Italy Hydrogen Production Steam

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Simulation of an hydrogen production steam reforming

industrial plant for energetic performance prediction

A. Carrara, A. Perdichizzi, G. Barigozzi*

Dipartimento di Ingegneria Industriale, Universita degli Studi di Bergamo, Viale Marconi 5, 24044 Dalmine (BG), Italy

a r t i c l e i n f o

Article history:Received 7 October 2009

Received in revised form

23 December 2009

Accepted 24 December 2009

Available online 6 February 2010

Keywords:

Gas separation

Methane steam reforming

Hydrogen

a b s t r a c t

This paper presents the results of a theoretical investigation whose aim was the devel-opment of a simulation tool for performance prediction of a steam reforming hydrogen

production plant, and particularly of its overall energetic efficiency. A 1500 Nm3 /h

hydrogen production plant was simulated. Field data coming from an industrial plant were

used for model validation in both design and off design operating conditions. To evaluate

the plant performances in terms of energetic efficiency, a particular attention was paid to

the simulation of all plant auxiliaries consumptions. Nevertheless the large uncertainty in

most of the field data values, the model was able to capture all the relevant phenomena

taking place in all the plant components, from reformer reactor up to CO2 sequestration

unit, in the investigated plant capacity range (40–100%).

ª 2009 Professor T. Nejat Veziroglu. Published by Elsevier Ltd. All rights reserved.

1. Introduction

Alternative energy sources and power generation technolo-

gies are required to face the declining of fossil fuel stocks as

well as the effects of carbon dioxide (CO2) emission on global

warming. From these points of view, hydrogen (H2) is a very

promising clean fuel,as no CO2 is produced by its combustion,

available as fuel for distributed power systems [1], for example

fuel cell systems [2]. Hydrogen is also an important raw

material for the chemical and refining industries, e.g. for the

production of ammonia and methanol.

In the last years an increasing interest from the energysector on hydrogen production techniques has been

observed. Today most of hydrogen is produced from fossil

fuel sources [3]. Clean production of hydrogen, for example

by water electrolysis using renewable energy, seems not

to be yet competitive with present-day renewable energy

technologies.

Hydrogen production from fossil fuel has as a conse-

quence CO2 generation. Application of CO2 separation and

sequestration systems to hydrogen production process has

often as a consequence a considerable reduction of process

energetic efficiency. For the near and medium term, the use

of hydrogen as energy vector and fuel in distributed power

plant systems needs to increase the energetic efficiency of

hydrogen production system from fossil fuels.

About 50% of hydrogen production in the world today is

based on methane steam reforming [4]. The methane steam

reforming process is based on two main reactions: the

reforming reaction

CH4 þH2O ¼ CO2 þ 3H2 DH298 ¼ 206 kJ=mol (1)and the water gas shift reaction (WGS):

CO þH2O ¼ CO2 þH2 DH298 ¼ 41 kJ=mol (2)

In the last years many research activities have shown the

potentiality of membrane technology [5]. Removing one or

more of the products with a membrane would cause a shift in

the reaction thermodynamic equilibrium, increasing the yield

of CH4 and CO conversion [6].

* Corresponding author. þ39 035 2052317; fax: þ39 035 2052077.E-mail address: [email protected] (G. Barigozzi).

A v a i l a b l e a t w w w . s c i e n c e d i r e c t . c o m

j o u r n a l h o m e p a g e : w w w . e l s e v i e r . c o m/ l o c a t e / h e

i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 3 4 9 9 – 3 5 0 8

0360-3199/$ – see front matter ª 2009 Professor T. Nejat Veziroglu. Published by Elsevier Ltd. All rights reserved.

doi:10.1016/j.ijhydene.2009.12.156

mailto:[email protected]

http://www.elsevier.com/locate/he

http://www.elsevier.com/locate/he

mailto:[email protected]





After sulfur removal, methane is mixed with superheated

steam with a design water to natural gas mass flow ratio

r ¼ 3.296 kg w /kg CH4, and then heated up by combustion

products to about 830 K (CC-2). The hot flow finally enters the

reformer unit (R), consisting in a radiant and a convectivesection. The radiant section includes eight reaction tubes,

with a Ni/CaO/Al2O3 catalyst, and two burners, making use of

methane and purge gas as fuel. The process gas leaves the

reformer at about 20 bar and at a temperature in the 1020–

1130 K range. Its design composition is reported in Table 2.

Before going to the shift reactor (SR), the process gas is cooled

to about 540 K, providing heat to the boiler (E1) for steam

generation and for superheated steam production (E2). The

mean temperature (570–670 K) adiabatic shift reactor makes

use of a Cu based catalyst, with a ZnO and Al2O3 support to

avoid sintering. The syngas then leaves the shift reactor at

about 18 bar and 620 K.

