Q & A > Hydrogen Technologies
Date  Replies
04/02/2021 Q: I work in a UOP licensed DHDS unit. It has one HP amine absorber to absorb H2S from recycle gas. At thr top of the column, a water wash facility with level tray is installed to wash lean amine from recycle gas. We are experiencing foaming. The water tray level quickly fills and goes to the next recycle gas knockout drum. Delta pressure of the column is also increasing and the level increases suddenly in the water tray even after isolating fresh water. The delta temperature between amine and process gas is 12oC. Drained liquid is milky in colour. Please suggest a remedy.

 		(3)
13/10/2020 Q: Which is the best technology (CPOX /SMR/PX/ATR/CR) for producing clean syngas from refinery off gas and PSA tail gas in terms of capex and opex ?  
07/10/2020 Q: Is there any commercial plant for producing clean syngas from refinery off gases (the tail gas from CRU units having no CO content)? (3)
04/10/2020 Q: In an energy conservation programme I'm focusing on steam calculation starting from the reboiler and the distribution lines in the refinery. Could you please tell me the first step to do this calculation, whether anyone has a relevant book or article, and can I use Hysys software? (1)
16/05/2020 Q: What are the causes behind the reformer outlet temperature dropping without any changes on the DCS? A few causes may be: 1.Burner fuel firing CV malfunction 2. Steam flow increased 3.Combustion air flow increased. Any other causes of a drop in the reformer outlet temperature? (1)
05/04/2020 Q:
After a UOP hydrotreating catalyst changeover an issue was observed in the product colour which became 15 instead of 30 while no particulates were present. I am seeking any explanation and solution. 		(5)
17/01/2020 Q: How can the effect of temperature increase in steam methane reforming over the catalyst lifetime be calculated?  
01/01/2020 Q: It is planned to replace the inlet system of the Reformer in a Hydrogen Generation unit with upgraded metallurgy from SS 304 to SS 347. This upgradation is warranted due to an increase in inlet temperature of the Reformer feed (Prereformer is being introduced as a part of a revamp).
We would like to know if anyone has carried out such modification in your hydrogen generation unit.
Please share the details including precautions to be taken.
 		 
12/11/2019 Q: Hello, In our hydrogen production plant we have our own steam generation system that produces 36 bar of steam at 390 degree C temperature. Steam produced in the steam drum has a continuous blow down vessel. We noticed that the conductivity analyzer goes so high that it reaches to NAN sometimes as it automatically reaches to its normal values without any external steps taken by us. We are curious about this automatic adjustment of conductivity in blow down. We have checked the instrument as well, but there is no issue with the analyzer. It is to be noted that the BFW is mixed with condensate separator before pressure swing adsoprtion. The condesate water first goes to degasifier and then mixed with demin water and then pumped to steam drum where steam is made. From couple of days pH of condensate is very low due to low inlet temperature due to dissociation of carbon dioxide, which made pH low. Kindly tell us if there is any reason for high conductivity as per our case. (1)
18/09/2019 Q: What is the reason for coke formation in the wet gas compressor and what is the solution? (1)
22/04/2019 Q: On what terms HT/ MT/LT shift converters are decided for Hydrogen unit?
 		(2)
26/01/2019 Q: Linden based Hydrogen unit, NG is using as feed for hydrogen generation. There is a hydrogenation reactor with COMOX catalyst where hydrogenation reaction takes place after that desulfurization reaction in next reactor with ZNO catalyst. NG feed having no sulfur and we operate the hydrogenation reactor in line for many hours. Since there is no S in NG feed no H2S formation will occur in that reactor. Catalytic reaction will not occur. As per requirement, we take sometime RFG (refinery fuel gas) , which having sulfur content for Hydrogenation reaction in COMox based reactor. My question is this type of feed variation, deactivated catalyst may work after taking RFG feed or how we can make activate COMoX catalyst when NG feed in line alone. (1)
22/11/2018 Q: Can anyone explain me what are all the parameters that affect H2 purity in Recycle gas in case of a Hydrotreater/Hydrocracker ? (4)
03/10/2018 Q: Is there any limitation for taking Hydrogen rich off gases as feed to reformer? My doubt is whether excess hydrogen in the feed to reformer can it cause reversible reactions as steam methane reforming is reversible in reaction? (4)
27/09/2018 Q: Is there any limitation for taking Hydrogen rich off gases as a feed to Reformer , If yes what is the limitation ? Is it because of Le Chatlier's principle if Hydrogen is more will it cause backward reaction? Or else is there any other limitation ?  
