Oxida tiv e coupling of methane: Resol utio n of the
surfa ce an d gas ph ase contr ib ution s to the
mec hanism of the ox ida ti v e cou pling of me than e at
Na 2 WO 4 - Mn - SiO 2 - catal y st
v orgele gt von
Dip lom - Ingenieur
Vi nz enz Fl ei sch er
geb. in Ber lin
Von der F akult ä t II - Math em ati k und Naturw issenscha ften
der T echnischen U nivers itä t Berlin
zur Er langung d es ak ademis chen Grades
Dok tor der I nge nieurwiss ensc haften
- Dr. - Ing. -
genehm igte Dis sert atio n
Prom otionsauss chuss:
Vorsit zender : Prof. Dr . Nedilijk o Budis a , TU Berlin
Gutacht er : Prof . Dr. rer . nat. R einhard Sc hom äcker , TU Ber lin
Gutacht er : Prof . Dr. rer . nat. R obert Sch lögl, FHI der M PG Berli n & Mühl heim an d er Ruhr
Gutacht er : Prof . Dr. - Ing. Ulr ich Niek en, Uni versität Stuttgart
T ag der wissens chaftlic hen Auss prache: 21.12.201 6
Berlin 201 7
II
A ck no w ledgements
T he present thes is was perf ormed and iss ued at the T echnical Chem istr y Dep artm ent of the
Chem istr y Ins titute at t he T echnisc he Univers ität Ber li n, i n the fr am ew or k of “Unicat” (Un ifying
Concepts in Cata lysis) project as a part of the su bprojec t D1/E1, “ Activat ion of Methane” .
First of all, I would like to thank m y su pervisor Prof. D r. Re inhard Schom äcker, f or admittin g m e
into his r esearch gr oup, for the super vi s ion of m y work and gi vi ng m e the c hance to achi eve this
m ilestone. I would a lso lik e to thank Prof . Dr. Ro bert S chlögl f or the sec ond su pervisi on. I t hank
Prof. Dr. - Ing. U lrich Niek en as my ex ternal ex aminer .
I thank also Dr. Patr ick Little wood and Samira Par ishan for gr eat discus sions about t he
devel opment of the setup, t he OCM and also experim ental tec hniques in g eneral . Further I lik e
to thank all co lleagues i n the res earch gro up of Prof. Dr. R . Schom äck er, especially:
Dr. Be njam in Beck
Dr. T orsten O tremba
Dr. Marc Schrö der
Maxim ilian Ne umann
Arno Z imm ermann
Luk as Thum
Marcel Sc hmidt
T obias Pogr z e ba
Dr. Mich ael Sch warze
I would also tha nk all s tudents who supp ort m e during this per iod, esp eciall y R olf Steuer , Maik a
Stöbe, T rung N go Than h and Maik Rudolf. I am also gratef ul for the s upport f rom Gabr iele
Vetter and Chr ista Löhr.
Also I h ave t o thank the internation al Max Planck resear ch schoo l (IMPR S) and all of its
m e m bers, for the initia l financ ial su pport of m y PhD and great discuss ions. I hav e learned a lot
about f undam ental researc h techni ques. Furt her m ore that sc hool gave m e the opport unity to
collabor ate w ith other people, w ho work als o in the O CM field .
I would lik e to thank Ulla Sim on and M aria G racia C olmenar es for the prepar ation,
charac terizat ion of sever al catal y s t m aterials and the c lose c ollaborat ion. I wou ld cordia lly lik e to
thank Dr. H.R. Go dini f or the disc ussions about t he OCM an d the proc ess en gineering as pects .
I would l ike to tha nk the Deuts che Fors chungs gemeins chaft ( DFG) for fundin g the Exc ellenc e
Cluster “ Unicat” (Unif ying Conc epts in C atalysis).
Forem ost I w ant to tha nk m y parents Brigitte and Klem ens for their supp ort of m y devel opm ent
during th e last y ear s and all m y life bef ore. T he sam e is true f or m y gir lfriend Lea, w ho
supporte d m e in all diff icult situati ons. Sh e was alwa y s th ere for m e an d I wo uld give also a
special Ack nowledgem ent to m y beloved d aughter a nd s on.
III
A bstract
Now aday s, ethe ne is prod uced through s team crack i ng of n aptha, which is a f raction of crude
oil o btaine d after its disti llation. Ethene is an im portant bas ic chem ical f or the che mic al industr y ,
but due to decreas ing rese rves of c rude oil an a lterna tive feeds tock is required . T he oxidative
coupl ing of m ethane (O CM) is a prom ising reac tion f or ethene pr oducti on. N at ural gas, whic h
contai ns high amounts of m ethane, of fers a new f eedstoc k f or the chem ical in dustr y . Further
m ethane contain ing gas es , such as s hale or bio g as , can be utili zed for that proc ess , too . Many
catal y s t s have bee n investigate d so f ar, but due t o sever al reasons a tech nical re alization on an
ind ustria l scale is hard to real ize. O n the on e han d a hig hly ac tive, sel ective and s table c atalyst
is neede d . O n th e other hand ther e is a s trong c ontrib ution of gas phase s ide rea ction s , due to
co - feeding of methane and ox y ge n at hi gh tem peratures , which requ ires an adequat e reactor
concept to ach ieve high et hene yields .
In the pr esented th esis, OCM r eactions were c arried out on Na 2 WO 4 /Mn/Si O 2 catal y st s in
absence of gas phase ox ygen to clar if y the s urfac e rea ction networ k. T e m peratur e program m ed
surf ace reac tion (TPSR) ex perim ents with m ethane, etha ne , and ethen e were c arried out t o
stud y th eir interac tion w ith the cata ly st m aterial. F urther m or e an activatio n energ y for the
selecti ve m ethane act ivatio n was calculat ed usi ng the R edhe ad m ethod. D y n am ic experim ents
at const ant reactio n tem perature in dicate tha t for the Na 2 WO 4 /Mn/SiO 2 ca talyst s trongl y b ound
ox y ge n is invol ved in the O CM process , which can sur vive se veral hours under OCM c onditions
at the c atalyst sur face. B ased on thes e res ults a chem ical loo ping set up was c onstructed, whic h
all o ws to dos e the reac tants s eparately i n defined a m ounts . In our st udies, we f ound that the
OCM y i eld is lim ited to 25 %, which is c aused by the ca talyst m ateria l itself. F urther
investi gations of diff erent ly loaded Na 2 WO 4 /Mn/SiO 2 catal y s ts b y r epetitive pu lse exper iments
reveal a c orrelati on between t he stored am ount of ox ygen and the m anganes e content. F or the
chem ical loop ing conc ept we f ound oper ation c onditi ons whic h enab le to conver t m ethane m ore
eff icient ly th an in t he classical co - feed m ode. Additi onall y to the kinet ic inves tigations an
upscali ng conc ept was pro posed.
IV
V
Erkl ärung
Ich erk läre h ierm it, dass ic h bislang an k einer an deren Hoc hschu le oder Fak ultät m eine
Prom otionsabsic ht beantr agt habe.
Die v orliegende Diss ertation wurde ber eits in For m von wi ssens chaf tlichen Publik ationen
veröff entlicht. B ei den e ntsprec henden Ver öff entlichunge n hand elt es sic h um folgende
Publik ationen:
Paper 1 : T herm al Reaction Ana lysis of Oxidati ve Coupl ing of Met hane
Ham id Reza Godi ni , Vin zenz Fleis cher, O liver Görk e, Stanis lav Jaso, Reinhar d Schom äck e r
and G ünter W oz n y
Chem ie Ingenieur Tec hnik, 2014 , 86(2), 19 06 - 1915 , D OI: 10.1002/c ite.20140 0080
Eigen anteil: Ic h bin Ersta ut or zusam men m it Ham id Reza Godini . In dies er Arbeit wurde der
Einfluss des Sa uerstof fpartia ldruck s in der oxidat iven K u pplung v on Met han (OCM) durc h
Sim ulation eines Reak tions netzwerk es für Gasphasen - Reak tionen und der I m ple m entierun g
eines einfac hen Katal ysatorber flächen r eakt ionsnet z werk es von m ir unters ucht. Die
Sim ulationen ha be ich in Chemki n ( Reaction Design ) für verschiedene R eaktorm odelle (PFT R,
CSTR) dur chgeführt . Anhand der Sim ulationserge bnisse wur de von m ir eine optim ierte
Sauers toff - Dosierungss trategie u nd die Auswir kung einer Ho t - Spot Aus bildun g auf di e OCM
disk utiert.
Paper 2 : Oxidative coup ling of m ethane — A com plex sur face/gas phase mec hanism wit h
strong im pact on t he re action e ngineering
Benjam in Beck , Vinze nz Fl ei scher , Sebast ian Arndt , Mig uel Gon zález Hevia , Atsus hi Urak aw a,
Peter Hugo and R einhard Schom äcker
Catal y s is Toda y , 2014, 228 , 212 - 218 , DOI: 10.10 16/j.c attod.20 13.11.0 59
Eigen anteil: Der Reak tionsverl auf der O CM R eaktion unter Ver wendung eines N a 2 WO 4 /Mn/SiO 2
Katal y s ators wurde v on mir in einem Hochdruck versuc hsstand unt ersucht. Aus den g ewonne n
Ergebnis sen (darg estellt in Abbi ldung 2) konnt e ich s chließen, dass höhere Absolut drück e die
Perf or m ance der OC M Reak tion in ein em F estbettreak tor ver bessern.
Paper 3 : Inves tigation of the surf ace reactio n netw ork o f the oxidati ve coupl ing of m ethane
over N a 2 WO 4 /Mn/SiO 2 c atalyst by tem perature pro gram med and d ynam ic experim ents
Vi nzenz Flei sche r , Rolf Steuer, S amir a Parishan a nd R ein hard Schom äcker
Journal of C at al ys i s , 2016, 341, 91 - 1 03 , DOI: 10. 1016/j .jcat.2016.0 6.014
Eigen anteil: Ic h bin Ers taut or. Die O CM Reak tion wurde vo n m ir unter Verwend ung eines
Na 2 WO 4 /Mn/SiO 2 Katal y s ator s m it Hilfe von temper atur - program m ierten und dy nam isc hen
Experim enten unters ucht und d isk utiert . Aus den Ergebn isse n wurde von mir ein
Reak tionsnetzwer k f ür die Oberf lächenreak tionen der OCM an ei nem Na 2 WO 4 /Mn/SiO 2
VI
Katal y s ators vo n m ir auf gestellt. Ferner wu rden die Sauer stoff speicherk apazität des
Katal y s ators un d deren Stabilitä t unter O CM Be dingun gen vo n m ir unters ucht und disk utiert .
Paper 4 : Chem ical loopi ng as reac tor conce pt for the ox idative co upling of metha ne over a
Na 2 WO 4 /Mn/SiO 2 cataly st
Vi nzenz Flei sche r , Patric k Litt lewood, S amir a Parishan , an d Rein hard Schom äcker
Chem ical Engi neering J ournal , 20 16, 306, 646 - 654, D OI: 10.101 6/j.cej .2016.07. 094
Eigen anteil: Ic h bin Erstau t or. M it H ilfe der g ewonnen Er kenntn issen a us Pap er 3 wu rde ei n
„ Sim ulated Chem ical Lo oping” - Versuchs stand k onzipiert und k onstruiert . Mit Hilfe di eses
Reak tionsführ ungskon z ept es wurde n von m ir Unter suchung en zur C 2 Ausbeut e M ax imierung
durchgef ührt. Ebens o führ te ich St abilitätst ests f ür den Na 2 WO 4 /Mn/SiO 2 K atalysator in
Chem ical - Loopin g Exp erim enten unter OCM B edingun gen durc h. Aus den g ewonne n
Ergebnis sen wur den von mir bestehende Pro zessk onzepte un d deren An wendung auf das
Chem ical Loop ing dis kutiert. W eiterhin folgte ein Ver gleich zwisch en s ta tionären und
dynam ischen Ex perim enten, der ebe nfalls a usführ li ch diskut iert w urd e.
Paper 5 : Investigation of the role of the Na 2 WO 4 /Mn/SiO 2 c atalyst c om position in th e oxid ative
coupl ing of m ethane b y che mic al looping exper im ents
Vinzen z Fleisc her , Maria G arcia Colm enares, Sam ira Par ishan, Ul la Sim on, Oliver Görk e,
Aleks ander Gur lo and R einhard Schom äcker
Journal of catal ysis, t o be subm itted
Eigen anteil: Ic h bin Ers taut or. In dies er Arbeit wur den verschiede ne, von Maria Gr acia
Colm enares zur Verf ügung ges tellte , Na 2 WO 4 /Mn/SiO 2 Kata lysatore n hin sic htlich ihr er
Sauers toff speicher - eige nschaf ten von m ir untersuc ht. H ierfür wurde v on m ir das „ Sim ulated
Chem ical Loopi ng” - K onze pt und r epetiti ve Puls - Ex perim ente verwe ndet . Di e repetit iven Puls -
Experim ente wurde von m ir m odelliert und es wur den Sim ulationen durc hgeführt zu r
Unters uchung des Sauers toff verbauchs. Aus de n Ergebn issen w ur de vo n m ir eine Kor relation
zwische n Sauers toffs peicherk apazität und Ma ngang ehalt un d Film dick e hergele itet. Fer ner
wurde ein Sc hema zur Funk tions wei se des Na 2 WO 4 / Mn/SiO 2 Kata lysator s in d y nam isc hen
Experim enten vo n m ir auf gestel lt.
Jewei ls ein Exem plar d er Publik ation ist d ieser D iss ertatio n beig efügt.
Ort, Datum Vi nzenz Flei sche r
VII
Table of c onte nt s
1 Introduc tion ……………… ………………………… ……………… …………………… ………... 1
1.1 Moti vation ...................................................................................................................... 1
1.2 Mec hanistic as pects ...................................................................................................... 1
1.2.1 G as phase re action n etwor k ...................................................................................... 1
1.2.2 C atalyst m aterials ...................................................................................................... 3
1.2.3 R eaction n etwork and kinet ic m odels ........................................................................ 5
1.3 Eng ineerin g aspect s ...................................................................................................... 6
1.3.1 Pr ocess eng ineerin g aspec ts .................................................................................... 6
1.3.2 T he chem ical loopin g proce ss conc ept ..................................................................... 8
1.4 O utline of this thes is .................................................................................................... 10
2 Metho ds and materi als …… …….……………… ………………………… …………… ……... 11
2.1 Sim ulation of the OCM gas phase reac tions an d catal yst surf ace reac tions .............. 11
2.2 C atalyst prepara tion .................................................................................................... 12
2.3 N itrogen adsor ption ..................................................................................................... 12
2.4 H igh press ure exper iments ......................................................................................... 12
2.5 Mas s spect rom etry ...................................................................................................... 13
2.6 T emper ature program med s urfac e reactions .............................................................. 14
2.7 D y nam ic ex perim ents .................................................................................................. 15
2.8 Sim ulated c hemic al loopin g ........................................................................................ 16
2.9 R esidence t ime anal y s es of the sim ulated chem ical looping s etup ............................ 18
3 Results an d discu ssion …………… ……………… ……………………… …………… …… ... . 20
3.1 C ontribution of gas ph ase r eactions to the OC M react ion .......................................... 20
3.2 Pr essure v ariation dur ing th e OCM re action ............................................................... 22
3.3 T emper ature program med s urfac e reaction exp erim ents ........................................... 24
3.4 D y nam ic ex perim ents .................................................................................................. 28
3.5 D evelopm ent of a s imulate d chem ical looping s etup .................................................. 30
3.5.1 Setu p constr uction ................................................................................................... 30
3.5.2 R esidence t im e analy s es of the sim ulated chem ical looping s etup ........................ 32
3.6 Pr ocess par ameter studies in s imulate d chem ical loopi ng exper iments .................... 35
3.6.1 Inf luence of flowra te and te m perature on th e OCM rea ction .................................. 35
3.6.2 C ontinuous operat ion of th e chem ical loop ing proces s and com parison with
stead y state proc ess m ode ..................................................................................... 37
3.7 Invest igation of the ro le of the c atal y s t com position in sim ulated chem ical loo ping
experim ents and r epetiti ve puls e experim ents ............................................................. 41
3.7.1 Var iation of the sur face c oncentrati on of Na 2 WO 4 and Mn o n COK - 12 s upport ..... 41
3.7.2 Var iation of the m anganese loading o n COK - 12 s uppor t ....................................... 44
3.7.3 Var iation of the supp ort m aterial ............................................................................. 47
4 Conc lusions ……………… ………………… ……………………… ……………… …………… 51
5 Li st of Refe rences ……… ………………………… ……………… …………………… ……..... 56
VIII
Nomencl ature
List of A bbreviations
A b breviati ons
Descri ption
CLC
Chem ical loop ing com bustion
CLR
Chem ical loop ing ref orming
CST R
Continu ous stirr ed tank reac tor
DFT
Densit y f unctional t h eor y
EPR
Electro n p aram agnetic r eso nance
FT IR
Fourier tr ansf orm inf rared spec troscop y
MFC
Mass f low contro ller
MS
Mass s pectrom eter
PFT R
Plug f low tube r eactor
OCM
O xidative coupli ng of m ethane
PV
Pulse v alve
SS A
Specific surf ace area
ST Y
S pace tim e y ield
TCD
T herm al conductivit y det ector
Td
T ransducer
T PSR
T emper ature program med s urfac e reaction
X AF S
X- r ay absorpt ion fine str uctur e
XPS
X- r ay photoelec tron sp ectrosc opy
XRD
X- r ay diffr action spectr osco py
List of sy mbol s
Symbol
Descri ption
Unit
A
Peak area
% ·s
A s pec
Specific surf ace area
m²/g
c
Concentr ation
m ole/ m l
D
Dif fus ion coef ficient
m²/s
d
Diam eter
m
d 32
Sauter mean di ameter
m
D ax
Axial d ispersio n coeff icient
m²/s
d h
Hydra ulic diam eter
m
d R
Reactor diam eter
m
E(t)
Reside nce tim e distribut ion f unction
1/s
E A
Activat ion energ y
kJ/ mole
k b
Bolt z m ann cons tant
J/K
k oo
Pre - f actor
1 / (s ·m ³ (x - 1) ·m o l e (x - 1) )
L
Length
m
M
Molar m as s g/mole
m
Mass
kg
n
Am ount of subs tance
m ole
N
Num ber of s eg m ents
-
p
Press ure
Pa
X
Symbol
Descri ption
Unit
Pe
Pecle t n um ber
-
R
Univers al gas c onstant
J /(m o le ·K)
Re
Reynol ds num ber
-
R o
O x y gen tr ansport ca pabi lit y
-
S
Selecti vity
-
Sc
Schm idt num ber
-
T
T e m peratur e
K
t
T im e
s
t delay
T i m e duratio n since gas puls e reaches th e reactor
inlet
s
t dosing
T i m e duratio n since the gas pulse is c ompletel y
dosed
s
t pulse
P ulse c ontact tim e
s /m
u
Flo w velocit y
m /s
V
Volum e
m³
V*
Vacanc y on th e catal y st s urf ace
-
𝐕𝐕
Flow rate
Nm l /m in
w
Mean f ree path l ength of a gas m olecule
m
X
Conver sion
-
O*
Ox y gen sp ecies boun d on t he catal y s t sur face
-
Y
Yield
-
z
Spatia l coordin ate
m
Δ R H
Reacti on enthalp y
kJ/ mole
ε
V oid f raction
-
ζ
D rag coef ficient
-
η
V isco sit y
Pa ·s
ρ
D ensit y
Kg/m ³
Introduct ion
1
1 Introduction
1.1 M otivation
Accor ding to the f orecas t by O wen et al ., the d em a nd f or crude o il wil l ra ise durin g the next
decades b y ab out 40 Bill . barrels . [1] C oncurr entl y , the rem aining and ne w ly discov ered oil
reser ves can cover th e dem and until the en d of this centur y. The m ain utilizat ion of c rude oil is
energ y generatio n (93 %) and m inor am ounts ( 7 %) are used for the pr oductio n of chem icals .
[2] Most of our d aily used pr oducts are based o n these chem icals. Es pec ia ll y plast ics , lacq uers ,
surf actants and bas ic che micals , w hich are further converted t o pharm ac euticals , f ertilizers and
textile - pu lps contri bute to t he dail y life in our h igh ly techno log i zed soc iety. On e of the m ost
im portant plast ics is pol y e th y l ene, w h ich is widel y used f or film s, pi pes , and f or wire or cab le
insulati on s. [3] T he pref erred pro duction of the mono m er ethylene is car ried ou t by the stea m
crack ing proces s. [4] T heref ore, a n alter native prod uction rout e based on a new feedst ock
seem s to be n ecess ary to over com e the high depen dence o n crude o il, co vering further ethene
request in the f uture .
Methane is the m ain compoun d in natur al gas , which is wide ly us ed for en ergy produc tion ,
nowada y s . [5] O ne pr omis ing reacti on to convert m ethane into C 2 buil din g bl ocks i s the
oxidat ive coup ling of m ethane (O CM) which was firs t published b y Keller and Bhasin in 1 982 .
[6] T he overall re action is given in equat ion (1 ).
x CH 4 + y O 2 → a C 2 H 6 + b C 2 H 4 + c CO 2 + d H 2 O + e CO
(1)
Th e r eaction requires methane , which is con verted b y a c atalyst into m eth y l r adicals . These
radic als c an coupl e to et hane, which c an be f urther deh y dro genated to eth ene. [7] Thus, t he
ox idat ive coupl ing of m ethane is a pr omis ing r oute to c onvert n atural gas into valu able
c om p ounds for c hem ical industr y , which is inde pende nt from crude oil .
1.2 Mechanistic aspects
1.2.1 Gas phase reacti on network
T he therm al coupling pr ocess of m ethane is therm odynam ically lim ited, beca use of its highl y
endother mic nature in abs ence of ox idizing age nts. T he use of oxidizing ag ents leads to an
exotherm al reac tion entha lpy, whic h does no t lim it the m ethane c onver sion b y th erm ody n am ics
anymore. [8] T he m ethane m olecule its elf has the str ongest C - H bond s trength in com parison
with oth er alk anes, w hic h m eans that a h igh acti vatio n barrier has to be overco m e for m ethane
activati on. T herefor e the OCM reac tion is t y p icall y per form ed above 700 °C. [9] U nfortunate ly ,
these hi gh temper atures caus e unselect ive react ions of m ethane and ox y gen, form ing deep
oxidat ion products . T hose s ide reacti ons ar e stro ngly exotherm ic and decreas e t he select ivit y of
C 2 com pounds in seve r a l w a ys as sho w n in T able 1 . T he use of oxygen as ox idizing ag ent has
the co nsequenc e that not onl y methane is con verted in to CO X . A lso the C 2 product s are af fected
Introduct ion
2
by consec utive oxi dation r eactions which lea d to the form ation of f urther deep ox idatio n
product s an d strong heat g eneratio n . Further, it has to be consi dered that t here is a com petition
between m ethane m olecule ac tivati on on the cat a l ys t s urface and in gas phase caused t roug h
the h arsh reac tion con dition s . T herefore not on ly the c atalyst cont ributes to the conversion of
m ethane .
T able 1 : Reaction enthalpi es of im portant r eaction s i n the OC M proces s [10]
No.
Reaction
𝚫𝚫 𝑹𝑹 𝑯𝑯 (kJ/mole)
1
CH 4 + 1
2 O 2 → C 2 H 6 + H 2 O
- 177
2
CH 4 + 2 O 2 → CO 2 + 2 H 2 O
- 803
3
C 2 H 6 + 1
2 O 2 → C 2 H 4 + H 2 O
- 103
4
C 2 H 6 + 7
2 O 2 → 2 CO 2 + 3 H 2 O
- 1429
5
C 2 H 4 + 3 O 2 → 2 CO 2 + 2 H 2 O
- 1323
As m entioned bef ore, the p ri m a ry r ole of the c atal y st m aterial is the c leavage of a C- H bond in
the m ethane m olecule s . T hese rad icals are high ly active and cou ple to et hane molec ules . Other
radic al int erm ediates form ed by activati on of dif ferent m olecules in gas phase or on the c atal y s t
surf ace are also invo lved . Doole y and Co worker s published a m echanism for m anifold gas
phase reac tion s , consid ering arou nd 160 0 reactions and 280 interm ediates . [11] Th eir reaction
m e chanism wa s v er ified b y Schwar z et al. in a pr ofile reactor whic h allo wed to s tud y th e loca l
gas c ompos ition of the gas f eed b y a sam pling cap illary. [1 2] A sens itivit y anal y s is of tha t
reaction netw ork was c arried out b y M avlyank ariev. [13] H e found that th e initia ting gas phas e
reaction steps are ind uced b y the reac tion s betwee n m ethane and ox ygen whic h form meth y l
and h y dr operox yl radicals . H ydroperox yl r adicals an d sim ilar oxygen radic als c ontribute to
fur ther unselect ive react ions b y ac tivatio n of alk anes or alk enes or the ir radical interm ediates.
Especia lly etha ne and ethene m olecules are highl y sens itive for f urther oxidat ion react ions.
Z an t hoff and Baerns po inte d out that a s cavenger mat erial for HO 2 • , OH • an d O: radic als seem
to be necess ary to pr eve nt tota l oxid ation rout es in gas phase after s uccess ful coup ling of
methy l radi cals. [ 14] A nother im portant interm ediate i s form aldeh y de which is hi ghly acti ve for
fur ther radical for m ation by its dec ompos ition into H • and HCO • r adic als . Th e s e radicals
c ontribut e to other rad ical reac tion st eps, w hich pr eferabl y lead t o deep oxidati on prod uct
form ation. T he origin of H 2 CO is basical ly the co upling proc ess of m eth y l rad icals with d iff erent
ox y ge n speci es. However , these gas phase re actions contribut e to th e limitati on of the C 2 y i eld,
which was pres ented by Zav yalova et al. f or sever al catal y st m aterials. [15] For the app lic ati on
of the O CM proc ess in ch em ical industr y , a C 2 yield abo ut 30 % is neces sar y for ec onomic al
ethene pr oductio n. [16,17 ] It has to be no ted that s uch cost estim ation s ar e strongl y de penden t
on the c urrent m ark et prices of m ethane a nd ethe ne.
Introduct ion
3
1.2.2 C atalyst mater ials
A large num ber of activ e m aterials f or the OCM r eaction was prese nted b y Bhas in and Ke ller ,
later rev iewed by Baerns and Hutchings et al. [18 – 20] T hey sug gested as m ost prom ising
m aterial MgO . O ne im portant f actor, bes ides cat alytic activit y for the OCM proc ess , is the l ong
tim e stability of the c atalys t m aterial. It was dem onstrated that several h ighl y act ive mater ials ,
perf orming the OCM reac tion under h arsh cond itions , tend to deactivate in ti me on s tream
experim ents. [21] O ne w ell - k nown c atal y s t mater ial is the Li/MgO s ystem , which was re vi ewed
in deta il by Arn dt et al. [8] T hat catal y s t m aterial tends to lo o s e Li in tim e on str eam experim ents
and its catalytic activit y c orrelates with t he num ber of steps or edges. [ 22] Theref ore the OC M
activit y of this mater ial has a strong s ensitivit y on its morpho logy. A propos ed mec hanism is the
m ethane adsorp tion and c oordi nation at a deh y drox y l ated step. T he m ethane m olecule is than
polari zed and the C - H b ond is cle aved by an O 2 m olecule f ro m gas phas e, w h ich f orm s an HO 2 •
radic al. [23] One import ant aspect is that the k nowledge a bout t he MgO m ateria l provid es the
opport unity of Densit y F unction al Theor y ( DFT ) calculations for a m echanistical ly understa nding
of the O CM proc ess. Kwapien et al. pr esented a mode l for a MgO clus ter , whic h invo lves
severa l ox y ge n speci es on the cat al y s t surfac e for m ethane act ivation . [24 ] That fact ind icates
that s im plified k inetic m odels, as Langm uir Hins helwo od or Ele y Ridea l, for that m aterial are not
able to pred ict the reac tion pr ogress in dif ferent reac tor s y stem s. This is because the ox ygen
do sin g strate gy of the feed and th e react ants disper sion m ay inf luence th e presenc e of th ese
species . Theref ore the contribu tion of the cata lyst mater ial to the O CM re action m ay diff er,
because the nature of dif ferent ox y gen ions or ra dical s pecies is ver y diff erent. T hat wil l be
discus sed later. U nfor tunatel y , Korf et al. dem onstr ated that the Li/MgO c atal y st i s not stab le in
tim e on stream experim ents. [25]
Another w e ll - known c atalys t m aterial is La 2 O 3 . T o reach high ca talyti c activ it y , l anthan - ox y -
carbonat es have to be f ormed b y the ads orpti on of CO 2 on the c atalyst s urface. T he insertion of
CO 2 m olecules d opes the topla y er str ucture o f th e ca talyst m aterial. T hat w as proven in
experim ental s tudies b y F ourier tra nsf orm inf rared spe ctrosc opy ( FT IR ) . [26,27] Howev er , ther e
is a deb ate abo ut the involvem ent of several oxyge n species on the catal y st s urface, w h ich
contribu te to th e OCM surf ace reac tion net work for that m aterial, t oo . The f irst di scuss ion about
m ethane interac tion with d iffer ent oxygen int erm ediates w as prese nted by Ka liaguine et a l. [28]
All discussed ox y g en spec ies on t he catal y s t sur face a re present ed in equ ation (2 ).
O 2
(g) O 2
(ads) O 2
- O 2 2- O - O 2-
+ e - + e - + e -
(2)
For lant hanum oxide it was propos ed that superox ide species (O 2
- ) are r espons ible for s elective
m ethane activat ion. [29] La combe et al. discuss ed the c ontribution of the “ w ork ing sur face” f or
the La 2 O 3 c atalyst. [30] T hey c onclu ded that especial ly the form ation of dee p oxidatio n products
is hig hly sens itive for l ow coor dinated s urfac e sites. T he origin of low co ordinat ed surf ace sites
is the selecti ve acti v ation of m ethane. Ther efore a s electiv e activati on of m ethane cr eates th e
active s ite for the uns elect ive reactio n. Palm er et al. dem onstrated b y DF T calc ulation s for a
lanthan um ox ide clust er that O - is the m ost acti ve ox y g en spec ies for sele ctive m ethane
Introduct ion
4
activati on. [ 31] Unf ortunatel y , O - is hard to d etect in Electron P aram agnetic Res onance (EPR )
spect roscop y s tudies , bec ause of its insta ble n ature in c om parison to the s uperoxo - and
peroxo - in term ediates. [32 ] Ferreira et al. d em onstrated b y do ping of the l anthanum ox ide with
ceria tha t the act ivity and selectivit y is enhanc ed in c o m parison to t he stand ard lantha num
oxide. I n their X - r ay p hotoelec tron spec troscop y ( XPS) s tudies the y propose d that O - is
respons ible for such fin ding s. It has t o be noted , that th e O - spec ies is pr eferab ly bound close to
the ceria ions. [33] A s imilar enha ncem ent eff ect was observe d by dopi ng of the MgO catal y s t
wit h iron ions . [34] I t was dem onstr ated through a n exper imental and theor etical s tudy that the
ox y ge n adsor ption o n do ped surf ace areas of C aO wit h mol y b denum follows t he react ion
pathwa y s ac cording to e quation (2) . [35] T hus, the in volvem ent of several ox ygen spec ies on
the cat alyst sur face of various active m ateria ls in th e OCM reac tion is an im portant as pect.
Althou gh the La 2 O 3 s ys t em allows for a s table perf ormance in tim e on str ea m experim ents, the
C 2 yi e ld is lim ited . [36]
Another well - know n high perform anc e catal y st m aterial for th e OCM r eactio n is t he
Na 2 WO 4 /Mn/SiO 2 s ystem . The act iv e p hase cont ains Na 2 WO 4 /Mn and is depos ited on a s ilica
support m aterial. T hat syst em was f irst pres ente d by Fang et al. [37] It was r eported b y d iff eren t
groups tha t the m aterial ha s a stable perf orm ance for several hun dred hours in tim e on str e am
experim ents. [38 – 4 1] Furth erm ore , a m aterial benchm ark was c arried out in a n OCM m ini - plant,
perf ormed in a f luidized b ed reactor, with dif ferent d oped lanth anum oxide sys tems and the
Na 2 WO 4 /Mn/SiO 2 c atal y s t. [42] In that s tudy it w as re ported that the Na 2 WO 4 /Mn/S iO 2 catal yst
sho w ed the best o verall perf ormance. This c ataly st is m ostly prepar ed b y inci pient wetness
im pregnation m ethod s , be cause of less exper imental ef fort c ompar ed to the so l - gel a nd t he
m ix ed slur ry m ethod. All p reparat ion methods sho w ed c ompar able catal y tic perf ormanc e and
stabi li t y . Major phases , f ound in X - ra y d iff raction spectr oscop y (XR D), are Mn 2 O 3 , Na 2 WO 4 and
α - cr istobal ite ( SiO 2 ). [43] Unfor tunatel y , t h is ca taly st has several com plex mater ial proper ties
and trem endous depende nc y o n com pound com position. It was report ed th at the ph ase
transiti on of the support m aterial into α - cris tobalite is induced b y Na + ions, which is absolut ely
necess ary f or a high per form ance of this catal y s t mater ial. [4 4] As cons equ ence t here is a
dram atic shr inkage of the specif ic surf ace area (SS A). Further m or e , t he active ph ase
(Na 2 WO 4 /Mn) undergoes surf ace recons truction an d has strong interac tions bet w e en the
support and th e trans ient m etal oxi de. [45 ] The h ighest catal y t ic p erform ance was report ed for a
com position of 5 wt - % Na 2 WO 4 and 2 wt - % M n(II). [39 ,45,46] K ou et al. st udied this cata lyst with
X- r ay absorpt ion fine str uctur e (XAFS) an d XPS. [4 7] T hey re porte d that the s urfac e of the f resh
catal y s t m aterial, enric hed b y ox ygen dur ing a c alcination proced ure, exists as an am orphous
phase which c ontains d ispers ed tungstat e and manganes e oxides. T heir findings were
conf ir m ed by Gho lipour et al., who carr ied ou t a variatio n of the s odium ion b y differ ent alk ali
m etal ions. [48] T hey concl uded that the alkali ion enabl es the str uctural f lexibilit y of the ac tive
phase on the c atalyst surf ace, which is high ly ac tive and selectiv e. Nipan calculate d a
theoret ical phase di agram , whi ch c onsiders se veral phases of the Na 2 O– WO 3 – Mn 2 O 3 – SiO 2
s y st em and its abil ity to f orm a quas i - li qu id acti v e su rfac e on the sup port m aterial. [49] T his
Introduct ion
5
aspect ena bles th e abi lity of an ox ygen sp illo ver reac tion betw een Mn 2 O 3 and Na 2 WO 4. In
princi ple, that re action is a redox m echanism between Mn( III+) an d W ( VI+) . [5 0,51] S uch spi ll -
over ef fec t enhances t he OCM ac tivity of the Na 2 WO 4 /Mn/ SiO 2 s y st em m ore than t hree t imes
com pared to the Na 2 WO 4 /Si O 2 s ys t em , which cann ot perf orm suc h a reaction . T he ox y gen -
spil lover eff ect is very sensiti ve to the m anganese ox ide oxidat ion state , w hic h s h ows also the
highes t activit y i n com pari son with other tr ansient metal ions. [52] I t was repor ted that al so a
hig h dispers ion of N a 2 WO 4 and Mn on t he supp ort m aterial en hance th e cata lytic ac tivity. [53] A
better d istributi on of these compounds could impr ove the reaction ra te of the oxygen - s pillo ver
reaction .
1.2.3 Reaction ne twork and kinetic mode ls
Lunsf ord and Cowor ker s prove d t hat the eth ane form ation proces s un dergo es m ethyl rad ical
coupl ing and m any groups contr ibuted to the de velop ment of a surf ace reaction network s with
big eff ort for several catal yst mater ial s. [54 ] One of th e firs t m echanism s for select iv e m ethan e
activati o n was post ulated by Luns ford and Co work ers for the Li/MgO catalyst. [55] The y
adapted t he assu m ption of Ito et al., who propos ed (Li + O - ) as active c enters. [ 56] T he center is
involve d in t he hydrog en abstr action reac tion of a methan e m olecule, w h ich is then co nverted
into 2( Li + O - )H an d fina lly r educed to Li + O 2- . T hereafter , it is reox idized by gas phase oxygen.