Syngas, H2 rich and CO poor, is then cooled down to about310 K for water condensation (V2) in three heat exchangers

(E3, E4, E5). The first two use the water process as cooling flow,

thus acting like an economizer providing pre-heated water to

the boiler. The third instead makes use of well water.

The dry syngas finally goes through the absorber unit for

CO2 sequestration. The remaining syngas stream, CO2 poor,

whose composition is reported in Table 2, enters the Pressure

Swing Adsorption (PSA) unit, for hydrogen separation. In the

PSA unit, thanks to the presence of molecular sieves and

a cyclic process, 82% of hydrogen is separated form syngas.

The remaining purge gas (Table 2) is used as fuel for the

burners in the reformer radiant section, while the hydrogen

flow is compressed from 16 bar to 200 bar.

2.2. Steam generation section

Condensed syngas water mixed with feed water is pressurized

to about 21 bar (P1 A/B), heated up (E4), de-aerated (DA) and

then pumped (P2 A/B) and heated up again (E3) by the process

gas before entering the LP drum (V1). From LP drum the feed

water goes through the boiler for steam generation at 26 bar.

Two boilers exist: the first one (E1) makes use of the processgas exiting the reformer reactor, while the second one (CC-1)

is located inside of the reformer convective section, recov-

ering heat from the combustion gas coming from the radiant

section. Steam is then superheated (E2) to about 570 K to be

finally mixed with the compressed natural gas and directed to

the reformer reactor.

2.3. Combustion line

The purge gas is burned with air in the reformer radiant

section, to generate the heat necessary for the reforming

process. Unfortunately, the purge gas mass flow rate is not

large enough to reach the right temperature in the reformerreactor, so natural gas is added. Purge gas combustion

supplies up to 80% of process heat, while natural gas allows

for bridgewall temperature control, that is the combustion gas

temperature at the exit of the reformer radiant section. In

design condition, the combustion products leave the reformer

radiant section at about 1230 K. The gas then enters the

convective section, where a boiler (CC-1) and three heat

exchangers (CC-2, CC-3, CC-4) exist: one for natural gas

heating before sulfur removal, one to heat the natural gas–

steam mixture before entering the reaction tubes and the

latter, located at the end of the convective section, to pre-heat

the combustion air. Combustion gases are finally released in

the atmosphere. The presence of two fans assures thecombustion products circulation through the reformer and

the heat exchangers.

2.4. CO2 absorption and liquefaction

After water separation, the syngas enters the absorption

tower (Fig. 2); the two stage absorption process is based on the

use of a mix of water and methyl-diethanolamine (MDEA).

While the syngas, CO2 poor, goes through the PSA unit for H2

separation, CO2, water and MDEA flow is regenerated in a sub-

atmospheric stripper tower. The pure CO2 flow then goes

through a two stages intercooled compressor before lique-

faction in a refrigerating thermodynamic cycle.

3. The model

The plant model was initially set up at design condition (100%

plant capacity) using the commercial code Aspen Plus

(version 2004.1). Aspen Plus software is a chemical engi-

neering process modeling tool suitable for steady state

simulation, design and performance monitoring. Aspen Plus

simulation software was used for the mass and energy

balance calculations. Plant model is structured, like the real

plant, in four sections: hydrogen production and compression,

water section, gas combustion line, and CO2 sequestration

Table 2 – Design gas compositions-plant capacity 100%.

Molar concentration %

Reformer exit PSA inlet PSA exit (purge gas)

CH4 3.10 4.91 18.64

H2 49.27 89.82 61.37

CO 8.98 2.534 9.63

CO2 5.60 2.06 7.81

H2O 32.84 0.34 1.30

N2 0.208 0.327 1.24

He 0.002 0.003 0.01

Table 1 – Design methane composition.