25/09/2018 Q: Our naphtha sample from crude distillation unit went off on chloride content with result showing 1.3 ppm.So we checked organic chloride in our crude tank. The result showed presence of organic chloride in that particular crude tank. So that tank is kept blocked and we resumed to our normal operation.
We did many brain-stroming and looked for solution.
Can the experts in this forum please suggest how to process this crude having high organic chloride. Is there any chemical treatment available. How can this problem be solved? 		(6)
02/05/2018 Q: We have three smr plant having 750nm3 capacity.one of the smr syngas header external paint was faded and peeled out from its position. Checked skin temperature on the surface, found higher than normal. What is the cause of burnt paint? Is there any special procedure for measuring skin temperature on syngas header like reformer tube temp measurement? (1)
18/03/2018 Q: In our HMU plant, we are going to do Pre-reformer reduction but we don't have any water collection system. Can anybody explain how to do the reduction and how to ensure that reduction of Pre-reformer is complete?  
13/03/2018 Q: We have kerosene hydro bone unit (desulphurization unit) , we faced technical problem which is non-improving (WSIM) water separometer index modified specification .
what is the mean thing that effect in (WSIM), and how can we improve it?
By the way, the reactor system of kerosene hydro bone unit is isolated.
 		(3)
10/03/2018 Q: In Hydrogen Generation Unit, Total Organic Carbon (TOC) is observed high in process condensate water collected from process condensate separators. What could be the reasons n how to rectify it.  
09/03/2018 Q: In Hydrogen generation unit, it is observed higher total organic content (TOC) in process condensate water collected from condensate separators. What could be the probable reason and how to mitigate/ rectify? (1)
01/03/2018 Q: Are there any differences between a naphtha hydrotreater and a diesel hydrotreater? (5)
26/01/2018 Q: In our refinery, C5/C6 isomerization process uses zeolite based catalyst and reaction loop pressure is maintained at 18 kg/cm2g. The ISOM unit uses H2 makeup from CRU and there is continuous purging of recycle gas from ISOM separator.
the design basis is 1635 kg/hr of purge gas (contains 187 kg/hr of H2) with 63 vol% H2 purity.
No HCL. 5-12 ppm of H2S.
Please advise how economical will it be to recover the H2 from purge gas using membrane separation process?
is there any refinery currently using membrane separation process for recovering H2 from purge gases at 18 kg/cm2g? 		(5)
09/12/2017 Q: I am interested in the forum's experience on (full) reusing of processing condensates from conventional Natural Gas / LPG steam reforming Units (from all major licensors). I have experience with an ex-KTI (Technip) based NG SMR whereby process condensates are fully a) fully reused in the BFW system post a simple air stripping for removal of dissolved CO2 and b) steam produced (at 40 barg) is used for both internal needs and export to other process user and a steam turbine / generator. The system operates as such successfully for many years. I understand that recent designs (and at higher steam pressures) call for either a) the steam not to be exported or b) a dual steam raising system one for internal consumption (where the process condensates can be recycled) and another for export steam on clean BFW. This, due to issues with organics difficult to be removed from the BFW and generated in the shift reactors.
Any experiences / views on this ?