Sinev d iscus sed sever al mec hanisms for C - H bond cl eavage o n the cat alyst sur face [5 7] and he
discus sed the reoxi dation pr ocess of the active sites in detail. He pr oposed th at the ox idative
deh y dr ogenati on in s tead y state exper iments of the form ed hydrox yl groups is the dom inant
reoxidat ion step. Korf et al. sugges ted th e particip ation of surf ace carbonates , whic h co ntri but e
to the se lective ac tivatio n of m ethane. [25] Hutchings and Co worker s studied M gO, Li/M gO and
Ag /Al 2 O 3 cat al y s ts material s by use of dif ferent oxidant age nts. [58] T h e y s ugges ted that O -
radic al io ns are th e oxygen species f or selec tive m ethane activati on , but due to a ve ry sho rt li fe
tim e under O CM condit ions the anal y t ical pro ve of th is spe cie s is d iff icult . Fur therm ore , the y
could sh ow that t wo active oxygen s pecies on t he Li/ MgO surf ace contribu te to selective and
unsel ective m ethane act ivation. [ 59]
As d iscu ssed ab ove , the c ontribution of gas p hase reac tions and the pres ence of several
ox y ge n interm ediates o n the cata lyst surf ace inc rease the exper imental ef fort and com plicate
the c alcul ations of reliabl e k inetic data . Base d on th e availab le data a variet y of s urfac e reac tion
network s and k inetic m od els ha ve been de veloped . [60 – 65] One w ell - known micr o kinetic
network was pres ented b y Sun and T hybaut for a Li/MgO catal y s t, which consi ders gas ph ase
reactions and reac tions on the catal y s t sur face . [6 6] A comparable m odel w as pre sented by Lee
et al. f or the Na 2 WO 4 /Mn/ Si O 2 catal y st. [6 7] B oth s urf ace reaction m odels ar e present ed in
T able 2 . It i s shown that both network s f ollow s im ilar reacti on pathw ay s despite the diff erent
nature of the c atal y st m aterials . Gas p hase oxygen i s adsorbe d by diss ociati on on a free ac tive
site ( V *) of t he ca talyst sur face a nd is con verted in t o a n activ e ox y ge n spec ies (O *) . Methane is
activate d b y C - H bon d cle avage by t hat ox y gen spec ies and m eth y l radica ls are pr oduced .
Further alkanes and alkane s are a ctiv ate d i n a si mila r reac tion s tep , wh ich pr oduces radi cal
interm ediates. Meth y l r adic als and CO ar e fur ther oxidi zed to alk oxides an d CO 2 .
Introduct ion
6
T able 2: Com parison o f reaction n etwork s on the c atal y s t surf ace of Li/MgO and
Na 2 WO 4 /Mn/SiO 2 catalyst
No.
Lee et a l. [67]
Sun et a l. [66]
1
O 2 + 2 ∗ V ⇄ 2 O ∗
O 2 + 2 V ∗ ⇄ 2 O ∗
2
CH 4 + O ∗ ⇄ CH 3 ∙ + OH ∗
CH 4 + O ∗ ⇄ CH 3 ∙ + OH ∗
3
C 2 H 4 + O ∗ ⇄ C 2 H 3 ∙ + OH ∗
C 2 H 4 + O ∗ ⇄ C 2 H 3 ∙ + OH ∗
4
C 2 H 6 + O ∗ ⇄ C 2 H 5 ∙ + OH ∗
C 2 H 6 + O ∗ ⇄ C 2 H 5 ∙ + OH ∗
5
2 OH ∗ ⇄ H 2 O ∗ + O ∗
6
CH 3 ∙ +3 O ∗ ⇄ HCO ∙∗ + 2 OH ∗
�
CH 3 ∙ +O ∗ ⇄ CH 3 O ∙∗
7
CH 3 O ∙∗ + O ∗ ⇄ CH 2 O ∗ + OH ∗
8
CH 2 O ∗ +O ∗ ⇄ HCO ∙ ∗ + OH ∗
9
HCO ∙ ∗ +O ∗ ⇄ CO ∗ + OH ∗
HCO ∙∗ +O ∗ ⇄ CO ∗ + OH ∗
10
CO ∗ + O ∗ ⇄ CO 2 ∗ + V ∗
CO ∗ + O ∗ ⇄ CO 2 ∗ + V ∗
11
CO + ∗ ⇄ CO ∗
12
CO 2 + ∗ ⇄ CO 2 ∗
13
2 OH ∗ ⇄ H 2 O+O ∗ + V ∗
4 HO 2 ∙ → 3 O 2 + 2 H 2
How eve r, the c ontribut ion of s everal oxygen interm ediates on the cat al y st surf ace was
propose d in recen t experim ental w or ks , by us e of diff erent oxi dizing agent (O 2 and N 2 O). [68,69]
It was d em onstrated b y us e of diff erent oxi dizing agents, w hic h were n itrous oxide and
m olecular ox ygen that the cata lytic activit y is dec reas ed but C 2 selectiv ity is dra stically
increas ed. I n dynam ic ex perim ents, the use of C O 2 as oxidant for the Na 2 WO 4 /Mn/SiO 2 c atal y s t
also s hows extrem ely high C 2 se l ectivity of about 95 % at 5 % m ethane c onversi on. [70]
Com parable ef fect s were sho wn by Tak anabe an d Iglesi a, who c o - fed w ater to th e reactant
m ixture at the reac tor inlet. [71] The y poi nted out th at also OH • r adicals c ontribut e as selecti v e
oxida nt s to the se lective m ethane activat ion on th e catal y s t sur face. B eck et al. dem onstr ated in
tem poral anal y s is of produ cts (T AP) experim ents that weak ly adsorbe d oxygen interm ediate s
con tribu te to the unselec tive activ ation of m ethane on M gO and Na 2 WO 4 /Mn/SiO 2 cataly sts,
while a strong ly bo und ox ygen spec ies is resp onsible f or the select ive m ethane ac tivation. [ 72]
T he contrib ution of sever al ox y g en inter mediates to th e catal y s t act ivit y is wel l - known f or
s y st ems as form aldehy de or eth y l ene oxid e produc tion. [73 ,74] , but not for t he OCM reac tion.
T herefore , experim ental metho ds are r equired to in vestigat e the co ntributi on of all ox y g en
species on the cata lyst s urface . Further mor e, it is necess ary to unders tand ho w all t hes e
species are form ed and interact w it h other c om pounds during the OCM r eactio n.
1.3 E ngineering aspects
1.3.1 Proc ess eng ineering asp ects
As m entioned in th e last s ection , o ne important aspect of the OCM react ion is the s trong
exotherm ic r eaction ent halp y and the c ontribut ion of gas phase re actions which leads to th e
form ations of hotspots . S e veral exp erim ental techn iques exist to av oid hot s pot f ormat ion or
handle the hotspot t em perature durin g OCM r eaction . Kooh et a l. opt imized th e stead y state
condit ions f or C 2 y ield ma ximiza tion with a La 2 O 3 cataly st , dea ling with h ot spot f orm ation . [75]
Introduct ion
7
T hey f ound that hig h linear space velocit ies and hig h temper atures are benef icial for the OC M
reaction . Such co nditions avoid back m i x ing of C 2 products to the cat alyst bed, which can
activate t hese pro ducts ag ain. Further oxidat ion react ions woul d lead to the f orm ation of deep
oxidat ion produc ts. Furth erm ore , ethane sh ould be quick ly deh y dr ogenate d to et hene, beca use
in com par ison to ethane it is mor e resistant ag ainst oxi dation re actions. W olf et al.
dem onstrated t hat a lso th e heat c onducti vity of the c atalyst bed is essent ial t o rem ove the
reaction heat from the cat alyst zone. [76 ] The c om plex inter play of gas p hase an d catal yst
surf ace reac tions is hig hly i nfluenced by the tem peratu re prof il e in the OCM r eacti on. T herefore ,
an up - s caling of an OCM r eactor is ver y c hallen ging. Ho wever, there ar e sever al publica tions
which pres ent proc ess en gineering as pects suc h as m odeling of reactor c oncepts and up -
sc aling . [40, 77 – 79] Furt herm ore , large am ounts of the catal y s t are nee ded fo r an industr ial
relevant scal e. An u p scal ed s ynthes is m ethod for large bat ches of Na 2 WO 4 /Mn/SiO 2 cat alyst s
was pr esented b y Sim on et al. [41] T he catal y s t was m ade by spra yed im pregnation i n a
fluidize d bed r eactor and s howed s table perf orm ance in t ime on st ream experim ents. T he
cluster of exc ell ence “U nif y i ng conce pts in catal y s is ” (UNICAT ) pres ented the succ essf ul
devel opment an d operat ion of an OCM mini - pla nt. [80] The OCM react ion was perform ed in a
fluidize d bed re actor w ith Na 2 WO 4 /Mn/SiO 2 and L a 2 O 3 /CaO cata lysts ach ievin g a C 2 y ie ld of
19 %. [42]
Another im portant aspect o f the O CM react ion is the s eparatio n process of the product s tream .
Godini et al. anal yz ed d iff erent separat ion conc epts a nd m ade cost es tim ations of a h y p othetica l
plant whic h pr oduces 240 000 t on s et hyle ne/year. [81 ] One of the m ost cos tly op erations is the
separati on of eth ylene from ethane, which is caus ed by ener gy intensive cr y o geni c rectif ication.
Severa l separ ation proc esses w er e discuss ed to se parate th e uncon verted m eth ane and d eep
oxidat ion produc ts fir st. That would r educe t he am ount of the total product s tream for c ry ogen ic
rectif ication dram atic ally . A lso th e adsor ption of m ethane, h y drog en and CO dir ectly aft er th e
OCM reac tor im proves the oper ation of that plant , w i th decre asing of the operati on costs .
Another idea for the s eparation of eth y l ene is a ch em ical reaction to co nvert it direct ly to
pol y m ers or to Ethy lbenz ene , which was present ed as the ACRO or OX CO pr oc ess . [82,83]
Besides these as pects, d iffer ent types of reac tor concepts were test ed to overc ome the
lim it ations of t he C 2 yield. Krug low a nd C o workers presented the operati on of a c ountercurr ent
m oving bed reactor . [84] T he idea of this concept is to rem ove C 2 produc ts from the feed s tream
during inc rease of methane c onversion . That was carried out b y reaction colum ns and
separati on colum ns. I n the fir st stage sm all am ounts of m ethane were con verted a nd C 2
products were sep arated i n the separ ation colum n. The rest of the feed was send to the nex t
reaction colum n , which pr oceeds the co nversion of m ethane. T he C 2 prod ucts of that ru n were
separate d again and the rest of the f eed wa s s end t o the next sta ge an d so on . The y wer e able
to reach m ethane c onversion s of 75 % with a C 2 y i eld of 55 %. Hu et a l. pres ented a du al
catal y s t, fix ed bed react or for a si m ultaneous pr oduct ion of ethen e and s y ngas . [85] T he yield
f or CO + C 2 H 4 was 18.5 % . God ini et a l. presen ted a m embrane react or conc ept f or the OCM
reaction . [8 6] Ox y g en is do sed in sm all am ounts thro ugh a poro us c eram ic reactor wall to th e
Intro duct ion
8
catal y s t fix ed bed sect ion, w hich is const antly fed with m ethane. T hat techni que all o ws
encoun tering gas p hase re actions, induced b y gas phase ox ygen, and 2 5.5 % C 2 yield were
reported .
Recent ly, Silur ia T echnolo gies pres ented a f ull oper ationa l plant, which is aim ed to co nvert
natura l gas into liq uid hydroc arbons econ om ically, perf ormin g the O CM react ion even wit h
lim ited C 2 yi e l d. [87 ] Howe ver, their process concept depends strongl y on ver y cheap m ethane
prices .
1.3.2 T he chem ical looping process concept
T he first publicati on abo ut c hemic al looping was report ed by Hurst in 1939 . [88] T he
Messer schm itt W assers toffgas - Gener ator was the pioneer setup to pr oduce h ydrogen by use of
steam and iron b y t hat con cept . Furth er deve lopment of that tec hnolog y offer s several t y pes of
applicat ion. The m ain r esearch f ields are chem ical loopin g com bustion (CLC) and Chem ical
loopin g reform ing (CLR). [8 9] The genera l proc ess conc ept is presen ted in Fig ure 1 . A r educed
ox y ge n carrier m ater ial Me X O Y-1 , t y p ica ll y a tr ansiti on metal ox ide, is oxid ized by air in t he firs t
reactor (Me X O Y ) . T he oxidized met al oxide is th en used in a s econd r eacto r to oxid ize a
reactant. The r educed oxygen c arrier is rec y cl ed bac k to the f irst reac tor, w her e it is ox idized
again. T here are m odified reac tor concep ts, w hich handle differ ent problem s durin g app lic ati on.
[90] T ypical c onfigur ations are two i nterconnec ted m oving or fluidi zed bed r eacto rs, altern ating
fix ed bed , fluidized bed rea ctors or rotating r eactors .
Figure 1: Sc heme of the c hem ical looping process concept
Introduct ion
9
T he CLC focus ses on energ y produc tion b y com plet e m ethane c ombus tion to CO 2 . T he
advant age of t hat tec hnolo gy, com pared to standar d co - feed of m ethane and ox y gen, is the
capture of CO 2 witho ut further separation proc ess es aft er the combus tion reactor . Further m ore ,
NO X emis sions are dras ticall y reduced, becaus e no nitr ogen is present in the com bustion
reactor . U nfortunat el y , the inves tment cost of a CL C plant is 71 % high er, c ompar ed to a
class ical natura l gas po wer plant. [9 1] The bi ggest C LC plant is operated by the T echnisc he
Univers ität Darm stadt with a nom inal po wer of 1 M W th . [92] T he CLR dea ls with t he oxidat ion of
m ethane to syngas . The m ain advant age of s uch applicat ion is th e heat ba lance , because the
oxidat ion react ion of the ox ygen carrier is exoth erm ic and the reform ing proces s endother m ic.
[93] T he ener gy of th e oxid ation re action is stored mos tly in th e ox y ge n carri er and tra nsporte d
to the r eform ing reactor, which a llows a g ood en ergy integr ation co ncept.
One of the m os t i m portant param eters in chem ical looping s y s tem s is the amount of stored
ox y ge n, which is tr anspor ted by the ox ygen car rier. T he ox y ge n transp ort capab ility R O is
descr ibed b y equat ion (3) , where m o is th e m ass of the oxidi zed carr ier m aterial an d m r is t he
m a ss of the reduce d one. The h igher t hat valu e is th e m ore ox y gen c an be s tored i n the ox y ge n
carr ier m aterial. Comm on ox ygen carri er m aterials were re viewed a nd d iscuss ed by Ad anez et
al. and Li et a l . [89,94 ] A high spec ific s urface ar ea (SSA) is r equired to prov ide enough
adsorpti on sites for gas phase ox y gen. F urther mor e , the oxidat ion and re duction rate shou ld be
high f or suf ficient res idenc e times in both re actors . In Addit ion, a h igh ava ilabilit y , low toxic ity,
good m echanica l strength and s uitable h eat capac ity of the ox y ge n carrier are als o necessar y.
[95]
R O = m O − m r
m O
(3)
T he mos t relevant ox y g en carr ier system s and their o xygen tra nsport cap abilit y are pr esented in
T able 3 . It has t o be not ed that the R O and the redox propert ies are s trongl y influence d b y the
support m ater ials. [96]
T able 3 : Most re levant ox ygen carr ier m aterials and t heir ox y gen transpo rt capab ility in
chem ical loopin g exper iments [89]
Redox s ystem
R O (-)
NiO / Ni 0.210
Mn 2 O 3 / MnO
0.100
Mn 2 O 3 / Mn 3 O 4
0.034
Mn 3 O 4 / MnO
0.070
Fe 2 O 3 / Fe 3 O 4
0.034
Co 3 O 4 / CoO
0.067
CuO / Cu
0.200
Another important aspect is the sup pression of cok e formation o n the carri er part icles . Ty p i cally
the or igin of cok e is the dispro portion of CO, ac cording to t he Boudo uard reac tion in
equatio n (4) .
Introduct ion
10
2 CO ⇌ C + CO 2
(4)
Ishida et al. dem onstrated t hat a ver y low conc entrat ion of water vapor in the re ac tant feed c an
com pletely suppr ess the f orm ation of c arbon cok e in CLC. [9 7] T he reason is the water gas shif t
reaction (Equat ion (5) ), whic h act as a co - oxida nt for the CO ox idation.
CO + H 2 O ⇌ CO 2 + H 2
(5)
Besides the produc tion of deep oxidat ion prod ucts and e nerg y or hydrogen , ot her gro ups
applie d the chem ical loopin g concept f or partial oxidation reac tions su ccessfully . T h e
challeng ing situ ation in th e product ion of eth y len e oxide is the prese nce of s everal ox y g en
species on the c atalyst s urface, c om parable to t he OCM s urface r eaction ne twork , as discus sed
in the prev ious s ection. [74] P ark and Cowork ers dem onstrated that t he cyc li c operat ion
im proves the se lectivit y to eth ylene oxi de dram aticall y. [98] I n another work the periodic
operat ion was ap plied for t he produc tion of m aleic acid f rom butadie ne. [99] Sim ilar sele ctivi ty
im provem ents for m aleic acid were repor ted by Lang e t al . [ 100]
T he cy c lic op eration of the ox idative deh ydrogen ation of propane was prese nted by Cre aser et
al. [101] T he y co uld dem onstr ate that the d y n am ic operation of the reactor for pr opene
form ation enhanc es the y i eld sig nific ant ly . On t he one hand that is c aused b y the avoi dance of
any gas ph ase reac tions , i nduced by oxygen . O n the other hand uns elective s ite reac tions on
the ca talyst m ateria l were m ini m ized. [1 02] T heir exp erim ents were ver ified b y Bal laniri et al.
who prepare d a variet y of supporte d vanadia cat al y s t s for this reaction, enhanc ing t he prope ne
y ield b y a f actor t wo . [103] T heref ore, the chem ical loo ping s y s tem seems to be a s uitable
reactor concept for the ox idati ve couplin g of m ethane , t oo .
1.4 Outline of this thesis
As disc ussed bef ore, the c ontributi on of gas p hase ox y g en an d several s urface o x y gen s pecies
to the pr oducti on of eth y le ne in the OCM reaction de creas e the C 2 yield . Such chal lenge co uld
be enco untered b y transie nt and c y c lic o peration ex perim ents. The Na 2 WO 4 /Mn/S iO 2 catal y st is
a we ll - k nown s y s tem , which pro vides lo ng tim e stabilit y and hig h activ ity. Further m ore, it can be
produce d in large scale s b atch. F inall y , m anganese o xide is well - k nown for its ox ygen stora ge
capacit y in diff erent c hemical loop ing s tudies and res ults from earlier pres ented transi ent
experim ents of the Na 2 WO 4 /Mn/SiO 2 c atal y s t sho w prom ising acti vity an d superi or C 2 select ivity
in com parison with co - feed experim ents. [1 04] It should be noted that the c ata ly s t mater ial has
to f ulf ill two r oles . On the o ne hand, it serves as oxyg en carrier mater ial, sim ilar as in the CL C
and CLR proces ses. On t he ot her hand, it f ulfills the role as a c ata ly st m ateria l, becaus e stored
ox y ge n on the catal yst su rface is c onverted to wate r by m ethane ac tivation, t o form m eth y l
radic als .
Paper 1 analyses the th erm al reaction sens itivit y of the OCM reac tion in c ase of hot s pot
form ation during s tead y s tate experim ents. Furtherm ore, the inf luence of gas phase and s urfac e
reactions in dif ferent ex perim ental and theoretica l stea dy stat e reactor s were investigat ed.
Methods and m aterials
11
Pape r 2 deals with the exper imental i nvestigat ion of the Na 2 WO 4 /Mn/SiO 2 cat alyst m aterial in
tem p oral anal y s is of produ cts (T AP) experim ents in absenc e of gas phase ox ygen. Further , the
catal y s t act ivity was studi ed in st eady stat e exp erim ents at pres sures u p to 10 bar and
calcula tions ab out reac tion eng ineerin g aspects of the OC M reaction in a stead y s tate f ixed bed
reactor w ere carr ied out.
Pape r 3 de als with temper ature pr ogramm ed and d y n am ic exper iments to study t he interac tion s
between OCM reacta nts such as CH 4 , C 2 H 6 an d C 2 H 4 i n abs ence of gas ph ase ox y g en with t he
Na 2 WO 4 /Mn/SiO 2 cataly st sur face . Addit ionally, th e ab ility to s tore and provi de strongl y bound
ox y ge n for the O CM reac tion un der non - s teady sta te cond itions was in v esti gated.
Paper 4 descr ibes the de velopm ent and oper ation o f a lab sc ale sim ulated c hem ical looping
setup f or the OCM pr oces s. Yield b oundar y s tudies in chem ical looping ex perim ents were
carr ied out an d com pared with s tandard stead y s tate experim ents. Based o n thes e results a
desig n proposal was m ade f or a hypothe tical chem ic al loop ing OCM plant.
Paper 5 investigat es the role of the c atalyst com pounds of the Na 2 WO 4 /Mn/SiO 2 m aterial . The
foc us is the f inding of a corr elation be tween t he acti ve com pounds an d the ox y g en s torage
capacit y. Furth erm ore , the inf luence of the catal yst ’ s SS A on the O CM reac tion in ch emic al
loopin g experim ents w as studied.
2 Methods and ma teria ls
2.1 Simulation of the OCM gas phase rea ctions and cataly st surf ace reac tion s
T he si m ulation of the O CM reac tion was perform ed in Chem kin ( Reaction Design / A ns ys) . For
the gas phase r eactio n mec hanism the Doole y m echanis m was us ed. [1 1] One of the k ey st eps
in OCM reac tion on th e catal y st sur face is the C - H bond cl eavage of m ethane , whic h i s the
origin f or me thyl rad icals that b y coupl ing f orm ethane . T he sim pl e st approach t o im plement a
chem ical reac tion, induc ed b y a c ataly st m aterial , is present ed in
T able 4 . A vacanc y (V *) on t he ca taly st sur face is fil led by dis sociati ve adsorpt ion of gas phase
ox y ge n. Met hane is activated in the selecti ve wa y by m ethyl radic al for mation. F inall y , the
form ed OH gr oup on the c atalyst sur face is rem oved by the f orm ation and desorptio n of water
m olecules. T hrough lack of reliable m icro k inetic data f or the Na 2 WO 4 /Mn/Si O 2 c atalyst, t he
k inetic param eters w ere set to the gener a ll y accepted fac t , of fast c atalyst reox idatio n and the
m ethane activat ion as rate limiting s tep. T he strong im pact of gas phas e activit y on prod uct
selecti vity was inves tigate d b y t he select ion of dif ferent i deal reac tors as the plug f low tub e
reactor (PFT R) with no b ack m ixing ef fects and the c ontinuous s tirred ta nk r eactor (CST R),
whi ch i s com pletel y m ixed. A detailed implem enta tion to R eaction Design Chem kin was
presente d in th e m odel -b ased a nal ysis part of Paper 1 .
Methods and m aterials
12
T able 4 : Sem i - empiric al catal y st reacti on netw ork of the OCM reac tion ( Paper 1 )
No .
Reaction step
k oo (1/s)
E A (kJ/mol e)
1
O 2 + 2 V ∗ ⇄ 2 O ∗
3· 10 6
40
2
CH 4 + O ∗ ⇄ CH 3 ∙ + OH ∗
3· 10 5
120
3
2 OH ∗ ⇄ H 2 O+O ∗ + V ∗
3· 10 6
40
2.2 Cataly st prepar at ion
T he referenc e Na 2 WO 4 /Mn/SiO 2 c atal y s t was s ynthesi zed and char acterized b y
Dr .- Ing . U lla S im on (Grouplea der: Pr of. Alek sander G urlo) . T he experim ental m ethods are
descr ibed in [41 ] . T he final catal y s t had a com positi on of 5 wt - % Na 2 WO 4 and 2 wt - % Mn( II)
ions. T he SSA was 1.86 m²/g and the p article s ize was 150 - 350 µm . This cata lyst was used in
all of our exper iments w hic h are prese nted Paper 1 - Paper 4 , and also as refe rence cataly st
m aterial in Paper 5 .
All other prese nted Na 2 WO 4 /Mn/SiO 2 c atal y s t s , suppor ted on C OK - 12, in Paper 5 we re
s y nth esized and ch aracter ized b y M aria G racia Col m enares (G roupleader: Pr of. Alek sander
Gurlo) . The X RD charac terization of the spe nd catal y s t s w er e carried o ut by Dr. - Ing. U lla
Sim on. It has to be note d that COK - 1 2 is a m esopor ous s il ica m ateria l.
2.3 Nitrogen adsor p tion
Nitrog en adsorpti on at 77 K is a w el l - know n techniqu e to determ ine the SSA of s olid mater ials,
accor ding to the theor y of Brunauer, Emm et , and T eller (BET) [105] . T he catal yst s were
degassed a t 120 °C and 0.15 m bar for 60 m inutes. T hereafter, th ey w ere anal yz ed by a
Microm eritics Gem ini III 2375 Sur face Area Anal yzer.
2.4 High pressu re experiments
A schem e of the high pres sure reac tor is pr esented in Figur e 2 . T he catal yst (100 m g) w as
place d on an inert be d of quartz gran ules i nside of a corund um inl a y . T he inla y was sur rounde d
by a ste el jack et to operat e at press ure up to 1 0 bar. T he feed was 9 5 vo l .- % methane and
5 vol% ox ygen wit h an overa ll gas f low rate bet ween 50 to 5 00 Nml / min and a cons tant
residenc e tim e of 0.16 s in eac h run. The r eaction temper ature was bet ween 70 0 – 8 00 °C and
the prod ucts were analy zed with a gas chrom atograp h GC - 20 1 4 b y Sh imad z u. More detai l s are
given in the experim ental p art of Pa per 2 .
Methods and m aterials
13
Figure 2 : Schem e of the high press ure re actor
2.5 Mass spectrometry
A schem e of the mas s spectrom eter (InProc ess Instr um ents, GAM 200) is pres ented in F i gure
3 . The g as m ixture f rom the r eactor out let w as se nd t o Inlet 1 of the m ass spec t r ometer (MS).
A sm all amount of the g as was s end con tinuous ly through a capi llar y by a r otary van e pum p,
while th e rest w as sent to t he vent. T he gas sam ple was ionized b y electro n impact io nization,
utilizi ng yttriat ed yttrium f ilament. A c ascade of a turb omolec ular pump, a d iaphr agm pu m p and
an ext ernal rotar y van e pum p generate s a pressure of 10 -6 m bar. The ion ized fr agments passed
a quadr upole, w hic h deflec ts the fr ag m ents by their spec ific ratio of Lorent z f orc e to inertial
for ce. Finall y a c hanneltro n s econdary elect ron m ultiplayer (S EM, 1200 V) was us ed to enhanc e
and an alyze the com position of the f ragm ents. T he softw are for c ontrol of the M S and
calcula tion of the com pound concen tration by bal ances of t he f ragm ent com position was
provide d b y In Process Instrum ents. The deta ils about compound c alibrati on and data r ecord ing
are pres ented in the exper im ental section and in th e support ing inf ormation of Pa per 3 .
Methods and m aterials
14
Figure 3 : Schem e of the quadrup ole m ass spec trom eter Model GAM 200 (InProc ess
Instrum ents )
2.6 T emperat ure p rogrammed surface reaction s
T he experim ental conc ept of a t em perature pro gramm ed sur face reac tion (T PS R) is pres ented
in F ig ure 4 . In a f irst ste p the cat al y s t was he ated u p (10 K/m in) under a flo w of an oxidi zing
agent ( e.g. synthetic air) . Aft er the oxidat ion treatm ent, the c atalyst was c ooled do wn again. T he
reactor w as purged b y an iner t gas ( e.g. N 2 , He ) t o rem ove ga s ph ase and weak ly bound
ox y ge n spec ies. T hereafter , a reactant ( e.g. hydrogen, m ethane…) was fed to the reactor, while
the cata lyst m aterial wa s heated up with a def ined tem perature r amp (1- 5 K/m in). T he
com pos ition at the outlet of the reactor w as cont inuo usly recorde d b y a det ector (e.g. therm al
conduct ivity d etector (TCD ), MS ). T he exact procedu re f or all T PS R ex perim ents and further
inform ation, e.g . tem perature pr ofile of the furnac e , are gi ven in P aper 3 and Pap er 3 - SI .
Methods and m aterials
15
Figure 4: Exper im ental concept of tem perat ure program med s urface reac tion (TPSR )
experim ents
2.7 Dy namic exper i ments
Dynam ic ex periments follow a s imilar experim ental c oncept as presented in th e last subsecti on
( 2.6 ) . T he catal y s t mater ial (1 g) was h eated u p to reac tion cond itions in a f low of an oxidi zing
agent ( e.g. s y n thetic air) . After reaching t he desir ed r eaction tem perature ( 700 – 800 ° C) , the
reactor was purge d by an inert gas (He) k eeping the tem perature cons tant. The reac tion was
carr ied out b y an im mediate rep lacem ent of the inert gas by the re actant gas ( e.g. met hane,
ethane, ethe ne) . T hat was realize d b y opening and c losing of s witching valv es (S wagel ok SS -
41 S2) in f ront of the g as suppl ies. Again, t he compos ition at the outlet was contin uousl y
m easured b y M S. The experim ental m ethod and al l param eters are presented in the
experim ental s ection of P aper 3 . T he sel ectivit y of the produc ts was deri ved accor ding to
equatio n (6).
S i = ν i ⋅ A Produc t , i
∑�ν i ⋅ A P roduct , i �
(6)
Further mor e , the con verted amount of stored oxygen n O* was der ived acc ording to t he ox y g en
balanc es given in the rat e equat ions (7) - (9) and final ly by equati on ( 10 ).
2 CH 4 + O ∗ → C 2 H 6 + H 2 O
(7)
2 CH 4 + O ∗ → C 2 H 4 + 2 H 2 O
(8)
CH 4 + 4 O ∗ → CO 2 + 2 H 2 O
(9)
n O ∗ = p ⋅ V
⋅ ∑�ν O ∗ , i ⋅ A i , product [% ⋅ s] �
R ⋅ T
( 10 )
Methods and m aterials
16
In addit ion , a m olecule specif ic oxygen ba lance was calculated ac cord ing to the react ion
network presente d in Fi gure 5 a nd the fol lowing as sum ptions :
• Ther e is no m ass trans port eff ect for the s tored am ount of oxygen
• All react ions ar e second order
• The r atio of k 3/k4 is betw een 2 an d 7 acc ording to the f indings of L unsford a nd
c owork ers [106,107]
For instanc e, for each dete cted ethen e m olecule, one O C2H6 and on e O C2H4 were converted t o
water. B ased on thes e ass u m ptions th e firs t four methane pu lses in re petitive p ulse
experim ents were f itted. Bas ed on th e fitt ing par ameter s (k1 – k 4) the r epetitive pul se
experim ent was simulated (15 m ethane pulses ) and the am ount of converte d oxygen was
determ ined ( Pap er 5 - SI )
Figure 5 : Ass umed OC M s ur face r eaction network of the Na 2 WO 4 /Mn/Si O 2 catal y st f or
m olecule specif ic oxygen balanc e ( Paper 5 )
2.8 Simulated chemical looping
T he adaption of the chem ical loo ping proc ess conc ept to the OCM reaction with
Na 2 WO 4 /Mn/SiO 2 catalyst s was reali z ed b y t wo pneu m atic six port valves . T hese va lves a llow
t he dosa ge of defined a m ounts (0.25 – 2 m l) of reactants to the catal yst fix ed bed. T he
sim ulated chem ical loop ing conc ept was c hosen to prevent c omm inution or abrasion of the
catal y s t par ticles whi ch is a sig nif icant problem in f luidized bed r eactors . A f unda mental process
sc heme is present ed in Figur e 6 . H elium (15 - 60 Nml /m in) was used as carri er g as to tr ansport
the r eactant gasses from the sam ple loop to t he reac tor, thr ough t he cata lyst f ixed b ed ( 0.5 – 1
g) , an d final ly t o the MS . In th e first ste p, oxy gen wa s dosed to the reactor for c atalyst oxi dation
at react ion tem perature, b y pulse valve 1 ( PV) . After complete dosage, t he posit ion of the fi rst
p ulse v alve was set bac k to re f ill the sam ple loop with ox ygen. The f low of the carr ier gas w as
used to p urge the reac tor, r emoving gas p hase and weak ly adsorbed ox ygen int erm ediates. A
second PV was utilize d to dos e m ethane to th e react or. T he filling pr ocedur e was carried out
paralle l to the dosi ng procedur e of ox y ge n. Final ly, non - convert ed m ethane was purge d out b y
the carr ier gas, t o avo id an ex plosive m ixtur e with the foll owed ox y ge n pulse.
Methods and m aterials
17
Figure 6 : Sim ulated chem ical l ooping c oncept f or the ox idative cou pling of m ethane ( Paper 4 )
T he construc tion plan of the chem ical loop ing setup is given in Paper 4 . T he control softwar e
for the pulse val ves was develope d in Micr osoft Visual Basic 6 . Fur ther, the experim ental
param eters and oper ation c onditions are also pres ented t he ex periment al part of Paper 4 . T he
basel ine correc ted p eaks were int egrated and m ethane con versio n (X CH4 ), th e se lect ivitie s of
d etecte d com pounds and t he C 2 y ield were der ived ac cording to e quations ( 12 ), ( 6) and ( 11 ).
X
CH 4
= ∑ � ν i ⋅ A Reaction pr oduc ts �
∑ � ν i ⋅ A Reaction pro d ucts � + CH 4 , unconverted
( 12 )
Y = X CH 4 ⋅ S C 2 H 6 + C 2 H 4
( 13 )
A m odification of that operatio n was t he re petitive pu lsing of m ethane. After oxidatio n (PV 1) of
the cat alyst m aterial a t 775 °C and 30 Nm l /m in, the re actor was purg ed by the carr ier gas. T han
the fir st m ethane puls e (v ia PV 2) was inj ected an d af ter com plete dos age the PV was set back
to fill ing p osition, imm ediatel y. The oxi dation pr ocedu re was spar ed out a nd th e reactor was
purge d by the car rier gas to rem ove unco nverted methane . Thereaf ter, an ad ditiona l m ethane
pulse was d osed to th e reactor . That proc edure was repeated s imilar to a titra tion exper iment.
T he experim ent ended, w h en no m ethane con version was obser ved an y m ore, in dicated b y no
detect ion of C 2 an d deep ox idation prod ucts. All experim ental param eters and condit ions are
presente d in th e experim ental sec tion of Paper 5 .
Methods and m aterials
18
2.9 Residence time analy s es of the si mulated chemical looping setup
T he residence t ime anal y s is w as carrie d out b y a ser ies of puls e mar king exper iments , using
PV 1 as injector f or diff erent tracer gases ( 0.5 - 2 m l pulse) and H e as car rier gas . First the
dosin g behav ior of the trac er puls e (A r) by t he pu lse val ve was anal yz ed at diff erent flo w rates
of 20 – 6 0 Nml /m in . A ther m al conductivit y det ector ( TCD) w as mount ed to t he pul se valv e pipe
instea d of the reac tor inlet. T his is pres ented in Figure 7 A . The T CD detected a signa l each 0.5
s. The r ecorded peak area for each experim ent wa s normali zed to one and t he results were
fitted in Berkle y Madonn a by square p ulse func tions ( Figur e 7 C , lef t par t ). T he f itting
param eters were the d elay tim e (t delay ) between det ect or and pu lse valv e, the pu lse width (t dosing )
and puls e am plitude. The squar e puls e functi on (s quar epuls e(t dos i ng , td elay )) in Berk ley Mado nna
gener ates a recta ngular s ignal w it h am pli tude high t of one. T hat sign al was m ultiplied b y an
am pli tude param eter whic h was relate d to the exper im ental trac er dispersio n. T he finall y f itted
rectangu lar sig nal s had pe ak area s of one, sim ilar to the reside nce tim e distributi on function E(t)
which is defin ed by equ ations ( 14 ) and ( 15 ).
