one f o r each

p i l e , col lect chips f rom the respect ive chain reclaimers and del iver them to

the pr imary screening sta t io n. The pr imary screening sta t i on consists of

equipment for rock and tramp iron removal and for rechipping of oversize chips.

Screened chips are transported by a t r i p p e r / s t a c k e r conveyor t o t he

secondary storage pi le . Stacking and reclai ming of the chips are id en t i ca l t o

the method used f o r p rimary storage. Two piles, each 1800 f t (550 m) lo ng and

80 f t (24 m) wide for m a 14-day storage. Also a prov is io n i s made t o bypass

secondary storage by using one of the two reclaiming belt conveyors.

The ch ip s fro m secondary storag e are screened t o remove any in ci de nt al

oversi ze tra sh and conveyed by chain conveyors t o the surge bins f o r the

dr yers . Two conveying st rands and two secondary screens are used t o ensure

uninter rupted ch ip supply .

Si x ro t a r y drum dry ers complete wi t h a burner, ash removal cyclone,

exhaust dust cyclone, duct ing, and a l l necessary appurtenances are i n s t a l l e d t oreduce the moisture content of green chips from 50 w t % o f t o t a l fe ed t o 10 w t% .

The by-product char f rom the ga s i f i e r i s used t o f ue l t he burners f o r t he

dryers. F ive dryers normal ly operate whi le the s ix th dry er i s on standby.



Dried ch ips are conveyed f rom the dryers to a surge b in of one hour

capaci t y . The ch ip i nven to ry o f t h i s b i n a l l ows a s ta r t up o f t he s tandby

dryer. The ch ips f rom t h i s b i n are conveyed t o th ree ga s i f i e r lock hopper sys-

tems equally spaced at 120'. The ch ips are then screw fed cont inuous ly to the

base o f the g as i f i e r a t an ad jus tab l e r a t e f rom each lock hopper.

Th ree gas i f i e r s a re p rov i ded fo r t he l a rge p l an t . Each g a s i f i e r i n t h e

l a r g e p l a n t i s c y l i n d r i c a l w i t h d im en si ons o f 15 f t (4.3 m) ID by 45 f t

2 2(13.7 m) to t a l he ight . Each i s prov ided wi th 440 f t (41 m ) o f 2- in. (5-cm)

tubing (RA-533 or Incoloy 800 H) f o r h ea t a d d i t i o n t o t h e f l u i d bed o f wood

char and catalyst. The opera t ing depth o f the f l u i d bed i s 10 f t ( 3 m). A

s ing le scaled-down v er s io n o f t h i s g a s i f i e r i s p ro vi de d f o r t h e 200-ton/day

p l a n t .

The wood ch ips are ga s i f i ed i n the presence of c at a l ys t w i t h steam to pro-

duce a gas co nt ai ni ng methane, carbon di ox id e, hydrogen, carbon monoxide, and

water . Gas i f i ca t ion i s no t comple te and there i s a char res idue by-product.

The ga s i f i e r operates at c ondi t ion s o f 10 atm (150 ps ia ) and 5 5 0 ~ ~1 0 0 0 ~ ~ ) .

The steam-carbon rea c t i on i s h i gh ly endothermic, wh i le the methanation reac t ion

i s exothermic . The ne t ga s i f i ca t i on reac t ion s a re s l i g h t l y exo thermic; how-

ever, a hea t inpu t i s needed t o ass is t heatup o f reac tan ts . The heat i s sup-

pl ied by hot combustion gases (up to 1 0 0 0 ~ ~ )f lowing through a bank of tubes

imnersed i n the re ac t i on bed. Char res idue by-pr od uct i s t h e f u e l f o r t h e

g a s i f i e r heater . Af te r leav ing the gas i f i er , the combust ion gases pass through

a se ri es o f heat exchangers t o recover heat. The gases are used t o superheat

s team t o the g as i f i er oper at in g temperature, t o generate 600 ps ig s team, and

t o p reheat the combus tion a i r t o the char burner .

The remainder o f the ga s i f i ca t i on a rea i s devo ted to hea t recovery, ca ta-

l y s t recovery, rege nera t ion and recy cle, and char recovery. The raw gas from

the gas i f ier passes through a ser ies of exchangers that recover heat by gener-

ating superheated steam. These exchangers a re a ser ie s o f c o i l s i n a re f r ac -

t o r y- l i ne d s h e l l through which th e raw gas flows. These exchangers are, i n

the order t h a t th e raw gas sees them, th e steam superheater, the bo i le r , and

the b o i l e r feedwater preheater . The gas i s cooled t o 1 7 5 ' ~ ( 3 50 ' ~) i n t hes e

exchangers. The gas con ta ins pa r t i c u l a t es i n the fo rm o f en t ra ined char and



c a t a l y s t . The gas i s then cleaned by a cyclone fo l low ed by a bag f i l t e r con-

ta in ing a h igh tempera tu re g lass fabr ic . The gas then f l ows t o th e next

processing area. The catalyst recovery area consists of screens and a magnetic

separato r and i s complete ly enc losed t o p reven t a i r con tact w i th th e ca ta l ys t .

Commercially av ai la bl e systems comprise the remainder o f the pl an t. Acid

gas removal i s by the Be nf ie ld process.

A l l l i qu i d was tewater s treams f o r the p lan t a re t r ea ted i n a neu t ra l i z i ng

basin and a subsequent three-stage b io log ica l t rea tment system be fo re be ing

al lowed to over f low to dra inage.

The cap i ta l cost es t imate fo r the 2000-ton/day plant as determined by Davy

McKee, Inc., i s sumnarized i n Table 33. The c a pi t al cos t estima tes (September

1980 basis) for the 2000-ton ( 1800 - t ) / day d ry wood p l a n t i s $95,115,000. E s t i -

mate accuracy i s rep or ted t o be -+ 25%.

Breakdown o f the t o t a l d i r ec t cos ts by p lan t a rea i s g i ven i n Table 34 f o r

the 2000-ton dry wood/day pl an t. The major co st areas are wood stora ge and

TABLE 33. Wood-To-Methane Capital Cost Sumnary--2000 ton/day Dry Wood

EquipmentDirect Purchase Mater ia lSubcontract : Mater ia l

Labor (385,000 MHR)D i r e c t H i r e Labor (685,000 MHR)

SIT Direct CostsF i e l d I n d i r e c t sPro-ServicesOther

9/12/80 Tot al I n s t al l ed Cost (T.I.C.)

Exclusions: Proper tySta r t-up Costs

Cost ($1,000)33,32010,690

59010,370

P l a n t ~ o a d w a y sDemoli t ion o f Underground Obstructio nsPremium TimeOperating and Maintenance CostsContingency



TABLE 34. D i r e c t Cost Sumnary for Wood-to-Methane Plant--2000 t o n l d a y Dry Wood

Cost ($1,000)

Wood St orage and Han dl in g 17,704

Wood Drying 6,821

Ga si f i ca ti on Area 11,986

Compression 1,926

S h i f t Conversion 2,998

Pri mary Methan ation 2,192

Acid Gas Removal 2,143

' F i n a l Methana tion and Produc t Gas Dr yi ng 2,964

Ca ta ly st Regeneration 1,512

Waste Water Treatment 1,801Raw Water and Co ol in g Water Treat ment 2,483

B o i l e r s and BFW Systems 7,290

Miscel laneous Ut i l i t y Systems 1,928

TOTAL $ 63,748

handl ing, ga si f i ca t i on and wood drying. The ga si f i ca t i on area includ es th e

g a s i f i e r s, heat reco ver y equipment, char and c at al y s t rec over y equipment, and

wood and catalyst feed hoppers.