(1)
25/10/2017 Q: What is the effect of free water in Naphtha feed on catalyst in Hydrogen generation unit? What % of water can be allowed in feed?  
08/09/2017 Q: I am working in UOP hydrocracker unit. In our feed filter backwash frequency is very high since one month. We have cleaned the filter elements one by one. After cleaning of filter some improvement seen for some days but then filter starts backwashing again. we have also checked CCR and asphaltene and all are within range. we process VGO and HCGO in our hydrocracker unit . It is recycle type with 97% conversion. please suggest remedy. (8)
31/07/2017 Q: What is meant by "False Air Damper" in reformer?
How does it work? 		 
17/07/2017 Q: What is the effect of reducing the system pressure in a gas oil/naphtha hydrotreater? Will this be able to reduce the hydrogen consumption given that there is still allowance on catalyst deactivation? What parameters do we need to consider before reducing the system pressure? (2)
14/07/2017 Q: How to find out the right/optimum flow of hydrogen in reductor in continuous catalytic reforming unit? What if I operate reductor at maximum capacity in my plant?  
12/01/2017 Q: We are currently designing a new grassroots unit for diesel hydrotreater (DHT). There are 2 different opinion when it come to hydrogen mixing point: either it is mixed before or after combined feed exchanger (CFE) .
The view for the mixing point to be after CFE have concern about polymerization or faster fouling inside the CFE while the view with mixing point before CFE saying the impact will be totally the opposite.
What is the basis/philosophy for DHT design on where to put the mix point? 		(8)
05/01/2017 Q: There are 2 hydrogen plants in our company, We use NG as feed(SMR Process) to produce high purity hydrogen. The raw hydrogen stream comes from steam reformer, after cooling down which is then sent to PSA to recover high purity hydrogen.
The problem is that we have a tube bend and the catalyst has last been changed in 2011, we don't need to run plant at full capacity as our need is not that much,
We are thinking of postponing the changing of catalyst for a year or so, but is it feasible??
What is the time when i should know that methane slippage is more now. Till how much percentage of methane slippage i can afford in plant which doesn't hamper CO Shift Catalyst and PSA Catalyst.
Also when i should know that its time for The Reformer catalyst to be change?
There are some comments from outsources about waiting for 6 months as we are getting purity and methane slippage is within limit.
Would you please advise on how much methane slippage is ok if maximum hydrogen production isn't concern but specific consumption should be low, which is the current scenario. 		(2)
14/09/2016 Q: We have replaced lead sulphur guard absorbent in Hydrogen unit. H2S is observed at lead and lag sulphur guard absorbent outlet. What can be the reason for it? What is the max H2S limit for pre-reformer I/L feed? Unit is operating on off gases and Natural gas Plant. (3)
15/07/2016 Q: IN our HCU Recycle Gas Compressor Turbine, make BHEL from one month Turbine Front & Rear journal bearing vibration suddenly increase from 15 micron to 80 micron for about 10 minutes and came back to normal during this period turbine bearing temperature also increase 3-9 degree centigrade. Please suggest the root cause. We do centrifuge of lube-oil once in 24 hors about 2-3 hours and lube oil also replace in december 2015. (1)
06/07/2016 Q: I am working in hydrocracker unit. Since two month LPG is getting failed due to positive H2S and in copper corrosion test. We are maintaining debutanizer bottom and top temperature 186 & 78 degree C and pressure 16.5 kg/cm2g. Our LPG r/d flow is10-12 m3/hr while lean amine flow is 20-25 m3/hr. When we check caustic strength found 20-22% which is quite normal. Generally we change caustic in 5-7 days.
Our system is like this LPG comes at top of debutanizer and goes in amine absorber at bottom and lean amine mix at top of absorber after amine wash goes to water wash amine settler where water circulation is done by pump before going to water wash LPG and water is mixed and mixed in mixer. After water wash LPG goes to caustic wash tank through nozzle but no circulation pump is there (Tank). After caustic wash LPG goes to sand filter and then goes to Deethaniser for FG removal and cooled in Water cooler and goes to LPG storage tank. Please suggest the solution.