E ( t ) = V
⋅ c ( t )
n tracer
( 14 )
∫ E ( t ) dt = 1
( 15 )
In the nex t experim ental set the r eactor, f illed with cat alyst (0 – 2 g) , was connec ted to the pi pe
of the puls e valve and t he T CD was place d at the rea ctor outle t, which is shown i n Figure 7 B. A
series of pulse m arking ex perim ents wer e carr ied out w here the reac tor tem peratur e was var ied
fr o m 23 °C to 8 00 °C a nd the flo w rat e was var ied betw een 20 – 60 Nm l / m in. Eac h experim ent
was c arried o ut with Ar an d repe ated with N 2 as tr acer s to cons ider possible trans port eff ects b y
diff usion.
T he calculated param eters t delay , t dosi n g and am plitude of the squarepu lse f unction w ere
im plemented to the axi al dis persion m odel. T hat m odel des cribes t he reac tor as an ideal p lug
flow re actor with d ispers ion eff ects in axia l directi on. The m aterial b alance is shown in
equatio n ( 16 ) . T he disper sion depe nds on an axial d ispersion c oeff icient (D ax ), which is relate d
to the B odenstei n num ber (Bo ) acc ording to e quatio n ( 17 ) and th e flo w velocit y (u) .
dc
dt = − u dc
dz + D ax d 2 c
dz 2
( 16 )
D ax = u ⋅ L
Bo
( 17 )
Methods and m aterials
19
Figure 7 : Sch eme of residenc e time anal yses exp erim ents , A : Ana lyses of Puls e dosi ng
behav ior of the PV , B : Re sidence t im e analyses of the heated ractor , C : Sche me of the fitt ing
procedur e
T he implem entation of the dispersi on m odel was car ried out b y the use of a high num ber of
segm ents , where each se gm ent represents a finitesi m al part of the reac tor. T he flow of the
tracer through the r eactor, passing each se gm ent, is affec ted by convect ion and disp ersion,
which is desc ribed ac cording to equations ( 18 ) an d ( 19 ) . The param eter N is the t otal num ber of
segm ents, which was set to 200 a nd L is the p hysical length of the r eactor which w as 55 cm .
T he implem entation of th e squar e - pulse sig nal to that m odel is s hown in eq uation ( 20 ).
dc [ 1. . N ]
dt = − u ( c [ i + 1 ] − c [ i − 1 ] )
2 ⋅ L i + D ax c [ i + 1 ] − 2 ⋅ c [ i ] + c [ i − 1 ]
( 2 ⋅ L i ) 2
( 18 )
L i = L
N
( 19 )
dc [ 0 ]
dt = − u ⋅ � − sq uarep ulse � t dose , t del ay � ⋅ ampli tude + c [ i + 1 ] �
L i
+ D ax c [ i + 1 ] − c [ i ]
( L i ) 2
( 20 )
T he recorded data fr o m the resi dence tim e anal y s es with he ated reac tor wer e fitted b y t his
m odel. The f itting param eters wer e the Bode nstein n um ber and the ef fec tiv e flo wrate ( Fig ure 7
C, right part ). T he eff ective flow rate a verages th e initial f low rate of the carrier gas and the
contribu tion of gas exp ans ion eff ects b y t he furnac e tem perature profil e. The results of the
reactor character izatio n are desc ribed i n section 3. 5.2 .
Results and discussi on
20
3 Result s and d i scu ssio n
3.1 Contribution of gas ph ase reactions to the OCM reaction
In Paper 1 , three dif ferent reactor types w ere inv estigated in which the Na 2 WO 4 /Mn/SiO 2
catal yz e d OC M re action was perf ormed: a f luidi zed bed reac tor, a stand ard fixe d bed reac tor,
and a m em brane reac tor. Al l exper iments w er e carri ed out under s tead y sta te condit ions bu t
with d iff erent oxygen dosi ng strat egies . T he flu idize d bed r eactor was fed with ox ygen and
m ethane at the i nlet. T he strong bac k m ixing behavior in the flu idized b ed reactor dist ribut ed the
ox y ge n quick ly in the whole reac tor cham ber. T he fixed b ed reac tor was f ed sim ilarly with t he
reactants , but d ispersi on eff ects and f ree gas spac e are lo wer than i n the flui dized be d reactor.
In the poro us m em brane reactor t he ox y g en was dos ed consta ntly in s ma ll fr actions thr ough a
ceram ic m embrane to the r eactor cham ber along th e catal y s t fixed bed . Stani slav Jaso an d
Ham id Godini f ound that t he m e m brane react or has a hig her C 2 yield and a better C 2 sele ctiv it y
in com parison to the f lui di zed bed re actor, d espite sim il ar initia l condit ions a s the cata ly st
m aterial, tem perature and gas hour s pace veloc ity ( Tabl e 5). The C 2 sele cti vity of the fixed bed
reactor w as in bet ween. T heref ore, a st rong inf luence b y chem ical reac tions i n the gas p hase
and he at trans port ef fect s, which also i nfluences rea ction rates , is as sum ed. To anal yze th is
situat ion in m ore detail , the diff erent oxygen d osing s trategies w ere sim ulated by us ing of three
diff erent reactor m odels .
T able 5: Ex periment al and sim ulation res ults of the OCM reacti on in a m embrane, s tandard Co -
fed f ixed bed and fluidi zed bed r eactor ( Paper 1 )
Reactor Paramete r
Experim ental
result (w ith
catal yst)
Simul ation of gas
phase r eaction s
(Dooley M ech.)
Simul ation of Gas
phase + semi -
emp. M od.
M embrane
reacto r
X(CH 4 )
0.33
≈
0
≈
0
S(C 2 )
0.65
0.45
0.99
S(CO X )
0.35
0.45
0.99
Fixed - bed
reacto r
X(CH 4 )
0.32
0.01
0.34
S(C 2 )
0.53
0.99
0.56
S(CO
X
) 0.47 0.01 0.44
Fluidize d bed
reacto r
X(CH 4 )
0.35
0.29
0.30
S(C 2 )
0.40
0.14
0.12
S(CO X )
0.60
0.86
0.88
T he f luidized b ed reactor has a sim ilar mixing beh avior as the c ontinuous stirre d tank r eactor
(CSTR) , w h ile the plug f low tube r eactor ( PFT R ) has no ax ial bac k m ixing behav ior, whic h is
com parable to th e standar d fixed bed reac tor. It w as dem onstrated t hat the gas phase r eacti on
network is activ e even in a bsence of t he c atal y s t mat erial ( T able 5) . T he conversi on of m ethane
corr elates with t he free g as s pace and the oxygen d ispers ion of the r eactor f or em pty r eactor
sim ulations . T he mem brane reac tor sho ws no m ethane co nversion , bec ause of low ox y g en
partial press ure and no ba ck m ixing, whi le i n t he PFTR the m ethane c onversi on is also ver y
low . T he CST R sho w s hi gh m ethane conver sion, w hich dem onstrates clearl y th at th e oxygen
Results and disc ussion
21
dos in g strateg y influenc es str ongly the res ults of our experim ents. T o proof our h y p othesis a
sim ulation of a s im plified surfac e r eaction m odel f or the cat al y s t mater ial was realized, as
presente d in
T able 4 and in Paper 1 . In that m odel m ethane is o n l y con verted in to m eth y l r adica ls on the
catal y s t surf ace a nd cata lyst reox idation steps were i mplem ented . Such dras tic sim plif icatio n
avo id s a com plex situ ation of deep ox idation product f orm ation in gas phase a nd on t he catal yst
surf ace , which c annot cl earl y be disti nguishe d .
T he sim ulation res ults of gas - phas e reac tion and sem i - empir ical catal y s t m odel ( Table 5 , la st
colum n) for the PFTR sho w ed hi gher m ethane conv ersion and higher C 2 selectiv it y in
com parison with t he CST R, whic h suppor ts the h y p othesis that a n ade quate ox ygen dosing
strateg y im proves the o verall OCM perf ormance in a reac tor. Hig h ox y g en par tial pressures and
fr ee gas space suppor t the form ation of dee p oxida tion pro ducts. On the o ne hand in the gas
phase th e molec ular oxyge n activates m et hane b y the form ation of perox o and meth yl rad icals .
T he presence of other ox yge n- spec ies or in term ediates m a y pre vent the c oupling proc ess of
m ethy l r adica ls and lead t o the for m ation of deep oxidat ion products by c oupling of m eth y l
radica ls with ox ygen - speci es in the gas ph ase . T he main r eason is that t he partial pr essure of
the ox ygen s pecies i n the early sta ge of the react ion progres s is muc h higher t han the part ial
press ure of m eth y l radical s. Further mor e , e ven after succ essf ul meth y l r adica l coup ling, th e
ethane m olecule c an be m ore easily activat ed in the gas phase or on t he catal y st s urface
because of its l ower C - H bond str ength. T heref ore, low ox y gen partial pressur es m ay i m prove
the OC M perf ormanc e in th e gas phas e b y suppres sing side react ions. O n the other hand t here
is a strong ef fec t of heat form ation. In the gas ph ase the heat rem oval is m uch s low er in
com parison to a sol id m aterial, beca use of the bi g differ ence in he at capac ity an d conduc tivit y .
Th eref ore , a conver sion of m ethane in the gas phase g enerates higher tem peratures in
com parison to the sam e reac tion on the catal y s t partic le s . Such eff ect enhances all activatio n
proces ses in the gas phas e, whi ch are m o s t l y the ori gin of deep ox idation pro duc ts. That eff ect
can be seen i n Figure 8 , where at lo w methane conversion , les s heat is genera ted an d theref ore
the r eactions o n the catal yst surf ace and the co upling process of meth y l radic als dom inate the
form ation of pr oducts, nam ely etha ne and et hene. If no heat is rem oved, the tem perature r aises
m ore and m ore, whil e the C 2 sel ectivit y decre ases s harpl y . The heat fo r mation l ead to ho t sp ot
form ation and the gas phas e reactio n network dominat es the overal l react ion progr ess . Th is ca n
be see n by the s harp decr ease of C 2 product f orm ation and t he simulta neous e nhancem ent of
deep ox idation product f orm ation.
Results and disc ussion
22
Figure 8: S imulated r eac tion pro gress of the fixed - b ed react or with the gas ph ase m odel and
sim ple m icr o - kinetic s urfac e model b y solvin g the f ree energ y equati on (hot s pot form ation) with
the initi al tem perature of 800°C ( 180 Nm l /m in N 2 ; 120 Nm l /m in CH 4 ; 60 Nml /m in O 2 ; 3. 5 g
catal y s t ) ( Paper 5 )
3.2 Pressure variation during the OCM reaction
In literat ure , the m os t OCM exper iment s were car ried out at am bient pres sure (1 atm ). T he
absol ute press ure se ems t o be a crucial param eter f or the OC M reac tion, bec ause it influences
the num ber of molec ules in th e gas phase and th e tr ansport ef fec ts to the c atal y st surf ace .
Pape r 2 deals with the investiga tion of the OCM reaction in a broad r ange of abs olute
press ure s. The res ults of the co - f ed stead y state exp eriments at high press ures are presented
in F igure 9. In gen eral, t he m ethane con version i ncreas es by an increase of t he absolut e
press ure. For th e exper iment s wi th an em pty r eacto r the C 2 sel ectivit y is belo w 0.1 and at
10 bar , no C 2 produc ts wer e observe d an y m ore. T hese res ults suppor t the h ypothes is
di scu ssed in se ction 3.1 . T he m ethane is activate d b y gas phas e ox ygen and a higher
concent ration ac celerat es t he form ation of deep oxidatio n products . W hen we per form ed the
s am e experim ent in a f ixed bed f illed with quartz gran ules, the C 2 sel ectivit y was muc h higher .
On the one hand, the q uartz gran ules m ay c ontribu te to the cata lytic acti vity. On the oth er hand
the inert bed decreases the f ree gas s pace s ignifi cantl y . T herefor e , the the rmal m ethane
activati on and coupl ing of m eth y l rad icals in the gas phase bec a m e the m ajor reac tion ro ute.
T hat is because at low m ethane c onversi on the de ep oxidat ion prod uct f ormation m ight b e
pref erred. T hat reaction r oute dec reases the partial pres sure of gas phase ox ygen m uch fas ter
because of the big dif ferenc e in the stoich iometr ic fac tors of m eth y l radical form ation an d
ox y ge n insert ion reac tions (Eq. (7) vs. (9) ). B y dec reasing the ox ygen par tial pres sure th e
Results and disc ussion
23
therm al coupling pr ocess o f m ethy l ra dicals becom es the mor e do m inant reac tion s tep . At high
m ethane conver sion f urther ox idation proc esses in the gas phase are not p ossible b ecaus e of
the low ox y g en conc entrat ion .
B y the use of a catal y s t fixe d bed, t he C 2 selec tivity inc reases f rom 0.52 to 0.57 only slig htl y w ith
press ure . T he C 2 produc t form ation benef its fr om tw o ef fec ts . Firstl y t he adsor ption r ates of
m ethane and ox y gen c ould be e nhanced . T heref ore the c atal y tic activit y is enh anced and t he
reoxidat ion of t he cat aly s t bed consum es the g as phase ox y g en m uch f aster than under
atm ospheric pr essure. Second ly , the c oupling proces s is m ore lik ely under higher pr essures .
T his is c aused b y t he eff ect that after suc cessf ul C – H bond cleava ge by the catal y s t, the m ean
fr ee pathwa y to t he next m eth y l rad ical is dec reased. Ther efore it is m ore likel y t hat two m ethyl
radica ls coup le and form an ethane m olecule.
Figure 9 : Influe nce of abs olute pres sure on C 2 + sel ectivit y at constant residen ce tim e 0.16s ,
100 m g catal y s t, 700 °C, 50 – 5 00 Nml CH 4 :O 2 (9 5:5) (conver ted fr om Pa per 2 )
In additi on a series of tem poral analysis of products (T AP) experim ents were carried out in
Pape r 2 b y B enjam in Beck . In suc h experiments , the absolute pr essure is in t he range of 1 –
100 m Pa and the react ants are inj ected as gas pulses to the f ixed bed react or, sim ultaneousl y
or de layed. A sim ultaneo us inject ion of methane and oxygen sim ulates th e stead y state
experim ent at low press ures , while a de layed r eactant injection s eparates the c atal y s t oxid ation
f rom the m ethane ac tivat io n b y th e cata ly st into tw o indepe ndent re action s teps. T hese reac tion
condit ions are c hosen c lose to the Knuds en - diff usion regime, were the c ollisio n of m olecules
Results and disc ussion
24
with eac h other is neg ligib le and the collis ion betw een gas phase m olecul es and sur faces is
dom inating. T herefore the press ure was set h igher, t o allow f or interac tions i n the gas p hase.
Other wise the coupling proces s of m ethyl radica ls woul d not be p ossib le. In the case of
sim ultaneous injection of th e reactant s, the m ethane convers ion was 0 .02 wit h a C 2 sele ctiv it y o f
0.95 f or the Na 2 WO 4 /Mn/SiO 2 c atalyst. Such h igh s electivit y indicat es that t he gas ph ase
reaction s , induce d b y t he gas phase ox ygen , is more or less in active and onl y the cou pling
proces s of m ethy l radic als is relevant . In the sec ond cas e, oxygen was f irst injec ted and the
m ethane pulse was injec ted dela yed in tim e . The delay t ime was between 0. 1 – 1 s . It was
observe d that a cons tant a mount of C 2 products was produc ed in each c ase, de spite the de la y
tim e bet w e en the two rea ctant p ulses . C ontrar y , the f orm ation of de ep oxid ation pr oducts
decreas es by an incr ease of the time dela y . We conclude d fr om t h e se exper iments that t wo
ox y ge n spec ies are prese nt o n the c atalyst surfac e at these c onditions : A strong ly and a weak ly
bound oxygen s pecies. W hile the s trongl y bound spec ies would rem ain on the c atal y s t surf ace
and its life tim e would b e long e nough t o interac t with t he dela yed m ethane puls e to f orm sim ilar
am ounts of C 2 products , th e weak ly bound oxygen s pec ies would desorb after the ox ygen pulse
left the reactor s y s tem . Theref ore, it would be res ponsi ble for deep oxi dation pro duc t form ation .
At long er dela y tim es between t he two reac tant puls es, m ost of the weak ly bound ox ygen
species w ould be desorb ed and t heref ore less deep ox idation pr oducts are produc ed.
3.3 T emperat ure p rogrammed surface reaction experiments
Pape r 3 dea ls with TPSR experim ents to stud y the sur f ace interact ion s of the OC M react ants in
the a bsence of the gas ph as e an d weakl y bou nd oxygen on th e Na 2 WO 4 /Mn/SiO 2 c atalyst. S uch
techni que fol lows the ide a of th e TAP experim ent s to s eparate the oxidation of the catal yst
m aterial from the activat ion reac tion of m ethane b y the c atal y st m aterial . T he r esults of CH 4 -
T PS R exper iments wit h dif fer ent heating r ates are pr esente d in Fig ure 10 A - C . It was found
that m ethane is co nverted into CO and ethan e. Inter estingl y first e thane is f ormed an d later CO,
which indicates t hat two d iffer ent active ox ygen spec ies c ontribute t o methane a ctivatio n. That
was c onclude d by th e fac t that t he oxid ation r eaction of meth y l radica ls w ould proc eed
imm ediatel y wit h surf ace bound ox ygen speci es. Our r esults sho w that the CO f orm ation starts
at higher t emper atures, which is a dela y of several m inutes, depen dent on th e heating rat e.
T hat time s pan is no t in th e range of m e t hyl rad ical li fetim es. Theref ore two d ifferent types of
m ethane activat ion induc ed by dif ferent ox y gen s pecies were sugges ted. T hese results are in
straight c ontrad iction w ith the m odel of Lee et a l., which was pres ented in t he introduc tion
sect ion ( T able 2) , where onl y on e oxygen s pecies c ontributes to all activa tion reactio ns of
alk anes [67] .
T he results of C 2 H 6 - TPSR ex perim ents are show n in Figure 10 D – F. In Paper 3 it was
dem onstrated tha t the reac tivit y of ethane is m uch higher than the met hane one. T heref ore it
was ex pected th at etha ne activat ion wou ld be obs erved at lower tem peratures , but inter estingl y
the pro duct form ation starts at higher temper atures i n co m parison w ith CH 4 - TPS R experim ents.
Further mor e, t he ethene form ation star ts in all exper im ents at t he sam e temper ature. T hat is
Results and disc ussion
25
also d ifferent to CH 4 - TPSR ex perim ents, where th e positio n of the eth ane peak depends on th e
heating rate. W e c oncluded t hat eth ane is pr eferabl y activated in gas phas e, whic h was
validate d in a blank exper im ent wi thout t he cata l y st . The activat ion barri er in the gas phase h as
to be m uch hi gher in comparis on to t he cat alytic on es. Other wise no catal ytic ac tivit y f or C – H
bond c leavage wo uld be obs erved for methane , accor din g to the C– H bond diss ociation
energ ies which ar e 439 k J/m ole for m ethane and 43 6 kJ/m ole for ethane [9 ] . In add ition w e
found the f orm ation of water , which indic ates an ox idative deh ydrogena tion reac tion of ethane
or the con versi on of form ed h y dr ogen b y th e cata ly st m aterial . T hat reactant could be for m ed by
therm al deh y dr ogenat ion reac tion of etha ne to ethen e. At higher t emper atures in C 2 H 6 - T PSR
experim ents C O 2 and CH 4 wer e observed, which was interpr eted as pro duct f ormation b y
ethene c onversi on.
Figure 10 G – I pres ent ou r findings of C 2 H 4 - TPSR ex periments . W e found in all exper iments
the for mation of CO 2 and CH 4 , starting at s im ilar tem peratur e ranges . In Pa per 3 we wer e
unable t o re solve the m echanism in detail f or these re action pathwa y s . Duri ng the experim ental
procedur e cok e form ation was observed , prove n by c atalyst re oxidati on and s imultaneo us CO 2
form ation. W e conc luded that both oxygen s pecies on the catal y s t surfac e m ay contribute to
eth y le ne acti vation. O n the one hand elec trophil ic oxygen cou ld be inserting t o the m olecule b y
an attac k of the eth ene do uble b ond. Suc h epox ide species would be ve r y uns table at OCM
condit ions a nd decom poses imm ediately to deep o xidation produc ts and in the absence of
ox y ge n to car bon. On the other han d nucle ophilic ox ygen m ay attac k one of the C – H bond s ,
whic h form s vin y l radicals. Thes e are well - k nown as pr ecurs or in carb on cok e form ation [108] .
Results and disc ussion
26
Figure 10 : T PSR produc t form ation peak s of m ethane, et hane an d ethene T PSR ex periments
at heat ing rates from 1 - 5 K/m in, 1 g N a 2 WO 4 /Mn/SiO 2 , A - C : m ethane T PSR, D-F : ethane T PS R
(He : C 2 H 6 , 95 :5), G-I : et hene T PSR (He : C 2 H 4 , 95:5), 30 Nml /m in ( Pa per 3 )
Accor ding to our findin gs by TP SR experim ents we pres ented in Paper 3 a r eac tion networ k,
which w as com pleted b y the findi ngs of Luns ford et al. an d Beck et al. [54, 72] . Th e react ion
network is sho wn in F igure 11 . Alkanes or A lkenes c an be co nverted to de ep oxidat ion produc ts
by an ox y gen adsorpt ion inter mediate. T he cata ly st s urface pr ovides e lectrophi lic and
nucle ophilic ox ygen spec ies. T he electrop hilic one activat es m ethane b y C – H b ond clea v age t o
form meth y l radic als. T hese can c ouple in the gas phase an d form ethane. Besides , m ethane
can be converte d b y n ucleoph ilic ox ygen spec ies into d eep ox idation produ cts. Etha ne is
convert ed in t he gas ph ase or on t he catal yst surf ace into e thene. T he form ed h ydrogen in gas
phase is also co nverted by the c atal y st m ater ial into w ater, w hen ox ygen o n the catal yst surf ace
is av ailable . Furth er m ore , ethene is con verted on the ca talyst sur face to deep ox idation
products , underg oing t he form ation of c oke.
Results and disc ussion
27
Figure 11 : P rop osed Reac tion netw ork by TPSR ex perim ents [54,72] ( Pap er 3 )
Accor ding to the work by Redhea d, the r ate lim iting step of a he terogene ous reaction of
m olecule s fr om a gas phas e on a c atal y st surf ace is sens itive to the tem perature ram p in TPSR
experim ents [ 109] . T ypi ca lly, tha t step c ould be t he adsor ption of the gas m olecu le on the
catal y s t surf ace, the acti vation of the m olecule, the f ormation of transit ion state or the
desorpti on of the pr oduct. I n OCM it is ac cep ted t hat t he C – H bond c leavage of m ethane is the
rate de term ining step [ 67,110] .
T herefore , the positi ons of the record ed ethane peak s in CH 4 - TPSR exp erim ents were f itted b y
an em pirical func tion. Ac cording to the R edh ead m ethod we c alculated a n activa tion ener gy of
275 k J/mole (E A,3 ). Suc h high value ref lects the natur e of the str ongl y bound oxy ge n speci es,
which m ay also invo lve the latt ice oxygen . A poss ible energ y pr ofile for the O CM reactio n on the
Na 2 WO 4 /Mn/SiO 2 c atal y s t is presented in Figur e 12 . Molecul ar oxygen u ndergo es the for mation
of an a dsorpt ion inter media te (O 2 , ads ). That is conv erted i nto the s trongl y bound ox y gen speci es
(O x *). W e assum e that the reac tion rate of the ads orbed ox y g en interm ediate with m ethane is
m u ch f aster in ste ad y s tate exp erim ents than th e reac tion rat e of nuc leophil ic ox y gen ( not
sho w n). T herefor e , in steady state e xperi m en ts E A,1 is m easur ed for the unsel ective reaction
pathwa y . T he appar ent activ ation energ y (E A,2 ) is obser ved in stead y state e xperim ents f or
selecti ve methane ac tivat ion bec ause the in teracti on betwee n O 2 , ads and O x * were not
consider ed so f ar. Howeve r, there is a b ig devi ation b etween o ur findings (E A,3 ) a nd those o nes
f rom the literature (E A,1 ). T hat was explained b y th e f act that the e nerg y l evel of O* is muc h
lower, com pared to O 2 , becaus e the O* species under goes ads orption and lattice inser tion. Bo th
reactions rel ease ener gy and theref ore m uch m ore energ y is r equire d to for m the transiti on
state f or selecti ve m ethane act ivation.
Results and disc ussion
28
Figure 12 : Pr oposed en erg y pr ofile f or the OCM rea ction at t he Na 2 WO 4 /Mn/S iO 2 c ataly st at
800 °C ( Pap er 3 )
3.4 Dy namic exper i ments
In Pap er 3 , d y n am ic experim ents at c onstant reac tion tem perature were carried out to
investi gate the oxygen sto rage cap acit y of the cata lyst and t he lif etim e of the stored ox y g en
species . Fir st the stab ility of the stron gly bou nd ox y g en was anal yz ed b y increa s ing t he purg e
tim e from 10 m inutes to 300 m inutes. The r esults ar e presented in F igure 13 . I n each
experim ent sim ilar pro duct dis tribut ion s w ere obser ved. The overal l C 2 se lectiviti es are in the
range of 0.9 and deep oxidatio n produc ts are onl y f orm ed in the firs t m inute s of each
experim ent. Our results i ndicate that under O CM condit ions the s tored ox y ge n can r emain on
the surf ace of the Na 2 WO 4 /Mn/SiO 2 cat alyst f or hours. W e concluded t hat the abi lity to store
ox y ge n over that per iod o f tim e ma y i nvolve lattice ox y ge n of the Na 2 WO 4 /Mn/SiO 2 cataly st
m aterial. Ther efore , a Mar s van Krevele n type of mechan ism may be a pplied for the OC M
surf ace reac tion network which is valid f or both, n ucleoph ilic and electroph ilic ox ygen sp ecies.
Figure 13 : OC M produc t signals f rom d y nam ic ex perim ents with m ethane f or increasing pur ge
duration at 750 °C , 20 Nm l /m in CH 4 for 10 min, 1 g Na 2 WO 4 /Mn/SiO 2 , A : 10 m in He pur ge, B :
180 m in He purge , C : 300 min He pur ge ( Pape r 3 )
Results and disc ussion
29
For eac h dynam ic ex perim ent an ox y g en balanc e was calcu lated accor ding to the m aterial
balanc es in sect ion 2.7 . T he resu lts are prese nted in T able 6 . W e found that the am ount of
consum ed ox ygen for the form ation of deep ox idation product s is m ore or les s cons tant in all
experim ents. F urther m ore , the amount of conver ted ox ygen to fo rm C 2 produc ts is als o alm ost
const ant. I n add itional ser ies of dynamic ex perim ents purge time w as reduced t o 15 s, which is
the exper im ental limitatio n of the setup. Even suc h shor t purge inter vals res ults in com parab le
C 2 sel ectivi ties as in the ex periments before . T hese res ults conf irm our previous hypotheses
about th e pres ence of two ox ygen sp ecies whic h are strongl y bo und on the cat alyst sur face.
T he ratio of ethane to et hene is stron gly af fected b y the f lowrate, wh ich suppor ts our h y p othes is
f rom section 3.3 that eth ane is pref erably form ed in th e gas p hase. Furt herm ore , a linear
depende nc y b etween the par tial press ure of methan e and c onvert ed am ount of ox y gen was
found. O ur f indings indicate t hat the Na 2 WO 4 /Mn/Si O 2 catal y s t is a suitab le ox ygen carr ier
m aterial for c hem ical looping exper iments , becaus e it fulf ills all b asic requ irem ents:
• Stable r edox pr operties u nder r eaction c onditi ons
• The s tored am ount of ox y gen is stabl e bound under re action c onditions
• Short pur ge tim es allo w ad equate o peratio n times
T able 6 : Com poun d base d ox y gen b alance for a b road ra nge of purge times in d y nam ic
experim ents f r om 15 s to 18000 s, 1 g Na 2 WO 4 /Mn/SiO 2 , 750 °C , 20 & 30 Nm l /m in CH 4 f or 10
m in ( Paper 3 )
Purge
time ( s)
Flow rate
( Nml /min)
CO
2
(#O/nm²)
CO
(#O/nm²)
C
2
H
6
(#O/nm²)
C
2
H
4
(#O/nm²)
Σ
(#O/nm²)
18,000
20
8.8
0.6
4.6
6.8
20.66
10,800
20
10.4
0.6
4.5
7.1
18.00
600
20
7.2
5.4
1.8
3.6
21.66
300
30
10.7
0
7.3
2.6
20.60
270
30
9.4
0
7.6
1.9
18.94
230
30
9.8
0
7.3
1.8
18.81
160
30
8.9
0
7.7
1.7
18.32
100
30
9.1
0
7.2
1.5
17.80
60
30
9.5
0
7.0
1.5
17.99
30
30
9.7
0
7.6
2.2
19.47
15
30
8.3
0
8.6
2.3
19.11
Calcu lation of the ox ygen t ranspor t ca pabil it y R O for t he Na 2 WO 4 /Mn/SiO 2 c atalys t accordin g to
equatio n (3) , c onsideri ng as m o only the w eight of activ e com pounds Na 2 WO 4 an d Mn gives
0.014 . I ncludin g the su pport m aterial in this ca lculati on for R O gi v es 1 ⋅ 10 -3 . Th ese val ue s are
m u ch lo wer in com pa rison to the m ateria ls listed in T able 3 . The mai n draw back i s the high
m olar m ass of WO 4 and the h igh am ount of silica, which of fers only a sm all SSA (1. 86 m ²/g) .
T he comm on range of the SSA of chem ical loop ing par ticles is 3 – 1 00 m ²/g. [95,111, 112]
Results and disc ussion
30
3.5 Dev elo pment of a s im u l ated chem i cal looping set up
3.5.1 Setup constr u ction
Based on our f indings in P aper 3 , a sim ulated c hemic al l oop ing setup was cons tructed b y the
use of tw o six - port valves. T he m ost impor tant informatio n of Paper 3 wa s the a mount of stored
ox y ge n i n orde r to do se adequate am ounts of th e rea ctant t o reac h high co nvers ion s . In Paper
4 th e detai led proces s con cept is describ ed and th e fina l flow char t is presente d in F igure 14 .
Mass f low controll er s (MFC) we r e used to c ontrol the f low rate s of carrier gas and reactant
gases . The reac tants wer e fed to the sam ple loops of the PV’s, which ha ve a volum e of
0.25 - 2 ml . Accor ding to the operatin g princi ple of the PV’s, which is presente d in the
supporti ng infor m ation of Paper 4 , the dosing and filli ng positi on was s witched by a p n eum atic
s y st em, which was contr olled b y solenoid valves ( SV). F urnac e tem perature, the solen oid
valves for the PV’s an d the operat ion of the MS was cont rolled b y a c omputer via transduc ers
(Td ). T he f low of carr ier gas was s pli t dir ectl y i n front of the pulse valves and re com bined at the
outlet of the PV’s. T he ov erall f low, whic h trans ports the injec ted reac tant s, was send to t he
reactor . After p assing t he catal yst fixed bed the r eactant pulse s w ere an alyzed by a MS.
Results and disc ussion
31
Figure 14 : Flow char t of the c hemic al looping se tup ( Pa per 4- SI )
Results and disc ussion
32
3.5.2 R esidence t ime anal yses of the simul ated chemi cal looping setup
T o analyze the disper sion of the gas puls e s b y t he pu lse valves and th e disper sion effec ts of the
reactor a wide series of res idence tim e anal yses exp erim ents were car ried out. Several ef fects
contribu te to the gas pulse dispers ion as the am ount of catal y s t, the tem peratur e profile of the
fur nace and the in itial ga s flow rate. T herefore a detailed s tud y w as requir ed. T he fitted
param eters f or the squar e - pulse signal s ar e liste d in T able 7 . T hese par am eters were
im plemented to th e resi dence tim e model f or the f ixed bed reactor .
T able 7 : F itted square - pu lse par ameter s for res idence tim e analy s es of gas puls es in Ber kle y
Madonn a
Flowrate
( Nml /min) Fitting p aramete r
Vol ume of sample loop (ml)
0.5
1
2
20
t delay ( s)
22
23
24
t dosing (s)
15
24
38
a m plitude (s -1 )
0.066 0.042 0.026
30
t delay ( s)
14
16
15
t dosing (s)
10
20
31
a m plitude (s -1 )
0.099
0.049
0.033
40
t delay ( s)
11
13
11
t dosing (s)
7
15
23
am plitude (s -1 )
0.135
0.066
0.044
50
t
delay
(s)
9 10 9
t dosing (s)
6
12
18
am plitude (s -1 )
0.171
0.083
0.056
60
t delay ( s)
7
8
8
t dosing (s)
5
10
15
am plitude (s -1 )
0.207
0.0998
0.067
T he f itted Bode nstein num bers of our pulse m ark ing ex periments with the fixed bed reactor for
diff erent am ounts of c atalys ts, tem peratures an d initia l flow r ates ar e pres ented i n Figure 15 . It
was f ound that the f low rate and t he tem perature hav e m inor influences on t he disper sion of the
dosed pu lses. W e couldn’t find dev iations of the gas pulse shap e by switc hing the tracer gas
from Ar to N 2 . Such result s ind icate that th e mas s transport is dom inated b y c onvec tion an d
dispers ion and not by mole cular dif fusion. B y incr easing t he f low rate or the am ount of c atalyst,
the Bod enstein n um ber increases . The op posite case is obs erved by increas ing the
tem perature, w hic h indicat es that higher t emper atures decr ease the d ispersio n eff ects of t he
fix ed bed reactor . T hese eff ects ar e well - k now n f or gas es, because of the ir stron g depende nce
on the P eclet num ber ( Pe).
Results and disc ussion
33
Figure 15 : Res ults of fitted Bod enst ein num bers for diff erent tem peratures an d initi al flo wra tes
at dif ferent am ounts of catal y st
T he Peclet num ber is def ined b y e quatio ns ( 21 ) - ( 23 ) .
Pe = Re ⋅ Sc
( 21 )
Re = u ⋅ d R
ν = u � ⋅ d 32 ⋅ ρ
η ⋅ ( 1 − ϵ )
( 22 )
Sc = ν
D
( 23 )
W e assum e a m onodisper sed fixed bed where th e Sauter m ean diam eter ( d 32 ) is equal to t he
m ean diameter of a catal yst par ticle (2 50 µm ).
d 32 = d pa rticle
( 24 )
T he densit y ρ of He at the highest tem peratur e in the furnac e (800 ° C) was der ived acc ording t o
equatio n ( 25 ).
ρ = n ⋅ M He ⋅ T 0
V ⋅ T max
( 25 )
T he mean flowrate u � was calc ulated for an aver age feed s tream , considering the highes t
tem perature an d a mean d iam eter with res pect to t he vo id fraction.
u � = T fixed be d ⋅ V
T 0 ⋅ A ⋅ ϵ
( 26 )
Results and disc ussion
34
Finall y the dif fusion co eff icient D was ca lculate d b y the fundam entals of the kinetic g as theor y
and the C hapm en - Ensk og equation f or a m ono atom ic , ideal gas (equati ons ( 27 ) - ( 29 ) ) . T he
cross se ctio n σ for a He atom is 0.26 nm . [113]
D = 1
s ⋅Λ⋅ w
( 27 )
Λ = ( k B ⋅ T fixed bed )
√ 2 ⋅ π ⋅ σ 2 ⋅ p
( 28 )
w = � 8 ⋅ k B ⋅ T fi xed bed
π ⋅ m
( 29 )
T he Peclet num bers for the ex perim ental setup are in t he range of 0.1 – 1. T hese low num bers
indicate th at the gas transpor t through the f ixed bed r eactor invo lves for ced diff u s ion effec ts by
convect ion. T he results of Mi y auch i and Kik uchi, w ho revie w ed a broad num ber of residenc e
tim e experim ents, conf irm our tren ds of the Bodens tein num ber s [114] .
T he results of the fitting pr ocess f or the effec tive flow rates ar e presente d i n Fi gure 16 . For the
blank reactor ( Figure 16 A) the s imulation of ide al gas la w, cons idering the tem perature prof ile,
was an adequa te approxi m ation. W hen the r eactor was op erated with a catal y s t f ixed bed ,
press ure drop ef fects m a y occ ur. T he pressur e drop ∆ p w as calc ulated ac cordi ng to equ ation
( 30 ).