The production cost of methane from wood i s ca lc ul at ed based upon t he

cap i ta l costs and opera t ing costs . The methods of ca lcu lat ing these costs are

those presented i n "Coal Gasification Commercial Concepts Gas Costs Guide-

l ines," a paper prepared f o r th e U.S. Energy Research and Development Adminis-

tration and the American Gas Association by C. F. Braun & Co. (NTIS 8463).

There a re two po ten t ia l methods o f f inanc ing a p lan t o f th is type , 1) u t i l i t y

f i n a n c i n g, and 2 ) p r iv a t e invest o r f inanc ing . Product ion costs are ca lcu lated

using both procedures.

The to t a l pl a nt investment i s estimate d t o be $95,115,000 (September 1980

bas is ) . The to t a l c ap i ta l r equ ir emen t f o r the p lan t i s ob ta ined by add i t i on o f

an al lowance fo r funds dur ing construct ion, sta r t- up cost s, and workin g



capital. These costs and the basis for their calculation are shown in

Table 35. The total capital requirement for this plant is $115,191,000.

The annual direct operating costs for wood conversion to methane are shown

in Table 36. These costs include raw materials, utilities, catalysts and

chemicals, labor, administration and general overhead, supplies, and taxes and

insurance, with a credit for by-product char. Total maintenance costs are cal-

culated as a percentage of plant investment as suggested by the guidelines. The

annual costs are $29,990,000. The most significant costs are wood, gasifier

catalysts, labor, and taxes and insurance. Labor costs are not very easy to

reduce significantly, while taxes depend upon local conditions and incentives.

The major variable costs are wood and catalyst usage in the gasifier. Studies

to improve catalyst life are in progress. At $20/dry ton for wood, which isthe value used for the base case shown in Table 36, wood costs are almost 45%

of the total direct costs and almost one-third of the total production costs

with utility financing. Thus, either lowering the wood cost or improving

yields from the wood have more impact on costs than any other single variable.

The production costs have also been calculated for wood costs of $5, $10, and

$40 per dry ton delivered to the plant, and the impact is i 1lustrated inFigure 53.

TABLE 35. Total Capital Cost Requirement For Wood-to-Methane Plant--2000 ton/day Dry Wood

Total Plant Investment

Allowance for Funds During Construction(Total Plant Investment x 1.25 yr x 0.09)

Start-up Costs (20% of Total Annual GrossOperating Costs)

Working Capital [Sum of (1) raw materialinventory of 14 days at full rate,(2) materials and supplies at 0.9% of totalplant investment, and (3) net receivablesat 1/24 annual gas and by-product revenueat calculated sales price]

Cost ($1,000)95,115

Total Capital Requirement



TABLE 36. Annual Direct Operating Costs For Wood-to-MethanePlants--2000 tonlday Dry Wood

Operating Factor: 330 dayslyr

Cost Component Annual Use

Raw MaterialWood

UtilitiesWaterElectricityDiesel Fuel

660,000 dry tons

332,640 M a193.25 x 10 kwh

108,900 gal

Catalysts and ChemicalsChemicalsShift Catalyst 1500lf 3Methanation Catalyst 1020/ft3Gasifier Catalyst 380,160 1b

LaborE s s OperatingMaintenance

Supervision

51 men @ 2080 h

@ 60% of totalmaintenance@ 20% of processoperating andmaintenance labor

Administration and General Overhead @60%

of total laborSupp iesOperat ng

Maintenance

Taxes and Insurance

Total Gross Operating Cost Per Year

By-Produc t Cred its

Char

Total Net Operating Cost Per Year

@ 30% of processoperating 1abor@ 40% of totalmaintenance cost

@ 2.7% of total plantinvestment

28,050 tons

Cost$/Unit $1000/yr

201dry ton 13,200



-

-

-

-

-

-

-

-

-

I

-

-

- -UTILITY F INANCING

A -PRIVATE INVESTOR FINA NCIN G-

I I I I I I I I I I I I I

WOOD, $1DRY TON

FIGURE 53. Effect of Wood Prices on Gas Cost for 2000 ton/day Dry Wood



Tables 37 and 38 g ive t he methods fo r ca lc u l a t in g p roduct ion costs based

on u t i 1i t y f i na nc i ng and p r i v a t e i nves t o r f i nanc ing , r espec t i ve l y . The cal cu-

l a t i o n s f o r t he base case of a wood cost of $ 2 0 / d r y ton are shown. For u t i l i t y

f inanc ing , the SNG pr od uc t io n co st s ar e $5.09, $5.56, $6.50, and $8.34 per m i l -

l i o n Bt u fo r wood pr ic es of $5, $10, $20, and $40 per d r y ton. For p r i v a t e

i n v e s t o r fi n a n ci n g , t h e SNG pr od uc t i on cos ts ar e $6.62, $7.11, $8.10 and $10.06

per m i 11i o n Btu f o r the cor responding wood costs .

The same design bas is was used t o determine ope rat in g cost s f o r a 200-dry-

t o n wood/day (180 t l d a y ) p l a n t f o r p r o d u c t i o n o f SNG. A1 1 y ie lds a re 10% o f

t he y ie lds f r om the l a r ge r p lan t . The on l y d i f f e ren ce i n th e sys tem was

on- si te ch ipp ing of the wood residue. De ta i l s a re p resen ted i n Appendix C and

sumnar ized below.

The to ta l i nves tmen t fo r the 200-dry- ton lday (180- t iday) p lan t was est i -

mated t o be $26,680.000 (September 1980 ba si s) . T o o b t a i n t h e t o t a l c a p i t a l

requ i rement f o r t he p lan t , add i t ion a l costs must be added t o the est imated

p lan t investment . These costs are an a1 1owance for funds dur ing construct ion,

s t a r t- up costs , and work ing ca p i ta l . These costs and the bas is fo r th e i r ca l -

cu l a t i on are shown i n Tab le 39. The to t a l c ap i ta l r equ i remen t f o r th i s p lan t

i s $31,805,000.

The annua l d i re c t opera t ing costs f o r t he smal l p la n t were ca l cu la ted and

a re shown i n Table 40. These cos ts i nc lude raw ma te r ia l s , u t i l i t i e s , ca t a l ys ts

and chemicals, labor, ad mi ni st ra ti on and general overhead, supp l ie s, and taxes

and insu rance, w i th a c r ed i t f o r by-product char. To ta l maintenance cos ts were

ca lc u la ted as percentage of p l an t investment as suggested by th e guide l ines.

These annual co st s are $6,833,600. The most s i g n i f i c a n t co st s are wood, gas i-

f i e r ca ta l ys t , labor , and taxes and insurance. Labor costs would not be ver y

easy t o re du ce s i g n i f i c a n t l y , w h i l e ta xe s w i l l depend upon local condit ions

and inc ent ives . The major v ar ia b l e costs are wood and ca ta ly st usage i n th e

g a s i f i e r . A t $20 / d r y to n f o r wood, wh ich i s the va lue used f o r t he base case

shown i n Table 40, wood costs are a lmost 20% of th e t o t a l d i r e c t costs and

a lm os t 15% o f t h e t o t a l p r o du c t i o n c os t s u s i n g u t i l i t y fi n a n ci n g . Thus, e i ther



TABLE 37 . Methane Cost-Utility Financing Method--2000 ton/day Dry Wood

Bases:

20-yr project life

5% year straight line depreciation on total capita1 requirementexcluding work capital

48% federal income tax rate75/25 debtlequity ratio10% interest on debt15% return on equity

Def nit on of Terms:

C = Total capital requirement, lo6$W = Working capital, 106$N = Total net operating cost, lo6 $/yrG = Annual gas production, 1012 Btulyrd = Fraction debti = Interest on debt, %/yrr = Return on equity, %/yrp = Return on rate base, %/yr

Equation for Return on 2ate Base:

p = (d) i + (1-d) r

General Gas Cost Equation:

Average gas cost, $/million Btu =

N + 0.05 (C-W) + 0.005 (p + 48/52 (1-d) r) (C + W)G

Calculation:

Average gas cost, $/mi llion Btu =



TABLE 39. To ta l Ca pi ta l Requirement f o r 200-ton/dayWood-to-Methane Plant

Total Plant InvestmentA1 1owance for funds During Construction(Total Plant Investment x 1.25 years x 0.09)

Star t-up Costs (20% o f T ot al Annual Grossoperat ing Costs)

Cost ($1,000)26,680

Working Capital (Sum of (1) raw materiali n v e nt o r y o f 1 4 days a t f u l l r a te , ( 2 ) m a t e r i a l sand suppl ies at 0.9% of total p lant investmentand (3) ne t rec ei va bl es a t 1/24 annual gas and

56 (1)240 (2)

by-products revenue at calcu 1ated sales pr ic e) 452 (3)

Total Capital Requirement $ 31,805

lowering the wood cost or improving yields from the wood would have more impact

on costs than any other single variable. The production costs have also been

cal cul at ed f o r wood costs o f $5, $10, and $40 per dr y ton deli ver ed t o the

p lan t .