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09/02/2016 Q: What will be the PONA content of vacuum gas oil which is used in hydrocracker? (1)
09/02/2016 Q: In hydrocracker unit what are the amount of heat of energy invloved for following reactions.
1) olefin saturaion
2) aromatic saturation
3) denitrification
4) desulfurisation
 		(1)
18/01/2016 Q: What is simplest method to Convert Nm3/hr to kg/hr
For gas mixture ( H2 +H2S)? 		(2)
10/12/2015 Q: Why does Hydrogen run counter to Joule-Thomson principles for gases? (3)
07/11/2015 Q: Why at low plant load S/C ratio is to be kept high? (3)
10/08/2015 Q: In Hydrogen plant steam flow is fixed at 30-70 % plant load in Haldor Topsoe unit, while in Linde hydrogen plant steam flow is fixed in between 30-50%. What are the reasons for this difference? (1)
20/07/2015 Q: In CCR unit we have 85% rich H2 coming from PSA as off gas. Currently we are using in the fuel gas mixing feeding our furnaces. I would like to ask whether it can be used in more fruitful ways?
Can it be combined in HGU at any stage as we already have 85% pure hydrogen??
(3)
16/06/2015 Q: We are observing high CS2 content in our straight run naphtha. This is not on regular basis but frequent and sometime it goes up to more than 20-25 ppm also. Please advise what can be source of such high CS2 content in naphtha intermittently.
The sources may be narrowed down to:
1. Presence of CS2 in Crude itself- Please suggest the probability of the same and if any known crude with high CS2?
2. Since CS2 formation requires very high temp, can it be formed in crude heaters? or any other process?
3. Though probability is less, can it come from recycle hydrotreated naphtha? 		(2)
12/06/2015 Q: In Hydrogen Generation Unit the pre-reformer reactor (having Ni based catalyst) differential pressure increases after every unit start-up by 0.1-0.2 kg/cm2. Before reformer feed cut, naptha vapor warmup line is kept lined up and reactor is kept at 470-490 deg C. Also, before naptha feed cut, catalyst re-reduction takes place under hydrogen+steam atmosphere. What is the reason for del P increase after every unit startup?  
11/06/2015 Q: On every "hydrogen generation unit" start up, the Pre-reformer reactor differential pressure increases by 0.1-0.2 kg/cm2. Before Reformer feed cut, the reactor catalyst temperature is maintained 470 - 490 deg C
Before feed cut, naptha warm-up line lined up. What is the reason for Pre-reformer del P increase? 		(2)
23/04/2015 Q: I was asked for evaluation of doing the leak test and pressurization step in the Hydrocracking unit start up with Nitrogen instead using Hydrogen. We have 1 reciprocant compressor for make up and one centrifugal compressor for recycle gas, I would like to know what do I have to consider to make this evaluation, what I know by now is that my recycle gas molecular weight is 4 and N2 is 28, so my centrifugal compresor could not be able to increase the pressure more than 250 psi (aprox). what should I take in account?. Is there any gain doing this? (4)
21/04/2015 Q: I would like to ask about required H2/HC ratio and coker naphtha processing in a naphtha hydrotreater.