Δ p = ζ ⋅ ρ
2 ⋅ u � 2 ⋅ L fixe d bed
d h
( 30 )
T he hydraulic di am eter of the c atal y st bed was d erive d by equat ion ( 31 ).
d h = 2
3 ⋅ ϵ
1 − ϵ ⋅ d parti cle
( 31 )
T he drag coeff icient ζ was derived b y th e E rg un equation ( eq. ( 32 )) .
ζ = 150
Re + 1, 75
( 32 )
T o calculate the dynam ic viscos ity η of He, the Hirs chfelder equ ation (eq . ( 33 ) ) was us ed ,
assum ing that He is not com pressible d uring the c ollision and theref ore the c ollisi on integr al Ω
was s et to one .
η = 5
16 ⋅ � π ⋅ m He ⋅ k b ⋅ T fixe d bed
π ⋅ σ 2 ⋅ Ω
( 33 )
Unfor tunatel y , our exper imenta l param eters are not in the v alidit y r ange of the Er gun equation.
T herefore , in the case of 2 g catal y st th e exper imental pr essure dr op is higher tha n the
sim ulated one ( Fi gure 16 B) .
Results and disc ussion
35
Another approxim ation t o calcula te the pres sure dr op is t he Koz eny – C arm an equation ( eq.
( 34 )). T hat equatio n is h ighl y s ensit ive to the v oid frac tion. As mention ed in th e int roduction, the
quasi liq uid surf ace of the c ataly s t mater ial tends to sinter the fixed be d and the vo id frac tion
changes under r eaction co nditio ns. Ther efore , for si mulat ions the use of the adjus ted fl ow rate
is requ ired.
Δ p = 180 ⋅η⋅ ( 1 − ϵ ) 2 ⋅ u
Φ s 2 ⋅ d particle
2 ⋅ ϵ 3
( 34 )
Figure 16 : Resu lts of the f itted eff ective flo w r ates f or differ ent temper atures and initial flo wrates
by the dispers ion m odel and ca lculated flowrate b y ideal g as law including pres sure drop by
Ergun eq uation A : em pt y t ube reac tor, B : 2 g catal y s t in f ixed bed r eactor
3.6 Process parameter studies in simulated chemical looping expe r imen t s
3.6.1 Influence of fl owrate and temper ature on the OCM r eac tion
In Paper 4 , w e c ompar ed several chem ical loopi ng experim ents for w hic h the ca lculation of the
pulse cont act tim e (t pul s e , c alculated ac cording t o equat ion ( 35 ) ) was neces sary . T his param eter
is s im ilar to the modif ied res idence tim e. T he higher t p uls e the longer the m ethan e pulse rem ains
in the c atal y st bed. T he main dif ferenc e s in our ex perim ents were dif ferent amounts of catal y s t
and flo w r ates.
t puls e = m cat ⋅ A spec
V
( 35 )
T he experim ental results o f the tem perature and f low r ate var iation are s hown in Figur e 17 . I n
our experim ents to s tud y t he yield boun dar y we obser ved a m ethane c onversi on range betwee n
0.05 - 0.8 ( Fi gure 17 A). Unf ortunate ly, the c orrelati on between m ethane c onvers ion and C 2
selecti vity is anti - pro portio nal ( Figur e 17 B) . At low m ethane conver sion, h igh C 2 selecti vities
Results and discussi on
36
were obs erved, while at hi gh m ethane conv ersion les s C 2 pr oducts were obser ved . Fi nall y th e
calcula tion of th e C 2 yield sho w s a m aximum of 0.25 ( F igure 17 C ). T hus, not o nl y th e ox y ge n
induc ed gas p hase reac tions contri bute to the yield m ax im um in th e OCM reac tion . Als o t he
catal y s t itself c ontribute s to an appar ent C 2 yie l d max i mu m .
In our T PSR experim ents in Pape r 3 , we p ointed out that two dif ferent ox y g en spec ies,
electrop hilic a nd nucleoph ilic ox y gen, o n the cat alyst s urfac e contribute to alk ane and alkene
act ivati on. Thes e species compete als o in chem ical looping experim ents with each other to
activate m ethane and other com ponents . Both s pecies ar e not r emovabl e durin g the pur ge step,
because both are s trongl y bound o n the c atal y st su rface, as dem onstrated i n our d y nam ic
experim ents in Pape r 3 . W e concluded t hat the presence of both ox y gen spec ies and th eir
interact ions with alkanes and alk enes are the origin of the C 2 y iel d ma xi mu m. T his is becaus e
one se lective r eactio n (m ethyl radic al form ation b y elec trophilic oxygen and m ethane) com petes
at least with three uns elec tiv e r eaction s on t he cat al y st surfac e . T hese are the re action of
m ethane with n ucleoph ilic ox ygen and the cons ecutiv e reaction of eth ene w it h both ox ygen
species .
Figure 17 : Resu lts of yield studies, 2 g Na 2 WO 4 /Mn/Si O 2 catal y s t, 15 - 50 Nm l /m i n, 700 - 80 0 °C,
1 m l CH 4 pulse , A : met hane c on version, B : C 2 selectivi ty , C : C 2 yi e ld ( Paper 4 )
Results and disc ussion
37
3.6.2 Continuous oper ation of the chemical looping process and c omparison with
stead y st ate pro cess mo de
W e also pres ent a c ontinu ous op eration of the s imulat ed chem ical l ooping c oncept in Paper 4 ,
which is show n in Figure 18 . The f irst ox y g en pulse ha d a h igher am plitude as the n ext on e,
because the c atal y st was alr eady oxidize d. W hen the m ethane puls e was d etected we obs erve d
sim ultaneous ly C 2 prod ucts and carbon d ioxide. W hen t he catal y s t wa s r eoxidized b y t he
fo llowing ox y g en puls e, w ater was rem oved fr om t he catal y st surf ace. T hese r esults ind icate
that the r edox pro perties of the Na 2 WO 4 /Mn/SiO 2 catal y s t m aterial are stabl e in dy na mic
experi m ents at r eacti on condit ions f or hour s. Further m ore , the cata l y st r eoxidat ion re action rat e
m u st be equ al or fas ter than th e OCM re action ra te , becaus e d ur ing the ex perim ent the
m ethane conversion was constant whic h indicates that ther e is no decrease in t h e stored
am ount of ox y gen. Oth erwise the m ethane conv ersion s hould be dec reased dur ing the
experim ental progres s .
Another aspect is the f ormation of water . It can be c lear ly se en that water is on l y removed f ro m
the cata lyst sur face b y th e ox y gen p ulse. T h erefore , w e conclude t hat water molec ules or OH
groups c an rem ain, eve n under these har d cond itions , on the cat alyst sur face. It is not c lear,
whether the y contr ibute to the OCM r eactio n or are involved in th e transpor t pro cess es of the
reactants .
In th e c ontinuo us oper ation of the ch em ical looping experim ent onl y one m ethane pulse was
injected before t he cata ly s t m aterial was reox idized. Theref ore it is poss ible th at even m ore
m ethane pulses can be c onverted a nd util ize the avail able am ount of st ored ox y g e n , bef ore the
catal y s t reoxi dation is requi red. Thes e repetit ive pulse experim ents will b e disc ussed later .
Results and disc ussion
38
Figure 18 : Earl y an d late reactan t and product peak s during c ontinuous sim ulated c hemic al
loopin g, 775 °C, 25 Nm l /min, 2 g c atalyst, pu lse s of 1 m l O 2 or 1 m l CH 4 ( Paper 4 )
Results and disc ussion
39
A com parison be tween a c o - feed s teady stat e and a s imulated chem ical l ooping experim ent is
also presented in Paper 4 , whi ch is show n in T able 8 . At s imilar met hane conver sion w e
observe d for the s imulated c hemic al l ooping ex perim ent a higher C 2 y i eld. Com paring both
experim ents at a sim ilar C 2 yie ld, the s imulate d chem ical loo ping exper iment s converts less
m ethane m ore selective.
Another im portant param eter whi ch was c ompar ed is the space time y i eld (ST Y). T he STY in
our c ontinuous ly o perated s imulate d chem ical loo ping ex periment s at 775 °C was
73 µm ol e (C 2 )/(h ⋅ g cat ) an d 83 µm ol e (C 2 )/(h ⋅ g cat ) at 8 00°C, resp ectivel y . In both s teady state
experim ents that p aram eter was muc h higher (529 and 970 µm ol e (C 2 )/(h ⋅ g cat ) ). Such a b ig
diff erence is c aused by the f act that the sim ulated c hemical loop ing setup h a d m uc h free gas
space (c a . 60 ml ), bec ause of construc tive asp ects f rom the pip ing s y st em and the furnac e
geom etry, whic h req uires a long t ubular reactor. T he free gas sp ace m ust be f illed w ith the
carr ier gas and als o the carr ier gas is neces sar y t o trans port the reac tants (1 ml ) through t he
setup . T yp ic a ll y the diff erence bet ween the dos ed re actant vol ume and the r eactor v olum e is
m uc h sm aller and th erefor e the STY would be m uch higher i n a larg e scale r eactor .
T able 8 : Com parison be tween st ead y s tate exp erim ent (800 °C, 25 & 5 0 Nml /min, 250 m g
catal y s t) and sim ulated che mical loop ing (2 g cat al y s t, 1 m l CH 4 pulse) ( Paper 4 )
Co - feed steady state experim en t
X (CH 4 )
X (O 2 )
S C 2
Y
0.29
1.00
0.63
0.18
0.18
1.00
0.81
0.15
Chemic a l l ooping – equal methan e conver sion
X (CH 4 )
T (°C)
Flowrate ( Nml /min)
S C 2
Y
0.29
775
25
0.74
0.21
0.18
750 50 0.89 0.17
Chemic a l l ooping – equal C 2 yield
X (CH 4 )
T (°C)
Flowrate ( Nml /min)
S C 2
Y
0.21
775
30
0.87
0.18
0.19
725
25
0.87
0.16
Based on our f indings in P aper 4 , a concept f or an OCM c hemic al l oopi ng pro cess plant was
propose d , whic h is shown in F igure 19 . The catal y s t o xidation reac tor has a sim il ar f unctionalit y
as an air sep aration unit c om bined with t he oxid izer r eactor ( F igure 1) , because no n itrogen is
transpor ted to t he OCM r eactor. In parallel , water i s also rem oved ther e and no c ondenser
behind the OC M reac tor is require d. T he other s eparat ion proces ses, lik e the CO 2 absorb er , the
CH 4 separa tor and the C 2 s eparator , were locate d accordin g t o the propos al of Godini et al . [ 80].
T he catalyst ox idation an d the OCM r eactio n could be perform ed in f lui di zed bed react ors,
because the c atal y s t see ms to be resis tant ag ainst c omm inution or abrasi on whic h sho w s
stable operat ion in a n OCM m ini - plant [11 5] . O ther wise, f our fixed b ed reac tors with alterna ting
feeds h av e to b e used, wh ich i s not shown here. As dem onstrated before, the c atal y s t
Results and disc ussion
40
reoxidat ion is very fas t and ther efore no tem perature swin g is r equired dur ing the operat ion ,
which is also be neficia l for a cont inuous op eration .
Figure 19 : Schem e of chem ical looping of t w o flu idized be d reactors [81] ( Paper 4 )
In the case of f ixed bed reactor s with alterna ting f eed , a p urge gas is r equir ed to rem ove
uncon verted m ethane or ox y ge n, sim ilar to th e sim ulated chem ical loop ing proc ess. O ne
attract ive com pound is c arbon dioxide, bec ause it could be m ore easil y s eparate d f rom the
uncon verted reac tants , than ni trogen or other inert g ases. T herefor e , we sw itched th e carrie d
gas i n the se tup fr o m He to CO 2 . Becaus e of th e hig h part ial pres sure of CO 2 , th e calc ulation of
the car bon ba lance was not s uff iciently c alculated . Th us, t he peak areas of C 2 product s were
com pared by the use of dif ferent c arrier gases . That par ameter is proportional t o the am ount of
the corr espond ing com pound. The r esults are pres ented in Fi gure 20 . Und er CO 2 atm osphere
le ss C 2 produc ts wer e obser v ed. T hese r esults indica te that th e cat aly s t mater ial is less act ive
or th e C 2 produc ts were c onverted i n gas phas e by CO 2 i nto H 2 and C O in a dr y r eform ing
reaction . Another altern ative sw eep gas would be water vapor.
Figure 20 : Eff ect of different carrier gases on C 2 pr oduction , 1 g c atalyst, 20 Nm l /m in, 1 m l CH 4
pulse , A : I nfluence on C 2 H 6 produc tion, B : I nfluence on C 2 H 4 pr oducti on,
Results and disc ussion
41
3.7 Investigation o f the role of the cata l y st composition in simulat ed chemical
looping experiments and repetitive pulse experiments
In Paper 5 , the com position of the cat al y st was s y s tem aticall y var ied to u nderstan d the
contribu tion of each part to th e catal ytic acti vity an d select ivit y . Further mor e , the or igin of the
two diff erent ox y g en speci es was inves tigated , as well as the t wo diff erent func tionalities of the
m aterial.
3.7.1 Variation of the surface c on ce ntration of Na 2 WO 4 and Mn on COK-12 support
T he first inv estigated para meter w as t he surf ace loa ding wit h Na 2 WO 4 and Mn 2 O 3 . Theref ore ,
severa l catal y s ts were pre pared, k eeping the to tal weig ht loading of 5 wt - % Na 2 WO 4 and 2 wt - %
Mn (II) c onstan t, s upporte d on silic a m ateria l (CO K - 12) which h ad diff erent specif ic surf ace
areas af ter calc ination. T he larger the SS A, the lower was the speci f i c surf ace concentrat ion of
both com pounds. T he scr eening resul ts at diff erent te mper atures are pres ented in F igure 21 . B y
decreas ing th e spec ific sur face concen tration of the c atal y s t com pounds an i ncrea se of c atalytic
activit y was obs erve d , b ut the C 2 select i vity dec reases . At the lowest s pecif ic surf ace
concent ration, t he catal y tic activit y decreas e s a gain. W e calc ulated a specif ic s urface
concent ration of 66 Mn /nm² and 33 W /nm² , w hic h had the hig hest cata lytic act ivity. Suc h effec t
was explain ed by two rela t ed phen omena.
First ly , it has to be cons idered that one of the c o mpoun ds of t he active film , the Mn X O Y or
Na 2 WO 4 , m ust be related t o the ox ygen st orage ca pacit y . Further , suc h high loadings of both
com pounds in dicate a multila y er ed acti ve film on a si lica s upport m ateria l, beca use the
m onolayer conc entratio n of several tr ansient m etal o xides sup porte d on silic a is r eached at
aroun d 1 atom/nm ² [116] . Ther efore , the film thick ness of the ac tive com pounds s ee m s to be
im portant. T he ratio of ox ygen atoms w hic h are ava ilable o n the cat alyst sur face f or m ethane
activati on (O s u rface ) an d those ones which ar e stored in the film of the active com pounds (O b ulk ) is
defined as O s u rface /O bul k . A thic k f il m would r esu lt in a lo w ratio, while a t hin f il m should inc rease
that v alue. T herefore , thi nner f il m s lead to high er activ ity, bec ause m ore ox ygen atom s are
a vail able f or m ethane ac tivation , w hen the weight lo adings wer e consta nt and the SSA was
increas ed .
Secon dly, the f il m thic kness of the active com pounds la y er influenc es their cr y s tall init y . A th in
film lea d s to sm aller cr y s tal s izes. On the o ne hand s maller c rystals can b e reduc ed in
tem perature progr amm ed desorption ex perim ents more eas ily [43] . T hat m eans that ox y g en
atom s are less stable bou nd on thi nner f ilms , w hic h m ay lead to a re lease int o gas phase,
without c ontribut ion to cata lytic acti vi t y . In add ition the c atalyst perf orm ance depends s trongl y
on the struc tural f lexibilit y of that f ilm [117,118] . A de crease of the cr y sta l size w ould incre ase
the int eractio n with th e suppor t m aterial and t he int erac tion betwee n Mn X O Y and N a 2 WO 4 w oul d
decreas e, which decreas es catal ytic ac tivity.
Results and disc ussion
42
Figure 21 : R esults of si ngle m ethane puls e exper im ents on Na 2 WO 4 /Mn/SiO 2 wit h dif ferent
specif ic surf ace areas , 20 Nml /m in, 0.6 g catal y st , 1 m l CH 4 pulse , A : m ethan e conversion, B :
C 2 selec tiv it y ( Paper 5 )
T he results of our re petiti v e pu lse ex perim ents for Na 2 WO 4 /Mn/SiO 2 catal y s ts with diff erent SSA
are pres ented i n Figure 22 . In al l cases high C 2 sel ectiv ities were obser ved an d sim ilar to th e
single p ulse ex perim ents, the h igher the methane c onversio n the lo wer was the C 2 s elect ivit y.
W e observed for the m ost a cti v e c atal y s t (3 .2 m ²/g) also t he highes t am ount of c onverted
ox y ge n ( Figure 22 C ). Su ch result vali dates our hypothes is about the cor relation bet wee n
activit y and the film thick ness as discuss ed befor e. A ccor ding to th e ca lculations of the m olecul e
specif ic ox y g en b alance, we obser ved no bi g diffe re nce in the am ount of oxyge n atom s f or
ethene f orm ation.
#
Results and disc ussion
43
Figure 22 : Res ults of repetit ive m ethane puls e experim ents on N a 2 WO 4 (5 wt - %) /Mn (2 wt -
%) /SiO 2 at dif ferent s pecifi c surf ace areas , 30 Nml /m i n, 775 ° C, 0.6 g cata ly s t; A-C : m ethane
convers ion, s electivit y and C 2 y ie ld for each pu lse; D : ox y ge n ba lance (m ethane b ased), E :
ox y ge n balanc e (m olecule s pecific, k3/k4 = 7) , F : ox ygen ba lance ( molecule s pec ific, k 3/k4 =2)
( Paper 5 )
Results and disc ussion
44
3.7.2 Variation of the manganese loa ding on COK - 12 supp ort
Our r esults of single pulse exper iments for cat aly st s with d iffer ent manganes e oxide loadin gs
are pres ented in F igure 23 . W e obser ve an optim u m between c atal y st activit y and m anganese
loa ding . N a 2 WO 4 r ich catal y s ts sho w high C 2 s elec ti vities but low methane c onversi on s . By
increas ing the m anganese oxide loading the ca talyst activ ity i ncreases , but t he C 2 s elect ivit y
decreas es. T he highes t perf ormanc e w as observed f or 5 wt - % Na 2 WO 4 and 2 wt - % Mn, which
is t he c lassica l com position f or th is catal yst. Furth er inc rease of the m anganes e oxide loa ding
decreas es the m ethane c onversi on and inc reases t he C 2 s elect ivit y again .
Figure 23 : R esults of che m ical looping exp erim ents on Na 2 WO 4 /Mn/SiO 2 catal ysts at d iff erent
Mn load ings and temper atures, 2 0 Nm l /m in , 1 g catal y s t, 1 m l CH 4 puls e , A : m ethane
convers ion, B : C 2 s elect ivit y ( Paper 5 )
T he results of our re petiti v e pu lse ex perim ents for the m anganese ox ide varia tion are pres ente d
in F igure 24 . T he stored ox y g en amount of the m o s t active cata l y st (1.58 wt -% Mn ) was
com pletely conv erted after 12 m ethane pulses . Higher loadings nee d m ore pulses , w hil e lower
loadin gs of m anganese oxide require only 8 m ethan e pulses unti l no m ethane is c onverte d .
T hese results ar e bas ed on the ox ygen bal ances ( Figure 24 C & D ) sho wing that t he
m anganese ox ide seem s to b e resp onsible f or th e storage cap acit y of the Na 2 WO 4 /Mn/ Si O 2
catal y s t in chem ical loop ing experim ents. T hat was conc luded fr om the fac t that Na 2 WO 4 r ic h
catal y s ts have s ignific antly lower ac tivit y and require l ess m ethane pulses to reac h non - catal y tic
activit y. The m ore m anganese was l oaded, the m ore met hane pulses were require d. The
am ount of st ored oxygen differ s by fac tor fiv e.
W e also obse rve d n o depe ndence for the am ount of o xygen in t he et hylene for mation r eaction.
T hat indicates t hat the ethylen e form ation is i ndepend ent fr om the c atalyst. T hat was concluded
because th ere is no i nfluen ce on the form ation rate ne ither by the ov erall surf a ce conc entration
of the Na 2 WO 4 /Mn phas e n or b y variatio n of the Na 2 WO 4 /Mn ratio. T heref ore , the f ormation of
ethene s eem s to ha ppen by g as phas e reac tion s teps . Dur ing our pulse experim ents, no
stoichi om etric amounts of hydroge n , c orrespond ing to the am ount of eth y le ne, were obs erved.
Results and disc ussion
45
W e assum e that the h ydrogen is quick ly consum ed b y r educt ion of the cata ly st , sim ilar to
tem perature progr amm ed reducti on exper iments wi th h y dr oge n .
By increasi ng the mangan ese oxide loading at c onstant tun gstate lo ading, the f ilm thick ness is
also inf luenced. The m ore m anganese is depos ited o n the sur face, t he thick er is the film of the
active c ompounds on the s upport m aterial . Theref ore m uch m ore m ethane puls es were r equir ed
to red uce the high ly load ed catal y s t until no cata lytic ac tivity was obser ved . In sec tion 3.7.1 w e
pointed out th at the cr y st allinit y of the acti ve phas e may cont ribu te als o to the c atalyst ac tivit y .
For s upported m anganes e oxides s uch ef fec t w as fou nd by dif ferent grou ps. J i et al. f ound by
Ram an spectr oscop y and X RD that at 2 w t - % Mn the M n 2 O 3 is f orm ed preferabl y , wher e Mn is
in oxid ation st ate + III [ 46] . Inc reasing of the mangan ese load ing f orms als o the m ixed ox ide
MnMn 6 O 12. T hat species s tores less amount s of oxy g en, wh ich was also d em onstrated b y
Stobbe et al. w ho perf or m ed T PSR experim ents with diff erent m anganese oxide species [119] .
Th ey also found t hat dif ferent m anganese ox ides are reduc ed b y m ethane to Mn O. It has to b e
noted th at Na 2 WO 4 can als o switch the oxi dation s tate b etween W +6 to W +5 , according to th e
findin gs of J iang et a l. , whi ch m ay also cont ribute to t he ox y g en stor age c apac ity [1 20] . I n our
repetit ive puls e experim ents s uch eff ect seems to be neglig ible, bec ause th e amount of stored
ox y ge n inc reases b y a fac tor of f iv e but not the lo ading of the Na 2 WO 4 . T herefore , a constan t
am ount of st ored oxygen would be expecte d.
Results and disc ussion
46
Figure 24 : Res ults of repetiti ve m ethane pulse experim ents on Na 2 WO 4 ( 5 wt - % )/Mn (var . wt -
%) /SiO 2 at dif ferent m anganes e loadin gs, 30 Nm l /min, 775 °C, 1 g c atal y s t; A-C : m ethane
convers ion, s electivit y and C 2 y ie ld for each pu lse; D : ox y gen balance (m ethane based) , E :
ox y ge n balanc e (m olecule s pecific, k3/k4 = 7) , F : ox ygen ba lance ( molecule s pec ific, k 3/k4 =2)
( Paper 5 )
Results and disc ussion
47
3.7.3 Variation of the support mater ial
T he results of diff erent sup ported Na 2 WO 4 /Mn/SiO 2 c atal y s ts in ch emical l ooping ex perim ents is
presente d in Fig ure 25 . W e observe d l owe r m ethane con versions f or the quar tz support ed
catal y s t com pared to the COK - 12 suppor ted one, b ut higher C 2 sel ectivi t y . Su ch diff erence in
acti vit y was al so repor ted by Yild iz et al. [53] . The y f ound sim il ar activ it y e nhan cement of the
catal y s t in ste ady state exp erim ents at SBA - 15 s uppor ted catal y s t mater ial s a nd conclu ded that
m ainly a better dis tribut ion of the active c ompoun ds contr ibutes to that im provem ent . In our
s tu d y ( Paper 5 ) we assum ed that the enhancem ent c omes not only from a better dis tribution of
the acti ve compounds . It h as to be consi dered that t he COK - 12 s upporte d catal y s t has also a
higher S SA and th erefor e the film thick ness and the c ry st allinit y pla y also an i m portant ro le, as
discus sed befor e.
Figure 25 : R esults of che m ical looping exp erim ents on N a 2 WO 4 /Mn/SiO 2 catal ysts at dif ferent
support m aterials and temper atures, 1 g catal y s t, 70 0 – 7 75 °C, 20 Nml /m in, 1 m l CH 4 pu lse
( Paper 5 )
T he results of our repetitiv e methane pu lse exper im ents for the two d iffer ent support ed
Na 2 WO 4 /Mn/SiO 2 c atal y s ts are pres ented in Figur e 26 . Two eff ects were obser ved. T he COK -
12 supp orted one is m ore active and s tores hi gher am ounts of oxygen, which c an be co nverted
by m ethane. F or the Q uart z supporte d one muc h m ore methane pulses were n eces sar y unt il no
m ethane conversi on was obs erved an y m ore. T hese f i ndi ngs support o ur hypoth esis about t he
influe nce of the f il m thick ness. T he COK - 12 s upported catal y s t ha d the s ame weig ht loadings of
Na 2 WO 4 and Mn b ut a m uch higher SSA th an the qu artz suppor ted one. Ther efore , the
O surface /O b ul k ra tio m ust be higher. As disc ussed bef ore , the h igher t hat rati o, the m ore ac tive is
the cat alyst beca use m ore ox y ge n is ex posed on t he catal yst surf ace to b e conver ted .
Results and disc ussion
48
Figure 26 : R esults of repetitiv e methane pu lse exp erim ents on Na 2 WO 4 /Mn/S iO 2 at diff erent
support m ater ials, 30 Nm l /m in, 775 °C, 1 g catal y st; A-C : m ethane co nversion , selecti vity an d
C 2 yield for each pulse; D : oxygen ba lance (m ethane based) , E : ox ygen bala nce (m olecule
specif ic, k 3/ k 4 = 7), F : ox y gen b alance ( molecule s peci fic, k 3/ k 4 =2) ( Paper 5 )
Results and disc ussion
49
T he findings of a ll repetit ive pulse experim ents are s ummar iz e d in Figur e 27 . T hree theoret ical
trend li nes for dif ferent r eduction steps of differ ent manganes e oxide species were calcu lated
and pr esented ( Figure 27 A) . T hese tr end lines cons ider t he Mn 2 O 3 , Mn 3 O 4 and MnO specie s,
which c oexist o n the c atalyst s urfac e [121] . W e also f ound that th e total am ount of stor ed
ox y ge n, which h as to b e released f or r eduction of Mn 2 O 3 in t o MnO , was ne ver rea ched. T his
could be exp lained by the fa ct that th e cat alyst m aterial can not be full y oxidi ze d und er OCM
condit ions or a certa in am ount of oxygen is not available f or m ethane con versio n and r emains in
the lattice of the catal yst active phase. I t has t o be co nsidered t hat at hig her tem peratures (ca .
775 °C) a p hase change of Mn 2 O 3 to Mn 3 O 4 was o bserved , which also i ndicates a loss of
ox y ge n [121] . T he inf luence of the f ilm thic knes s on th e ox y ge n stor age c apacit y is pr esented in
Figure 27 B. W e c alculated a theor etical f ilm thick ness of the active com pounds on the su pport
m aterial accor ding to eq uations ( 36 ) and ( 37 ).
d = m ( Na 2 W O 4 )
ρ ( Na
2
WO
4
) ⋅ A
spec , cat , as prepared
+ m ( Mn ( II ) )
ρ ( Mn
2
O
3
) ⋅ A
spec , cat , as pre pared
( 36 )
# Layer = d ( Na 2 WO 4 )
L ( W − O ) + d ( Mn 2 O 3 )
L ( Mn − O )
( 37 )
As disc ussed bef ore thin ner f il m s (below 25 t heoretical La yers) le ad to low er catal y t ic
perf ormance. The increas e of the film thicknes s sho w ed s imilar r esults at 43 La y er s, whic h
sho w s clear ly, that t he film thick ness is a ver y s ensiti ve param eter f or a high perf ormanc e
Na 2 WO 4 /Mn/SiO 2 c atalyst.
Figure 27 : T ota l stored a m ount of ox y gen at diff erent ly pr epared Na 2 WO 4 /Mn/SiO 2 c atalyst
m aterials, A : correl ation bet ween ox y g en stora ge capac ity and m anganese ox ide loading, B :
corr elation bet ween ox ygen stor age cap acit y a nd f il m thic kness of Na 2 WO 4 and Mn 2 O 3 ( Paper
5 )
Results and disc ussion
50
As pres ented in the in troduc tion , manganes e oxide s ystem s are wel l - k now n as o xygen stora ge
m aterial in se veral ot her c hem ical looping studies [89] . It is c omm only use d in m ethane
com bustion or s yn thes is gas product ion. In the c ase of the OCM reac tion on Na 2 WO 4 /Mn/SiO 2
the ox idation s tate of m ethane changes onl y f rom ( – IV) t o (- III). T herefor e , th e role of the
m anganese oxi de is sim ilar, but t he role as catal y s t i s v er y differ ent.
It was dem onstr ated that Na 2 WO 4 r ich catalysts ar e less active in chem ical loo ing exper iments ,
but high ly s elect ive for C 2 products. As dem onstr ated , the ox y ge n stora ge capac ity d epends o n
the m anganese ox ide. T herefore , both c ompounds seem to have a diff erent role . W e conc luded
that on the one han d the tungstat e is r espons ible f or the s elect ive acti vation of metha ne; m ay be
by in v olv ement of a redox i nterm ediate W 5+ . O n the other hand , the m angan ese oxid e shows
clear l y redox ac tivit y but with less C 2 selectivit y. That was dem onst rated by Jon es et al . who
studie d the OCM perf orm ance of Mn X O Y /SiO 2 c atalysts [122] . T he coop erat ion ef fec t betw een
both com pounds was repor ted as ox ygen spi ll - over b y Jiang et al. [50] . T hey propos ed a redox
m e chanism accor ding to equations ( 38 ) - ( 42 ) . Sim ilar f indings were re ported b y Li [51] . The
m ethane is acti vate d on a tungst ate spec ies b y C - H bond c leavage. Thereb y m eth y l rad icals
are f orm ed, w hic h can co uple to form ethane , while tungst ate is r educed f rom W 6+ to W 5+ . T he
m anganese ox ide is able t o re oxidize the tu ngstat e, which enhances the c atal yst acti vity. T he
catal y s t reoxi dation r ate str ongly dep ends on t he reox idation rat e of the m a ngan ese oxi de
species . Howev er, the f indings of Jones et a l. indicate that also th e m anganese o xide is abl e to
fo r m active sites f or the s elective ac tivatio n of m ethane [122] . As rep orted by other gro ups,
supporte d m anganese ox ides t end to form also deep oxidat ion produc ts [121, 123,11 2] ,
T herefore , the pro posed r edox m echanism of J iang et al. has to be ex tended ac cordi ng to
equatio ns ( 43 )-( 45 ) , whic h cons iders f urther re dox activ ity of manganes e oxide , as t he tota l
oxidat ion react ion of m ethane a nd also a selec tive activation . H oweve r, due t o the fact tha t even
sm all amount s of Na 2 WO 4 enhances the m ethane con versio n and the C 2 se lectivit y , t he re actio n
rates of equa tions ( 38 ) - ( 42 ) m ust be higher c om pared to t he reacti on rates in e quations ( 44 )-
( 45 ) . Anot her pos sible ex planation is the num ber of s elective, activ e sites , wh ich cou ld be lo wer
for the manganes e oxide system . [46]
∗ − O ∗ − W 6 + + 𝐂𝐂𝐇𝐇 𝟒𝟒 → H + − O ∗ − − W 6 + − C H 3
−
Selecti ve CH 4 act ivation
on WO Z
( 38 )
H + − O ∗ − − W 6 + − CH 3
− → OH − , ∗ + 𝐂𝐂𝐇𝐇 𝟑𝟑
+ W 5+
( 39 )
W 5 + → W 6+ + e −
Redox m echanis m
(s pill - over )
( 40 )
Mn 3+ + e − → M n 2+
( 41 )
Mn 2+ + O 2 → Mn 3+
( 42 )
a Mn 2 O 3 + b 𝐂𝐂𝐇𝐇 𝟒𝟒 →→ 2 a M nO + b 𝐂𝐂𝐎𝐎 𝐱𝐱 + c H 2 O
Unselect ive CH
4
activati on ( 43 )
∗ − O ∗ − Mn 3+ + 𝐂𝐂𝐇𝐇 𝟒𝟒 → H + − O ∗ − − Mn 3+ − CH 3
−
Selecti ve CH 4 act ivation
on Mn X O Y
( 44 )
H + − O ∗−− Mn 3 + − CH 3
− → OH − , ∗ + 𝐂𝐂𝐇𝐇 𝟑𝟑
+ Mn 2+
( 45 )
Conclus ions
51
W e conclude d that t he pro posed oxygen spi ll - over effect bet ween Na 2 WO 4 and Mn se em s to be
a s uitab le mec hanism to e xplain our findings . As disc ussed bef ore, the film thicknes s of the
active c ompounds trigg ers m ainly the ac tivit y and the ox y ge n stora ge capacit y of the c ata ly st.
T he Mn X O Y s pecies d epen ds on both param eters and theref ore tunes the m ethane ac tivation
potenti al , which c ould b e explai ned b y equation s ( 40 )-( 42 ) . The thick er the f il m the m ore Mn 3 O 4
is poss ibly form ed. That effec t leads to a lo wer am ount of stor ed oxygen and m ay also cha nge
the red ox potent ial of the cat aly s t m aterial. Another point is the phase ch ange of Mn 2 O 3 to
Mn 3 O 4 , whic h is sim ilar to an auto - reduc tion st ep . Su ch a r eaction w o uld also c ontribut e to the
catal y s t acti vity.
4 Conclusi ons
In this work we investi gate d the inf luences of gas phase r eactions o n the cat al y s t perf orm ance
of the O CM react ion. T hat was don e by ex perim ents and s imulat ions of d ifferent reactor t y p es,
which have dif ferent m ixing prop erties . In our ex peri mental and also in sim ulation st udies we
found th at the gas phas e reac tion net work is ver y sensiti ve to the m ixing properti es of th e
reactor , indicate d by differ ent convers ion s of m ethane and ox y g en in flu idized bed ( CSTR),
fix ed bed (PFT R) and m embr ane reactor und er sam e reac tion conditions . T he addition of a
s em i - em pirica l m odel of t he cata lyst f unctiona li t y, whi ch onl y genera tes m ethyl r adicals s howed
that the f ina l select ivi ti es of the reaction produc ts depend s trongl y on the o xygen dos age
strateg y. H igh oxygen partial pr essures and stron g m ixing beha vior lead to a sign ificant
decreas e of t he am ount of C 2 produc ts. Furt herm ore, the sam e ef fect leads to th e form ation of
Hot Spots in the g as ph ase, which are res ponsib le for a s ignific ant increas e in the react ion rate.
T he m ain reasons are the lo w hea t c onduct ivit y and heat c apacit y of the gases , which leads to
heat tr ansfer problem s. T h us, a reactor s ystem is req uired which im proves th e c atalyst oxid ation
without t he presenc e of high ox y g en part ial press ures.
T o improve the catal y s t ox idation r ate , t he influ ences of high er pr essure i n OCM ex perim ents
w as in v esti gated. Our results indic ate that the ads orptio n process on the ca talyst s urfac e w as
enhanc ed , which incre ase s the m ethane convers ion and C 2 s elec tiv it y . F urther , t he enha nced
oxidat ion rate of the cat alyst m aterial, c aused b y enhanc ed ox y g en adsor ption, decreases the
am ount of gas phas e ox y gen. T hus , less gas p hase ox idation re actions lead to le ss f or mation of
deep ox idation products.