I n Tables 41 and 42 the methods for ca lc ul at in g produc t ion co sts are giv en

based upon u t i l i t y f inanc ing and pr iv at e inves tor f inanc ing, respect ive ly . The

calculat ions for the base case of a wood cost of $20/dry ton are shown. Foru t i l i t y fi na nc in g, t he SNG product ion co st s ar e $14.34, $14.83, $15.86, and

$17.84 per m i l l i o n Btu for wood pr ices o f $5, $10, $20, and $40 pe r dry ton.

For p r i va t e inves to r fi nancing, t he SNG pr od uc ti on co st s ar e $18.76, $19.26,

$20.28, and $22.31 per m i l l i o n Btu f o r the corresponding wood costs.

Wood-to-Methanol Plant

The design basis for the wood-to-methanol p l a n t was developed i n the la b-

oratory and PDU stud ies and adjusted t o opera t ion a t 10 atm pressure. The main

adjustment was an increase i n the methane conc entrat i on i n the gases from thegas i f i e r . Overal l gas y i e ld used i n t he des ign bas is i s l ess t han recen t oper-

a t i n g r e s u l t s from the PDU as presented above.

The design basis for the 2000- ton (1800- t ) dry wood/day p la n t i s shown i n

Table 43. A schematic f l o w diagram showing process areas i s giv en i n



TABLE 40. Annual Direct Operating Costs for 200-ton/day Wood-to-Methane Plant

Operating Factor: 330 days/yr

Cost Component

Raw Mater a1Wood


Catalysts and ChemicalsChemicalsShift Catalyst

Methanati on CatalystGasifier Catalyst

LaborProcess OperatingMai ntenance

Supervision

Administration and General Overhead

Supp iesOperating

Maintenance

Taxes and Insurance

Total Gross Operating Cost per Year

By-Product CreditsChar

Cost

Annual Use $/Unit $1000/yr

66,000 dry tons 20/dry ton 1,320.0

33,264 Mgal O.SO/Mgal 16.67.76 x 106 kwh 0.03/kWh 232.810,890 gal l.OO/gal 10.9

44 men @ 2080 h 10.70/h 979.3@ 60% of totalmaintenance 760.8@ 20% of processoperati ng andmaintenance labor 348.0

@ 60% of total labor 1252.9

@ 30% of processoperating labor@ 40% of totalmaintenance

@ 2.7% of totalinvestment

2805 tons 15.35/ton (43.1)

Total Net Operating Cost per Year



TABLE 41. Gas Cost - U t i l i t y F inanc ing Method--200 t o n l d a y Dry Wood

Bases:

20-

y r p r o j e c t l i f e5% year s t ra ig h t l i n e deprec i t a t i on on t o ta l c ap i t a l r equi rementexc lud ing work ing cap i ta l

48% fed er al income tax r a t e75/25 d e b t l e q u i t y r a t i o10% i n te re s t on debt15% re tu rn on e qu i t y

Defin i tion o f Terms:

C = Total Capital Requirement, l o 6 $W = Working Capital, l o 6 $N = Total Net Operating Cost, l o 6 $ / y r

G = Annual Gas Production, 1012 B tu /y rd = Frac t io n debti = Interest on debt , % / y rr = Return on equity, % / y rp = Return on rate basis, %/yr

Equation f o r Return on Rate Base:

p = (d ) i + ( 1 -d ) r

General Gas Cost Equation:

Average Gas Cost, $/m i l 1ion B tu =

N + 0.05 (C-W) + 0.005 (p + 48/52 (1-d) r ) (C + W)G

Calcu la t ion:

Average gas cost, $/mil l ion Btu =



TABLE 42. Gas Cost - Equity Financing Method--200 t on /day Dry Y a o d

Bases:

20-yr p r oj e ct f i l e16-yr sum-of- the-years' - d i g i t s d e p r ec i at i on on t o t a l p l a n t

investment100% equ i t y cap i t a l

12% DCF r e t u r n r a t e48% fe der al income tax r a t e

Def in i t ion o f Terms:

I = Total p lant investment, l o 6 $s = Star t-up costs, 106 $W = Working Capital, l o 6 $N = Tota l ne t opera t ing cos t , 106 $/yr

G= Annual gas production, 1 0 1 2 B t u l y r

Gas Cost Equat io n a t 12% DCF Return:

Gas cost, $ / l o 6 B t u =

Calcu la t ion :

Gas cost, $/lo6 B tu =



TABLE 43. Wood-to-Methanol Plant Design Basis--2000 t o n l d a y Dry Wood

Wood-to-Methanol

Capacity, Dry Tons Woodlday ( t lday) 2000 (1800)

Loca tion Pa c i f i c Nor thwest

Produc t Spec i f i ca t ions Methanol: 99.3% purity

Gas i f ie r Cond i t ions

Temperature, OC 750

Pressure, atm 10

Steam Rate, w t l w t Dry Wood 0.75

Char Production, w t l w t MAF Wood 0.15

Char Heating Value, B t u l l b ( k J / k g ) 14,000 (33,000)

Gas Production, w t / w t MAF wood 1.60

Cold Gas Eff ic iency, % 88


CH4

C02co

Figure 54. The des ign bas is fo r the smal l p la n t i s s imp ly sca led t o a

capaci ty of 200 t o n l d a y dr y wood (180 t l d a y ) . Wood fee dst ock i s t h e same as

presented i n Table 32.

The wood-to-methanol pla nt s are designed t o process 2000 tons (1800 t)

per day o f d ry wood and 10% of t h i s capaci ty . Produc t ion f o r t he la rge p l an t



FLUE GAS

WOOD DELIVERED

TO PLANT

LOSSES

FIGURE 54. Wood-To-Methanol Process Area

-WOOD RECEI VING

AND STORAGE_ A C I D G A S

REMOVAL

WOOD

DRYING

ICON DENSATE

v

a A A

CHAR TO DRYER

FRESHA I R T O D R Y E R CATALYST

CATALYTIC

GASIFICATION

STEAM

CATALYST

II CHAR AND

CATALYST CATALYSTRECOVERY

7

-CHAR AND 4v

CO SHIFT

-

WOOD

COMPRESSION

PURGE GAS

REFORMING

CHAR -BOILER

4

TO HEAT RECOVERY

9

METHANOL

SYNTHES lMETHANOL

PRODUCT

SYSTEMSv

RAW WATER

WASTE

WATER

TREATMENT-RAW WATER

L I M E 1PRODUCT

S H I P P I N GTREATMENT

a

bSTORAGE AND

SLUDGE 1TREATED WATER

COOL ING .WASTE

TOWER WATER

1METHANOL

PRODUCT



i s 997 tons (900 t ) per day o f methanol w i t h a HHV o f 9784 Btu per pound

(22,700 kJ / kg ) . A l l process and support f a c i l i t i e s necessary t o convert wood

t o methanol are inc luded i n the est imate.