We have a unit processing a mix of straight run and coker light naptha. Unit consists of two reactors, one for diolefin saturation and one for HDS and olefin saturation, both use regenerated NiMo catalyst. Colleagues intend to raise coker naphtha ratio, which is currently maximized in 12%. I made some calculations which resulted, that coker naphtha has around 90 Nm3/m3 chemical H2 consumption, and the units H2/HC ratio is around 60-100 Nm3/m3 depending on throughput. 12% naphtha results in ~16 Nm3/m3 chemical H2 consumption. If I remember well, the H2/HC ratio should be at least 5 times the chemical consumption, in this case 5*16=80. Am I right, or can this value safely be reduced? Does anything else restrict the max ratio of coker naptha processing? Temperature raise is about 25-28 °C on HDS reactor with an inlet temperature of 290 °C. 		(4)
21/01/2015 Q: We have a kettle type reboiler with weir plate and its liquid outlet is going regenerator in benfield process of co2 removal in ammonia plant. Is it necessary or required to put vortex breaker in the reboiler liquid outlet? Which design is applicable for designod reboiler? (1)
22/12/2014 Q: What happens to the catalyst if water goes to naphtha or diesel hydrotreater reactor along with feed which is having Nickel molybdenum catalyst. (2)
19/11/2014 Q: What are the Pros & Cons in case of Hot start-up of Hydrogen generation unit? Why it is generally not preferred and also why is there no detailed procedure given in operating manual ? (1)
19/11/2014 Q: I am working in Hydrogen generation unit. Our naphtha vaporiser in HDS section got fouled frequently. What shall be the reason behind choking of naphtha vaporiser? (1)
11/11/2014 Q: What is typical specific natural gas consumption [(NG Feed+ NG Fuel)/Hydrogen; wt basis] for hydrogen generation unit?
 		(1)
08/11/2014 Q: I am working in an HGU unit. I want to know if olefins or unsaturated compound increases, what will happen in prereformer catalyst.


 		(2)
14/10/2014 Q: I am working in Hydrogen generation unit. In hydrogen export line one Low point drain flange caught fire due to minor leakage of hydrogen but no source of ignition was there. We could not find reason why auto ignition happens without source. If anybody know give some reason. (5)
25/09/2014 Q: Can we process FCC's Clarified oil (CLO) or Decant oil as feed to Hydro cracker? My question is that Unconverted oil from Hydro cracker is usually good feed to FCC, So I would like to know if we process FCC CLO in hydrocracker then how much of it will it to convert to Unconverted oil in Hydrocracker? We will use filters to reduce catalyst content in CLO so that hydro treater won't get affected. (2)
24/07/2014 Q: I am working in Hydrogen generation unit. I want to know whether if naphtha preheater tubes got a leak and super heated HP steam went to naphtha side then would superheated HP steam go to hydrogenerator (Co-Mo catalyst)? What is the effect of steam on Co-Mo catalyst life? (4)
11/07/2014 Q: In our Hydrogen Generation Unit HP steam silica level is running high at about 0.1 PPM against design value of 0.045 PPM. We maintain BFW Ph-9.5, excess Phosphate - 3 PPM, Hydrazine excess 0.1 PPM and continuous blow down Gestra valve is 100% open. Conductivity and TDS is normal. How can we reduce silica? There are no BFW exchanger leakages
 		(2)
27/06/2014 Q: Obj: Increasing recovery of H2 by slightly compromising on H2 purity (from 99.99 to 99.90)
Present status: In one of our refineries, the recovery of H2 through pressure swing absorption is around 89.5%. Purity obtained in 99.99% whereas 99.90% is sufficient. 		(3)
25/06/2014 Q: Concern: Trade off between purity and recovery of H2 from PSA
Obj: Increasing recovery of H2 on slightly compromise on purity (99.99 to 99.90)
Present Status: In our one of Refineries, the recovery of H2 through PSA is around 89.5%. Purity obtained in 99.99% whereas 99.90% is sufficient. Any changes to the operation still gives 99.99% purity thereby reducing the recovery to almost 89-89.5% 		(1)
13/04/2014 Q: We have a Steam methane reformer having side fired self respiratory burners. To attain the correct O2 in flue gas of primary reformer, burner dampers are being adjusted. What is the correct sequence for throttling the burners? Should the bottom most burners should be throttled more than the top ones or vice versa?  
11/04/2014 Q: In case of side fired self respiratory burners in reformers what is the correct sequence of adjusting the air?