Element Molar concentration %

CH4 96.34

C2H6 1.79

C3H8 0.41

n-C4H10 0.08

i-C4H10 0.06n-C5H12 0.01

i-C5H12 0.01

C6H14 0.01

CO2 0.09

N2 1.19

He 0.01

S (mg/Nm3) 5

i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 3 4 9 9 – 3 5 0 8 3501





the reformer reactor exit section. The design specification for

the CH4 molar concentration has been again obtained from

measured data (4.0%).

In the shift reactor of course the only possible reaction is

the shift reaction (2). A major difference with respect to the

reformer is that the shift reactor has a fixed working pressure

of 18.3 bar and it is an adiabatic reactor. The chemical equi-

librium is anyways calculated defining the pressure value and

the already mentioned possible products, with CH4 defined as

an inert. Table 4 summarizes the main input data for both

reformer and shift reactor simulations.

All process gas coolers are simulated using classical heatexchangers, imposing the gas outlet temperature in the

design condition and computing the recovered heat. Pres-

sure losses are evenly distributed between the heat

exchangers in such a way to obtain the design PSA operating

pressure (17 bar). The PSA process is finally simply simulated

using a splitter, fixing the percentage of hydrogen separated

from the dry syngas at about 79%. This value is slightly

different from the design specification and was derived from

field data; moreover it was practically unaffected by changes

in plant capacity. The hydrogen is then pressurized using a 3

stage intercooled compressor, from which it is delivered at

200 bar and 297 K. A 50% overall compressor efficiency

was used. Table 5 summarizes the main input parametersfor the design of heat exchangers, PSA and hydrogen

compressor unit.

3.2. Steam generation section

The required water mass flow rate depends on the water to

natural gas mass flow ratio r, the fixed plant capacity and

the natural gas mass flow rate entering the reformer

reactor. Once the need of water for the reforming process is

known, the system calculates the feed water mass flow,

which depends on the condensed water and on the de-

aerator requirements. Fig. 4 shows all the components thewater/steam flow encounters before mixing with the

methane.

It has to be observed that, in the real plant, two boilers

operating in a parallel configuration do exist. In the present

analysis the evaporative section was instead modeled using

a serial configuration. In this way, not a complete evaporation

takes place in the two boilers. In fact, the heat recovered from

E1 heat exchanger, already computed in the previous section,

is used to produce wet steam, that becomes saturated only at

the exit of CC-1 heat exchanger.

Finally, to correctly estimate the feed water pumps power

consumption, manufacturer supplied performance curves

were included into the model. All input parameters aresummarized in Table 6.

3.3. Combustion line

Fig. 5 shows the combustion products line while model input

parameters are summarized in Table 7. Three input flows do

exist (purge gas, natural gas and air) and one output flow (the

combustion gas discharged into the ambient at the stack).

The required natural gas mass flow is computed knowing the

purge gas energy input and the reformer reactor energy

requirement. The necessary air mass flow is then computed

specifying the excess air 3 in the model input data. A 10%

excess air applied to the 100% plant capacity designcondition.

The reformer radiant section is simulated using 2 units:

an RSTOIC reactor, for combustion simulation, and a heat

exchanger (CC-R) to simulate the heat exchange between

combustion gas and reformer reactor. An RSTOIC unit is

suitable to simulate a reactor when reaction kinetics are

unknown or unimportant and when stoichiometry and the

molar extent or conversion is known for each reaction. The

reformer reactor (R) already introduced in the hydrogen

production and compression section thus calculates the

heat necessary for the reforming process while CC-R

removes this heat from combustion gas, knowing the

combustion gas exit temperature, i.e. the bridgewall

Table 3 – Reactions taking place in the Reforming reactor.

CH4 þ H2O4 CO þ 3H2 NC5H12 þ 5H2O4 5CO þ 11H2

C2H6 þ 2H2O4 2CO þ 5H2 IC5H12 þ 5H2O4 5CO þ 11H2

C3H8 þ 3H2O4 3CO þ 7H2 C6H14 þ 6H2O4 6CO þ 13H2

NC4H10 þ 4H2O4 4CO þ 9H2 CO þ H2O4 CO2 þ H2

IC4H10 þ 4H2O4 4CO þ 9H2

Table 4 – Input data and main assumption for reformerand shift reactor.