The the rmal e f f i c i en cy o f t he p la n t i s de f ined as

Methanol, HHV* = looWood, HHV + E l e c t r i c i t y + Diesel Fuel

The thermal e f f ic ie nc y i s 53%.

Wood storage, handl ing, and dr yi ng are i d e n ti c al t o th e systems described

prev ious ly f o r t he wood-to-methane pl an t, The small pl an t inc lud es on - s i t e

ch ipp ing of the res idue.

Three gasi f iers are provided for the 2000 ton (1800 t ) / d a y p lan t . Each

gas i f i e r i s c y l i nd r i c a l w i t h dimensions o f 15 f t (4.3 m) I D by 50 f t (15 m)2 2

t o t a l height . Each i s provided wi th 3900 f t (360 m ) o f 2- in. (5-cm) tubing

(RA-533 or Inco loy 800 H) f o r heat a ddi t i on t o the f l u i d bed of wood char and

c a t a l y s t . The operat ing depth o f the f l u i d bed i s 10 f t (3 m). A s ing le ,

scaled-down vers ion o f t h e i r des ign i s p rovided f o r t he 200 t o n l d a y p lan t .

F ines i n t he char- c a t a ly s t m i xt u re ar e p r im a r i l y char. Very l i t t l e c at a-

l y s t i s i n the f i ne s which are removed by screening. The char and ca ta ly st i nthe coarse f ra ct io n can then be separated by a magnet ic r o l l separator . Cata-

l y s t r ecove ry i s 90% f o r t he wood-to-methanol plant. No regene rat ion o f cat a-

l y s t i s considered.

A l l other operat ions i n the schematics shown i n Figure 54 are based on

commercial ly avai lable technology. A b r i e f d e s c ri p t i on o f t he p l a n ts i s p re-

sented i n the fo l l owi ng d iscussion.

Wood chips f rom the g as i f ie r feed bin s are cont inuou sly fed by screw con-

veyors t o t he f l u id i zed- bed ga s i f i e r . H igh pressure steam i s fed i n t o each bedt o gas i f y the wood as wel l as f l u i d i z e the so l ids . Contained wi th in the bed i s

the ca ta ly s t requ i red f o r t he reac tion , a 318- in. (9.5-mm) diameter spherical

ba l l o f n i c k e l c a t a l y s t on a s i l i c a - a lumina s t ruc tu re . The gas i f i e rs i n t he



wood-to-methanol plants operate at a temperature of 7 5 0 ' ~ ( 1 3 8 0 ' ~ ) and a pres-

sure of 10 atm (150 psi a). A t these condit ions most hydrocarbons are cracked

t o g ive a crude synthesis gas co ns is t i ng pr i ma r i l y o f hydrogen and carbon

oxides with only a small amount of methane.

About 285 tons per day ( dr y ba si s) (259 t / d a y ) ( o r 10% o f t h i s f o r t h e

smal l p l an t ) o f t he wood supp l ied to th e gas i f i c a t io n a rea a re burned t o supply

the necessary heat inpu t t o the ga s i f i e r . Th is i s accomplished by f i r s t a l low-

i ng the ho t combust i on gases--approximately 9 8 0 ' ~ ( 1 8 0 0 ~ F ) - - to c i r c u l a t e

through tube bundles located i n each g as i f i er bed. These gases upon exiting

the tube bundles are then used t o preheat the g as i f i er f l u i d i z i n g steam and

combustion a i r f o r th e wood burners.

The remainder o f the g as i f i c a t io n a rea i n the p lan t i s devo ted t o hea t

recovery, c at a ly s t reco ver y and recy cle , and char recove ry. The raw gas fr om

the gasi f ier passes through a ser ies of exchangers that recover heat by gener-

ating superheated steam. These exchangers a re a ser ies o f c o i l s i n a re f ra c -

t o r y l in e d sh el l through which the raw gas f lows. These exchangers are, i n the

orde r t h a t th e raw gas sees them, t he steam superh eater, t he CO s h i f t p re-

heater, the boi l er , and the bo i l er feedwater preheater . The gas i s cooled t o

1 7 5 ' ~ ( 3 50 ' ~) i n these exchangers. The gas con ta ins par t i cu la te i n the fo rm o f

ent rain ed char and ca ta ly st . The gas i s then cleaned by a cyclone fol l owe d by

a bag f i l t e r co n ta in ing a h igh temperatu re g lass f abr ic . The gas then fl ow s t o

the next processing area. The cata lyst recovery area consists o f screens and

a magnetic separator and i s complete ly enclosed t o prevent a i r conta ct wi th the

c a t a l y s t .

Commercial ly avai lable systems comprise the remainder of the plant. Acid

gas removal i s by th e Be nf ie ld process. The I C I low pressure methanol process

i s used f o r the methanol synthesis. A l l l i q u i d wastewater streams fo r the

p lan t a re t r ea ted i n a neutra l iz ing basin and a subsequent three stage b io log i -

ca l t reatment system before being a l lowed t o over f l ow t o drainage.

The capital cost estimate as determined by Davy McKee, Inc., i s summarized

i n Table 44 fo r the 2000- ton dry wood/day p lan t . The cap i ta l cost es t imate

(September 1980 basis) for the 2000- ton dry wood ( 1800 - t ) / day p l an t i s

$120,830.000. Esti mate accuracy i s rep ort ed t o be + 25%.



TABLE 44. Wood-to-Methanol Capital Cost Sumnary--2000 ton/day Dry Wood

Equipment

Direct Purchase M a t r i a lSubcontract: Material

Labor (403,000 MHRS)D i r e c t H i r e Labor (665,000 MHRS)

SIT Di rec t CostsF i e l d I n d i re c t sPro-ServicesOtherMETHANOL PLANT (By Lake1and-T.I.C. )

9/12/80 Tota l Ins ta l led Cost (T.I C.)

Cost ($1,000)

35,390

10,840650

11,6008,530

67,01014,15015,920

1,750

Exclusions: Property

Star t-

up CostsP lan t ~ o a d w a y sDemolit ion of Underground Obstruct ionsPremium TimeOperating and Maintenance CostsContingency

Breakdown of the total d i re c t cost s by p lan t a rea i s g i ven i n Table 45 f o r

the 2000-ton ( 1 8 0 0 - t ) d ry wood/day p la n t . The major cost areas are wood stor -

age and hand1ing, g a si f i c at i on and wood drying . The gas i f ica t ion area inc ludes

the g a si f i e rs heat recovery equipment, char and cat al ys t recovery equipment,

and wood and catalyst feed hoppers.

The product ion cost of fuel-gra de methanol fr om wood i s c al cu la te d based

upon the capi ta l costs and operat ing costs. The methods of calculat ing these

costs are those presented i n "Coal Gasification Commercial Concepts Gas Cost

Guidelines" , a paper prepared f o r the Uni ted Sta tes Energy Research and Devel-

opment Administration and the American Gas Association by C. F. Braun & Co.

(NTIS 8463). Product ion cost s are cal cul at ed usi ng two po te nt ia l methods o f

f i nancing a p lan t o f t h i s t ype, ( 1 ) u t i l i t y f i na n ci n g, and ( 2 ) p r i v a t e i n v es t o r

f inanc ing.