From bottom row burners to top row burners in increasing trends:
in 1st row 40%, 2nd row 40%, 3rd row 40%, 4th row 30%, 5th row 30% & 6th row 30%
OR
in 1st row 30%, 2nd row 30%, 3rd row 30%, 4th row 40%, 5th row 40% & 6th row 40%.
This flue gas is going to convection section for heat recovery. 		 
13/01/2014 Q: We have steam methane reformer. The outer surface of the tube is having deposits and leading to high fuel consumption as well as high temperature in flue gas side in waste heat section. During turn around we want to clean the outer surface of reformer catalyst tube so that we can reduce the fuel consumption and reduction in waste heat section temperatures. Is there any standard method available to clean the outer surface of the tubes? (5)
18/09/2013 Q: How can I calculate the optimal velocity in furnace tubing? At our gasoil/kero hydrotreater we operate usually at low throughput, but we keep the recycle gas at a higher value than needed for the reaction, to prevent the coking of furnace tubes. I guess that the optimal recycle gas amount could be calculated, but I don't know how to do it.

Some additional info: It's the unit manager's explanation that he doesn't want to decrease recycle gas to prevent heater coking. We are usually running on low throughput with 4-500 Nm3/m3 H2/CH ratio. In the last cycle we had pressure drop problems on our reactor, we found solid deposit on top of the bed. We performed a furnace coke burning process during the last turnaround, and found that there was some significant coking in the furnace. Our licensors suggestion is, that H2/CH ratio should be approx. 5 times the H2 consumption. Based on this, 100-200 Nm3/m3 would be enough, but we are running often at 400-500 ratio, which is way higher than suggested. 		(3)
29/07/2013 Q: We are facing the problem of almost all flange leakage in reformer inlet after shutdown (normal or emergency) plant licensor is Haldor Topsoe (4)
26/03/2013 Q: We are going to commission a new refinery which includes Hydrogen Unit also. The Naphtha feed specication for the Hydrogen unit is <1ppm Sulfur and boiling range of IBP-95 degC. But during commissioning we cannot suppy the above naphtha spec. So it is agreed to supply naphtha with boiling range of IBP-160 degC and Sulfur <150 ppm for initial one month. The Feed Desulphurised catalyst vendor says it cannot handle beyond 30 ppm Sulfur. Can anyone share such experience and can advise how to manage the situation. Also what will be the Hydogen yield with changed specfication of feed Naphtha? (6)
17/02/2013 Q: Please advice about the operation practice of the reciprocating compressor in order to increase life time of discharge valves in the compressor cylinders.
We're operating a Hydrogen plant using water electolyser cells - using alkaline which is KOH, operating temp 70C and production rate 170 nm3/hr where reciprocating machines are used to suck hydrogen from a gas holder directly after the cells and discharge it to the high pressure line to the customer (200 bar).
The Gas Compressor comprises of 5 stages, where first stage is consits of 2 plate valves.
The problem always happens to first stage Gas Compressor. Just 15~30 days after starting, the discharge valves in the cylinder always develops a leak and fails. Upon dismantling, we used to find some debris and gum particle (deposits) in between valve plate and valves top. The problem occurs over and over again.
The deposits particle is KOH which comes with the hydrogen gas vapor passing through the suction filter and rest on the valve.
(2)
12/11/2012 Q: What exactly potash fixation means for reformer catalyst and does over steaming have some impact on catalyst performance? What factors will lead to leaching of potash from reformer catalyst? (2)
26/10/2012 Q: We have hydrogen reciprocating four cylinder compressor. Every time after over hauling we take No Load trial for 2 hrs with discharge valve removed condition. Then we purge with nitrogen, then twice with Hydrogen. Now During Load trial Motor trip on Over load, (0% capacity), after barring compressor around 2 rotation, the compressor started Normal.