In T AP e xp erim ent s it was dem onstr ated that weak ly b ound oxygen, a poss ible a dsorption
interm ediate of gas phase oxygen, influenc es stro ngly the met hane act ivation pr oces s. W e ak l y
bound ox ygen l eads to d eep oxi dation pr oducts , while s trongl y boun d ox y gen seems to be
respons ible f or selec tive m ethane activati on. T herefor e , a stu dy of the c atalyst under stead y
state c onditio ns is af fec ted b y both ox ygen species and the reliab ility of kinetic studies under
stead y state c onditions is ques tionable, becaus e the c ontribut ions of both spec ies c an hardl y be
resolv ed.
Conclus ions
52
T hus , dynamic experim ental techni que s, such as T PSR , w ere used to investi gate the c atalyst .
B y the use of suc h technique s, the ox idation pr ocess of the catal y s t m aterial b y gas phase
ox y ge n was se parated f rom the OCM reac tion. T hat allo w ed th e inves tigation of the c a ta l yt ic
surf ace reac tion network without the inf luence s by gas phase ox ygen and its adsorpt ion
interm ediate s on the c atalyst . I n tem perature progra mm ed surfac e reaction experim ents, w e
found that t wo strong ly bound oxygen spec ies are i nvolved in the OCM reac tion on the c atal y s t
surf ace. Elec trophilic ox y g en seem s to be respo nsible for th e selec tive m ethane ac tivatio n,
while n ucleoph ilic ox y g en is r esponsible f or the for mat ion of deep oxidat ion produc ts. T he
form ed ethane is quick ly c onver ted to ethen e in th e gas phase b y ther mal de hydrog enation or
oxidat ive deh y dro genation on the cata lyst surf ace. Bo th oxygen s pecies m a y a lso contri bute to
unsel ective eth ene act i vation. B y the var iation of the he ating ra tes in m etha ne TPSR
experim ents , an act ivation energ y of 275 k J/m ole for the s electi v e methan e activati on was
derived accordi ng to th e Redhe ad method. Th is high energ y barrier ref lects t hat the str ong ly
bound ox yge n s pecies ha ve a com parab le nature as lattice ox y ge n. Fur ther, thes e res ults
indicate a Mars - van - K revel en lik e reactio n mec h anism for the OC M reaction on the
Na 2 WO 4 /Mn/SiO 2 catalyst.
D yn am ic experim ents at const ant reacti on tem perature s sho wed that the str ongl y boun d oxygen
species are st ored at OC M cond itions for h ours o n th e catal y st surf ace. W e also found a linear
depende nce of t he m ethane part ial press ure. An oxyg en balance of the f ormed pr oducts g ave
an ox y g en s torage cap acity of 20 O/ nm ², which ind icates the in volvem ent of sever al s ubla y er s
fr o m the Na 2 WO 4 /Mn phase . As mentio ned in th e i ntroduction, a st ructural flexibili ty of the
Na 2 WO 4 /Mn ph ase is nece ssar y for a highl y active c atal y st. T heref ore it seem s to be pos sible
that suc h an am orphous , fle xible s tructure allo w s t he involv ement of m any sublayers f or
m ethane activat ion react ions.
Based on our f indings we deve loped a s imulate d chem ical looping setup, des pite of a l ow
ox y ge n trans port capab ilit y R O, whic h was 0.014 – 0.0 01 . That tec hnique al lowed the dosag e of
defined am ount s of m ethane. The m ain princ iple of op eration w as an oxidat ion step, f ollowed b y
a purge step t o rem ove gas phas e ox y ge n from the reactor by the help of a car rier gas . Then , a
m ethane puls e was dose d and u ncon verted m ethane and O CM pr oducts w er e tr ansported by
an add itional p urge s tep t o the det ector . W e could sh ow that t he Na 2 WO 4 /Mn/S iO 2 cat alyst has
a stabl e perform ance in contin uous chem ical loo ping oper ation . In our yi el d o ptim iz at ion studi es
we found a n apparent C 2 yield m aximum of 0.25 . T hat seem s to be c aused by the contr ibution
of stored electro philic and nucle ophilic ox ygen on th e catal y s t surf ace to the OCM reac tion ,
whic h acti vate m ethane as well as ethen e. A com pari son bet ween s imulated ch em ical looping
experim ents and stead y state ex perim ents sho w ed that m ethane can b e convert ed mor e
eff iciently in d ynam ic experim ents due to the h igher selectivi ty than in s tead y s tate m ode . I n
Figure 28 , we com pared th e chem ical loo ping resu lts with the liter ature in ord er to h ighlight th e
advant age of t his techniq ue f or the OCM re action. At m ethane convers ions lo wer than 0.3 m ore
C 2 prod ucts are f orm ed in chem ical looping ex perim ents tha n in stea dy state experim ents. T hat
diff erence ref lects the in volvem ent of s everal oxyge n interm ediates, which cont ribute in gas
Conclus ions
53
phase and on the c atal y s t sur face to alk ane and a lkene ac tivation in stead y st ate oper ation .
Further we dem onstrated that a lso hig h methane con v ersio ns a re re achable in c hem ical lo oping
experim ents , w h ich can not be invest igated u nder stead y sta te condit ions. T he required r atio of
CH 4 :O 2 in th e feed m ixt ure is higher than 0.5 , but for this com posi tio n t he explo sive regim e is
reached .
Due to t he fact tha t no g as phase ox ygen is pr esent in chem ical loopi ng experim ents , the
m ethane is com pletel y ac tivat ed by ox ygen f rom t he catalyst sur face and t he most heat is als o
gener ated ther e, whic h can be m uch eas ier rem oved than in g as phase. It was al so highlig hted
that the us e of diff erent reac tor techniques and f eed strategies impr ove the OCM perf ormanc e.
A ccor ding to the f indings o f Takanabe et a l. the ad dition of water im proves the C 2 selec tivity at
lo w m e than e convers ions add itional ly . [71] Increas e d m ethane con versions lead to the
form ation of m ore water and ther efore th at ef fect is not obs ervable anymor e. Go dini an d
Cowork ers dem ons trated the succ essf ul operation of a m em brane reactor , perf orm ing the OCM
reaction [86] . The y d em onstrated th at the dosa ge of low amounts of ox y ge n along th e cata ly st
fix ed bed ha d a s im ilar, benefici al effec t as the sim ulated chem ical loo ping ex periments .
Kruglo w et al. pres ente d a count ercurr ent movin g bed re actor, w h ere m ethane is partial ly
convert ed and th e C 2 prod uc ts were s eparated f rom the res t of the f eed [84] . Unc onvert ed
m ethane was fed to t he ne xt reactor sect ion, whic h w as repeate d four tim es. T hey reporte d a C 2
yield of 0.55.
T he conversion of lo w amounts of a reac tant at high s electiv ities is car ried out in the productio n
of eth y l ene oxid e in c hem ical industr y . [ 124] T he m ost impor tant fact or is the high pric e
diff erence bet ween et hene and eth y l ene oxi de , wh ich m akes suc h a proces s eco nomic al. Thus,
the chem ical loop ing seem s to have th e potent ial for industrial a pplication , due to its potenti al to
increas e the se lectivit y f or t he desired products .
Conclus ions
54
Figure 28 : Com paris on of literat ure resu lts for Na 2 WO 4 /Mn/ SiO 2 cat aly s t in stan dard fixed b ed
reactor s with s imulated c hem ical loop ing and ot her reac tor conc epts ( Paper 4 )
One of the m ost i m portant param eters in chem ical looping ex perim ents is t he oxygen stora ge
capacit y, which was inves tigated b y th e s y nt hesis of dif ferent N a 2 WO 4 /Mn/SiO 2 catal y s ts. W e
found a re lations hip bet ween the s torage c apacit y of ox y gen a nd the m anganes e oxide l oading.
W e also f ound an inf luence of the film thick ness of Na 2 WO 4 /Mn on the support m ateria l on OCM
act ivit y. In ad dition the f orm ation of ethene seem s to be m ostl y inde pendent f rom the c ataly s t
and tak es pl ace on ly in the gas phase b y therm al deh ydrogenat ion of et hane. A s um m ar y ab out
the func tional ity of t he Na 2 WO 4 /Mn/SiO 2 c atal y s t in chem ical loopi ng exper im ents are presen ted
in Paper 5 and schem aticall y shown in Figur e 29 .
T he active c ompounds Na 2 WO 4 and Mn X O Y are dep osited o n SiO 2 as supp ort m ater ial . T he
sodium ions induc e a phas e transition of the SiO 2 in to α - cr istoba lit e [44] . That phas e change
allo w s a str uctura l flexib ility of the N a 2 WO 4 /Mn X O Y m ixture under OCM c onditions [48] . Su ch
am orphous proper ty enab les hig h OCM per form ance for that mater ial b y ox ygen sp ill - over
reaction between Mn and Na 2 WO 4 [50,51] . Dur ing th e oxidation of the catal y st m a terial with g as
phase ox ygen an a dsorbed ox ygen inter mediate is form ed, which is abl e to convert m ethane
into d eep oxi dation pr oduc ts [72] . T he oxi dized cat al y st stores electr ophilic a nd nucle ophilic
ox y ge n speci es. Thos e can in teract in diff erent wa y s . Electr ophilic ox ygen (O elec. ) r eacts with
m ethane a nd leads to m ethyl ra dical f orm ation. Thos e radic als can c ouple, poss ibly in gas
phase, and f orm ethane [54] . Af terwards , it is d eh ydrogenate d in th e gas phase t o ethe ne,
which ca n be ac tivated by electr ophil ic ox y g en, wh ich form s ox y genat es. T hese ox y g enates
tend to form deep oxidati on pro ducts. Nucleop hilic o xygen ( O n ucl . ) i nterac ts with m ethane and
leads to deep oxid ation produc t form ati on. T hat species can be reduce d b y m olec ular hy droge n
Conclus ions
55
to water and can convert ethen e into deep oxidat ion pr oducts by the f orm ation of c arbon cok e
as int erm ediate, which is further oxidize d . T he role of Na 2 WO 4 seems to be t he ac tive site f or
selecti ve activation of m e t hane . T he m anganese o xide contr ibutes b y its com plex r edox
propert ies. O n the one ha nd it is ab le to oxid ize the tungstate s pecies , w hich enha nces the
fo r m ation rate of m eth y l ra dicals [ 50] . O n the oth er hand it also c ontribut es to the f ormation of
deep ox idation pro ducts. T he film thickness of the active p hase is anoth er im portant f actor for a
hig h ly ac tive Na 2 WO 4 /Mn/SiO 2 catalyst in chem ical loo ping exper im ents. T hic k films store m ore
ox y ge n in the l attice, which is not acc essible for m ethane activat ion. Un der O CM reaction
tem peratures the Mn 2 O 3 could tra nsform also in to Mn 3 O 4 [121] . That m a y suppre ss the c ataly st
perf ormance. Thinner film s show a hi gher perform ance, bec ause m ore ox ygen is acces sible on
the ca talyst surfac e for m ethane act ivation. Furth er decre ase of the f ilm thick ness m a y
destab ilize the str uctural f lexibilit y , d ecreas ing the ox y g en spi ll - over eff ect and f inally , decreas e
the cata lytic acti vi t y . In a ddit ion oxygen c ould be less stable bound a nd could be m ore eas ily
desorb i nto the gas phas e, caus in g again unf avorable r eactions.
Figure 29 : Prop osed OCM surf ace and gas phase r eac tion netw ork , and wor king f unctional ity
of the N a 2 WO 4 /Mn/SiO 2 c atalyst i n simulated c hemic al looping exp erim ents [44,48, 50,54,72]
( Paper 5 )
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[120] Z. - C. Jiang, L. - B. Feng, H. Gong, H. - L. W a ng, Methane and Alk ane Conversion
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Paper 1
Thermal Reaction Analysis of Oxidative Coupling
of Methane
Hamid Reza Godini
1,
*
, {
, Vinzenz Fleischer
2, {
, Oliver Görke
3
, Stanislav Jaso
1
,
Reinhard Schomäcker
2,
*, Günter W ozny
1
DOI: 10.1002/cite.201400080
Dedicated to Prof. Dr . Manfred Baerns on the occasio n of his 80th Birthday
F or more than three decades, the oxidative coupling of methane (OCM) process has been investigated as a promising alter -
native approach for ethylene production. S imulations of different sets of surface mechanisms over the Na
2
WO
4
/Mn/S iO
2
catalyst and the gas phase reactions that come along with the OCM reaction were analyzed in a fixed-bed, membrane, and
fluidized-bed reactor . The results were compared with the experimental data generated in an OCM mini-plant. It was
observed that the gas phase reactions are crucial in reducing the overall selectivity , especially in the fluidized-bed reactor .
K eywords: Gas phase reactions, Oxidative coupling of methane, Reaction kinetic, Reaction network, Reactor design,
Thermal reaction engineering
Received: M ay 28, 2014; revised: A ugust 12, 2014; accepted: A ugust 15, 2014
1 Introduction
Global ethylene production capacity was over 140 million
ta
–1
in 2013 [1]. Depending on the easier access to various
hydrocarbon resources and their local prices, different feed-
stocks are exploited for producing ethylene in different parts
of the world. F or instance, in Asia-P acific and in W estern
Europe, where around half of the world ethylene-production
facilities are located, mainly naphtha is used as the feed-
stock of cracking processes while in North America and the
Middle East, accounting for 40 % of the global ethylene pro-
duction, mainly ethane and natural gas are used as the feed
for crackers. In general, ethane and natural gas are becom-
ing the dominant feedstock for ethylene production. This is
part of a major trend to substitute the oil-based feedstocks
by natural gas for producing chemicals, especially in the US
after the shale gas revolution. Oxidative coupling of
methane (OCM) uses methane as the main component of
natural gas to directly produce ethylene. This highlights the
significance of the OCM process in this trend.
The systematic research on the OCM process started in
DOW Chemical in the early 80s and the first report was
published by Keller and Bhasin [2]. Since then and during
the last three decades, several catalysts and reactor concepts
have been proposed for OCM [3 – 5]. Pioneering work by
Prof. Baerns and his research groups at Ruhr -Universität
Bochum and later at the Institute of A pplied Chemistry
Adlershof and F ritz H aber Institute of the Max Planck
Society Berlin are considered to be fundamental and vital
for OCM research in the last three decades [6 – 8]. UniCat
(Unifying Concepts in C atalysis, www .unicat.tu-berlin.de), a
research group coordinated by the Berlin Institute of T ech-
nology and funded by the German Research F oundation
(Deutsche F orschungsgesellschaft), also has the honor of
his associated membership and benefits from his expertise
in the field of OCM.
F urther advancing the understanding of the OCM reac-
tion mechanism and improving the performance of the
OCM catalyst, reactor and process have been the main mo-
tives of Prof. Baerns and UniCat’s OCM research group. In
this paper , the issue of the selectivity is addressed in the
context of catalyst and reactor performance analysis with
the view on the OCM reaction mechanism. The effect of
homogeneous gas phase reactions and catalytic activity in
different reactors are investigated in details.
1.1 Reaction Mechanisms and Important Aspects
In the early eighties, Lunsford and co-workers showed that
the methyl radical formation by hydrogen abstraction and
radical coupling are the fundamental mechanisms of the
www .cit-journal.com © 2014 WILEY -VCH V erlag GmbH & Co. KGaA, W einheim Chem. Ing. T ech. 2014, 86, No. 11, 1906–191 5
–
1
Dr . H amid Reza Godini (hamid.r [email protected] n.de),
Dr . Stanislav J aso, Prof. G ünter W ozny , Berlin Institute of T echnol-
ogy , Chair of Process Dynamics and Operation, Straße des 17. J uni
135, Sekr . KWT -9, 10623 Berlin, Germany ;
2
Vinzenz Fleischer , Prof.
Reinhard Schomäcker (Schomaeck [email protected] ), Berlin Insti-
tute of T echnology , Institute for T echnical Chemistry , Straße des
17. J uni 124, 10623 Berlin, Germany ;
3
Dr . Oliver Görke, Berlin
Institute of T echnology , Department of Ceramic M aterials, Institute
for Material Science and T echnologies, Hardenberg straße 40,
10623 Berlin, Germany .
‡First and second author contributed equally to this paper .
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OCM reaction [9]. There are some macro [7, 10, 11] and micro-
kinetic models [12 – 15] that comb ine methane activation via
adsor bed oxygen, methyl radical formation , coupling, oxida-
tion of metha ne and the products, reforming, and dehydr o-
genation react ions to represent the OCM reactions netwo rk.
An OCM kinetic model should consider the contribution
of both gas phase and catalytic reactions in order to predict
the methane conversion and product selectivity . The major
challenge, however , is to determine the kinetic parameters
for the modeling of the surface reactions, which are strongly
influenced by the gas phase contribution.
Dooley and co-workers published an extended micro-
kinetic mechanism of the methane oxidation in homo-
geneous gas phases [16]. However , simulations with the
Dooley model and comparing the results with the experi-
mental data enable predicting the conversion and selectivity
with high accuracy [17]. Therefore, this model was used
here in the simulation part of this investigation to represent
the gas phase contributions to the OCM reaction [16].
1.2 Limitations of Catalyst, Reactor and Operating
Conditions
The OCM net energy balance is highly exothermic and with
increasing conversion of methane, the operating tempera-
ture continually increases until the temperature gradient
between the reaction temperature and the temperature
in the external surrounding environment becomes high
enough to transfer the excess generated heat out of the cata-
lytic bed. A pproaching the reactor end, both methane and
oxygen (in reactors with co-feeding structure) are progres-
sively consumed and the operating temperature decreases.
Being able to contr ol the react or’s operating temperature
not only secu res a safe reactor opera tion, but also improves
the reactor perf ormance in terms of selectiv ity and yield of
ethylene. This is mainly because of the cruc ial effect of tem-
perature on gas phase reaction s. Signif icance of the effect of
gas phase reactions [18 – 21], however , is different in different
react or feeding policies. Therefore, choo sing a proper reactor
and set of operating condi tions can significant ly enhance the
OCM react or performan ce with respect to all these aspects.
A wide range of the reactor concepts has been proposed
for the OCM application. Each type of reactor offers an
advantage with respect to one or some of the OCM reactor
performance indicators. Usually the fixed-bed reactor is con-
sidered to be a standard reactor concept and offers a simple
construction and operation in industrial-scale operation.
Although this is not completely the case for the OCM reac-
tor , it is reasonable to investigate the performance of a fixed-
bed reactor for the OCM reaction. On the other side, flui-
dized-bed reactors seem to be one of the advantageous and
ultimate choices for the OCM reaction unit due to their iso-
thermal performance [22 – 24].
It is also a fact that low concentrations of oxygen are
advantageous for high C
2
selectivity [25 – 27]. The partial
pressure of oxygen can be reduced by either a diluted gas
stream or an inorganic membrane to distribute the oxygen
along the bed and keeping its local concentration low . This
is the operating concept of the OCM membrane reactor .
Among the membrane reactor structures applied for the
OCM reaction so far , the porous packed-bed membrane
reactor (PBMR) offers a fine oxygen-dosing potential and
provides a proper permeation and contact-volume ratio.
Therefore, it allows achieving a significant amount of
methane conversion which ensures a high level of C
2
yield.
The performance of the fixed-bed [28, 29], fluidized-bed
[23, 24] and membrane reactors [29 – 32] for OCM were
experimentally investigated in UniCat’s OCM mini-plant.
The observed performances will be discussed here based on
the mechanisms of the OCM reaction simulated with the
detailed micro-kinetic model. Obviously , this analysis will
be performed in the context of the whole OCM process ana-
lysis. F or instance, injecting an inert diluting gas such as
nitrogen, which is prescribed to overcome the challenge of
hot spot formation in the fixed-bed catalytic reactor , will
itself increase the costs of down-stream units markedly .
Performing current experimental and model-based analy-
sis aims to assess the contribution of gas phase reactions on
the OCM reactor performance. A simplified catalytic model
which represents a solely radical generation mechanism is
exploited in this analysis as a key tool.
2 Experimentation
The performed experimental activities in this investigation
are briefly reported here. The catalyst preparation and
characterization have been described elsewhere in details
[33].
The Na
2
WO
4
/Mn/S iO
2
catalyst has shown a promising
potential [34 – 38] and was used in the experimental part
of this project. In most of the reports, this catalyst has
been prepared by the incipient wetness impregnation
method according to the reported recipe by W ang et al.
[35]. Detailed experimental aspects for testing the
Na
2
WO
4
/Mn/S iO
2
catalyst in fluidized-bed, fixed-bed and
membrane reactors have also been provided in our pre-
vious publications [23, 24, 31].
2.1 Experimental Setup
The original 600 mm long tubular a -alumina microfiltration
membrane was purchased from F raunhofer Institute for
C eramic T echnology and has a 7 mm inner diameter (ID),
1.5 mm thickness and 3 l m average pore size. This commer -
cially available membrane was coated with glassy materials
(BOTZ) to provide the desired permeation range of
1 – 15 cm
3
cm
–2
min
–1
bar
–1
[31]. The modified ceramic mem-
brane is implemented inside the metal reactor module and
separates the shell side with the inner diameter of 20 mm
Chem. Ing. T ech. 2014, 86, No. 11, 1906–191 5 © 2014 WILEY -VCH V erlag GmbH & Co. KGaA, W einheim www .cit-journal.com
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and the tube side with the outer diameter (OD) of 10 mm.
Inside the ceramic membrane, a WIKA k-type multipoint
thermocouple (with 3 mm OD stem) is located along the
reactor , which enables measuring the reaction temperature
in 10 points.
The temperature of the fluidized-bed reaction zone was
also recorded via multipoint thermocouple in 8 points. In
the fluidized-bed reactor setup the operating temperature
can be controlled in the range of 10 K deviation along the
bed. It has a 40 mm inner diameter . The block-flow dia-
grams of the reactor setups are shown in Fig. 1.
The membrane reactor module was implemented inside
an electrical tubular furnace where the operating tempera-
ture is varied using a two-zone electrical heating. T wo sepa-
rate thermocouples are located behind the metal shield
inside the furnace and measure the temperatures in each
zone of the furnace. By assigning the set points to these
temperatures and controlling them, the heat duty of each
electrical heating element is tuned. Therefore, the applied
high-frequency on/off heating/non-heating mechanism in
this furnace allows controlling the temperature along the
reactor . A similar split tubular furnace was used for the flui-
dized-bed reactor . It should be highlighted that the recorded
operating temperature inside the catalytic bed is affected
mainly by two factors, namely 1) the reaction rate and che-
mically generated heat, and 2) the value of the set tempera-
ture in the surrounding electrical tube furnace (TW) [32].
The second factor also affects the reaction rate inside the cat-
alytic packed bed. The values of the applied temperature of
the electrical heaters, the set points and the measured
values of the feed flow rates and operating pressures were
set and monitored online.
An IR gas analyzer monitors the concentration of ethy-
lene, ethane, methane, carbon dioxide, and oxygen in the
reactor outlet gas stream. The precision of the measure-
ments performed using the IR gas analyzer was confirmed
using GC sampling. Having considered the precision of the
measurements and control devices, the observed selectivity ,
conversion and yield in average have ± 10 % margin of error
in reference to their reported values in this paper .
www .cit-journal.com © 2014 WILEY -VCH V erlag GmbH & Co. KGaA, W einheim Chem. Ing. T ech. 2014, 86, No. 11, 1906–191 5
Figure 1. The block-flow diagram of the investigated reactor setups; (a) flui dized-bed reactor setup; (b) fixed-bed and membrane
reactor setup.
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3 Model-Based Analysis
A comprehensive micro-kinetic model of the Na
2
WO
4
/Mn/
SiO
2
catalyst has been published by Lee and coworkers [39].
This mechanism represents two adsorption steps and
14 surface reactions. However , the residual of each kinetic
parameter (uncertainty for calculating the value of para-
meters) is between 0.33 – 1. This is due to the fact that not
considering the gas phase reactions leads to substantial
errors and the conclusion that the effect of gas phase reac-
tions should be included in the kinetic model. On the other
hand, the mechanism suggested by Dooley et al. includes
269 species and 1583 reactions and efficiently represents
the gas phase reactions.
In this research, a semi-empirical micro-kinetic surface
reaction model of three reaction steps (T ab. 1) is proposed as
a constant source for methyl radical formation to represent
the catalytic surface reaction, which can interact with the
gas phase reactions along the reactor .
Considering dummy thermodynamic data and the reac-
tion order of zero for oxygen enable the assumption of the
steady state conditions at the catalyst surface and neglecting
the influence of the generated heat at the catalyst surface on
the adsorption processes. Kinetic properties were chosen in
a way that the reoxidation (rate equations 1 and 3 in T ab. 1)
is not limited and it is considered to be much faster than
the methane activation. This reduced surface model allows
the simulation of an ideal catalyst which enables analyzing
the effects of gas phase reactions. The relative rates of the
gas phase and surface reactions and their dependency on
the operating temperature and activity of catalyst has been
investigated before [40]. There, it has been reported that the
methyl radical generation over the catalyst surface is a
strongly temperature-dependent reaction and is dramati-
cally slower than the temperature-independent radical cou-
pling reaction in the gas phase.
Pre-exponential factors and activation energies reported
in T ab. 1 were fitted to predict the performance of the fixed-
bed reactor . F or accurate modeling, both the thermal and
reaction performance indicators were exploited to determine
the kinetic parameters. For instance, it was tried to capture
the observed axial temperature profiles of the fixed-bed and
membrane reactor in the simulation. No radial profile was
considered, because no adequate experimental data with
regard to the radial temperature profile were available.
Moreover , enormous amounts of numerical calculations are
required in the case of using two-dimensional models which
leads to a long time simulation. Side reactions on the reac-
tor wall and on the quartz granulates were neglected, since
they have been found to be negligible also in former experi-
ments [31]. The specific surface area was set to 2 m
2
g
–1
for
the catalyst, which is consistent with the BET analysis of this
catalyst [33]. The number of active centers was set to
4O *n m
–2
, which is in the range of reported values for var -
ious supported oxidation catalysts [41, 42].
The software Reaction design Chemkin was used for the
simulation. Pressure drop was calculated by the Ergun
equation for the fixed-bed and membrane reactors. The
Reynolds number for the reactor was calculated to be
around 5700, indicating a turbulent flow regime. Consider -
ing these facts makes it possible to simulate the fixed-bed
reactor with the PFTR (Plug Flow T ubular Reactor) model.
Simulations of the membrane reactor were carried out by
dividing the reactor into a system of 7 equal PFTR’s, follow-
ing the same assumptions as mentioned for the fixed-bed
reactor . Between each PFTR pair , a zero dimensional mixer
with an oxygen inlet was placed, as oxygen source and for
representing the membrane reactor . The fluidized-bed reac-
tor was modeled by a CSTR (Continuous Stirred-T ank Reac-
tor) model which advocates an isothermal performance. All
reactor and catalyst properties were kept for simulations as
they were in the experimentation.
4 Results and Discussion
In the context of this research, the observed performance of
OCM reactors in miniplant-scale operation, especially the
effects of methane-to-oxygen ratio and gas dilution in all
three reactors, were analyzed.
Increasing the methane-to-oxygen ratio usually increases
the C
2
selectivity mainly because of reducing the partial
pressure of oxygen in the reaction mixture. Such a trend is
typical as shown in Fig. 2 for the co-feed as well as the oxy-
gen dosing reactors. Nevertheless, using inert gas dilution
reduces the partial pressure of both oxygen and methane in
the gas phase and therefore also reduces the rates of the un-
desired gas phase reactions. Using a proper amount of gas
dilution also can improve the OCM reactor performance as
a result of establishing a better thermal operation. H ere, it
is attempted to distinguish between these thermal and reac-
tion effects. This can be achieved by analyzing the effects of
dilution in fixed-bed, membrane, and isothermal fluidized-
bed reactor . Special attention is devoted to highlight the
effect of dilution on the intensity of gas phase reactions in
an OCM react or . F or instance, under the isot hermal perfor -
mance of the OCM fluidized -bed reactor , it has been obser ved
that introducing higher amounts of nitrogen signif icantly
affects the react or perfor mance by improvi ng the C
2
selec tiv-
ity and suppress ing the undesire d gas phase reactions [24].
In all cases, nitrogen dilution significantly enhances the
C
2
selectivity . Therefore, usually the lowest C
2
selectivity is
Chem. Ing. T ech. 2014, 86, No. 11, 1906–191 5 © 2014 WILEY -VCH V erlag GmbH & Co. KGaA, W einheim www .cit-journal.com
T able 1. Pseudo micro-kinet ic for surface reactions to describe
the generation rate of methyl radicals.
No. Reaction step k
pre
[s
–1
] E
A
[kJ mol
–1
]
1O
2
+2V * → 2O * 3 · 1 0
6
40
2C H
4
+O * → CH
3
.
+ OH* 3 · 10
5
120
3 2 OH* → H
2
O + O* + V* 3 · 10
6
40
Research Article 1909
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achieved when the lowest nitrogen dilution and methane-to-
oxygen ratios are employed. In the case of membrane and
fixed-bed reactors, however , the effects of dilution (using
excess methane or nitrogen dilution) are coupled with the
thermal heat transfer effect.
In an OCM membrane reactor , as is seen in Fig. 2, despite
of the observed moderate reduction in methane conversion,
C
2
yield usually is improved by increasing the nitrogen
dilution. By increasing the methane-to-oxygen ratio in a
membrane reactor , C
2
yield usually is decreased due to the
reduction in both methane conversion and operating tem-
perature. Therefore, as has been demonstrated in Fig. 2c,
due to the thermal effect of gas dilution (using excess
methane or nitrogen dilution), the trend of C
2
yield
might show a local maximum in membrane reactor op-
eration.
It can be generally concluded that the quantitative impact
of the inert dilution on improving the C
2
yield in the OCM
reactor depends on the current range of C
2
-selectivity , the
intensity of the generated reaction heat and the possibility
to control the reaction temperature and maintain the de-
sired level of methane conversion. The last thermal factor is
not reflected in the isothermal fluidized-bed reactor .
Similar trends were observed for the performance of the
Sr -La
2
O
3
/C aO catalyst in a fluidized-bed reactor but with
lower C
2
selectivity and yield. Beside these two important
stable catalysts, many other types of OCM catalysts have
been reported in literature. F or instance, LiMgO catalyst has
been used in several experimental researches. However , this
catalyst and most of other catalysts so far reported for OCM
have not proven to be stable enough for continuous indus-
trial application.
4.1 Simulation Results
After reviewing the experimental observations, it is tried to
distinguish these aspects in the model-based analysis of the
OCM performance in the different reactors based on differ -
ent sets of catalytic and gas phase reactions. All obtained
experimental and simulation results are shown in T ab. 2.
H ere, X , Y , S represents the conversion, yield and selectiv-
ity , respectively . All experimental data were selected at simi-
lar methane conversion for better comparison. T wo major
differences between these reactor types are the mixing beha-
vior and the oxygen feeding policy . On the one hand, the tur -
bulent gas fraction is responsible for a well-mixed isother -
mal fluidized bed reactor . On the other hand, a lower void-
fraction in the fixed-bed or membrane reactor results in
local mixing and non-uniform temperature distribution.
www .cit-journal.com © 2014 WILEY -VCH V erlag GmbH & Co. KGaA, W einheim Chem. Ing. T ech. 2014, 86, No. 11, 1906–191 5
Figure 2. Effect of gas dilution on the performance of the investigated reactors; (a, b) flui dized-bed; (c, d) membrane reactor .
1910 Research A rticle
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This also means that more space is available for the gas
phase reactions in the fluidized-bed reactor . F urthermore,
the initial partial pressure of oxygen is relatively higher for
the fixed-bed and fluidized-bed reactor in comparison to its
low value in the membrane reactor . These differences can
explain the following observations.
The experimental data show a higher C
2
selectivity in the
membrane reactor in comparison to the fluidized-bed reac-
tor . All different models describe this trend with different
accuracy . Simulations of the fixed-bed and membrane reac-
tor with the Dooley mechanism alone (gas phase) show very
low methane conversion in comparison to the fluidized-bed
reactor . F or the fluidized-bed reactor , significant combustion
products were predicted. This is the first indication of domi-
nant gas phase side reactions for this reactor due to its high
initial oxygen concentration and a totally mixed system.
The combination of gas phase reactions and the micro-
kinetic model for the catalyst surface reported by Lee et al.
shows more complicated results. Simulation results for the
membrane reactor in this case show no methane conver -
sion. This is mainly due to the presence of high amounts of
nitrogen, which results in an extremely low oxygen partial
pressure, low space for the gas phase reactions and a short
residence time. This is an indication that the membrane
reactor model has a strong quenching behavior . The high
amount of nitrogen and low oxygen partial pressure prevent
any reactions in the gas phase and consequently over the
catalyst surface. T esting the initial conditions for the fixed-
bed reactor in this case also showed no methane and oxygen
conversion. F or the fixed-bed reactor the precise prediction
of methane conversion and C
2
selectivity was not possible
via the model-based analysis using the aforementioned
model combination. Moreover , the C
2
selectivity for the
simulated fluidized-bed reactor drops slightly in comparison
to the case where a pure gas phase model was used. This is
due to the effect of catalytically formed ethane and/or ethy-
lene which are less stable than methane and might be oxi-
dized by gas phase reactions. The oxidation of the formed
methyl radicals may also be a plausible
reason for these observations.
Implementation of the Dooley gas
phase reaction network and the simple
micro-kinetic surface reaction model
reported in T ab. 1 show a similar situa-
tion for prediction of conversion and
selectivity in all reactor types. It is seen
that the proposed simple micro-kinetic
surface reaction model is at least as
good as the model which used the de-
tailed gas phase and surface micro-ki-
netic reported by Lee et al. In case of
the membrane reactor , the observed
performance indicators are not differ -
ent from the other model-based analy-
sis. Therefore, it can be concluded that
the observed very low values of
methane conversion are due to the quenching behavior of
this system.
A detailed prediction of the reaction progress in terms of
methane conversion and C
2
selectivity for the fixed-bed reac-
tor is shown in Fig. 3. The proposed simplified model pre-
dicts the methane conversion and C
2
selectivity for the fixed
bed reactor very satisfactorily . Its prediction potential in this
case is better than the model proposed by Lee et al. The C
2
selectivity at the reactor inlet is very high and drops con-
stantly as methane and oxygen conversion increase along
the reactor , while CO
X
selectivity increases. The products of
the gas phase reactions in this case determine the perfor -
mance of the fluidized-bed reactor and the accuracy of the
selectivity is not similarly high.
The fluidized-bed reactor shows significant contribution
of the gas phase reactions in all three models. This is espe-
cially highlighted in the simulation results based on the
Chem. Ing. T ech. 2014, 86, No. 11, 1906–191 5 © 2014 WILEY -VCH V erlag GmbH & Co. KGaA, W einheim www .cit-journal.com
T able 2. Comparing the observed performances of various reactor types and their simulation
results to identify the contributions of the g as phase and surface reactions.
Reactor Parameter Exp. result Dooley Mech. Dooley + Lee Dooley + semi-emp.
M embrane
reactor
X
CH
4
0.33 ≈ 0 ≈ 0 ≈ 0
S
C
2
0.65 0.55 0.99 0.99
S
CO
x
0.35 0.45 0.01 0.01
F ixed-bed
reactor
X
CH
4
0.32 0.01 0.02 0.34
S
C
2
0.53 0.99 0.65 0.56
S
CO
x
0.47 0.01 0.35 0.44
Fluidized-bed
reactor
X
CH
4
0.35 0.29 0.29 0.30
S
C
2
0.40 0.14 0.11 0.12
S
CO
x
0.60 0.86 0.89 0.88
Figure 3. Prediction of the reaction-progress in the fixed-bed re-
actor using the gas phase model and pseudo surface kinetic mod-
el with measured temperature profile (3 nmL s
–1
N
2
; 2 nmL s
–1
CH
4
;
1 nmL s
–1
O
2
; 3.5 g catalyst; 3.5 g inert quartz).
Research Article 1911
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pure gas phase reactions. Including the surface kinetic mod-
els into the model also shows similar results. The simu-
lation results of the fixed-bed reactor show relatively low gas
phase reactivity . Including the surface reactions in this
simulation in the form of the proposed simple micro-kinetic
model results in an acceptable prediction of the methane
conversion and C
2
selectivity . For the case of membrane
reactor , which has the lowest gas phase reactivity , all models
fail to predict the experimental results. All these facts lead
to the conclusion that there is a very complex situation at
the surface of OCM catalysts, which cannot be easily de-
scribed by a standard model and considering a single active
site. In fact, catalytic activity can be divided into two parts.