Reformer Shift Reactor

Inputs p ¼ 19.7 bar p ¼ 18.3 bar

Assumptions – Adiabatic

Specifications Restrict chemical equilibrium.

Specify temperature approach or reaction.

Possible products CH4, H2, CO, CO2, H2O, N2, He, Ar

Property method SRK IDEAL

Inert – %CH4 ¼ 1

Restricted

equilibrium

See Table 3 Individual reaction:

CO þ

H2O¼

CO2þ

H2

Table 5 – Design data for heat exchangers, PSA and H2

compressor.

Natural gas supply temperature 350 K

Natural gas supply pressure 21 bar

CC-2 exit Temperature 830 K

CC-3 exit Temperature 650 K

E2 exit Temperature 540 KE5 exit Temperature 310 K

PSA pressure 17 bar

PSA %H2 separation 79 %

H2 compressor exit pressure 200 bar

H2 compressor efficiency 50 %

1 Intercooler

exit temperature

306 K

2 Intercooler

exit temperature

307 K

3 Intercooler

exit temperature

297 K




temperature. The air pre-heater CC-4 has been designed in

such a way to heat the air up to 650 K. Finally, a constant

fans overall efficiency of 50% was used all over the consid-ered operating range.

3.4. CO2 absorption and liquefaction

The CO2 sequestration section is made of the ammine cycle

(Fig. 2) and the CO2 liquefaction cycle. The ammine cycle has

two input flows, the dry syngas and a mix of water and

ammine, and two output flows, the CO2 poor syngas and the

ammine CO2 rich flow. A mix of water, 50%wt, diethanolamine

(DEA), 25%wt, and methyl-diethanolamine (MDEA), 25%wt

was used [15]. The ammine mass flow delivered to the

absorber tower was set so to have the design CO2 concentra-

tion in the syngas CO2 poor flow. A constant efficiency of 60%

was used in the water and ammine mixture circulation pump

model.

The CO2 rich flow is regenerated in the stripper tower.

Before entering the stripper the CO2 rich flow is heated up by

the hot regenerated flow. Both absorber and stripper towers

are modeled with a RADFRAC unit. RADFRAC is a rigorous

model for simulating all types of multistage vapor-liquid

fractionation operations. Simulation model calculates re-

boiler duty so to have an exit ammine flow temperature of

about 380 K.

About 90% of CO2 is sequestrated; at a plant capacity of

1500 Nm3 /h this results in 750 kg/h of pure CO2 generated. A

two stages intercooled compressor (50% overall efficiency)

allows the pure CO2 flow to reach 17 bar and 295 K before

liquefaction in a refrigerating thermodynamic cycle. In the

refrigeration cycle, the R507 working fluid flow is

compressed at 13 bar and 340 K, condensed and finally

laminated before entering the evaporative section for CO2

liquefaction. Every detail of this cycle has been also included

in the model, thus allowing to compute the electrical power

consumption. In particular, the compression section has

been simulated using two compressors, in parallel configu-

ration, having an 83% isentropic efficiency. Table 8 summa-

rizes the input parameters to the CO2 absorption and

liquefaction section model.

4. The plant control strategy

A high degree of automation characterizes the plant. The

operator in fact only selects the desired plant capacity,while the control system automatically sets all the relevant

parameters. In particular the plant capacity is allowed to

vary within the range 30–100%. The plant capacity specified,

the control system modifies the steam over carbon ratio r

and the bridgewall temperature Tbw following the values

reported in Fig. 6. It has to be observed that r has to be

always maintained larger than 1.7 to avoid coke formation.

To control the bridgewall temperature, the system in turns

modifies the methane and combustion air mass flow rates,

knowing the required excess air 3 (Fig. 6).

The combustion computation is performed in an iterative

way: the model initially calculates the purge gas mass flow to

combustion line and sets the air mass flow according to theprescribed excess air assuming no natural gas to the burner.

The CC-R combustion gas exit temperature and the bridgewall

temperature are then computed and the latter is compared

with the prescribed value. In case a difference between

computed and prescribed values does exist, a methane mass

flow is added to the burner so to reach the correct bridgewall

temperature.