The total investment for the 2000- ton ( 1 8 0 0 - t ) / d a y dry wood p lan t i s e s t i -

mated t o be $120,830,000 (September 1980 bas is ) . The to ta l cap i ta l requ i rement



TABLE 45. Direct Cost Sumnary For Wood-to-Methanol Plant--2000 tonlday Dry Wood

Wood Storage and HandlingWood DryingGasification AreaShift ConversionAcid Gas RemovalCompress onMethanol Synthesis(T. I.C. by Davy Powergas-Lakeland, FL)Purge Gas ReformingWaste Water TreatmentRaw Water and Cooling Water TreatmentBoilers and BFW SystemsMiscellaneous Utility SytemsStorage and Loadi ng-Product and Ut 1it es

TOTAL

Cost ($1,000)

17,7046,82117,889

8504,5395,412

for the plant is obtained by addition of an allowance for funds during con-

struction, start-up costs, and working capital. These costs and the basis for

their calculation are shown in Table 46. The total capital requirement for

this plant is $145,571,000.

The annual direct operating costs for the 2000-ton/day dry wood plant have

been calculated and are shown in Table 47. These costs include raw materials,

utilities, catalyst and chemicals, labor, administration and general overhead,

supplies, and taxes and insurance. Maintenance costs are calculated as a per-

centage of capital investment, as suggested by the cited guidelines and are

$36,464,000. The most significant costs are wood, gasifier catalyst, labor,

and taxes and insurance. Labor costs are not very easy to reduce signifi-

cantly, while taxes will depend upon local conditions and incentives. The

major variable costs are wood and catalyst usage in the gasifier. Studies on

improvement of catalyst life are in progress. At $20/dry ton for wood, whichis the value used for the base case shown in Table 47, wood costs are about

one-third of the total direct costs and one-fourth of the total production

costs. Thus, either lowering the wood cost or improving yields from the wood

have more impact on costs than any other single variable. The production costs



TABLE 46. Total Capital Required For Wood-to-Methanol Plant--2000 tonlday Dry Wood

Total Plant Investment

Allowance for Funds During Construction(Total Plant Investment x 1.25 yr x 0.09)

Start-up Costs (20% of Total Annual GrossOperating Costs)

Working Capital (Sum of (1) raw materialinventory of 14 days at full rate,(2) materials and supplies at 0.9% of totalplant investment, and (3) net receivables

at 1/24 annual methanol revenue atcalculated sales price)

Cost ($1,000)

120,830

Total Capital Requirement

have also been calculated for wood costs of $5, $10, and $40 per dry ton, and

the impact is illustrated in Figure 55. These prices for wood include delivery

to the plant.

In Tables 48 and 49 methods are given for calculating production costs

based upon utility financing and private investor financing, respectively. The

calculations for the base case of a wood cost of $20/dry ton are shown. For

utility financing, the methanol production costs are $0.45, $0.48, $0.55, and

$0.69 per gallon for wood prices of $5, $10, $20, and $40 per dry ton. For

private investor financing, the methanol production costs are $0.59, $0.62,

$0.69, and $0.83 per gallon for the corresponding wood costs.

The same design basis was used to determine operating costs for a 200-ton

dry wood (180-t) per day plant for production of fuel grade methanol. All

yields are 10% of the yields from the larger plant. Details are presented in

Appendix D and sumnarized below.

The total plant investment was estimated to be $34,830,000 September 1980

basis. To obtain the total capital requirement for the plant, additional costs

must be added to the estimated plant investment. These costs are an a1lowance



TABLE 47. Annual Direct Operating Costs For Wood-to-Methanol Plant--2000 ton dry wood/day

Operating factor: 330 dayslyr

CostCost Component Annual Use $/Unit $1000/yr

Raw Mater a1Wood


Catalysts and ChemicalsGasifier Catalyst

Shift catalystvChloride Guard CatalystSulfur Guard CatalystMethanol CatalystReformer Catalyst

LaborProcess OperatingMaintenance

Supervision

660,000 dry tons 20/dry ton 13,200

526,522 M a1I 0.50lMgal 2631.76 x 10 Kwh 0.03/kWh 5,280108,900 gal l.OO/gal 109

330,660 I b 8.5111 b 2,814

150 ft3 107/ft3 163,300 ft3 151/fj3 4982,500 ft3 75/ft 188Confidential -- 360140 ft3 235/ft3 33

55 men @ 2080 h 10.70/h 1,2248 60% of total 3,160maintenance@ 20% of processoperating andmaintenance labor 877

Administration and General Overhead @ 60% of total labor

Supp iesOperating

Maintenance

Taxes and Insurance

@ 30% of processoperating 1abor@ 40% of totalmaintenance cost

@ 2.77% of totalplant investment

Total Operating Cost Per Year $ 36,912



-

-

-

-

-

-

-

-

-

-

-

-

-

--UTILITY FINANCING

A -PRIVATE INVESTOR FINANCING

I I I I I I I I I I I I I

WOOD $ /DRY TON

FIGURE 55. Effect of Wood Prices on Methanol Cost for 2000 ton lday Dry Wood



TABLE 48. Methanol C o s t - Ut i l i t y Financing Method--2000 t on /day Dry Wood

Bases:

20-y r p r o j e c t l i f e5%/y r

s t ra igh t l i ne deprec ia t ion on t o ta l cap i t a l r equ i rementexc luding working cap i ta l48% fe de ra l income t ax r a t e75/25 deb t l e q u i t y r a t i o10% i n te re s t on debt15% re tu rn on eq ui ty

Def in i t ion o f Terms:

C = Total capi ta l requirement , 106 $W = Working capital, 106 $N = Total net operat ing cost , l o 6 $ / y r

G = Annual fuel product ion, l o 6 g a l l y r

d = Frac t ion debti = Interest on debt, %/ y rr = Return on equity, % / y rp = Return on rate base, %/ y r

Equation f o r Return on Rate Base

p = (d) i + ( 1 -d ) r

Average methanol cos t, $/g al =

N + 0.05 (C-W) + 0.005 (p + 48/52 (1-d) r ) (C+W)G

Calculat ion:

Average methanol cost =



TABLE 49. Methanol Cost-Equity Financing Method--2000 t on /day Dry Wood

Bases:

20-y r p r o j e c t l i f e16 - y r sum

-of-the years' - d i g i t s d e pre ci a ti o n on t o t a l p l a n t

investment100% equ i t y capi t a l12% DCF r e t u r n r a t e48% fe de ra l income tax r a t e

De f in i t i o n o f Terms:

I = Tota l p lant investment, l o 6 $S = Start-up costs, 106 $W = Working capital , l o 6 $N

= Tota l net operat ing cost ,l o 6

t / y rG = Annual methanol production, 10 g a l / y r

Methanol Cost Equati on a t 12% DCF Re tu rn

Methanol cost, $/gal =

Calcu la t ion :




for funds during construction, start-up costs, and working capital. These

costs and the basis for their calculation are shown in Table 50. The total

capital requirement for this plant is $41,221,000.

TABLE 50. Total Capital Requirement for 200-ton Dry Wood/Day Methanol Plant

Total Plant InvestmentCost ($1,000)

34.830Allowance for Funds During Construction 3;918(Total Plant Investment x 1.25 yr x 0.09)

Start-up Costs (20% of Total Annual Gross 1,559Operating Costs)

Working Capital (Sum of (1) Raw MaterialInventory of 14 days at full rate, 56 (1)

(2) Materials and Supplies at 0.9% of 313 (2)Total Plant Investment, and (3) Net 545 (3)Receivables at 1/24 annual methanolrevenue at calculated sales price)

Total Capital Requirement 41,221

The annual direct operating costs were calculated and are shown in

Table 51. These costs include raw materials, utilities, catalyst and chemi-

cals, labor, administration and general overhead, supplies, and taxes and

insurance. Maintenance costs are calculated as a percentage of capital invest-

ment, as suggested by the cited guidelines. These annual costs are $7,794,800.

The most significant costs are wood, gasifier catalyst. labor, and taxes and

insurance. Labor costs would not be very easy to reduce significantly, while

taxes will depend upon local conditions and incentives. The major variable

costs are wood and catalyst usage in the gasifier. At $20/dry ton for wood,

which is the value used for the base case shown in Table 51, wood costs are

about 17% of the total direct costs and are 10% of the total production costs.