This follows the idea of Ahari et al. [43], which is presented
in Fig. 4.
On one hand, the activity and conversion rates correspond
to Eqs. (1) and (2) in T ab. 1. On the other hand, the selectiv-
ity is dominated by deep oxidation mechanisms over the cat-
alyst surface and the methyl radical coupling process in the
gas phase. Both of these interactive phenomena together
are responsible for product formation. After this step, the
product is completely decoupled from the catalytic system
and follows the basics of thermodynamics and kinetics of
gas phase reactions. Improving the yield can only be
achieved by controlling these phenomena.
The first part of the catalytic cycle is started by the funda-
mental phenomenon of adsorption. In all kinetic models,
the molecular oxygen is adsorbed and interacts directly with
an active center of the catalytic material. The newly formed
species are responsible for methane activation as well as
deep oxidation processes over the surface. The stoichio-
metry of the active oxygen species is set to molecular or
atomic oxygen, but still unclear . In our previous paper , it
was shown that a more complex adsorption phenomenon
takes place on the Na
2
WO
4
/Mn/S iO
2
catalyst. It was shown
that two different oxygen species determine the type of cata-
lytic activity . On the one hand, there is a weakly bound oxy-
gen species, which is responsible for deep oxidation reac-
tions. On the other hand, there is a stronger adsorbed
species, which activates methane in a selective way [44].
This fact makes it clear that the characteristics of oxygen
adsorption play a crucial role on the surface activity . How-
ever , another parameter is the number of active sites. It is
very difficult to determine this number , since it requires
identification as well as the quantification.
The methane activation process follows the properties of
the different adsorbed oxygen species at the catalyst surface.
Right at this point, each of the investigated reactors will be
differently affected by this behavior because of the different
oxygen dosing strategies. The membrane reactor has an
excess of nitrogen in the axial inlet position. The oxygen par -
tial pressure increases along the axial position and promotes
more and more oxygen adsorption. In the fixed-bed reactor ,
the initial oxygen partial pressure is on its maximum and
drops by catalytic and gas phase reactions. At the end of the
reactor , the oxygen adsorption limits the catalytic activity .
F or the fluidized-bed reactor , the situation is totally differ -
ent. The strong mixing of reactants and products in this
reactor provides an almost constant partial pressure of oxy-
gen in the whole volume.
The second point is that the selectivity is controlled in the
beginning by the process of coupling of methyl radicals in
competition to the methane combustion reaction. This step
can later be influenced by the adsorbed oxygen, which leads
to deep oxidation of the products. Besides observing these
surface phenomena, the same situation is present in the gas
phase. In addition to this point, each reactor system influ-
ences this sensitive reaction network also by different contri-
butions of void fractions and oxygen dosage. In the mem-
brane reactor the coupling process is the dominant step in
the gas phase. The deep oxidation takes place at the catalyst
surface. In contrary , the fluidized-bed reactor provides a lot
of gas phase oxygen which makes it easier to follow the
route of products combustion in the gas phase reactions net-
work. This is clearly shown according to the simulation
results reported in T ab. 2 and Fig. 3 where the fluidized-bed
reactor model using only Dooley’s mechanism, which repre-
sents only the gas phase reactions, shows a high level of
methane conversion and a very low level of selectivity . The
predominant process can be recognized from F ig. 3. H ere it
can be explained that the formation of ethane is secured via
coupling of methyl radicals. This step is later followed by
combustion processes as higher conversions and lower
selectivity are obtained.
The reaction routes in the fixed-bed reactor change over
the length. At the inlet of this reactor , the coupling process
dominates to provide a high selectivity at low conversions,
which changes with the formation of thermodynamically
less stable products like ethane and ethylene. Along with
intensive methyl radical coupling, the C
2
-products are
exposed to the gas phase reaction network. The easier activa-
tion of ethane compared to methane plays the major role
now and starts facilitating the competitive oxidation. As a
result of this situation, further oxidation of ethylene is also
encountered, ending with the significant concentration of
deeply oxidized products. At this point, the simple micro-
www .cit-journal.com © 2014 WILEY -VCH V erlag GmbH & Co. KGaA, W einheim Chem. Ing. T ech. 2014, 86, No. 11, 1906–191 5
Figure 4. Sch eme of the OCM surface reaction mechanism [43].
1912 Research A rticle
Chemie
Ingenieur
T echnik
kinetic surface mechanism shows already that the gas phase
reactions significantly contribute to the CO
x
production
since no deep oxidation reactions are included in this mod-
el. Beside the effect of oxygen partial pressure, temperature
also has the strongest influence on this network (Fig. 5).
The initial feed temperature was set to 800 °C and the heat
transfer potential was not considered in the model. This sce-
nario leads to a hot spot formation. The combustion of
alkanes strongly increases the local heat formation. At this
point, the reaction system favors the combustion routes,
which is accelerating itself. It should be emphasized, how-
ever , that the selectivity is affected by the level of hot spot
and conversion.
The membrane reactor quenches most of the gas phase
steps, as shown in the simulation. Reviewing the reported
data in T ab. 2 for the model which uses Dooley mechanism
shows no gas phase activity , which is due to quenching
effect. This is the main explanation for the highest observed
C
2
selectivity in the experiments. The fluidized-bed reactor
shows significant potential for the gas phase reactivity ,
which was reflected in the experimental and simulation
results as the high selectivity towards deep oxidation pro-
ducts. F or the fixed-bed reactor , the fraction of C
2
products
causes more combustion products, because the gas phase
oxidation has a strong influence at the outlet, but shows a
minor oxidation potential at the reactor entrance where
more stable reactants exist.
It became clear that oxygen adsorption, activation phe-
nomena, and the gas phase reaction network influence each
other strongly , which limits the methane conversion as well
as the C
2
selectivity . This becomes most evident by the fact
that a simple methyl radical formation mechanism in com-
bination with a comprehensive gas phase model predicts
most of the observed results in three different reactors in a
satisfactory way .
It can be clearly concluded that the reactor design has a
strong influence on the ratio of gas phase and catalyst con-
tributions. In order to overcome the yield limitation, as a
consequence of restricted conversion and declining selectiv-
ity , the reactor engineering should clearly be the main focus
of the further OCM process development.
4.2 Limitations of Kinetic Modeling
The simplified model implemented in this research consists
of an individual surface-activation step and several gas
phase reactions and was clearly able to predict the general
trends of the reactor performance. As a result, it was high-
lighted how important it is to consider the contribution of
the gas phase reactions and it became clear that the gas
phase chemistry is not only a minor side reaction, but an
essential part of the main reaction path. However , this is
only a first step and further detailed mechanisms and inter -
active effects should be considered in order to improve the
predicting potential of kinetic models.
The used minimal surface mechanism is not complete,
but due to the lack of data and considering the nature of
consecutive reactions, they have not been included in this
simulation. The modeling approach in this manuscript is
completely different from the ones reported in literature.
H ere, it was aimed to see how the contribution of the gas
phase reactions can be predicted using this simplified mod-
el. Another aim is the identification of missing steps in the
surface model.
F or instance, the Lee mechanism has several surface spe-
cies included, but shows only low activity , which indicates
the restrictions with regard to predicting the activation of
hydrocarbons. This fact makes it clear that this micro-
kinetic model does not cover all necessary reactions steps or
its kinetic parameters have not been determined adequately .
Moreover , the implemented simplified model in this
research fails to predict the performance of the reaction
atmosphere with low partial pressure of oxygen because of
too low values of oxygen binding parameters of the surface
of the catalyst. These values should be properly tuned in
order to ensure realistic levels of catalytic activation of all
involved species.
Another aspect which should be considered is the forma-
tion of hydrogen or carbon deposition, as was observed in
our experiments. However , similar to several other available
kinetic models [10, 11], these aspects have not been consid-
ered in the modeling due to difficulty of establishing the full
mass balance of the species in separate streams.
Chem. Ing. T ech. 2014, 86, No. 11, 1906–191 5 © 2014 WILEY -VCH V erlag GmbH & Co. KGaA, W einheim www .cit-journal.com
Figure 5. Reaction progress of the fixed-bed reactor with the gas
phase model and simple micro-kinet ic surface model by solving
the free energy equation (hot spot formation) with the initial
temperature of 800 °C (3 nmL s
–1
N
2
; 2 nmL s
–1
CH
4
; 1 nmL s
–1
O
2
;
3.5 g catalyst; 3.5 g inert quartz).
Research Article 1913
Chemie
Ingenieur
T echnik
4.3 General Thermal Engineering Observations
In the context of OCM reactor engineering, it is not an exag-
geration to claim that the thermal engineering remains to
be the main challenging task in designing an efficient OCM
reactor . It can be concluded that due to the applied heating/
non-heating procedure in the electrical furnace, the cooling
effects of the inert gas dilution might not be completely
tracked down in the observed experimental trends. Most of
the experimental studies reported for the OCM reactors so
far have also been performed using electrical furnaces and
have exploited similar mechanism to control the thermal
performance of the OCM reactors. Therefore, in the re-
ported trends of the reactor performance indicators in these
cases, also the effect of changing the reaction temperature
is not clearly separable from the effect of introducing the
inert dilution. This thermal control mechanism acts like
implementing a constant reactor wall policy . In case of
implementing a constant heat flux, this effect will be better
highlighted and distinguished. However , this procedure also
has its own limitations and in any types of the applied con-
trol mechanisms using electrical furnaces, there is a need
for fine-tuning the frequency of the electrical power to
maintain a stable thermal and reaction performance in the
OCM membrane reactor .
The procedure of heating also has a crucial impact on the
reactor performance. In order to clarify this, the thermal
reaction performance of the OCM reactor was investigated
under the co-feeding scenario. It was observed that when
the operating temperature is rising, the C
2
yield and C
2
H
4
yield are also rising up to a certain level. In this experimen-
tation, even a C
2
yield of 29.5 % and a C
2
H
4
yield of 23.6 %
were temporarily recorded. After a short time, the yield
decreases again. This is assumed to be closely related to the
rate of methyl radical generation over the surface of the
OCM catalyst and the intensifying combustion effect in the
gas phase due to temperature rise. This phenomenon was
observed repeatedly in numerous experiments.
5 Conclusions
According to the performed experimental and model-based
analysis, it was shown that the proposed simple micro-
kinetic surface model coupled with a comprehensive gas
phase model is capable of predicting the performance of a
fixed-bed and a fluidized-bed reactor with the same preci-
sion as a detailed formal kinetic model does.
The fraction of weakly or strongly adsorbed oxygen over
the catalyst surface is a crucial factor for the selectivity . Oxy-
gen feeding policy plays an important role in this regard.
F or instance, in the membrane reactor where the oxygen
partial pressure increases gradually along the bed, more oxy-
gen adsorption and possibly stronger adsorption will be
observed at the end of reactor . It is opposite to the case of
the co-feed fixed-bed reactor . At the front end of the fixed-
bed reactor , weak oxygen adsorption limits the catalytic
selectivity . In the fluidized-bed reactor , however , a signifi-
cant portion of gas phase reactions is responsible for the ob-
served low selectivity .
Acknowledgment
The authors acknowledge the financial support from the
Cluster of Excellence „Unifying Concepts in Catalysis“ coor -
dinated by the T echnische Universität Berlin and funded by
the German Research F oundation (Deutsche F orschungs-
gemeinschaft).
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Research Article 1915
Chemie
Ingenieur
T echnik
Paper 2
Catalysis
Today
228
(2014)
212–218
Contents
lists
available
at
ScienceDirect
Catal y sis
T oda y
j
o
ur
na
l
ho
me
page:
www.elsevier.com/locate/cattod
Oxidative
coupling
of
methane—A
complex
surface/gas
phase
mechanism
with
strong
impact
on
the
reaction
engineering
Benjamin
Beck a ,
Vinzenz
Fleischer a ,
Sebastian
Arndt a ,
Miguel
González
Hevia b ,
Atsushi
Urakawa b ,
Peter
Hugo a ,
Reinhard
Schomäcker a , ∗
a Technische
Universität
Berlin,
Institut
für
Chemie,
Straße
des
17.
Juni
124,
10623
Berlin,
Germany
b Institute
of
Chemical
Research
of
Catalonia,
Avgda.
Països
Catalans
16,
43007
Tarragona,
Spain
a
r
t
i
c
l
e
i
n
f
o
Article
history:
Received
13
July
2013
Received
in
revised
form
14
November
2013
Accepted
23
November
2013
Available
online
9
January
2014
Keywords:
Mn/Na 2 WO 4 /SiO 2
Oxidative
coupling
of
methane
High
pressure
Temporal
analysis
of
products
Maximum
yield
a
b
s
t
r
a
c
t
The
oxidative
coupling
of
methane
over
Mn/Na 2 WO 4 /SiO 2 has
been
investigated
at
reaction
conditions
suitable
for
industrial
applications
up
to
10
bar
in
a
fixed
bed
reactor
as
well
as
by
temporal
analysis
of
products
(TAP)
by
admitting
pulses
of
methane,
ethane
and
ethene
oxygen
mixtures.
The
influence
of
pressure
on
selectivity
is
investigated
and
a
concept
for
optimizing
it
is
derived.
A
maximum
yield
is
estimated
from
the
ratios
of
the
involved
main
reactions
of
the
reaction
network
assuming
that
the
undesired
parallel
reactions
can
be
suppressed.
©
2013
Elsevier
B.V.
All
rights
reserved.
1.
Introduction
Oxidative
coupling
of
methane
(OCM)
to
ethylene
offers
great
industrial
potential,
because
it
would
broaden
the
feedstock
basis
for
chemical
industry.
Today
crude
oil
derived
olefins
and
aromatic
hydrocarbons
via
steam
cracking
of
naphtha
are
still
the
crucial
raw
materials
for
the
majority
of
value
added
chains
in
chemical
industry.
The
importance
of
alternative
carbon
sources
for
chemical
processes
becomes
greater
as
oil
price
increases
[1] .
For
the
development
of
an
OCM
process
Mn/Na 2 WO 4 /SiO 2 has
been
highlighted
as
catalyst
from
the
rich
literature
on
OCM
[2,3] ,
but
little
is
known
about
its
structure
and
the
reaction
mecha-
nism
at
this
catalyst.
Its
reported
stability
and
high
yield
were
the
important
motivations
to
select
the
catalyst
for
detailed
studies
[4–6] .
However,
the
obtained
yield
still
needs
to
be
improved
for
industrial
application.
In
most
reports
the
focus
is
mainly
directed
to
the
catalyst
mate-
rial,
and
reactor
setups
and
conditions
are
designed
and
selected
in
a
way
that
no
reaction
of
methane
is
observed
in
an
empty
reactor.
One
has
to
mention
that
for
industrial
application
high
pressures
are
necessary
to
make
the
process
economically
viable,
but
detailed
studies
of
empty
reactors
at
higher
pressures
are
∗ Corresponding
author.
Tel.:
+49
30
31424973.
E-mail
address:
[email protected]
(R.
Schomäcker).
missing
in
literature.
Furthermore,
the
consecutive
reactions
of
the
targeted
C 2 products
has
the
strongest
negative
impact
on
the
yield
[7–9] .
Reactivity
data
of
ethane
or
ethylene
at
increased
pressures
are
almost
not
existent
in
the
literature.
At
typical
OCM
conditions
of
800 ◦ C
virtually
no
material
typically
used
for
these
reactors
or
packing
is
absolutely
chemically
inert.
For
the
development
of
an
industrial
scale
reactor
one
should
expect
that
these
factors
contribute
even
more
to
the
reaction,
due
to
use
of
materials
like
stainless
steel.
The
purpose
of
this
paper
is
to
point
out
important
factors
which
have
to
be
considered
for
the
reaction
engineering
of
OCM
including
the
interaction
of
gas
phase
and
surface
reactions.
The
design
of
an
industrial
scale
reactor
requires
knowledge
about
the
reaction
kinetics.
If
a
reaction
network
shares
strong
con-
tributions
of
gas
phase
and
surface
reactions,
it
is
very
difficult
to
derive
this
knowledge
from
experimental
series
at
typical
reaction
conditions
with
the
use
of
a
fixed
bed
reactor
only.
This
is
due
to
inseparable
kinetic
data
of
gas
phase
and
surface
contribution.
OCM
is
known
to
function
in
a
very
complex
reaction
network
containing
a
variety
of
surface
and
gas
phase
reactions,
which
is
depicted
in
a
simplified
scheme
in
Fig.
1 .
Obviously,
the
yield
of
ethane
and
ethyl-
ene
is
strongly
dependent
on
the
parallel
and
consecutive
reactions
to
carbon
oxides.
Obtaining
more
information
about
the
surface
catalyzed
activation
of
methane,
ethane
and
ethene
requires
sup-
pression
of
the
gas
phase
activation
of
these
components,
which
can
be
attained
at
the
very
low
pressure
such
as
the
condition
of
temporal
analysis
of
products
(TAP)
reactors.
0920-5861/$
–
see
front
matter
©
2013
Elsevier
B.V.
All
rights
reserved.
http://dx.doi.org/10.1016/j.cattod.2013.11.059
B.
Beck
et
al.
/
Catalysis
Today
228
(2014)
212–218
213
Fig.
1.
Simplified
reaction
network
of
the
oxidative
coupling
of
methane
at
oxides.
2.
Experimental
2.1.
Catalyst
synthesis
Mn/Na 2 WO 4 /SiO 2 catalyst
material
was
prepared
by
fluidized
bed
processing
of
the
constituent
substances
as
described
else-
where
[4] .
The
final
catalyst
contains
2
wt%
Mn(II)
and
5
wt%
Na 2 WO 4 .
A
detailed
characterization
was
also
performed
by
Simon
et
al.
[4] .
2.2.
Fixed
bed
reactor
OCM
was
investigated
in
a
stainless
steel
reactor
(65
cm
length
and
4
mm
i.d.)
equipped
with
an
inlay
made
of
corundum
to
min-
imize
catalytic
reactions
at
the
reactor
walls.
This
reactor
can
be
operated
up
to
10
bars.
The
reaction
temperature
was
between
700
and
800 ◦ C.
The
product
gases
were
analyzed
with
a
gas
chro-
matograph
GC-2014
by
Shimadzu,
which
was
equipped
with
a
methanizer
and
thermal
conductivity
and
flame
ionization
detec-
tors.
The
reactor
was
filled
with
a
fixed
bed
of
quartz
granules
of
400
to
800
m
diameter
up
to
height
of
38
cm
to
end
in
the
isother-
mal
zone
of
the
reactor.
On
the
top
of
the
quartz
bed,
a
small
piece
of
quartz
wool
was
placed
to
keep
the
100
mg
Mn/Na 2 WO 4 /SiO 2
in
position.
The
feeding
mixture
contained
95
vol%
methane
and
5
vol%
oxygen
with
an
overall
gas
flow
rate
between
50
to
500
N
ml/min,
in
such
a
way
that
the
residence
time
is
kept
con-
stant
with
varying
pressure.
The
shift
of
the
explosion
region
to
lower
oxygen
fractions
with
increasing
pressure
prohibits
the
usage
of
high
oxygen
fractions.
2.3.
Temporal
analysis
of
products
TAP
experiments
was
conducted
in
a
TAP-2
device,
which
is
described
in
detail
elsewhere
[10] .
In
a
small
reactor
(71
mm
length
and
4.6
mm
i.d.)
made
of
quartz,
fixed
bed
50
mg
of
cat-
alyst
(2–3
mm
thickness)
was
placed,
sandwiched
by
two
layers
of
quartz
particles.
Both
catalyst
and
quartz
particles
were
in
the
200–300
m
range.
Before
the
TAP
measurements
the
catalyst
was
pretreated
for
30
min
at
650 ◦ C
in
oxygen
flow.
Measurements
were
performed
at
temperatures
between
500
and
800 ◦ C
with
pulse
sizes
large
enough
to
be
affected
by
intermolecular
collisions
(ca.
10 17 molecules
per
pulse)
due
to
the
absence
of
desired
products
at
small
pulse
sizes.
Using
the
large
pulse
size,
the
mass
transport
is
generally
characterized
as
molecular
diffusion
regime.
Two
types
of
experiments
were
conducted:
simultaneous
puls-
ing
and
sequential
pulsing
of
oxygen
and
the
specific
hydrocarbon;
i.e.
methane,
ethane
or
ethylene.
The
pulsed
gas
consisted
of
a
hydrocarbon
to
oxygen
to
neon
(internal
standard)
mixture
of
2/1/4.
In
the
case
of
sequential
pulsing
the
methane
to
oxygen
ratio
cannot
set
accurately
due
the
usage
of
two
pulse
valves
and
is
adjusted
individually
to
obtain
accurate
TAP
responses.
In
the
case
Fig.
2.
Oxygen
conversion
with
resulting
product
distribution
at
constant
residence
time
at
various
pressures
at
700 ◦ C.
of
MgO
a
methane
to
oxygen
ratio
of
roughly
6
to
1
is
obtained.
In
the
case
of
Mn/Na 2 WO 4 /SiO 2 the
methane
to
oxygen
ratio
was
roughly
1
to
1.
The
response
pulses
of
each
component
were
calculated
by
averaging
ten
pulses
per
investigated
atom
mass
unit
(amu).
Since
carbon
monoxide,
carbon
dioxide,
ethene
and
ethane
share
some
relevant
amus,
it
was
necessary
to
use
the
fragmentation
patterns
of
these
compounds
for
quantification.
For
the
sake
of
compar-
ison,
we
have
also
investigated
magnesia
by
TAP
as
reference
beside
Mn/Na 2 WO 4 /SiO 2 .
It
is
well
known
that
doping
magnesia
with
lithium
results
in
an
enhanced
catalytic
performance,
but
the
instability
of
this
system
prevents
a
closer
investigation
by
TAP
experiments.
3.
Results
and
discussion
3.1.
Catalyst
characterization
A
detailed
characterization
of
the
used
Mn/Na 2 WO 4 /SiO 2 is
given
by
Simon
et
al.
[4] .
The
catalyst
support
was
homogenously
coated
with
Na/WO 4 and
Mn
precursors
resulting
in
crystalline
phases
of
Mn 2 O 3 and
Na 2 WO 4 .
The
support
is
transformed
into
the
stable
cristobalite
phase
with
a
macroporous
surface
morphol-
ogy.
The
final
catalyst
contains
2
wt%
Mn(II)
and
5
wt%
Na 2 WO 4 and
has
a
surface
area
of
1.86
m 2 /g.
The
catalyst
was
chemically
and
mechanically
stable
during
OCM.
Magnesium
oxide
was
purchased
from
Sigma-Aldrich
(99.9%)
and
has
a
surface
area
of
78
m 2 /g.
3.2.
Pressure
experiments
The
results
of
the
OCM
reaction
at
up
to
10
bar
and
700 ◦ C
are
shown
in
Fig.
2 .
Experiments
were
performed
with
an
empty
reac-
tor
(left),
a
fix
bed
of
quartz
granules
(middle)
and
a
fixed
bed
of
quartz
granules
with
a
layer
of
catalyst
on
top
of
it
(right).
The
tint
inside
the
columns
specifies
the
yield
to
carbon
oxides,
C 2 and
C 3 components,
respectively.
The
carbon
balance
for
the
pressure
experiments
did
not
deviate
more
than
0.5%.
Strikingly
the
empty
reactor
results
in
the
highest
conversion
of
oxygen
over
the
full
range
of
examined
pressure,
reaching
almost
complete
conversion
at
8
bar.
Only
a
small
increase
in
C 2+ yield
is
observed
above
5
bar,
which
is
caused
by
the
increasing
influence
of
the
total
oxidation
of
the
C 2+ components.
The
sigmoid
characteristic
of
the
oxygen
conversion
is
typical
for
reactions
with
orders
higher
than
one.
In
contrast,
the
fixed
bed
of
quartz
granules
results
in
a
reduced
res-
idence
time
that
leads
to
the
lowest
conversion
over
the
complete
range
of
pressure.
The
behavior
seems
to
be
almost
the
same
as
214
B.
Beck
et
al.
/
Catalysis
Today
228
(2014)
212–218
Fig.
3.
Influence
of
absolute
pressure
on
C 2+ selectivity
at
constant
residence
time.
in
the
empty
reactor
with
a
shift
to
higher
pressures.
Addition
of
Mn/Na 2 WO 4 /SiO 2 on
top
of
the
quartz
fixed
bed
regains
higher
oxygen
conversions.
The
absence
of
a
clear
sigmoid
behavior
with
increasing
pressure
indicates
a
more
surface
controlled
reaction
sequence,
with
a
much
higher
yield
of
C 2+ components
compared
to
the
reactor
loaded
only
with
quartz.
Fig.
3
shows
the
corresponding
selectivities
to
C 2+ at
the
investi-
gated
pressures.
The
empty
reactor
as
well
as
the
fixed
bed
of
quartz
granules
passes
through
a
maximum
selectivity
to
C 2+ at
4
and
8
bar,
respectively.
In
contrast,
the
addition
of
Mn/Na 2 WO 4 /SiO 2 results
in
a
steady
increase
of
selectivity
to
C 2+ with
increasing
pressure.
Its
maximum
is
not
reached
in
the
investigated
range
of
pressure.
The
C 2 selectivity
in
OCM
is
strongly
affected
by
consecutive
reac-
tions.
This
results
generally
in
a
loss
of
selectivity
with
increasing
conversion
due
to
total
oxidation
[11] .
The
observation
of
a
higher
selectivity
in
combination
with
a
higher
methane
and
oxygen
con-
version
implies
a
clear
benefit
of
higher
pressure
toward
better
C 2
selectivity.
For
detailed
understanding
of
the
homogeneous
gas
phase
as
well
as
the
heterogeneous
surface
catalyzed
reactions,
it
is
neces-
sary
to
quantify
their
contributions
to
the
overall
product
mixture
in
order
to
maximize
the
C 2+ yield.
Zanthoff
and
Baerns
developed
a
kinetic
model
for
the
gas
phase
contribution
including
33
species
and
192
elementary
steps,
which
describes
the
product
distribu-
tion
and
conversion
at
the
end
of
the
reactor
accurately
as
long
as
the
reactor
temperature
is
higher
than
600 ◦ C
[12,13] .
A
recent
investigation
of
the
spatial
compositional
profile
along
a
reactor
has
exhibited
discrepancies
with
the
simulated
profiles
with
respect
to
conversion
and
composition
[13] .
Better
consistency
could
be
achieved
by
extension
of
the
gas
phase
model
to
more
species
and
reactions.
Marin
et
al.
[14]
coupled
a
gas
phase
reaction
model
to
a
heterogeneous
catalysis
one.
However,
the
gas
phase
part
alone
of
this
model
does
not
explain
the
observed
increase
in
C 2+ selectivity
with
increasing
pressure
at
low
oxygen
conversions
with
quartz
granules
packing
( Figs.
2
and
3 )
[15] .
Nevertheless,
it
is
necessary
to
couple
the
gas
phase
model
to
a
model
describing
heterogeneous
surface
reactions.
At
typical
reac-
tion
conditions
their
kinetic
data
are
obviously
compromised
by
gas
phase
contributions.
This
unusual
approach
is
necessary
due
to
the
high
complexity
of
OCM.
Therefore
we
employed
the
TAP
tech-
nique
with
the
aim
to
directly
obtain
kinetic
data
of
the
CH
bond
activation
reaction
based
on
the
assumption
of
the
reaction
as
the
rate
determining
step.
Gaining
the
surface
reactivity
of
C 1 and
C 2
without
the
influence
of
gas
phase
reactions
is
expected
to
reveal
factors
maximizing
the
yield
of
desired
products
while
minimizing
the
formation
of
the
undesired
CO x by-products.
Fig.
4.
TAP
response
of
CH 4 /O 2 /Ne
(=4/1/2)
pulse
over
magnesia
at
700 ◦ C.
3.3.
Temporal
analysis
of
products
Generally
TAP
measurements
are
performed
with
small
pulse
sizes
resulting
in
gas
transport
by
Knudsen
diffusion,
in
which
gas
phase
intermolecular
collisions
are
negligible
[16–19] .
As
depicted
in
Fig.
1 ,
the
recombination
of
methyl
radicals
takes
place
in
gas
phase
as
reported
by
many
authors
[20–24] .
The
fact
that
we
did
not
detect
C 2 products
in
the
Knudsen
diffusion
regime
supports
this
hypothesis.
Figs.
4
and
5
show
the
time
resolved
pulse
responses
at
700 ◦ C
and
Table
1
provides
the
quantitative
evaluation
by
the
zeroth
moment
of
the
pulse
response
(i.e.
integrated
area
of
the
pulse)
for
both
catalysts.
The
product
responses
except
CO 2 are
noisier
than
the
reactant
responses
and
were
smoothened
for
better
illustration
in
Figs.
4
and
5 .
The
major
differences
in
the
reactivity
of
the
two
materials
are
the
lower
methane
conversion
and
higher
selectivity
to
ethane
of
Mn/Na 2 WO 4 /SiO 2 compared
to
those
of
magnesia.
The
carbon
balance
for
the
TAP
measurements
did
not
deviate
more
than
3%.
The
positions
of
the
peak
maxima
give
evidence
about
the
reac-
tion
sequence.
In
Fig.
4
one
can
clearly
see
that
carbon
monoxide
is
formed
in
parallel
to
the
formation
of
ethane
from
methane.
The
origin
of
CO
formation
via
a
consecutive
oxidation
of
methyl
radicals
is
reasonable
( Fig.
1 ).
Methyl
radicals
are
produced
from
methane
at
the
catalyst
surface
by
direct
interaction
with
surface
adsorbed
oxygen
[16] .
They
can
be
oxidized
before
desorption
or
by
re-adsorption
at
the
catalyst
or
quartz
surface;
if
not
they
combine
to
ethane
in
the
gas
phase.
The
delayed
peak
maxima
observed
for
magnesia
compared
to
those
of
Mn/Na 2 WO 4 /SiO 2 indicate
more
Fig.
5.
TAP
response
of
CH 4 /O 2 /Ne
4/1/2
at
Mn/Na 2 WO 4 /SiO 2 at
700 ◦ C.
B.
Beck
et
al.
/
Catalysis
Today
228
(2014)
212–218
215
Fig.
6.
TAP
responses
of
sequential
pulsing
of
first
oxygen
and
then
methane
with
a
delay
t
at
800 ◦ C.
stable
surface
intermediates
over
magnesia,
easing
the
consecu-
tive
oxidation
of
methyl
radicals,
which
could
be
the
major
reason
for
the
inferior
selectivity
to
C 2 products
( Table
1 ).
For
magnesia
the
response
time
of
carbon
dioxide
is
strongly
delayed,
indicating
that
it
is
produced
by
species
strongly
adsorbed
onto
magnesia
or
simply
carbon
dioxide
interacts
strongly
with
the
catalyst
surface.
Analogous
behavior
can
be
observed
when
using
ethane
or
ethyl-
ene
as
feed
gas
(not
shown).
No
conversion
could
be
observed
by
pulsing
mixtures
of
hydrocarbon
(methane,
ethane
or
ethene)
and
oxygen
in
a
reactor
only
filled
with
quartz
particles.
Therefore
a
gas
phase
activation
of
these
components
can
be
excluded.
Interest-
ingly,
the
oxygen
balance
cannot
be
closed
for
Mn/Na 2 WO 4 /SiO 2 .
Pulse
experiments
with
pure
oxygen
result
in
a
loss
of
oxygen,
which
can
be
explained
by
the
oxidation
of
the
material
partially
reduced
under
the
TAP
vacuum
conditions.
Furthermore,
sequential
pulse
experiments
were
used
to
inves-
tigate
the
reactivity
of
adsorbed
oxygen
at
800 ◦ C.
First
oxygen
is
pulsed
and
next
a
pulse
of
methane
is
given
with
an
offset
of
0.0,
0.5
and
1
s.
The
oxygen
response
for
magnesia
is
presented
in
Fig.
6 A.
The
response
of
the
oxygen
pulse
has
a
similar
shape
with
an
abrupt
decay
shortly
after
the
pulse
of
methane.
This
rapid
oxygen
consumption
is
attributed
to
weakly
adsorbed
but
highly
reactive
oxygen
species.
Due
to
decreasing
concentration
of
these
species
with
time,
a
decrease
of
the
reaction
products
with
increasing
the
delay
of
the
methane
pulse
is
expected
to
be
observed.
Carbon
diox-
ide
is
identified
as
product
showing
the
expected
behavior
( Fig.
6 B).
In
contrast,
in
Fig.
6 C
the
pulse
intensity
of
ethane
is
not
decreasing
by
delayed
methane
pulses.
This
indicates
that
the
presence
of
a
second
species
of
more
strongly
adsorbed
oxygen
is
responsible
for
the
methyl
radical
formation.
It
should
be
mentioned
that
no
Table
1
Reactivity
data
of
the
two
catalysts
determined
from
TAP
data
at
700 ◦ C.
Catalyst
Conversion
CH 4 (%)
S
CO
(%)
S
CO 2 (%)
S
C 2 H 6 (%)
S
C 2 H 4 (%)
Mn/Na 2 WO 4 /SiO 2 ∼ 2
3.6
1.4
95
0
MgO
16.0
83.4
14.6
2
0
216
B.
Beck
et
al.
/
Catalysis
Today
228
(2014)
212–218
Table
2
Integrated
MS
responses
of
the
components
presented
in
Fig.
6 .
Catalyst
t
[s] MgO
Mn/Na 2 WO 4 /SiO 2
O 2 CO 2 C 2 H 6 O 2 CO 2 C 2 H 6
0
2.15E
−
2
8.93E
−
2
1.15E
−
2
7.83E
−
1
8.59E
−
3
1.33E
−
2
0.5
1.58E
−
1
5.28E
−
2
1.16E
−
2
1.15
8.23E
−
3
1.28E
−
2
1
2.23E
−
1 4.50E
−
2 1.16E
−
2 1.21
6.77E
−
3
1.38E
−
2
product
formation
occurs
by
pulsing
first
methane
and
then
oxy-
gen.
The
oxygen
response
for
Mn/Na 2 WO 4 /SiO 2 is
shown
in
Fig.
6 D.
When
the
two
pulses
are
made
at
the
same
time
the
total
pulse
size
is
bigger,
farther
from
Knudsen
diffusion
regime,
and
convec-
tive
flow
is
more
important,
rendering
the
responses
narrower.
Immediately
after
methane
enters
the
reactor,
the
oxygen
response
passes
through
a
small
maximum,
which
is
not
followed
by
an
abrupt
decay.
The
small
maximum
is
an
artifact
resulting
from
the
pulsing
of
methane
and
is
also
visible
in
the
inert
gas
response
(not
shown).
The
missing
fast
decay
is
attributed
to
the
lack
of
highly
reactive
adsorbed
oxygen
species
as
observed
for
MgO.
The
integrated
quantities
of
the
sequential
pulse
experiments
are
pre-
sented
in
Table
2 .
Due
to
the
different
methane
to
oxygen
ratios
the
obtained
quantities
are
not
directly
comparable
with
the
simulta-
neous
pulse
experiments.
Nevertheless,
Mn/Na 2 WO 4 /SiO 2 shows
a
similar
response
to
that
of
magnesia
at
varying
delay
of
methane
pulse
for
carbon
dioxide
and
ethane
as
presented
in
Fig.
6 E
and
F,
revealing
the
presence
of
a
second
species
of
more
strongly
adsorbed
oxygen.
Therefore
it
is
reasonable
that
the
ratio
between
strong
and
weak
adsorbed
oxygen
species
is
responsible
for
the
overall
much
higher
C 2 selectivity
of
Mn/Na 2 WO 4 /SiO 2 compared
to
that
of
magnesia.
Simultaneous
pulsing
of
methane
and 18 O 2 was
also
examined
to
get
closer
insight
into
the
nature
of
the
oxygen
species
responsi-
ble
for
methyl
radical
formation.
Unfortunately,
no 18 O
containing
products
were
detected
unless
a
very
large
amount
of 18 O 2 was
used.
Interestingly,
the
response
at
amu
36
is
almost
replaced
by
responses
at
the
amu’s
32
and
34
indicating
a
very
fast
exchange
between
gas
phase,
adsorbed
and
lattice
oxygen.