Why compressor started smoothly after free rotation? We already taken No load trial one day before. 		(2)
31/07/2012 Q: Want to know Approach to Equilibrium calculation for the naphtha steam reformer of refinery having reformer exit design methane slip of 2.85 mol% dry basis. I have information on calculation of approach to Equilibrium (ATE) if wet base composition at reformer and shift converter is available. In this case steam/gas ratio can be calculated directly based on moles of H2O available. Normally this is not available from lab. They are giving dry analysis at reformer and shift converter outlet. Based on reformer inlet feed flow, steam flow, C/H of feed, recycle gas ratio, inlet pressure, PDI accross reformer and dry composition at exit of reformer and shift converter; is it possible to calculate ATE. If anyone have develop such formula/corelations please share with me so that I can know the catalyst activity and present ATE of reformer and HT shift converter. (1)
24/04/2012 Q: Why give Top Fire Burners in Reformer in some H2 Plants? What is advantage of that? (3)
16/03/2012 Q: if we want to reduce the Hydrogen purity in DIESEL HYDROTREATER the current H2 purity is 99.99 fro PSA Unit now we want to take from other plant (Rheniformer units),During hydrogen plant turnaround, PSA’s are not in operation and only the off gas from Rheniformer units, low purity hydrogen, is available. This make-up gas can be used as hydrogen for the DHT to keep it running for the duration of the whole hydrogen plant outage.
the hydrogen from PSA :
HYDROGEN ------> 99.99
C1------> 0.1
CO + CO2 -----> 20 MAX
H2S ----> ZERO
HCL < 1
***** new hydrogen make-up with a reduced purity, coming from Rheniformer units
H2 87.5
C1 5.8
C2 3.6
C3 1.6
C4 0.4
C5+ 1.1
- what is the side effect of low purity for all the plants, recycle compressor, make up compressor and the load in addition the make up it is suitable for for such low purity?
What is the impact on
-Reaction section
-Product quality
-Recycle compressor?






 		(3)
12/12/2011 Q: We use two block valves with one blind for isolation at boundary limit of each process unit in our refinery. Gate valve is selected for block valve mainly. For hydrogen system, we select one ball valve (Orbit Valve) installed at main header side, and one gate valve at process side for block valve service. Would you please advice if Rising Stem Ball Valve is better for hydrogen system, and what condition should be used? (1)
06/12/2011 Q: In my refinery we have Diesel Hydrotreater unit which has undergone an emergency shutdown, following the tripping of recycle gas compressor. The compressor was tripped off due to the false ESD activation of high level switch on the suction K.O drum. Recycle compressor started up and started increasing the reactor inlet temperature . During the DHT unit start up, higher hydrogen consumption of 1.5 vs normal 1.0 MMSCFH and lower system pressure of 620 vs normal of 670 PSIG was observed.
My question is where is go the higher hydrogen consumption?
 		(4)
22/09/2011 Q: Could someone explain the significance of the H2 to hydrocarbon ratio in Naphtha reactors please? What is the effect on reactions when you increase / decrease the ratio and are there any other effects?  
20/06/2011 Q: We need to revamp our NHT. Before revamp: 23500 bpd SR Naphtha 100ppm S, Naphtha product for CCR feed has 0.5 wt ppm. After revamp: 30 000 bpd (90% vol SR Naphtha, 10 % coker Naphtha) with 0.1 wt ppm in product for our new regulation. We have 1 reactor (R1) with 1 bed of catalyst (18m3 catalyst in 27m3 reactor). I think we should install one more reactor. But I don't know which case is better between: Case 1: Feed-R1-R2-Stripper-splitter and Case 2: Feed-R1-Stripper-Splitter-R2 (recycle bottom product from splitter to R2)-R1. May you have any advice for our revamp?