Thus, either lowering the wood cost or improving yields from the wood have more

impact on costs than any other single variable. The production costs have also

been calculated for wood costs of $5, $10, and $40 per dry ton. These prices

for wood include delivery to the plant.



TABLE 51. Annual D i r e c t ~ p e r a t ' i n gCosts for 200- ton Dry Wood/Day Methanol Plant

Operating Factor : 330 d a y s l y r

CostCost Component Annual Use $ /Un i t $1000/yr

Raw Materi 1Wood

U t i l i t i e sWaterE l e c t r i c i t yDiesel Fuel

Catalysts and ChemicalsG a s i f i e r Ca t a l ys t

S h i f t c a t a l y s t vChlor ide Guard CatalystSul fur Guard Cata lystMethanol CatalystReformer Catalyst

LaborProcess OperatingMaintenance

Supervis ion

66,000 Dry Ton 20/Dry Ton 1,320.0

52,652 Mgal 0.50/Mga 1 26.32.43 x 107 kwh 0.03/kwh 729.610,890 gal l .OO/gal 10.9

33,066 1b 8.51/7b 281.4

15 f t 3 1 0 7 /f t 3 1.6330 f t 3 1 5 1 / f t 3 50.0250 f t 3 7 5 / f t 3 18.8Conf identi 1 ------ 36.015 f t 3 2 3 5 /f t 3 3.5

48 Men @ 2080 h 10.70/h 1,068.3@ 60% o f To ta l 873.6Maintenance@ 20% o f proc ess 213.7operating andmaintenance labor

Administrative and General Overhead 8 60% o f To ta l Labor 1,293.4

Suppl iesOperat ing

Maintenance

Taxes and Insurance

@ 30% o f ProcessOperatin g Labor8 40% of TotalMaintenance Cost

@ 2.77% o f T ot alPlant Investment

Total Gross Operating Costs Per Year ( a > 7,794.8

(a ) No cr ed it s were taken f o r any by-product so that Total Net Operating Costsare the same as the Total Gross.



Tables 52 and 53 gi ve the methods f o r c al cu la t i ng prod uct ion cos ts based

upon u t i l i t y f i nanc ing and p r i va te inves to r f i nancing , respec t i ve ly . The c a l -

c u l a t i ons for the base case of a wood cost of $20/dry t on are shown. For

u t i 1i t y f in anc ing , th e methanol pr od uc ti on co sts are $1.20, $1.23, $1.30 and

$1.44 per gall on f o r wood p r i c e s of $5, $10, $20, and $40 per dr y ton. For

p ri v at e in ve sto r f inanc ing , the methanol pro duc t io n cost s are $1.60, $1.63,

$1.70 and $1.84 per g al lo n f o r t he corresp onding wood cos ts.

COMPUTER MODELLING STUDIES

Computer mode ll in g work i s underway t o develop a program t h a t w i l l

describe commercial plants for producing methane and methanol by catalyt ic

g a si f i ca ti o n o f biomass. The model w i l l do heat and material balance calcula-

t i o n s and then economic eva lua tio ns. The e f f e c t s on o v er a l l economics o f

changing process parameters should be easily determined.

The hea rt o f th e modell ing e f f o r t i s t he ASPEN computer code being devel-

oped a t Massachusetts In s t i t u t e of Technology. ASPEN i s an acronym f o r

Advanced Simulat ion for Process Engineering analysis. One of the primary

obj ect iv es o f the ASPEN pr oj ec t i s t o develop a proto typ e simu lator and demon-

s t r a t e t h e a b i l i t y t o s im ulat e s p e c i f i c f o s s i l f u e l c on ve rs io n processes o f

i n te r e st t o DOE.

Th is s imu la tor i s f l ex ib le and w i l l model a var ie ty o f un i t operat ions .

It performs heat and material balance calculat ions, does crude equipment sizing

calcu 1a t i ons and computes several economic parameters. It i s considered a

s ta te- o f- the- ar t s imula t ion code and i s s t i l l be ing rev ised and tested. It i s

the only known publ ic ly avai lable code that w i l l handle st reams with sol ids or

model solids handling equipment.

Several chemical reactor models are avai 1abl e i n ASPEN, bu t a11 are idea l

de sc ri pt io ns such as an eq ui li br i um model, a pl ug fl o w model and a backmix

model. None o f these adequately descri bes th e batch rea ct or used i n our la b-

ora to ry s tud ies o r t he f l u id - bed r eac tor i n the PDU. Fo rt un at el y th e ASPEN

code i s designed t o al lo w t he user t o sup ply h i s own models when necessary.



TABLE 52. Methanol Cost - U t i l i t y F inanc ing Method--200 ton Dry WoodlDay

Bases:

20-y r p r o j e c t l i f e5% year s t r ai gh t l i n e depre ciat io n on Total Capi ta l Requirement

excluding Working Capital48% fed er al income ta x r a t e75/25 d e b t / e q u i t y r a t i o10% i n te r es t on debt15% re tu rn on e qu i t y

Defin i t ion o f Terms:

C = Total Capital Requirement, 106 $W = Working Capital, 106 $N

= Total Net Operating Cost, l o 6 $ / y rG = Annual Fuel Production, 1 0 6 g a l / y rd = Fract ion Debti = Interest on Debt, % / y rr = Return on equity, % / y rp = Return on Rate Base, % / y r

Equation f o r Return on Rate Base:

Average Methanol Cost, $./gal =

N + .05 (C-W) + .005 ( p + 48/52 (1-d)r ) (C+W)G

Calcu la t ion:

Average Methanol Cost =



TABLE 53. Methanol Cost - Equity Financing Method--200 ton Dry Wood/Day

Bases:

20-yr project life

16-yr sum-of-the-years' digits depreciation on Total PlantInvestment100% equity capital12% DCF return rate48% federal income tax rate

Definition of Terms:

I = Total Plant Investment, lo6 $S = Start-up Costs, 106 $W = Working Capital, 106 $N = Total Net Operating Cost, 106 i/yr

G = Annual Methanol Production, 10 gal/yr

Methanol Cost Equation at 12% DCF Return

Methanol Cost, $/gal =

Calculation:




The re ac to r models e x i st in g i n t he ASPEN program are no t accurat e descr ip-

t i on s o f our exper imenta l u n i ts p r im ar i l y because mass and hea t t r an sf e r

e f f ec t s a re i gno red . Thus, much ti me has been spent an aly zi ng our d at a t o

q u a n t i f y any d i f f u s i o n a l l i m i t a t i o n s .

Some o f t he lab or ato ry work showed an incr ease i n rea ct io n r a t e as t he

steam f l o w r a t e i nc re as ed . T h i s sug ges ts a t l e a s t t h e p o s s i b i l i t y o f d i f f u -

sio nal l i mi t a t i o n s i n th e system. Furthermore, some of t he h ig h temperature

( 8 5 0 ' ~ ) d a t a d i d n o t f i t wi th t he res t on a s tanda rd ac t i va t i on ene rgy p l o t .

The apparent act ivat ion energy decreased at the h igher temperatures. Once

agai n t h i s suggest s d i f f us i o na l l i m i t a t i ons . These obse rva t i ons were j u s t i f i -

c a t i o n f o r t r y i n g t o q u an t i f y r e l a t i v e k i n e t i c and d i f f u s i o n a l r a t es .