Determination
of
apparent
activation
energies
requires
a
constant
amount
of
active
sites
within
the
studied
temperature
range.
Due
to
the
reduction
of
Mn/Na 2 WO 4 /SiO 2 in
vacuum,
that
is
temperature
dependent,
a
constant
amount
of
active
sites
is
not
provided.
A
reduction
of
magnesia
in
vacuum
was
not
exper-
imentally
observed.
Therefore
the
temperature
dependence
of
the
surface
reactions
of
methane
and
the
desired
intermediates
can
only
be
provided
for
magnesia.
Rate
constants
were
determined
by
the
zeroth
moment
of
the
reactant
response
assuming
a
first
order
rate
expression
of
the
hydrocarbon
(methane,
ethane
or
ethene)
and
oxygen
[19] .
Therefore,
apparent
activation
energies
could
be
determined
from
the
Arrhenius
plot
in
Fig.
7 ,
as
presented
in
Table
3 .
A
loss
of
active
sites
by
agglomeration
was
not
observed.
Inter-
estingly,
the
apparent
activation
energies
of
methane
and
ethane
are
the
same
within
the
experimental
error.
The
ratio
between
the
pre-exponential
factors
of
ethane
to
methane
is
1.5,
which
is
in
accordance
with
the
ratio
of
C–H
bond
quantity.
Buyevskaya
et
al.
found
a
significant
difference
between
methane
with
70
kJ/mol
and
ethane
with
29
kJ/mol
for
samarium
oxide,
which
was
assigned
to
the
different
strength
of
the
C–H
bonds
involved
in
the
rate
Table
3
Kinetic
data
of
magnesia
in
the
TAP
reactor.
CH 4 C 2 H 6 C 2 H 4
E a ,app [kJ/mol]
113
±
2
115
±
6
82
±
3
k ∞ [mol/l/s]
2.27E
+
07
3.40E
+
07
1.18E
+
06
limiting
step
[16] .
A
similar
value
for
the
apparent
activation
energies
of
methane
and
ethane
could
be
caused
by
the
conditions,
i.e.
the
bond
cleavage
is
not
the
controlling
parameter.
From
the
kinetic
data
it
is
now
possible
to
estimate
the
max-
imum
yield
which
would
be
achievable
assuming
that
methane,
ethane
and
ethene
are
only
activated
by
surface
catalyzed
reaction
pathways
and
each
C–H
activation
successfully
leads
to
the
forma-
tion
of
the
desired
product.
We
used
a
simplified
reaction
network
of
consecutive
reactions
assuming
a
constant
oxygen
partial
pres-
sure
as
shown
below
and
simulated
it
by
application
of
a
batch
reactor
model
2CH 4
O 2
−→ C 2 H 6
O 2
−→ C 2 H 4
O 2
−→ 2CO x
Fig.
8
shows
the
maximum
yield
of
C 2 components
as
a
function
of
the
methane
conversion
at
750 ◦ C.
The
maximum
yield
of
C 2 is
around
60%,
at
a
methane
conversion
of
75%.
Additionally,
we
sim-
ulated
the
C 2 yield
at
atmospheric
conditions
and
constant
oxygen
partial
pressure
including
gas
phase
reactions
by
application
of
a
Fig.
7.
Arrhenius
plot
of
different
substrates
from
TAP
experiments
with
magnesia.
Fig.
8.
Simulated
yields
of
C 2 products
as
a
function
of
methane
conversion
for
mag-
nesia
calculated
using
the
kinetic
parameters
obtained
from
the
TAP
measurements
(solid
line),
the
micro
kinetic
model
of
Sun/Marin
[14]
(dotted
line),
in
comparison
to
the
reported
C 2 yield
from
literature
[26–29]
(cross
symbol).
B.
Beck
et
al.
/
Catalysis
Today
228
(2014)
212–218
217
Fig.
9.
Simulated
yields
of
C 2 products
as
a
function
of
methane
conversion
for
Mn/Na 2 WO 4 /SiO 2 calculated
using
the
kinetic
parameters
obtained
from
the
OCM
reaction
in
a
recirculating
batch
reactor
[11]
(solid
line)
in
comparison
to
singular
reported
C 2 yields
from
literature
[30–39]
(cross
symbol).
batch
reactor
model
as
predicted
by
the
model
of
Sun
and
Marin,
which
is
the
only
available
microkinetic
model
for
magnesia
based
catalysts
[14] .
The
maximum
C 2 yield
is
predicted
to
be
around
42%
at
a
methane
conversion
of
60%.
A
comparison
with
the
experimen-
tal
data
of
other
research
groups
with
this
model
leads
to
a
good
prediction
of
the
methane
conversion,
but
a
large
overestimation
of
the
C 2 selectivity
[25] .
Takanabe
et
al.
[11]
simulated
the
C 2 yield
dependent
of
the
methane
conversion
( Fig.
9 )
for
Mn/Na 2 WO 4 /SiO 2 in
a
recirculating
batch
reactor
at
atmospheric
pressure
including
direct
oxidation
pathways
form
methane
and
ethane
to
carbon
oxides.
The
maxi-
mum
C 2 yield
is
located
around
24%
at
a
methane
conversion
of
55%.
Comparison
of
the
simulated
C 2 yield
with
those
available
in
litera-
ture
using
various
reactor
setups
and
reaction
conditions
exhibits
a
conspicuous
consistency.
It
is
reasonable
that
suppression
of
uns-
elective
gas
phase
activation
of
methane,
ethane
and
ethene
will
lead
to
notable
higher
C 2 yields
as
already
shown
for
magnesia.
Independent
of
the
type
of
reactors,
different
kinetic
models
used
(TAP
or
normal
pressurized
reactor),
and
the
type
of
cata-
lysts,
a
methane
conversion
between
55%
and
75%
seems
to
be
optimal
to
maximize
the
C 2 yield.
Due
to
the
explosion
limit
of
methane/oxygen
mixtures
and
the
suppression
of
undesired
gas
phase
reactions,
the
oxygen
partial
pressure
must
be
kept
at
a
low
level.
Therefore,
a
distributed
oxygen
feed
along
the
reactor
would
be
a
good
alternative
which
could
be
implemented
by
means
of
a
membrane
or
stage-type
reactor.
Additionally,
the
formation
of
hot
spots
by
the
strongly
exothermic
total
oxidation
reaction
has
to
be
prevented.
From
this
qualitative
insight
into
the
reaction
mechanism
and
the
thermodynamical
data
some
indication
for
the
reaction
engi-
neering
and
reactor
design
can
be
drawn.
The
complexity
of
the
reaction
mechanism
prevents
accurate
simulations
of
the
derived
reactor
concepts.
Therefore
these
have
to
be
tested
and
investigated
in
detail
before
further
upscaling.
3.4.
Thermodynamics
and
reaction
engineering
aspects
Industrial
application
of
heterogeneous
catalyzed
gas
phase
reactions
are
often
realized
in
tubular
reactors.
In
the
case
of
exothermic
reactions
the
heat
production
can
lead
to
hot
spot
formations
and
in
the
worst
case
to
a
thermal
runaway
of
the
reactor.
The
critical
tube
diameter
is
a
safety-related
parameter
giving
the
upper
limit
to
assure
the
necessary
heat
transfer.
The
thermodynamic
data
from
the
NIST
database
[40]
are
used
to
calculate
the
reaction
parameters.
The
enthalpy
of
reaction
for
the
OCM
in
the
temperature
range
of
500
to
1000 ◦ C
is
− 176
kJ/mol
with
an
equilibrium
constant
of
1.01E
+
06.
Therefore,
OCM
is
highly
exothermic
and
not
thermodynamically
limited
in
practice.
The
heat
capacity
of
the
gas
mixture
was
calculated
with
Eq.
(1) .
C p =
x 0 , CH 4 C p, CH 4 +
x 0 , O 2 C p, O 2 (1)
For
a
mixture
of
5%
oxygen
and
95%
methane
the
heat
capac-
ity
is
71.81
J/(mol
K).
The
adiabatic
temperature
increase
is
507
K
under
the
assumption
of
a
selectivity
of
70%
ethene
and
30%
to
total
combustion
products
and
is
calculated
with
Eq.
(2) .
T ad = − H R c A, 0
v A ¯
c p
(2)
To
estimate
the
critical
tube
diameter
of
the
fixed
bed
reactor
Eqs.
(3)
and
(4)
from
[41,42]
are
used.
d crit = 8
Ge
L
R
T 2
w
Da
E a T Ad
(3)
Ge
= 0 . 125 d p /d
2
− 1
−
2 d p /d 2 (4)
In
Eq.
(3)
the
Damköhler
number
( Da )
was
added
under
the
assumption
of
a
first
order
reaction,
caused
by
a
quasi
stationary
methane
concentration.
The
geometry
factor
( Ge )
describes
the
vol-
ume
properties
of
particles
in
a
tube.
Typical
dimensions
of
tubular
reactors
are
given
by
the
length/diameter
ratio
of
100
and
a
particle
size
ratio
to
tube
diameter
of
0.1.
With
a
wall
temperature
of
700 ◦ C
the
critical
tube
diameter
is
8
mm.
An
increase
of
the
oxygen
mole
fraction
to
0.15
leads
to
a
reduc-
tion
of
the
critical
tube
diameter
to
2
mm.
Taking
into
account
that
upscaling
leads
to
similar
problems,
the
fixed
bed
reactor
is
unsuitable
for
OCM
reactions
at
industrial
conditions.
The
reason
is
the
fast
heat
production
and
the
low
thermal
conductivity
of
the
catalyst
bed.
The
absolute
value
of
the
critical
tube
diameter
is
dom-
inated
by
the
product
distribution.
Side
reactions
resulting
in
the
formation
of
carbon
monoxide
and
hydrogen
will
shift
it
to
slightly
higher
values.
Consecutive
reactions
resulting
in
the
formation
of
carbon
dioxide
and
water
will
shift
it
to
lower
values.
4.
Conclusion
The
experiments
at
elevated
pressures
point
out
that
the
yield
of
C 2+ components
increases
with
increasing
pressure
at
the
same
residence
time.
Without
a
catalyst
the
yield
is
passing
through
a
maximum,
which
is
caused
by
the
consecutive
oxidation
of
C 2+ .
This
occurs
in
homogenous
gas
phase
reaction
as
well
as
heteroge-
neously
catalyzed
surface
reaction
at
the
catalyst
itself,
the
reactor
wall,
and
quartz
packing.
It
is
evident
that
the
unselective
gas
phase
reactions
are
more
accelerated
by
higher
pressures
than
surface
catalyzed
reactions.
Thus
it
is
reasonable
that
each
specific
reactor
setup
will
pass
through
a
maximum
yield
of
C 2+ components
with
increasing
pressure.
The
methane
conversions
to
maximize
the
C 2
yield
are
mostly
in
the
60–75%
range.
Sufficient
oxygen
has
to
be
provided
at
a
constant,
but
relatively
low
oxygen
partial
pressure.
Kinetic
models
for
OCM
are
developed
for
plug
flow
reactors
with
a
relatively
high
oxygen
feed
fraction.
Therefore
for
reactor
simula-
tions
it
is
necessary
to
extrapolate
these
models
to
low
fractions
of
oxygen
leading
to
inaccuracies
in
the
prediction
of
conversion
and
selectivity.
Nevertheless,
it
is
not
possible
to
prevent
the
formation
of
hotspots
inside
the
catalyst
bed,
due
to
the
strongly
exothermic
total
combustion.
Ja ˇ
so
et
al.
[30]
showed
that
the
use
of
a
fluidized
bed
reactor
can
overcome
this
flaw
by
providing
isothermal
condi-
tions
even
at
high
oxygen
to
methane
ratios.
On
the
other
hand,
the
broad
residence
time
distribution
of
the
gas
will
favor
consecutive
reactions,
which
are
lowering
the
C 2 yield.
Additionally,
the
phys-
ical
strain
could
destroy
the
catalyst
material
during
fluidization,
but
this
was
not
observed
in
the
case
of
Mn/Na 2 WO 4 /SiO 2 .
Another
218
B.
Beck
et
al.
/
Catalysis
Today
228
(2014)
212–218
potential
approach
was
done
by
Carr
et
al.
by
using
a
simulated
countercurrent
moving-bed
chromatographic
reactor
to
prevent
total
oxidation
products
by
continuous
separation
of
the
reaction
products
[43] .
C 2 yields
of
more
than
50%
were
obtained
[44,45] .
Concerning
the
different
strength
of
the
C–H
bonds
between
methane
and
ethane
one
could
expect
a
lower
activation
energy
for
ethane,
which
is
obviously
negative
for
the
C 2 yield.
Interest-
ingly
magnesia
does
not
show
this
behavior
indicating
a
subsequent
higher
activation
barrier
to
overcome,
which
could
be
related
to
the
hydroxyl
formation
on
the
catalyst
surface
due
to
the
C–H
bond
activation.
Even
though
Mn/Na 2 WO 4 /SiO 2 was
not
suited
for
the
determination
of
apparent
activation
energies
at
vacuum
conditions
due
to
vacuum-induced
surface
reduction,
it
was
pos-
sible
to
find
some
significant
differences
between
this
catalyst
and
magnesia,
explaining
the
higher
C 2+ selectivity
observed
for
Mn/Na 2 WO 4 /SiO 2 .
Oxidation
to
carbon
oxides
proceeds
over
mag-
nesia
and
Mn/Na 2 WO 4 /SiO 2 mainly
by
reaction
of
weakly
adsorbed
oxygen.
The
formation
of
methyl
radicals
proceeds
by
a
second
more
strongly
adsorbed
oxygen
species.
Mn/Na 2 WO 4 /SiO 2 pro-
vides
a
much
higher
ratio
of
strongly
adsorbed
oxygen
to
weakly
adsorbed
oxygen,
which
highly
improves
the
selectivity
to
C 2 .
Also,
very
fast
oxygen
exchange
between
adsorbed
and
lattice
oxygen
was
observed
for
both
materials,
but
at
present
it
was
not
possible
to
assign
the
methyl
radical
formation
to
a
specific
type
of
oxygen
such
as
lattice
or
strongly
adsorbed
oxygen
by
studies
using
labeled
oxygen.
There
are
still
open
questions
about
the
nature
of
the
oxy-
gen
species
responsible
for
the
methyl
radical
formation
and
the
rate
and
selectivity
determination
step.
Nevertheless
qualitative
suggestions
for
the
reactor
design
and
operation
can
be
drawn.
Acknowledgment
We
thank
the
Deutsche
Forschungsgemeinschaft
for
financial
support
of
our
work
in
the
cluster
of
excellence
UniCat
Berlin.
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Paper 3
Investigation of the surface reaction network of the oxidative coupling
of methane over Na
2
WO
4
/Mn/SiO
2
catalyst by temperature programmed
and dynamic experiments
Vinzenz Fleischer, Rolf Steuer, Samira Parishan, Reinhard Schomäcker
⇑
Technische Universität Berlin, Institut für Chemie, Straße des 17. Juni 124, 10623 Berlin, Germany
article info
Article history:
Received 1 December 2015
Revised 26 May 2016
Accepted 20 June 2016
Keywords:
Oxidative coupling of methane
Na
2
WO
4
/Mn/SiO
2
Temperature programmed experiments
Dynamic experiments
abstract
In this work a series of temperature programmed experiments were carried out on a Na
2
WO
4
/Mn/SiO
2
catalyst. In TPR experiments we tested the reducibil ity of this catalyst and O
2
desorption behavior was
investigated by TPD. TPSR experiments in a flow of methane, ethane or ethene gave information about
the reaction network of OCM on the catalyst surface, without the presence of gas phase reactions, induced
by gas phase oxygen. We found indications of involvement of two different active oxygen species on the
catalyst surface. Furthermore an activation energy of 275 kJ/mole for selective methane activation was
determined. Dynamic experiments were performed to determine the amount of avai lable oxygen species
for the OCM reaction. Variation of methane partial pressure and flow rate showed a linear correlation
between methane partial pressure and surface oxygen conversion in dynamic experiments.
Ó 2016 Elsevier Inc. All rights reserved.
1. Introduction
A well-established process for ethylene production is the steam
cracking process, which cracks naphtha to olefins and other hydro-
carbons. Shortage of crude oil reserves has attracted attention
toward alternative processes, which use more available feedstocks.
The high availability of methane in natural gas makes it a suitable
feedstock alternative for short-chain olefins [1,2] . The oxidative
coupling of methane (OCM) is a promising reaction for ethylene
production. One of the most stable catalysts described for OCM
in the literature is Na
2
WO
4
/Mn/SiO
2,
which has good performance
and stability as shown in several publications [3–5] .
One of the major challenges in experimental studies of OCM is
the parallel reaction network of gas phase and surface reactions,
which have a strong influence on each other. For gas phase reac-
tions, Dooley and coworkers published an extended micro kinetic
model [6] . This micro kinetic model considers a network of 1582
reactions and 269 species, which are mainly radical reactions.
Unfortunately this complex reaction network allows only simula-
tion of ideal reactors, because of the large set of reactions.
However, the gas phase network is well described by the Dooley
model, which was shown by Schwarz and Coworkers [7] . Formal
kinetic models of OCM were published by several groups for
different catalysts [8–16] . All of these proposed mechanisms show
similar pathways for reactants and products. All groups who carried
out these experiments with extensive experimental efforts, were
using similar reactors operated in the classical steady state mode.
The surface reaction network, which is not fully explored, is
strongly influenced by gas phase reactions, and its kinetic param-
eters offer several constellations to fit experimental results. Such
complexity is caused by the mentioned involvement of various
radical species which are formed via gas phase reactions or during
reactions on the catalyst surface [17–20] . In addition for several
OCM catalysts different oxygen intermediates contribute to selec-
tive and unselective activation of methane on the catalyst surface.
The presence of these different species is also influenced by gas
phase oxygen [21–23] .
A micro kinetic surface reaction model for the surface reactions
was published by Lee and Coworkers for the Na
2
WO
4
/Mn/SiO
2
cat-
alyst [24] . Another one is the model published by Sun and Thybaut
for a MgO catalyst [25] . Both models were developed for different
catalysts but have similar elementary surface reaction steps which
are shown in Table 1 . Their models assume dissociative adsorption
of oxygen on a free site of the catalyst (
⁄
), which is in equilibrium
with gas phase oxygen. Methane activation occurs by a surface
reaction with dissociated oxygen (O
⁄
) and releases methyl radicals
to the gas phase. The coupling reaction of the radicals happens in
gas phase close to the catalyst surface where excess heat is
released, which is not shown in Table 1 . The same activation
http://dx.doi.org/10.1016/j.jcat.2016.06.014
0021-9517/ Ó 2016 Elsevier Inc. All rights reserved.
⇑
Corresponding author.
E-mail addresses: [email protected] (V. Fleischer), schomaecker
@mailbox.tu-berlin.de (R. Schomäcker).
Journal of Catalysis 341 (2016) 91–103
Contents lists available at ScienceDirect
Journal of Catalysis
journal homepage: www.else vier.com/locate/jcat
may also happen to the formed ethane and even ethene, which
leads to the formation of C
2
H
5
or C
2
H
3
radicals. An interesting fact
is that no ethene combustion on the catalyst surface is considered
in both literature models. A typical formation route for deep oxida-
tion products is the reaction of methyl radicals and surface bound
oxygen which forms a HCO
⁄
radical which is bound on the catalyst
surface, which is further oxidized to CO
⁄
and surface bound OH
groups (OH
⁄
) in a parallel route to the coupling reaction. The origin
of CO
2
on the catalyst surface (CO
2
⁄
) is the oxidation of CO
⁄
by O
⁄
.A
new aspect of the catalyst surface reaction network was published
by Beck at al. for the Na
2
WO
4
/Mn/SiO
2
as well as for MgO. In tem-
poral analysis of products (TAP) experiments, they could show the
presence of two different oxygen species, which exist in parallel on
the catalyst surface. Furthermore they could show that both spe-
cies have different reaction pathways [26] . One of them is weakly
bound (O
2,ads
) and opens the route to deep oxidation products. The
other, stronger bound oxygen species (O
x
⁄
) is responsible for the
selective methane activation. Following these results a plausible
surface reaction network is also presented in Table 1 .
In addition, there is also a discussion about the formal kinetic
mechanistic aspects of oxygen and methane activation described
by an Eley-Rideal, Mars-van-Krevelen or dual site Langmuir Hin-
shelwood mechanism [12,27,28] . On the one hand, in the
Eley-Rideal mechanism gas phase oxygen is required for selective
activate methane in the OCM process. On the other hand, in a
Mars-van-Krevelen type mechanism lattice oxygen is involved
during C A H bond cleavage of methane to form a methyl radical.
A possible interaction of lattice oxygen with methane offers the
opportunity to avoid the presence of gas phase oxygen in temper-
ature programmed experiments and dynamic experiments. During
these unsteady state experiments the catalyst oxidation and the
methane coupling reaction can be separated into two different
steps. One of the first overviews about these techniques and exper-
imental results from several groups was published by Falconer and
Schwarz [29] . Another review about these techniques was pre-
sented by Niemantsverdriet [30] .
One important factor in temperature programmed reduction
experiments is the bond dissociation energy of the reactants. For
methane (439 kJ/mole) and hydrogen (436 kJ/mole) these energies
are similar, and for ethane (423 kJ/mole) it is lower than those of
hydrogen and ethene (464 kJ/mole), which has the highest bond
dissociation energy in this group of compounds [31] . Therefore it
is possible to reduce the catalyst in the same manner as in a
H
2
-temperature programmed reduction (TPR) experiments using
OCM reactants to study their pathways in the reaction network
in absence of gas phase oxygen. These types of experiments were
introduced as temperature programm ed surface reactions (TPSR)
by McCarty and Wise some decades ago [32] . Heating rate varia-
tions in TPSR experiments allow the determination of the activa-
tion energy of desorption or surface reaction steps. This is known
as the Redhead method [33] . In this work we want to study the
OCM surface mediated reaction network in absence of gas phase
oxygen, to understand the role of surface bound oxygen on the
Na
2
WO
4
/Mn/SiO
2
catalyst material. Temperature programmed
reaction experiments utilizing methane, ethane and ethene as
reactant give qualitative insights to their interaction with the sur-
face bound oxygen species. In addition we performed a series of
kinetic TPSR studies for the selective activation of methane. The
focus of further dynamic experiments at constant temperature is
to quantify the converted amount of strongly surface bound oxy-
gen, the stability of the oxygen intermediates on the catalyst sur-
face and the nature of active sites for that material.
2. Experimental
2.1. Catalyst preparation
The details of the preparation and the characterizati on of the
catalyst are described elsewhere [34] . The final catalyst contains
5 wt.% Na
2
WO
4
, 2 wt.% Mn(II) ions and has a specific surface area
of 1.86 m
2
/g. The catalyst material was analyzed by nitrogen
adsorption and X-ray diffraction analysis after oxidation pretreat-
ment and after dynamic experiments. The results are presented
and discussed in the supporting information.
2.2. Experimental setup and mass spectrometer
All experiments were carried out in a fixed bed reactor made of
quartz. The catalyst was placed on a quartz frit (200
l
m pore size)
in the isothermal zone. The isothermal zone is above the frit and
has a length of 5 cm. A scheme of the reactor, a construction plan
of the setup and analysis of temperature profile aspects are shown
in the supporting information. The inner diameter is 9 mm. The
type K thermocoupl e (NiCrNi) is covered by a quartz-m ade capil-
lary ( d
in
= 4 mm), which seals the reactor on top. The reactants
come through the upper inlet. The bottom part of the reactor
shrinks in diameter and is connected to a mass spectrometer or
thermal conductivity detector. The feed composition was con-
trolled by mass flow controllers (MFC) and switching valves were
installed, enabling the ability to interrupt reactant flow immedi-
ately. Detection was carried out with a quadrupole mass spectrom-
Table 1
Comparison of micro-kinetic surfac e reaction models from Lee et al. and Sun et al. ( x , y , z = stoichiometric factors) [24–26] .
Lee et al. [24] Sun et al. [25] Beck et al. [26]
O
2
þ 2
¢ 2O
O
2
þ 2
¢ 2O
O
2
þ ¢ O
2 ; ads
O
2 ; ads
¢ yO
x
CH
4
þ O
¢ CH
3
þ OH
CH
4
þ O
¢ CH
3
þ OH
CH
4
þ O
¢ CH
3
þ OH
C
2
H
4
þ O
C
2
H
3
þ OH
C
2
H
4
þ O
¢ C
2
H
3
þ OH
C
2
H
4
þ O
¢ C
2
H
3
þ OH
C
2
H
6
þ O
¢ C
2
H
5
þ OH
C
2
H
6
þ O
¢ C
2
H
5
þ OH
C
2
H
6
þ O
¢ C
2
H
5
þ OH
2OH
¢ H
2
O
þ O
2OH
¢ H
2
O
þ O
CH
3
þ 3O
¢ HCO
þ 2OH
CH
3
þ O
¢ CH
3
O
CH
3
O
þ O
¢ CH
2
O
þ OH
CH
2
O
þ O
¢ HCO
þ OH
HCO
þ O
¢ CO
þ OH
HCO
þ O
¢ CO
þ OH
CO
þ O
¢ CO
2
þ CO
þ O
¢ CO
2
þ
CO þ
¢ CO
CO
2
þ ¢ CO
2
2OH
¢ H
2
O þ O
þ 4HO
2
! 3O
2
þ 2H
2
xO
2 ; ads
þ CH
4
¢ CO
y
þ zH
2
O
xO
2 ; ads
þ C
2
H
6
¢ 2CO
y
þ zH
2
O
xO
2 ; ads
þ C
2
H
4
¢ 2CO
y
þ zH
2
O
92 V. Fleischer et al. / Journal of Catalysis 341 (2016) 91–103
eter (QMS, IPI GAM 200) with channeltron detector. The QMS was
equipped with yttriated filaments. For O
2
detection in temperature
programmed desorption experiments m / e = 32 was chosen and
data points were recorded each 5 s. The m / e values for OCM com-
pound calibration are listed in Table 2 . Each compound was cali-
brated by utilization of a calibration gas bottle filled with
5 ± 0.005 vol.% of one compound and the rest was filled with
Helium. The calibration of m / e = 28 for ethane and ethene was nec-
essary to prevent false-signals for CO or other compounds. As indi-
cator for ethane the m / e = 30 and for ethylene m / e = 27 were
chosen, while the other listed masses were calibrated by their
specific, relative intensities for these molecules. Therefore the feed
compositions of these compounds were determined by three cali-
bration factors for ethane and ethene to enhance detection
accuracy.
The final mole fraction ( x
i
) of each compound was calculated
according to Eq. (1) :
x
compound
ð x ... z Þ¼ S
m
e
ð i ... j Þ
c
i
ð 1 Þ
S
m
e
ð i Þ – relative intensity matrix of compound i ,
c
ð i Þ –calibration
factor matrix for compound i .
The relative intensity was derived by
S
m
e
; i
.
.
.
S
m
e
; j
0
B
B
@ 1
C
C
A
¼
S
compound
X
S
m
e
; i
S
compound
Z
S
m
e
; j
.
.
. .
.
. .
.
.
S
compound
X
S
m
e
; i
S
compound
Z
S
m
e
; j
0
B
B
B
B
@
1
C
C
C
C
A
x
compound
x
.
.
.
x
compound
z
0
B
B
@ 1
C
C
A
ð 2 Þ
S
m / e , i
– Sensitivity for a specific m / e ratio;
S
compound
X
S
m
e
; i
– Signal of compound x on m / e , i ;
x
compound
x
– molefraction of compound x .
2.3. Temperature programmed reduction (TPR)
The fresh catalyst (60 mg) was placed in the reactor, which was
heated up under synthetic air (N
2
:O
2
– 4:1, 99.9%, 30 sccm/min)
with a temperature-ramp of 10 K/min to 1023 K, keeping this tem-
perature for an hour and cooling down over night under constant
air flow. To remove gas phase and adsorbed oxygen, a constant
flow of nitrogen (30 sccm/min) purged through the reactor at room
temperature for 30 min. The catalyst reduction was carried out by
a feed stream of N
2
:H
2
(9:1, 60 sccm/min) with a heating rate of
2 K/min, heating up until 1093 K. The analysis of the effluent gas
was done with a thermal conductivity detector (Messkonzept,
FTD 200) monitoring the stream composition each 0.5 s.
2.4. Temperature programmed desorption (O
2
-TPD)
For temperature programmed desorption of oxygen, 1 g of fresh
catalyst was used. The catalyst was heated up in a flow of He
(99.999%, 20 sccm/min), with a temperature ramp of 5 K/min to
1093 K. In a second experiment 1 g of fresh catalyst was oxidized
according to the method as mentioned in the TPR experiment.
After the oxidation, the catalyst was treated as in the first O
2
-
TPD experiment. The detection was carried by mass spectrometer,
which was described before.
2.5. Temperature programmed surface reaction (TPSR)
Each TPSR experiment was carried out with 1 g Na
2
WO
4
/Mn/
SiO
2
fresh catalyst. The catalyst oxidation in all experiments was
the same procedure as described in the TPR experiment. To remove
gas phase and adsorbed oxygen, the catalyst was purged for 30 min
by a constant flow of He (99.999%, 30 sccm/min) at room temper-
ature. For methane-TPSR, a flow of 30 sccm/min methane (99.99%)
was chosen. Ethane and ethylene-TPSR were carried out in a flow
of He:C
2
H
X
(95:5, 30 sccm/min, 99.98% C
2
H
X
). All TPSR experi-
ments had an initial temperature of 298 K and the catalyst material
was heated up under the reactant feed with a temperature ramp of
1 K/min with a final temperature of 1073 K. Each TPSR experiment
was repeated by variation of the temperature with 3, 4 (only for
methane) and 5 K/min. For product detection a QMS was used, as
described above.
2.6. Dynamic experiments
The reactor, filled with 1 g fresh catalyst, was heated up to
1023 K under a constant flow of He:O
2
(9:1, 30 sccm/min). The
final temperature was kept constant during the whole experiment.
After 10 min at constant reaction temperature, the oxygen flow
was stopped by deactivation of the mass flow controller (MFC)
and closing the switching valve. A flow of Helium (20 sccm/min)
was used as purging step to remove gas phase and weakly
adsorbed oxygen for 5 min at constant temperature. Helium feed
was immediately replaced by CH
4
(20 sccm/min) for the reaction
mode, keeping the temperature constant. That was done by closing
and opening of the switching valves, simultaneously. The methane
flow reached the reactor after pressure fluctuation had already
been compensated for in the feed tube. In the first series of
dynamic experiments the purge time was varied from 10,
180 min to 300 min without using new catalyst material for each
run. The methane step time was 10 min. After these experiments
the catalyst material was not reoxidized and cooled down in a
stream of He (30 sccm/min). The catalyst materials were analyzed
by nitrogen adsorption (BET) and XRD. This was followed by a ser-
ies of experiments reducing the purge time interval from 5 min
down to 15 s, using fresh catalyst material again. During these
set of experiments, the methane flow rate was 30 sccm/min and
the methane time step was 10 min.
In addition, we also varied the methane partial pressure during
the reaction sequence in another experimental series with fresh
catalyst material. That was realized by a lower flow rate of purge
gas down to 10 sccm/min or 15 sccm/min, while methane flow rate
was set in parallel to 20 sccm/min or 15 sccm/min. In total a feed of
30 sccm/min reached the reactor. Here the reaction step for
methane was set to 5 min.
2.7. Steady state experiment
The reactor was filled with 1 g fresh catalyst and heated up to
1023 K under He:O
2
(8:2, 30 sccm/min). The feed was changed to
CH
4
:O
2
(95:5, 180 sccm/min) for 1 h, reaching the steady state con-
ditions. The analyses of products were carried out with a QMS.
Table 2
Calibrated masses for compound detection in dynamic experiments.
Molecule m / e
CO
2
44
CO 28
He 4
H
2
2
H
2
O 18; 17
CH
4
16; 15
C
2
H
6
30; 29; 28
C
2
H
4
28;27;26
O
2
32
V. Fleischer et al. / Journal of Catalysis 341 (2016) 91–103 93
3. Results and discussion
3.1. TPR
The result of the H
2
-TPR experiment is presented in Fig. 1 . The
oxidized catalyst showed one reduction pattern at 950 K. Further-
more the reduction curve shows early reduction by a slight
decrease of the H
2
signal with respect to the baseline. This is also
observed in the work of Shahri et al. [35] . They proposed that the
early reduction is related to manganese oxide.
Peak integration was carried out and the amount of converted
hydrogen was calculated according to Eq. (3) . For the integration
an adapted baseline was set to the corresponding hydrogen signals
before and after the reduction pattern. The hydrogen consumption
is equivalent to 17 O-atoms/nm
2
, which is an astonishing high
value. Typically expected monolayer coverages of transition metal
oxides on various support materials are between 0.1 and 10 atoms/
nm
2
[36] . It was found that various manganese oxide species and
supported manganese oxides on silica can be reduced to MnO
under hydrogen atmosphere [37,38] . Furthermore it was shown
in EPR experiments that only manganese ions are reduced at
1073 K in the Na
2
WO
4
/Mn/SiO
2
catalyst, while tungsten ions had
a constant valency of +VI [39] . In addition, the active phase of this
material has a very flexible structure [40] . Therefore it is possible
that also sublayers are involved in the reduction reaction during
TPR experiments, considering the small specific surface area of that
catalyst.
n
H
2
¼
A
H
2
; peak
½ % s _
V
m
3
s
hi
p ½ Pa
R
J
mol K
T ½ K ð 3 Þ
3.2. O
2
-TPD
Experimental results of O
2
-TPD experiments are depicted in
Fig. 2 . The ‘‘as calcined” catalyst, without further oxidation treat-
ment, shows an O
2
peak at 550 K. For the oxidized catalyst such
peak is missing. Radhakrishnan et al. reported such early desorp-
tion patterns for different supported Mn
x
O
y
catalysts [41] . Further-
more this phenomenon is also observed for bulk manganese oxides
[42] . This may explain also the early reduction of the catalyst,
observed in TPR experiments. However, in both cases stronger O
2
desorption starts at 1000 K. TPD experiments for that catalyst
material have already been published by different groups. In the
review article by Arndt et al. it is mentioned, that O
2
-TPD experi-
ments published by Fang et al. show no oxygen loss of the catalyst
[43,44] . This is in clear contradictio n to the report by Liu et al. [45] .
They observed two oxygen signals at 1070 K and 1150 K, which
were interpreted as surface layer oxygen and bulk oxygen. Espe-
cially the desorption peak at high temperatures in our experiments
is in good agreement with their results.
3.3. TPSR
The results of the TPSR experiments are shown in Fig. 3 . Differ-
ent symbols represent different compounds. Symbols do not repre-
sent data points, which are too many, to show due to the fast
detection rate. They represent an interpolation of the recorded
data. The results of methane feed TPSR experiments are shown in
Fig. 3 A–C, with increasing temperature ramps.
Only CO and ethane formation is observed in this set of exper-
iments. In each case first ethane is formed and followed by detec-
tion of CO. Both peaks shift at higher heating rates to higher
temperatures. The ethane formation shifts slightly and forms no
clear peaks at heating rates above 4 K/min. Such spectra can be
interpreted in different ways. The most important step for the for-
mation of ethane is the selective activation of methane by C–H
bond cleavage. These radicals combine and form an ethane mole-
cule. Na
2
WO
4
/Mn/SiO
2
catalyst material was verified by Jiang
et al. [46] According to the results of Lee et al. these radicals are
also converted by surface bound oxygen into deep oxidation prod-
ucts [47] . On the other hand methane or ethane can be activated by
another surface side, which leads also to the formation of deep oxi-
dation products. We assume a short lifetime for the methyl radicals
and therefore an immediate coupling or adsorption process on the
catalyst surface. As mentioned before, in our TPSR experiments CO
is formed much later and therefore the coupling process seems to
be preferred. Furthermore the CO formation has to be observed
parallel to ethane formation, when the oxidation of ethane is the
origin of CO formation. Thus, the unselective activation of methane
by another oxygen species on the catalyst surface is responsible for
the formation of CO.
To analyze that situation in more detail, a series of simulations
with the discussed reaction mechanisms in TPSR experiments were
carried out and are presented in the supporting information. Our
simulation results confirm that two active oxygen species on the
catalyst surface in absence of gas phase oxygen contribute to
OCM activity. Therefore not only gas phase and oxygen adsorption
intermediates contribute to unselective methane activation. The
nature of these species could be electrophilic and nucleophilic.