A modification in the plant capacity will also imply

changes in pressure and CH4 molar concentration at the exit

of reactor. Pressure and CH4 rate variations were then

provided to the model (Fig. 6), allowing to calculate the reactor

gas exit temperature and the reactor energy balance for vari-

able plant capacity.

Fig. 4 – Water section.

Table 6 – Water section input data.

Feed water temperature 288 K

E4 exit temperature 368 K

E3 exit temperature 488 K

Evaporation temperature 499 K

E2 exit

temperature (superheated steam)

570 K




The laws of variation of all those parameters as a function

of plant capacity were derived fitting field data with poly-

nomial interpolation curves.

5. Model validation

Figs. 3–5 report pressure, temperature and mass flow rate

values of most significant streams provided by the model at

design condition. Table 9 reports the overall mass balance of

the steam reforming plant at design condition. The mass

balance of re-boiler of the stripper tower in the CO2 seques-

tration process is not considered. All the simulation outputs

are in a reasonable agreement with field data. In particular,

one can observe that both reactor and shift reactor exit

temperature values are correctly computed (1048 K and 617 Krespectively), as well as the steam production (1429.5 kg/h).

With a methane input of 433.7 kg/h, an hydrogen production

of 135.9 kg/h is achieved as well as a CO2 production of

845.6 kg/h. The purge gas to the reactor (222.5 kg/h) is

combined with a small methane mass flow (17 kg/h), giving

rise to combustion gases exiting the stack at 570 K. The

659.2 kg/h condensed water flow coming from the water

separation unit is combined with feed water to restore the

required mass flow for steam production.

The overall energy balance at design condition is reported

in Fig. 7. Besides to reactor and burner, methane is also

provided to the re-boiler of the stripper tower (75.2 kg/h) to

regenerate the CO2 rich flow. Steam reforming section, CO2

sequestration and liquefaction section and H2 compression

section all contribute in a similar way to the auxiliaries power

consumption. The quite high loss level (about 3000 kW,

roughly corresponding to 40% of energy input) can be ascribed

to losses in energy conversion and heat transfer processes, as

well as in the heat rejection to the ambient through

compressor intercoolers, in the liquefaction cycle and in re-

boiler stack. Moreover, it has to be underlined that, even if the

energy value of CO2 is zero, its sequestration is very expen-

sive. The same stands for the energy related to the hydrogen

output: in theenergy balanceonlyits LHVis considered, but to

be safely stored it needs to be compressed and, again, this

requires an extra energy input.

Once the model has been set up in design condition, off

design runs were performed for validation purposes,

comparing the code results against experimental data. Field

data also including chemical composition of syngas in

different plant locations as well as electric auxiliaries

absorption were provided over a wide range of plant capac-

ities, from 45% up to 100%.

Before presenting the validation results, the reader has to

be awarded that field data here reported for comparison are

mean values, as they have been obtained averaging, at

constant plant capacity, the plant control panel data recorded

over a one month time period. Deviations as large as 5% in

some of the recorded parameters were observed. Moreover,

the model input data for the off design computations have

been also obtained by interpolating the corresponding aver-

aged field data. Finally, sensors precision is not always

adequate for simulation purposes, even if it is so from

a control point of view. All these aspects will influence both

Fig. 5 – Combustion gas line.

Table 7 – Combustion line input data.

3 10 %

CC-R exit temperature (Bridgewall Temperature) 1230 K

Ambient temperature 288 K

CC-4 air exit temperature 650 K

Table 8 – CO2 sequestration and liquefaction input data.

Absorber top stage pressure 16.5 bar

Nr of absorption stages 2

Stripper top stage pressure 0.2488 bar

Nr of stripper stages 2

Ammine flow exit temperature 380 K

Circulating pumps efficiency 60 %

CO2 compressor exit pressure 17 bar

CO2 compressor exit Temperature 295 K

CO2 compressor overall efficiency 50 %

Refrigeration cycle working fluid R507

Refrigeration cycle condensing pressure 13 bar

Refrigeration cycle compressor efficiency 83 %






investigated domain. Larger differences only take place at

the lowest plant capacity.