The an al ys is s ta rt ed simple by assuming steam was th e prima ry r ea c t in g

species. Furthermore, i t was assumed th at a g lobal k i n e t i c expres sion could

d e sc r ib e t h e a ct u a l g a s i f i c a t i o n r e a c t i o n. For a s ing le pa r t i c l e we then ge t

t he f o l l ow i ng by equat i ng t he ra te o f steam t rans fe r t o t he sur f ace t o t he

rate of steam consumption by chemical react ion:

- r = mA ( C - C ) = k CP B S R S

wherer =

Rate o f s team reac t io nkm = Mass t ra ns f e r co e f f i c i en t

A = S ur fa ce a re a o f p a r t i c l eP

C B = Bulk steam concentrat ion

CS = Par t i c l e su r f ace s team concen t ra t i on

kR = F i r s t o rd er k i n e t i c r a t e c on st an t

The mass t r an sf e r c oe f f i c ie n t i s a fun ct i on o f t he Schmid t and Reyno lds

numbers. I t was evaluated using the j f a c t o r .



dpu.where: j = 0.357 For f ix ed bed w i th Re =-

RE10.359' ubetween 3 and 2000

p = Mass density

p* = Molar dens i ty

u = Gas veloci ty

p = Gas v iscos i ty

D = D i f f u s i o n c o e f f i c i e n t

d = P a r t i c l e d ia me te rP

A d i f f us i o n coe f f i e c i e n t f o r steam i n t he o the r gases wa's evaluated us ing

techniques i n Reid, Pr aus ni tz, and Sherwood (1977).

By rear rang ing th e r a t e equat ions we get a form for th e r a t e express ion

th a t inc ludes mass t rans fe r e f fe c t s .

The l i m i t i n g cases a re wor th no t ing

I f k A >> kR then r = k R CBm P

I f kR >> k A then r =k A CBm P m PIf R = kmAp then r = 112 kR CB

K i n et i c s l i m i t ov e r a l l r a t e

Mass t rans fe r l i m i t s ra teBoth mass transfer and

k i n e t i c s i n f lu e n ce r a t e

A problem ar i ses i n de te rmin ing kR f rom the ba tch reac to r da ta . Exper i-

menta l da ta were co l lec ted fo r the express ion .

where Wc = Weight o f carbon. I n terms of carbon convers ion t o gases (Xc )

t h i s becomes



Where Wc = i n i t i a l w ei gh t o f c ar bo n p re se nt .0

NO s imp le r e la t i onsh ip ex i s t s be t w een k c and kR because the s to ich iomet ry

f o r s team rea c t i ng w i t h wood i s no t w e l l de f ined f o r va r y i ng s teamlwood r a t i o s

and temperatures. Furthermore, the concent r a t ion o f steam i n th e re ac to r a t

any time was not measured. From mater ia l balance considerat ions a steam con-

cent ra t ion was ca lcu la ted and a "pseudo sto ichiometry " proposed a t each tem-

pera t u re. Th is a l low ed ca lc u l a t i o n o f kR valu es from th e measured k c r a t e con-

s t an t s . Then one can compute a r a t i o o f kR/kmAp.

Some pre l i m i na ry re s u l t s are

Temperature, OC

550

A l though t he re a re many unc e r t a in t i es i n t he ca lcu l a t ed va lues f o r kR, t h e

t r end shown i s i n t e re s t i ng . These r e su l t s sugges t su rf ace d i f f us io n cannot be

ignored and i t t en ds t o l i m i t r a t e s a t h i g h t em pe ra tu re s i n t h e b at c h r e a c t o r .

The p r i ma r y r ea so n f o r t h i s i s b e l i e v e d t o be t h e l ow gas v e l o c i t i e s ( t 2 cm l s ) .

A s i m i l a r a n a l y s i s of t h e PDU dat a showed th e mass t ra ns fe r r a t e was hi gh

enough a t 5 0 0 ' ~ so k i n e t i c s a re p r o ba b l y l i m i t i n g t he r e. Gas v e l o c i t i e s a t

5 0 0 ~ ~were ,100 cmls . No good PDU d a t a a t 8 5 0 ' ~ a r e a v a i l a b l e t o check t h e

model at h igh temperatures.

But t and Weekman (1974) have t a b ul a te d a s e r i e s of t e s t c r i t e r i a t o de te r-

mine w he ther hea t and/ o r mass t r an s f e r e f f e c t s a re impor t an t f o r ca t a l y t i c

r eac t ions . The c r i t e r i a check f o r su rf ace d i f f u s io n as w e l l as po re d i f f us io n

res i s t ances . The f o rmu la t i ons can app ly equ a l l y w e l l t o wood pa r t i c le s i f t h e



pore charac te r i s t i cs a re known. Info rmat ion on pore s t ru ct ur e i n wood was

taken from Boyd e t a l . (1979). R e su l ts from a p p l y i ng t h e c r i t e r i a o f B u t t and

Weekman supported the fo l l ow in g conc lus ions :

1) Mass t ra ns f e r does a f f e c t ov e ra l l r e ac t i on ra tes i n t he ba t ch

reac to r .

2) Hea t t r an s fe r ra te s i n t he ba t ch reac to r ra te a re l ow enough t ha t

pa r t i c l e t empera tu res (a t s teady s ta te ) a re 20-40°c ( a t 5 5 0 ' ~ and

850°c, re sp ec t i ve ly ) lower than th e gas temperatures. (Rad iat i on

ef fe ct s are ignored i n Weekman's a nal ys i s) .

3 ) Ne i ther hea t nor mass t r an s f e r l i m i ta t i on s occur i n the PDU (a t le as t

n o t a t 5 0 0 ~ ~ ) .

4 ) Pore d i f f u s i o n l i m i t a t i o n s do n o t e x i s t i n e i t h e r t h e ba tc h r e a c t o r

o r th e PDU.

An impor tan t qu a l i f i e r he re i s t ha t seve ral of Weekman's c r i t e r i a requ i r e t he

same kR value ment ioned ear l ier . Unce r ta i n i t i es i n ge t t i ng t h a t number have

already been mentioned.

Ef fo r t s t o der i ve a pure l y k i n e t i c model a re based on desc r ib in g the PDU

sinc e mass and heat t r an sf er l im i t a t i on s do no t appear t o be important . A f a i r

amount of information i n t he l i t e r a t u r e e x i s t s f o r d es cr ib in g t h e py r o l y s i s o f

coa ls and ce l l u l o se and then gas i f i c a t i on o f a w ide va r i e t y o f cha rs (Lowry

1963) . Un fo r tuna te l y nea r l y a l l t he ra te exp ress ions a re in tended to p red i c t

only carbon conversions. Few go on t o g i ve methods f o r pr ed ic t i ng gas y ie ld s.

No one has suc ces sfu l ly model le d ga s i f i ca t i on of coal or wood by pr ed ic t i ng

both conversions and gas composit ions. Our pl an i s to :

1) Assume pyr o l ys is occurs ra p i d l y and i s not a f f ec t ed by steam. We

a l lo w t h e p y r o l y s i s r e a c t i o n t o g i v e e q u i l i b r i u m y i e l d s o f carbon

and produce known amounts o f gases. Pure carbon i s th e assumed so l i d

produc t . Tar fo rmat ion i s ignored .



2 ) Take the pyro l ys i s p roducts as the s t a r t i ng mix tu re fo r the fo l low ing

se t o f r e a c t i o n s

Not a1 1 o f these reac t io ns are mutual ly independent. React ion a) - b ) w i l l

g ive e ) on a pure ly s to i ch iome tr ic bas is . However, re ac t i on e) i s cata lyze d by

some o f th e ca ta ly st s used i n our te st s. Hence, in cl ud in g i t was th e ea si es t

way t o a ccount f o r t h i s ca t a l y t i c e f f e c t .

The eq ui l i br iu m pred ic t i on s appear reasonable. Rate d ata f o r t h e so l i d

reac t i ons (a,b,c) come from Lowry but are f o r coconut sh el l charcoal and are

based on u n i t mass o f carbon i n th e system. Coconut s h el l charc oal was th e

c lo ses t th in g t o wood f o r wh ich da ta were ava i lab le . React ion ra te constan ts

for the gas phase reactions have also been found based on unit catalyst weight.