On the one hand electrophilic oxygen is responsible for selective
methane activation, which forms methyl radicals. On the other
hand the nucleophilic species contributes to deep oxidation
Fig. 1. H
2
TPR of oxidized Na
2
WO
4
/Mn/SiO
2
catalysts. 60 mg catalyst, N
2
:H
2
(9:1,
60 sccm/min), 2 K/min.
Fig. 2. O
2
TPD of different pretreated Na
2
WO
4
/Mn/SiO
2
catalysts. 1 g catalyst, He
(99.999%, 20 sccm/min), 2 K/min.
94 V. Fleischer et al. / Journal of Catalysis 341 (2016) 91–103
product formation. In the literature similar TPSR experiments were
carried out with Sr
x
Ti
y
MgO, confirming our results [48] . The peak
shift at different heating rates in TPSR experiments correlates with
the slowest reaction in the reaction network. For OCM, this is
assumed to be the methane activation reaction [49,50] . That effect
is used in the next section to calculate the activation energy of
selective methane activation.
It has to be noted that the variations in the peak areas in these
experiments are caused in two different ways. On the one hand
slow heating rates result in a longer reaction time for oxygen con-
sumption for each temperature during the experiment. Therefore
low amounts of methane can be converted already at lower tem-
peratures. Thus, the peak amplitude is different for slow heating
rates, because less oxygen is available on the catalyst surface at
that time step. In addition, fast heating rates shift the methane
activation by oxygen to higher temperatures and therefore the
reaction rate is much higher, because of higher temperature and
more oxygen which was not converted so far. On the other hand,
as shown in the last section, oxygen can desorb during the exper-
iment. Therefore it is also possible that oxygen is released in gas
phase, while methane adsorbs on the catalyst surface.
In the same manner as methane TPSR experiments, a series of
TPSR experiments for ethane are shown in Fig. 3 D–F. We observe
ethane conversion to ethene starting at 890 K in all experiments.
At 1050 K most of the ethane in the feed has been converted to
ethene but conversion does not change further. The initial temper-
ature of ethane consumption is higher than for methane activation
in TPSR experiments. According to the different C–H bond
strengths of methane and ethane, a lower activation temperature
would be expected. However our TPSR experiments indicate a dif-
ferent activation mechanism for this molecule, which is not
effected by C–H bond cleavage on the catalyst surface, which will
Fig. 3. TPSR product formation peaks of methane, ethane and ethene TPSR experiments at heating rates from 1 to 5 K/min, 1 g Na
2
WO
4
/Mn/SiO
2
, A–C: methane TPSR, D–F:
ethane TPSR (He:C
2
H
6
, 95:5), G–I: ethene TPSR (He: C
2
H
4
, 95:5), 1 g Na
2
WO
4
/Mn/SiO
2
, 30 sccm/min.
V. Fleischer et al. / Journal of Catalysis 341 (2016) 91–103 95
be discussed later. Furthermore, the formation of CO
2
and CH
4
can
be observed at 950 K. These products indicate deep oxidation reac-
tion steps as well as cracking reactions. Some carbon deposition as
consecutive reaction product is also observed during reactor
purging.
In TPSR experiments utilizing ethene as reactant only CO
2
and
CH
4
formation at 920 K can be observed. That is in good agreement
with our results of ethane TPSR experiments. CH
4
is a typical crack-
ing product, which is formed from ethene during coke formation
on the catalyst surface [51,52] . Carbon deposition was verified by
oxidation of the catalyst subsequent to the TPSR experiment. Inter-
estingly the formation of CO
2
and CH
4
starts in these experimental
sets at the same temperature, which indicates a fast consecutive
reaction from ethene to carbon deposition followed by CO
2
forma-
tion. Another possible explanation is a parallel reaction of ethene
to cracking products and to carbon dioxide. The formation of deep
oxidation and cracking products from ethene in absence of gas
phase oxygen and the adsorption intermediate confirms our
hypothesis that two active oxygen species seem to be responsible
for ethene activation. On the one hand electrophilic oxygen can
interact with the double bond of ethene, which could form ethy-
lene oxygenates. These intermediates are further oxidized to deep
oxidation products. On the other hand nucleophilic oxygen can
activate ethene by C–H bond cleavage, which results into the for-
mation of deep oxidation and cracking products. Such effect is
well-known for ethylene oxide synthesis and for OCM on silver
catalysts [53,54] .
In Fig. 4 A the product formation from ethane TPSR is presented
in more detail for further discussion, than in Fig. 3 . At 890 K small
amounts of ethene can be observed, while water and hydrogen for-
mation starts at the same temperature. At 950 K the amount of
water decreases, while the amount of hydrogen still increases. As
mentioned before, with increase of ethane conversion, CO
2
and
methane signals can be observed. For monitoring the water forma-
tion in a consecutive reaction, the detailed product signals for
ethene TPSR are shown in Fig. 4 B. We observe at 920 K small
amounts of water and CO
2
formation. At 1050 K the water amount
decreases to the baseline level, while the CO
2
signal stays constant
and methane formation can be observed. We observe also hydro-
gen formation at this temperature and much stronger ethene
decomposition. Both product spectra show that the oxygen rich
surface provides enough potential for oxidative dehydrogenation
reaction steps, indicated by water formation. When that surface
oxygen species is consumed, the reaction pathway changes to ther-
mal decomposition reactions. For ethane TPSR experiments it is not
clear whether thermal and oxidative dehydrogenation follows con-
secutive or parallel reaction routes.
To analyze this situation we calculated the free reaction energy
for different coking and oxidation reactions at 1073 K by Eq. (4) .
The thermodynamic data for each compound was presented by
McBride et al. [55] . The results are shown in Table 3 .
D
G
R
¼ X
i
m
i
D
H
f ; i
T X
i
m
i
D
S
f ; i
ð 4 Þ
The oxidative dehydrogenation (ODH) is preferred compared to
the thermal dehydrogenation (TDH) ( Table 3 , reaction A and B).
Our ethane TPSR experiments in an empty reactor show thermal
dehydrogenation in the same temperature range. According to that
result we conclude that this phenomenon is a parallel route in the
gas phase. The coupling process of two methyl radicals in gas phase
at that temperature has a free energy of 328 kJ/mole. Thus, the
energy must be transferred from the ethane intermediate for stabi-
lization. Typically, that could be done by a collision with another
molecule or the splitting of ethane into new molecules as ethene
and hydrogen [6] . Decomposition of ethane to coke and hydrogen
is not favored, but the backward reaction to methane by reduction
may happen in the gas phase ( Table 3 , reaction C and D). The split-
ting reaction of ethane to methane hydrogen is necessary. As main
hydrogen source the TDH of ethane may be responsible. Finally the
ODH reaction of ethane is thermodynamically preferred but lim-
ited to the amount of stored oxygen on the catalyst surface. For
TDH no adsorption of ethane on the catalyst surface is required.
The ethene decomposition to carbon and methane is slightly
preferred, compared to the route of carbon and hydrogen forma-
tion only ( Table 3 , reaction E and F). The formation route of carbon
dioxide cannot be resolved exactly, because the TPSR method is not
sensitive enough for this reaction and the thermodynamic data for
the Na
2
WO
4
/Mn/SiO
2
catalyst are unknown. Isotopic labeling
Fig. 4. Detailed ethane and ethene TPSR spectra, A: Ethane, B: Ethene, 3 K/min, 1 g Na
2
WO
4
/Mn/SiO
2
, 30 sccm/min flow (He:C
2
H
X
, 95:5).
Table 3
Calculated free energy for oxidation and coking reactions from ethane and ethene at
1073 K.
No. Reaction
D
G (kJ/mole)
AC
2
H
6
+½ O
2
? C
2
H
4
+H
2
O 187.4
BC
2
H
6
? C
2
H
4
+H
2
6.9
CC
2
H
6
? 2C + 3H
2
32.0
DC
2
H
6
+H
2
? 2C H
4
79.0
EC
2
H
4
? 2C + 2H
2
107.4
FC
2
H
4
? C+C H
4
118.9
96 V. Fleischer et al. / Journal of Catalysis 341 (2016) 91–103
experiments with
13
C in cofeed for that catalyst material are
required to understand the origin of deep oxidation products.
To compare our findings, a well-established reaction network
for the OCM is shown in Fig. 5 A. It must be considered that the
reaction network combines oxygen induced gas phase reactions,
further gas phase activity by water and oxygen containing prod-
ucts, gas phase interactions of short lifetime intermediates as rad-
icals and surface reactions in each step. Methane undergoes a
parallel reaction to ethane or the deep oxidation product CO. The
first route leads to ethene which can be converted to CO
2
, which
is in equilibrium with CO [56] . There are some modifications of this
model, which were used for formal kinetic fitting procedures for
the Na
2
WO
4
/Mn/SiO
2
catalyst [8,9] .
The micro kinetic model from the introduction part is presented
in Fig. 5 B for further comparison. Interpretation of the TPSR results
leads to a similar reaction model which is presented in Fig. 5 C. The
methane activation on the catalyst surface is followed by coupling
process or by methyl radical absorption and the formation of oxy-
genates. In contrast, our TPSR results indicate that methane activa-
tion is performed by two different surface oxygen species. Similar
to the results of Sun et al. the ethane molecule is activated by oxy-
gen on the catalyst surface or in gas phase by different radical spe-
cies, which is common for the TDH reaction. The absence of gas
phase oxygen leads to hydrogen production instead of water for-
mation. Similar findings in these models show that the TDH is a
serious parallel reaction route in the oxidative coupling of
methane. Such information is typically not obtainable in co-
feeded studies [57] . Furthermore this route reveals the complexity
of different gas phase contributions to OCM, considering that the
TDH is also a radical chain mechanism, as the selective methane
activation on the catalyst surface is. Both reactions are initialized
by methyl radical formation [50,58] . The mechanism of ethane
activation on the catalyst surface by TPSR experiments cannot be
revealed. In contradiction to the literature model, which requires
gas phase oxygen for the formation of deep oxidation products
by ethene oxidation, the TPSR results for the Na
2
WO
4
/Mn/SiO
2
cat-
alyst reveal that also lattice oxygen is able to form such kind of
products. Therefore it is possible that a Mars-van-Krevelen type
mechanism should be considered for ethylene oxidation. In addi-
tion, it must be considered that also the cracking mechanism of
ethylene is based on radical reactions and coke formation is pre-
ferred at higher ethene partial pressures [59] .
However, the proposed model is similar to the key steps of the
reaction network for the MgO catalyst. It shows that the MgO cat-
alyst and the Na
2
WO
4
/Mn/SiO
2
catalyst have several similarities,
despite their chemical difference.
3.4. Arrhenius parameter of methane TPSR
In the last section we pointed out, that the rate-determining
step in TPSR experiments is strongly influenced by the heating rate.
Redhead showed that desorption or the rate-determining reaction
step follows an Arrhenius type function [33] . The rate constant of
the surface reaction is proportional to the ratio of heating rate
( b ) and peak amplitude temperature ( T
A
) of the formed product,
which is shown in Eq. (5) . Following this method an Arrhenius plot
can be constructed. Therefore we carried out methane TPSR exper-
iments with heating rate variation from 1 to 4 K/min. Higher heat-
ing rates show no clear ethane peak as can be seen in the last
section. For a more exact determination of the peak temperature
from the observed ethane peaks, each peak was fitted to an empir-
ical amplitude of the Gaussian function, which is shown in Eq. (6) .
All fits had an accuracy of 0.85 or better. The peak temperatures T
A
,
derived by this method are shown in Table 4 .
k b
T
2
A
! ð 5 Þ
x
i
¼ x
i
0
þ A exp 0 : 5 T T
A
w
2
! ð 6 Þ
Fitting parameters:
A – peak amplitude (ordinate).
T
A
– temperature of peak (abscissa).
x
i 0
–offset of compound baseline.
w - half width of peak signal.
The inverse temperatures of the ethane peaks were plotted to
the Napierian logarithm of heating rate divided by the square of
Fig. 5. A: Classical OCM reaction mechanism, including gas and surface reaction
steps [8,9] , B: OCM reaction network according to Table 1 , C: Proposed reaction
route for OCM surface reaction based on TPSR experiments for Na
2
WO
4
/Mn/SiO
2
catalyst.
Table 4
Ethane peak temperatures from methane TPSR experiments with heating rates from 1
to 4 K/min fitted by an empirical Gaussian function, 1 g Na
2
WO
4
/Mn/SiO
2
, 30 sccm/
min methane.
Heating rate b (K/min) Temperature of ethane peak amplitude T
A
(K)
1 964 ± 3
2 983 ± 1
3 985 ± 1
4 1003 ± 1
V. Fleischer et al. / Journal of Catalysis 341 (2016) 91–103 97
temperature of the peak amplitude, following the Redhead method
[33] . The result is shown in Fig. 6 .
Deriving the activation energy from this plot results in
275 ± 29 kJ/mole. The error of this astonishingly high value is
strongly influenced by the ethane peak amplitude temperatures
and the fitting results. Especially slow heating rates show a noisy
and weak ethane signal. Other influences are CO formation and
carbon deposition, which can block surface sites.
Comparing this activation energy to other apparent formal
kinetic model parameters shows a high discrepancy between these
values. The apparent activation energy for methane activation in
steady state experiments varies from 133 kJ/mole up to 212 kJ/
mole [35,60,61] . The different results show clearly that the deriva-
tion of a reliable result for this step is very challenging, considering
that different oxygen species with different lifetime contributes to
methane activation on the catalyst surface. However, Lunsford and
Coworkers found similar value of 270 kJ/mole [62] .
One reason for such a difference is the assumptions for the
derivation of the reaction rate laws. In formal kinetic models,
which were discussed in the last section, the activation of oxygen
undergoes direct dissociation to an adsorbed species (O
x
⁄
), which
is shown in Eq. (7) . In our work, we assume an involvement of
strongly bound oxygen which may also be lattice oxygen (O
x
⁄
). This
may happen by an adsorption process (O
2, ads
) and a consecutive
dissociation reaction, which is shown in Eqs. (8) and (9) .
O
2
þ 2 ¢ 2O
x
ð 7 Þ
O
2
þ ¢ O
2 ; ads
ð 8 Þ
O
2 ; ads
¢ x O
x
ð 9 Þ
To discuss this situation in more detail, a possible energy profile
of the OCM consecutive reaction network is constructed in Fig. 7 .
We assume a two step activation process for oxygen, starting by
adsorption (O
2, ads
) followed by dissociation (O
x
⁄
), which was postu-
lated by Beck et al. [26] . The difference between these two species is
their interaction with the surface. The adsorbed, electrophilic oxy-
gen species O
2, ads
is responsible for deep oxidation products and is
weakly bound on the catalyst surface. The dissociation of this spe-
cies to the nucleophilic O
x
⁄
involves the catalyst material, which
causes a strong interaction. Both oxygen species must have differ-
ent energy levels, because of their binding situation. The activation
of methane to deep oxidation products undergoes a transition state
with O
2, ads
( Fig. 7 ,E
A1
) in steady state experiments. Assuming that
the conversion of methane, ethane and ethene is initialized by
C–H bond cleavage at the same oxygen species results in similar
activation energies in the range of 120–180 kJ/mol. The much
higher activation energy obtained from the TPSR experiments
( Fig. 7 ,E
A3
) clearly indicates the presence of the strongly bound,
presumably dissociated oxygen species at a much lower energy
level. Since this species is already formed during the pretreatment
of the catalyst, a higher activation barrier has to be overcome.
Due to the course of the experiments such a more stable species
is required that survives the pretreatmen t of the catalyst. In addi-
tion this is supported by the findings of dynamic experiments,
which are discussed in the next section. Similar findings and con-
clusions were made by Sofranko et al. and Jones et al., who carried
out several dynamic experiments with catalyst materials including
alkali promoted Mn/SiO
2
catalysts of different compositions
[63–65] . They concluded that different surface oxides are involved
in the OCM process, which confirms our results. Furthermore they
have mentioned that non acidic oxide species are responsible for
unselective methane activation. Several methane activation mech-
anisms were reviewed and discussed by Sinev [66] . He showed that
a heterolytic proton or hydride ion abstraction of a C–H bond has
similar energy expenditures as the molar energy of the lattice for
stable oxides such as MgO or Al
2
O
3
.
3.5. Dynamic experiments at constant temperature
3.5.1. Comparison of dynamic and steady state experiments
The results of our steady state experiment are shown in Table 5 .
Comparing the selectivities of both experimental sets, steady state
and dynamic experiment ( Table 6 ) shows similar selectivities of C2
and deep oxidation products. It has to be noted that gas phase oxy-
gen contributes to the results of steady state experiments, but not
to the dynamic ones. The major difference in selectivities is the
ratio of ethane to ethene. In the steady state experiment, ethane
is the major product while in dynamic mode ethene dominates.
The low methane conversion and high C2 selectivity in the steady
state experiment indicate less gas phase activity, compare d to the
catalytic activity. Therefore we determined the initial reaction rate
of methane conversion from Eq. (10) :
r
CH
4
¼ _
n
CH
4
X
CH
4
m
cat
ð 10 Þ
Fig. 6. Arrhenius type diagram of methane TPSR experiments using the ethane peak
amplitudes, with heating rate of 1–4 K/min, 1 g Na
2
WO
4
/Mn/SiO
2
, 30 sccm/min
methane.
Fig. 7. Proposed energy profile for the OCM reaction at the Na
2
WO
4
/Mn/SiO
2
catalyst at 1073 K.
Table 5
Results of steady state experiment, CH
4
:O
2
95:5, 180 sccm/min, 1023 K, 1 g catalyst.
XC H
4
XO
2
SC
2
H
6
SC
2
H
4
SC O
2
Y
0.032 0.393 0.676 0.114 0.210 0.025
98 V. Fleischer et al. / Journal of Catalysis 341 (2016) 91–103
For the dynamic experiment, we assume that the activation of
methane is the rate determini ng step. All reaction steps should
have a reaction order of one, because of the activation by adsorp-
tion on the catalyst surface. We calculated the formation rates of
all products during the first 20 s, using Eq. (11) . It has to be noted
that the methane partial pressure is not constant and still increas-
ing in this time span. The experimental conditions and results for
the dynamic experiment, with a purge time of 10 min, are shown
in Table 6 . A detailed discussion for the dynamic experiment fol-
lows in the next section.
r
CH
4
¼ X
products ; i
n
CH
4
ð dt Þ
m
cat
m
i
dx
i
dt
i
ð 11 Þ
n
CH
4
– Moles of methane in 20 s.
x
i
– Mole fraction of detected compound.
m
i
– Stoichiometric factor.
In the steady state mode (242
l
mol/g
cat
min
1
), the rate is much
higher compared to the dynamic mode (5.19
l
mol/g
cat
min
1
). The
difference of almost two orders of magnitudes between both reac-
tion rates indicates several effects.
On the one hand, in steady state mode the methane partial pres-
sure is constant at 0.95 bar, while in the early period of the
dynamic experiment methane partial pressure reaches only
0.1 bar. On the other hand, in steady state the major product is
ethane, while in dynamic mode ethane and ethene have the same
ratio. Therefore the formed ethane, which is even more reactive,
competes in the dynamic experiment with the fed methane for
the oxygen and therefore has also a substantial effect on the mea-
sured rate of methane consumption. We assume, that the low flow
of 20 sccm/min in the dynamic experiment compared to 180 sccm/
min in the steady state experiment allows more consecutive oxida-
tive dehydrogenation of ethane. In Addition the catalyst in steady
state mode is continuously reoxidized by gas phase oxygen, which
also enhances the reaction rate. Therefore not only the strongly
bound oxygen intermediate is present on the catalyst surface.
The weakly bound intermediate contributes also to the reaction
rate. The reoxidation rate of the catalyst is fast compared to
methane activation. Thus, the catalyst activity is not limited by
the available amount of oxygen [47] .
In contradiction, in dynamic experiments only small amounts of
oxygen are stored on the catalyst surface, which are rapidly con-
sumed. Further effects may limit the reaction rate by mass trans-
port limitation. Possible is the formation of OH groups on
Na
2
WO
4
/Mn/SiO
2
, which can even block sites for methane adsorp-
tion. An additional limiting effect is the formation of coke, which
can also block or convert active sites. All these effects cannot be
fully excluded.
However, the large difference in methane partial pressures in
steady state and dynamic experiments and the big differences in
availability of oxygen in both experimental modes seem to have
the strongest effects on the reaction rate. Thus, we assume to
observe comparable initial activity of the active oxygen sites in
both types of experiments.
3.5.2. Influence of purge time and flow rate in dynamic experimen ts
In the first series we carried out a variation of the purge time
with Helium from 10 min to 5 h. The results for these experiments
are shown in Fig. 8 . At time 0 s, the methane flow reached the cat-
alyst bed after purging with inert gas for different time intervals. In
all experiments we obtained a C2 selectivity of 85% or better. The
amplitudes of the product peaks are much higher compared to
the TPSR experiments. In the TPSR experiments the time span
between oxidation and reaction is dependent on the heating rate.
Furthermore, in TPSR experiments the oxygen conversion starts
at lower temperatures and the amount of stored oxygen decreases
in a longer time span. Slow increase of temperature enhances this
effect. Reaching high temperatures, most of the oxygen will be con-
verted and the peak amplitude is low. In the TPD experiment we
could observe an oxygen loss at these temperatures, which may
also contribute.
During the first minute in all dynamic experiments we observe
deep oxidation products, indicated by CO and CO
2
formation. In
this time interval the partial pressure of methane is still increasing,
because it replaces the He from the purging process. At high
methane partial pressures and reaching the ethene peak ampli-
tude, the CO signal vanishes. The same phenomenon is observed
for the CO
2
signal, which has a longer time span than CO. The early
decrease of the CO
2
signal indicates that the available amount of
oxygen for OCM reaction decreases rapidly and only small amounts
are consumed for ethane and ethene formation. The preferred
route of deep oxidation product formation, from methane or
ethene, during dynamic experiments cannot be revealed. Isotopic
labeling experiments for the Na
2
WO
4
/Mn/SiO
2
catalyst reveal no
kinetic isotope effect (KIE) of CH
4
/CD
4
for deep oxidation product
formation [67] . It was concluded, that a different site on the cata-
lyst surface is responsible for CO
X
formation, which was suggested
to be MnO
X
.
To quantify our results an oxygen balance for each compound
was calculated. To do so, the peak areas for each compound ( A
i , peak
)
were integrated with respect to the baseline level. The amount of
each product was calculated as shown in Eq. (12) and the corre-
sponding moles of consumed oxygen with respect to the stoi-
chiometry were determined. We assume that surface bound
oxygen (O
⁄
) is an atomic species and quantitatively converted to
water. The corresponding reactions are shown in Eqs. (13)–(16) .
n
i
¼
A
i ; peak
½ % s _
V
m
3
s
hi
p ½ Pa
R
J
mol K
T ½ K ð 12 Þ
CH
4
þ 4O ! CO
2
þ 2H
2
O ð 13 Þ
2CH
4
þ O ! C
2
H
6
þ H
2
O ð 14 Þ
Table 6
Species based oxygen balance for the dynamic experiments of methane at purge process duration from 10 min to 300 min, 1023 K, 20 sccm/min CH
4
for 10 min, 1 g Na
2
WO
4
/Mn/
SiO
2
.
Purge interval 10 min 180 min 300 min
CO
2
CO C
2
H
6
C
2
H
4
CO
2
CO C
2
H
6
C
2
H
4
CO
2
CO C
2
H
6
C
2
H
4
Aver. S 0.08 0.03 0.44 0.45 0.14 0.01 0.48 0.37 0.12 0.01 0.50 0.37
Aver. S 0.11 0.89 0.15 0.85 0.13 0.87
#O/nm
2
7.2 5.4 1.8 3.6 10.4 0.6 4.5 7.1 8.8 0.58 4.6 6.8
#O/nm
2
18.00 21.66 20.66
V. Fleischer et al. / Journal of Catalysis 341 (2016) 91–103 99
2CH
4
þ 2O ! C
2
H
4
þ 2H
2
O ð 15 Þ
CH
4
þ 3O ! CO þ 2H
2
O ð 16 Þ
The results of these oxygen balances are listed in Table 6 . The
calculated oxygen amounts are similar to those from the H
2
-TPR
experiment. Increasing the purge time has no substantial influence
on the available oxygen amount. As mentioned before, most oxy-
gen was converted to form deep oxidation products. Only increas-
ing of the purge time interval to 5 h decreases the peak amplitudes
of all formed products slightly. Another effect is the decrease of the
peak tailing of ethene, whereas the tailing of CO
2
is increased. This
implies that the catalyst structure could have changed. As demon-
strated in our TPD experiment the catalyst material desorbs oxy-
gen at high temperatures. That could lead to a change of the
oxidation state of manganese oxide from Mn
2
O
3
to Mn
3
O
4
, for
instance. Both species have different activity for unselective
methane activation [38] .
With this experiment, we can show that the provided oxygen is
stored under harsh conditions even for hours. This implies that the
active center on the Na
2
WO
4
/Mn/SiO
2
for OCM is bound or inte-
grated oxygen in the lattice structure, because simply adsorbed
oxygen species should not survive such treatment without addi-
tional stabilization. From the mechanistic point of view these
results support the theory of a Mars-van-Krevelen like mechanism
for the OCM reaction on Na
2
WO
4
/Mn/SiO
2
.
Similar step-change experiments in the OCM with this material
were carried out by Salehoun and coworkers [68] . They could
show an increase of the ethane peak by an increase of reaction
temperature, indicating larger amounts of available oxygen. An
interesting difference is the observed ratio of ethane and ethene,
because their results show strong ethane formation and less
ethene formation, but overall C2 selectivity is close to our results.
They also concluded that C2 products may come from lattice oxy-
gen by selective methane activation, where CO is formed by unse-
lective reaction of methane with gas phase oxygen. Both
interpretations confirm our results. They mentioned also that the
Fig. 8. OCM product signals from dynamic experiments with methane with increasing of purge duration at 1023 K, 20 sccm/min CH
4
for 10 min, 1 g Na
2
WO
4
/Mn/SiO
2
,A–
10 min He purge, B – 180 min He purge, C – 300 min He purge.
Fig. 9. Selectivity for purge time reduction of dynamic experiments from 5 min to
15 s, 1 g Na
2
WO
4
/Mn/SiO
2
, 1023 K, 30 sccm/min CH
4
for 10 min.
100 V. Fleischer et al. / Journal of Catalysis 341 (2016) 91–103
reduced catalyst will be reoxidized by gas phase oxygen, while the
catalyst reoxidation rate is much faster than its rate of reduction by
bond cleavage of alkanes or alkenes.
The selectivities observed in short purge time experiments are
shown in Fig. 9 . A detailed oxygen balance for each compound
can be found in Table 7 . The C2 selectivity changes substantially.
The major product is ethane, whereas ethene decreases and is in
the range of the CO
2
signal, which is a minor product. During these
experiments no CO was detected by decrease of residence time. On
the one hand that could be interpreted, that no methane was acti-
vated in an unselective way and the CO
2
results from ethene acti-
vation, according to the discussed reaction network before. On the
other hand it must be considered that CO could be also oxidized to
CO
2
on the catalyst surface. Therefore it is not possible to distin-
guish between methane and ethene activation to form deep oxida-
tion products.
The number of available oxygen atoms stays constant at
20 O/nm
2
. A more or less constant number of converted oxygen
atoms from the catalyst surface indicate that no significant coke for-
mation from ethene disproportionation blocks available surface
oxygen for the OCM reaction, in both sets of experiments. These
experiments imply that the feed contact time with the catalyst
bed has a strong influence on the C2 product ratio, but not on deep
oxidation products. That supports our statement from the TPSR sec-
tion that a different oxygen species must be responsible for deep oxi-
dation product formation, which can interact with methane and
ethene in a different, independent route. In our TPSR experiments
we discussed that the parallel reactions ODH on the catalyst surface
and TDH in gas phase are responsible for C
2
H
4
formation. It has to be
noted that we cannot reveal how strong these routes are involved in
product formation.
Another observed phenomenon is that there is still no influence
of the purge interval on the OCM itself. Even for 15 s purge
Table 7
Compound based oxygen balance for short He purge times in dynamic experiments from 300 s down to 15 s, 1 g Na
2
WO
4
/Mn/SiO
2
, 1023 K, 30 sccm/min CH
4
for 10 min.
Purge time (s) CO
2
(#O/nm
2
) CO (#O/nm
2
)C
2
H
6
(#O/nm
2
)C
2
H
4
(#O/nm
2
)
R
(#O/nm
2
)
300 10.70 0 7.33 2.57 20.60
270 9.41 0 7.64 1.89 18.94
230 9.77 0 7.27 1.77 18.81
160 8.94 0 7.72 1.66 18.32
100 9.14 0 7.19 1.47 17.80
60 9.46 0 7.02 1.52 17.99
30 9.74 0 7.55 2.19 19.47
15 8.27 0 8.55 2.30 19.11
Fig. 10. Total number of oxygen atoms per nm
2
at different purge time durations in dynamic experiments, 1 g Na
2
WO
4
/Mn/SiO
2
, 1023 K, 20–30 sccm/min CH
4
for 10 min.
Fig. 11. Selectivity of the varied methane feed concentration in dynamic experi-
ments, 1 g Na
2
WO
4
/Mn/SiO
2
, 1023 K, 30 sccm/min CH
4
for 5 min.
V. Fleischer et al. / Journal of Catalysis 341 (2016) 91–103 101
duration, there is no shift in the selectivities observed. This result is
only plausible, when short purge processes are sufficient for
removing all gas phase and adsorbed oxygen from the catalyst sur-
face and the reoxidation rate of the catalyst is fast. A sum of the
available oxygen on the Na
2
WO
4
/Mn/SiO
2
is shown in Fig. 10 that
shows no significant change in the available oxygen amount on the
catalyst surface for purge time variation.
3.5.3. Methane partial pressure variation in the reactant feed of
dynamic experiments
The response of the selectivities to methane partial pressure
variation in the reactant feed is shown in Fig. 11 . The ethene selec-
tivity increases slightly when reducing the methane partial pres-
sure. CO
2
is only observed for pure methane feed and only small
traces of CO appear. Therefore the C2 selectivity is around 0.9. A
decrease of the methane partial pressure shifts the overall C2
selectivity for these experimental conditions to 0.96. By reducing
the methane fraction in the feed, the converted oxygen amount
from the catalyst surface decreases in the same ratio ( Table 8 ).
These numbers indicate that less oxygen is consumed, compared
to the dynamic experiments before, which showed 20 O/nm
2
.I n
addition, reducing the reactant feed duration from 10 to 5 min
leads to a decrease of the converted oxygen, in pure methane feed
experiment, too.
This indicates again, that there is linear methane partial pres-
sure dependence for the overall OCM process on Na
2
WO
4
/Mn/
SiO
2
catalysts. Our results are in good agreement with isotopic
labeling experiments of Burch et al., who found a similar correla-
tion [67] .
4. Conclusion
In our H
2
-TPR experiments we found reduction of Na
2
WO
4
/Mn/
SiO
2
catalyst at 950 K. The O
2
-TPD shows weak early oxygen des-
orption at 540 K and a clear oxygen desorption peak at 1070 K.
Both experiments show that the catalyst structure provides oxygen
which interacts with different oxidation potentials.
In TPSR experiments we found that methane can be converted
into CO or ethane. That is caused by two different oxygen species
according to our simulation results. The formed ethane can be oxi-
dized further to ethene in a parallel reaction, by oxidative or ther-
mal dehydrogenation on the catalyst surface or in gas phase. This
process is dependent on the available oxygen amount at the cata-
lyst surface and a serious route in the OCM reaction. Ethene under-
goes disproportion to coke and methane formation. We also
observe CO
2
formation in a parallel pathway. All these pathways
were only obtainable in non-cofeed experiments. Methane TPSR
experiments at heating rates of 1–4 K/min result in an activation
energy of 275 kJ/mole for methane activation and ethane forma-
tion by catalyst bound oxygen. This value is common by the
assumption that lattice oxygen or a strongly adsorbed oxygen spe-
cies is involved in the rate determining step of the OCM process on
the Na
2
WO
4
/Mn/SiO
2
catalyst. The higher activation energy
obtained from a TPSR experiment compared to methane activation
energies from steady state experiments in the literature can be
explained by an energy profile including such an activated oxygen
species, which is shown in Fig. 7 . We conclude that strongly
adsorbed or lattice oxygen is responsible for methane activation,
which has a much lower energy level, compared to adsorbed
molecular oxygen. In our dynamic experiments we found that
around 20 O/nm
2
are available on the catalyst surface for the
OCM process. This number is independent of purge times between
15 s and 300 min and even of the flow rate of methane, which has a
strong effect on C2 selectivity but not on the total product amount.
Such a high number of available oxygen atoms per surface area
indicate involvement of lattice oxygen and may be from sublayers
of the catalyst material, caused on its flexible structure properties.
A linear decrease of the converted oxygen atoms on the catalyst
surface is only observed when reducing the methane feed duration.
Partial pressure variation of methane showed a linear dependence
on the converted number of oxygen atoms, too.
Finally, the nature of oxygen adsorption and conversion by that
catalyst material has strong similarities with other well-known
partial oxidation reactions as the synthesis of formaldehyde or
ethylene oxide from methanol or ethene on silver catalysts
[53,54,69,70] . The function to store a selective oxygen species for
chemical reactions under harsh conditions for different periods of
time might be useful for dynamic reactor concepts, such as chem-
ical looping. Such a concept has the advantage of excluding oxida-
tion reactions in gas phase. Typical material properties for such a
concept are fast storage of oxygen and fast oxygen release for
chemical reactions [71] . In our work we could show that a reoxida-
tion of the material is fast compared to the OCM reaction. A direct
conversion of the strongly bound oxygen species on the catalyst
surface by methane might be a proper manner for a modified
chemical looping concept. Such a concept was already successfully
tested for propane dehydrogenation and is reviewed in detail
elsewhere [72,73] .
Acknowledgment
This work is part of the UNICAT excellence cluster, which is
coordinated by the Technische Universität Berlin. The funding of
this cluster by the German Research Foundation (DFG) is gratefully
acknowledged.
Appendix A. Supplementary material
Supplementary data associated with this article can be found, in
the online version, at http://dx.doi.org/10.1016/j.jcat.2016.06.014 .
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1
I n ves t i ga t i o n o f t h e s ur fa c e re a ct i o n ne t wo rk
o f t h e o x i d a t i ve c o u p li n g o f m e t h an e ov e r
Na 2 WO 4 / M n / S i O 2 c at alys t by te m p era t u r e
p rog ram m e d a nd dy na m i c e xp e r i m en t s -
S up p o r ti n g i n for m a ti on
1. Experiment al
1.1. Setup f or TPSR and Dy namic experiments
The set up for dynamic experiments is illustrated in Fig ure 1 . Different gases can be transported by
mass flow controllers (MFC). In front o f each MFC a switching v al v e (SV) was installed. The
connection pipe between these valves and the reactor entrance has a length of 2 m to handle
pressure fluctua tio n by switchin g the tw o wa y val ves. All co nne ction s were tested by leakage spray
to te st for micro gas leaka ges. That was verified by mass spectr o sc opy to ensure that even no traces
of gas phase oxygen by air leakage co n tribute to dyn am ic ex peri ments. The methan e M FC has two
inlets, which can be s elec t ed by a thr ee way val ve. That allows the dose of ethan e and ethylene. Al l
MFC’s were calibrated by an external flow meter (Gil ian, Gilibrator -2 Calibrator) co nsid ering the
different gases . A detailed analysis about the furnace temperature profile is reported in the next
section. Finall y the outlet of the reac to r was conn ected to a quadrupole mass spectr o m eter (IPI QMS ,
GAM 200).
2
Figure 1 - Constructio n scheme of the dynami c setup
1.2. Reac to r and F u rnac e
The quartz made react or is p resented in Figure 2. The t ype K ther mocouple is plac ed in a quartz
channel. The tip p o siti o n o f that channel is insid e of the catal y st bed. That position was used for
temperature control. The c atalyst material is placed on a quartz frit, which is in th e isothermal zon e
of the reactor.
[Document text truncated for crawler view.]
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