Finally, the model prediction capability of auxiliaries

(pumps, fans and compressors) electrical absorption was alsovalidated against field data. Unfortunately, beside the global

plant auxiliaries power absorption, only the electrical power

absorbed by the CO2 sequestration section is available from

the plant control panel. Fig. 12 reports the comparison

between simulation and field data. A general good agreement

between simulated and field data characterizes both overall

and CO2 section auxiliary consumption. The small difference

existing between measured and computed global auxiliary

power consumption is mainly related to the different opera-

tion strategy of some pumps: continuous in the model,

discontinuous because coupled with storage tanks in the real

plant.

Finally, to allow for a comparison between different plantsolutions, for example without and with the introduction of

permeable membranes, an energetic efficiency for the

hydrogen production plant was introduced. It has been

defined as the ratio between the energy of the produced

hydrogen and the energy associated to all plant inputs: the

methane mixed with steam introduced in the reformer,

burned in the reformer radiant section and in the ammine

regeneration process. Two definitions were used: a gross

efficiency, not considering the electrical auxiliaries power

consumption (4) anda net efficiency,also including it (5).Tobe

consistent with the other terms in the efficiency formulation,

the auxiliaries electrical power Pel,aux has been first converted

into its equivalent methane consumption, by means of

a conversion efficiency hconv of 43%, the average conversion

efficiency from natural gas to electricity suggested by the

Italian Regulatory Authority for Electricity and Gas (www.

autorita.energia.it) for 10–25 MW sized power plants oper-

ating in the Italian scenario.

hgross ¼mH2

LHVH2mCH4Reformer þmCH4 Burners þmCH4 Stripper

LHVCH4

(4)

hnet ¼mH2

LHVH2mCH4 Reformer þmCH4 BurnersþmCH4Stripper

LHVCH4 þPel;Aux=hconv

(5)

Fig. 13 compares plant gross and net efficiencies computed

from the model against the same quantities derived from field

data for different plant capacities. First of all a very stable

behavior of the hydrogen production plant can be observed,

for variable plant capacity. An average gross and net plant

energetic efficiency of 63.5% and 52.1% respectively have been

measured, with deviations of about 2.5% for both definitions.

Moreover, the model prediction capability of the overall plant

conversion efficiency is quite good, as maximum differences

of about 3.5% do exist between field data and simulation

results.

Fig. 10 – CH4 mass flow rate to the burner for variable plant

capacity.

Fig. 11 – Reformer exit CO2 molar concentration for variable

plant capacity.

Fig. 9 – Sequestered CO2 mass flow rate for variable plant

capacity.

Fig. 12 – Electrical need of global plant and CO2 section for

variable plant capacity.


http://www.autorita.energia.it/






6. Conclusions

The present paper presents a simulation tool for design and off

design performance prediction of a hydrogen production

industrial plant based on methane steam reforming. This gavethe opportunity to deeply analyze the off design behavior of the

whole plant, in the meanwhile providing a useful tool for future

investigation on the effects of the introduction of Pd–Ag

membranes on theenergeticperformances ofthe plant,whichis

being the topic of a following paper. To deeply investigate

different plant solutions, a model was developed in AspenPlus

environment. Many details have been included in the model, in

order to simulate as close as possible all plant operational

features, from thermodynamics properties, to chemical

composition of main streams to electrical auxiliaries consump-

tion. An energeticefficiency hasbeen also introducedto quantify

the methane to hydrogen conversion process quality.

The model, developed under design conditions, has been

successfully validated over a wide range of off design opera-

tions, through a comparison against field data. Some differ-

ences between real plant data and simulation results have

been evidenced, but they have been mainly ascribed to a lack

of accuracy in the field instrumentation. Anyways, the model

was able to correctly capture the trends of variation with plant

capacity of all relevant parameters.

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Nomenclature

LHV: lower heating value, J/kg m: mass flow rate, kg/sp: pressure, barPaux: auxiliaries electric power, kWr: water to methane mass flow ratio

T: temperature, KTbw: bridgewall temperature3: excess air [%]h: energetic efficiency [%]

Fig. 13 – Gross and net plant electrical efficiency forvariable

plant capacity.


Italy Hydrogen Production Steam

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