When a1 1 ra te s are adjus ted t o a common m o l e s l t i m e bas is, th e gas phase reac-

t i on ra tes a re abou t 3 orders of magnitude higher than the slowest carbon reac-

t i o n r a t e ( a t 5 5 0 ' ~ ) . This means no more carbon disappears after pyrolysis

so gas phase compo siti ons are changed on ly by re ac ti on s d and e. The r e s u l t i n g

carbon conversions are lower than experimentally observed and gas compositions

do no t agree e i t her . Ef fo r t s now are d i re cte d toward f in d in g be t te r r a t e da ta

and check ing fo r any e r ro r s i n conve r t i ng a l l i n fo rma t ion t o a cons is ten t se t

o f u n i t s .

FUTURE MODELLING EFFORTS-A f te r a co n s i s t en t se t o f k i n e t i c d a ta i s deter mi ned t h a t w i l l agree wi th

the PDU data , the gas i f ie r model w i l l be coded t o fit i n t o t he ASPEN computerprogram. The ASPEN program can then generate heat and mater ial balance infor -

mat ion f o r se lected changes i n opera t ing cond i t ions. A v i su a l i n sp e ct i o n o f

t he se r e s u l t s w i l l show whi ch equipment needs t o be r e s i z e d (based on a com-

pa r i son w i t h Davy McKee base-1 in e s tud ies) . T h e o r e t i ca l l y t h e ASPEN code can

aut omat ica l ly go from heat and mat er ia l ba lance cal cu l at i ons t o equipment



sizing to equipment costing and then to economic evaluations. The manual

inspection is desirable if only one or two pieces of equipment will change

size. The savings in computer run time will be significant.

Once equipment costs are defined ASPEN generates a report on the economicsof the process. The calculations and format of the report follow ESCOE

guidelines.



ACKNOWLEDGMENTS

The au thors a re espec ia l l y g r a t e f u l t o T. J. Kendron, 0. A. Kuby,

M. McCl in tock , and J. H. Rooker o f Davy McKee, Inc., who completed the economic

eva luat ions. Spec ia l apprec ia t ion i s a lso extended t o J. E. Leonard,

W. F. Riemath, G. L. Roberts, E. D. Smith, and W . A. Wilcox, o f PNL, whose

e f f o r t s have been v i t a l t o the success of the exper imental program.



REFERENCES

Andrew, P.? G. Martin and A. Noher. 1971. "The Cracking of Butyl-BenzenesOver Silica-Alumina Catalysts." Journal of Catalysis 21:255-262.

Araki, M., and V. Ponec. 1976. "Methanation of Carbon Monoxide on Nickel andNickel-Copper Alloys." Journal of Catalysis 44:439-448.

Boyd, M., et al. 1979. "Pyrolysis and Gasification of Hybrid Poplar SPP." InProceedings of the 1979 Annual AIChE Meeting, paper 78C, San Francisco,California.

Butt, J. B., and V. W. Weekman, Jr. 1974. "Characterization of the ActivitySelectivity and Aging Properties of Heterogeneous Catalysts." AIChESymposi um Series 70(143):27-41.

Gardner, N., E. Samuels and K. Wilks. 1979. Advan. Chem. Ser. 131:217.

Gumz, W. 1950. Gas Producers and Blast Furnaces. Wiley and Sons, New York.

Johnson, J. L. 1974. Advan. Chem. Ser. 131:145.

Lefrancois, P. A., K. M. Barclay and G. T. Skaperdas. 1967. Advan. Chem. Ser.69:68.

Lowry, H. H., ed. 1963. Chemistry of Coal Utilization, Supplementary Vol.Chapter 20, John Wiley and Sons, New York.

Mitchell, D. H., et al. 1980. "Methane/Methanol by Catalytic Gasification of

Biomass." -CEP 76(9) :53-57.

Mudse, L. K., L. J. Sealock, Jr., and S. L. Weber. 1979. "Catalyzed Steamasi ification of Biomass." Journal of Analytical and Applied pyrolysis1:165-175.

Mudqe, L. K., et al. 1979a. "Catalytic Gasification of Wood Residues forethanol en era ti on." In proceedings of Mid-American Wood CombustionConference, MASEC-CF-80-005, Traverse City, Michigan.

Mudqe, L. K., et al. 1979b. "Catalytic Gasification of Biomass," p. 351-357.1i Proceedings of the 3rd Annual Biomass Energy Systems conference,SERIITP-33285, Golden, Colorado.

Mudge, L. Key et al. 1979c. "Investigation of Gasification of Biomass in thePresence of Catalysts," PNL-SA-7939. Presented at the 8th ThermochemicalConversion Contractor's Meeting, Seattle, Washington.

Mudge, L. K., et al. 1979d. "Investigation of Gasification of Biomass in thePresence of Catalysts," PNL-SA-7939. Rev. 1. Presented at the 9th Thermo-chemical Conversion Contractor's Meeting, Rolla, Missouri.



Mudge, L. K., e t al . 1980. " I nves t iga t ion o f Gas i f i ca t ion o f Biomass i n t hePresence of Catalysts," PNL-SA-7939, Rev. 2. Presented a t t he 10 th Thermo-chemical Conversion C o n t r a c t o r ' s Meeting, Berkeley, California.

Reid, R. C., J. M. Prausni tz and T. K. Sherwood. 1977. 'TheProper t ies o fGases and L i q u i d s . " 3rd ed,. McGraw H i l l , New York.

Robertus, R. J., e t a l . 1979. " I n v e s t i g a t i o n o f th e Gas i f ica t io n of Wood i nthe Presence of Mixed C ata l ys t s . " Presented at the Annual AIChE Meeting,San Fransisco, California.

Sealock, Jr., J., S. L. Weber and L. K. Mudge. 1980. "Kinetics and Mechanismso f Steam Ga si f i ca t i on o f Biomass i n t he Presence o f Al ka l i Carbonates."Biosources Digest 2(1 ):12-26.

Shafizadeh, F. 1975. " I n d u s t r i a l Pyro lys is o f Ce l lu los ic Mater ia ls ." Appl iedPyro lys i s Symposium No. 28, pp. 153-174. Wiley and Sons, New York.

Shafizadeh, F., P. S. Chin and W. DeGroot. 1975. J. F i r e and F l a m n a b i l i t y lFire Retardant Chem. 2:195.

Walkup, P.C., e t al . 1978. " I n v e s t i g a t i o n o f Gas i f i ca t ion o f Biomass i n t hepresence of Mu1 t i p l e Cata lys ts I ' 1n Proceedings of the Second AnnualSymposium on Fuels from Biomass. CONF-7805107-PI, W. W. Shuster, ed.,Renssel aer Poly tec hni c I n s t i t u t e , Troy, New York.

Weber, S. L., e t al . 1980. Ga si f i ca t i on o f Biomass i n the Presence o fMu l t ip le Cata lys t s f o r t he D i rec t P roduc t ion o f Spec i f i c P roduc ts .PNL-SA-7763, P a c i f i c Northwest Laborat ory, Rich land, Washington.



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Washington, DC 20545

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Washing ton, DC 20545

DOE R i ch l and Opera t ions Of f i ce

H. E. RansomD. R. Segna

26 Pa ci f i c Northwest Laborator yC. M. DevarvD. H. ~ i t c h G l lL. K. Mudge (10)R. E. N i g h t i n g a l e

R. J. RobertusG. F. S c h i e f e l b e i n ( 2)L. J. Sealock, J r .P. C. WalkupS. L. WeberPub l i sh ing Coord ina t ion ( 2 )Techn ica l In fo rmat ion (5 )

27 DOE Tec hni ca l I nf or ma ti on Center

Methanol II

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