Additional info:
Of course that Case 1 is traditional process revamp. But I have just read an article from Chevron, about their process revamp as Case 2. It called SSRS Isocraking (single stage reverse sequencing), licensed by Chevron Lummus Global. In that article, they said that the revamped unit can run at 133% of original design capacity with the existing recycle gas compressor. I think in case 2, R2 is existent reactor and R1 is new one (because R1's volume needs to be bigger than R2) This article named "Hydroprocessing upgrades to meet changing fuels requirement", Jay Parekh and Harjeet Virdi. Unfortunately, It's not for NHT, It's Hydrocracking. Is it O.K if I use Case 2 for my NHT revamp?
 		(9)
11/05/2011 Q: What is the standard value of sox/nox in atmosphere if emitting from hydrogen generation unit reformer for fg/naphtha/off gas firing? (1)
11/05/2011 Q: Our de-aerator conductivity is running high while de-aeration pressure is 0.3 kg/cm2g and temperature is 107 to 110 degree centigrade. Any thoughts on reasons and solutions? (3)
24/03/2011 Q: Does a prereformer require an automated bypass? Some catalyst vendors insist on auto bypass; some say it is not required. What is best practice? (2)
23/03/2011 Q: How do you calculate steam-to-carbon ratio in H.G.U.? (1)
03/11/2010 Q: What are the likely benefits of processing Methanol in shift reactors of existing NG based hydrogen Plants? What modifications would be required..?  
09/10/2010 Q: My company aims at further processing the atm. distillation residue (Mazot); and a hydrocracker unit has been chosen for this task. We need to estimate the cost of the unit and its facilities like the vacuum tower and the vis-breaker. How would you suggest we get a rough initial estimate of the costs involved? (5)
11/07/2010 Q: There are 3 hydrogen plants in our refinery. All of them are steam reforming process. We use OG, LPG and naphtha as feed to produce high purity hydrogen. The raw hydrogen stream comes from steam reformer, after cooling down then sent to PSA to recover high purity hydrogen.
The design recovery efficiency of PSA is 89%, we found the actual efficiency at high throughput condition is 85% or less only. This is a bottleneck for hydrogen plants.
There are some comments from outsources. Someone said the operation life of molecular sieve used in PSA is very long, 10 years or more, we don’t need to replace it. But another had opposite comments. Would you please advise. 		(6)
03/07/2010 Q: What are the benefits of adding process steam in pre reformer inlet and reformer inlet separately? In some hydrogen plant it is mixed only in reformer inlet. What is the advantage of that? (2)
13/05/2010 Q: What are the benefits of a top fired reformer versus a sided fired one? (5)
03/04/2010 Q: Kindly throw some light on catalyst loading in Reformer tubes by Spiral loading technique. Does this technique provide much stable distribution of catalyst and improvement in any process performance like uniform tube temperature, pressure drop across each tube constant, and reduced loading time? (2)
08/03/2010 Q: In our Once Through Hydrocracker Unit, the Recycle Gas Compressor is surging from 100% opening of the anti-surge valve to 0% without any change in process parameters. It was also observed that just prior to surging the total flow at the inlet of the RGC was also increasing. We have got an amine column at the inlet of RGC suction after HP separator to reduce sulphur loading. But now due to some constraints the amine flow had to be reduced. Can anybody explain the phenomenon? (3)
26/12/2009 Q: We have our Hydrogen generation unit installed. Compared to all other unit furnaces/heaters, the reformer of HGU is top fired. Why it is so? (2)
15/11/2009 Q: What is PRD mode in automatic process control? (5)
01/11/2009 Q: What is main purpose of putting sealing steam in a turbine? (1)
01/11/2009 Q: In our DHDT recycle gas compressor primary seal vent flow at non driver end side has reduced to zero while it was previously 5 Nm3/hour. Driver end side flow is running between 30 Nm3/hour. What is the possible reason behind flow reduction? (1)
23/10/2009 Q: We have 4 hydrogen gas cigars (reservoirs). On the inlet and delivery line there are valves which stock is limited. Now we want to buy some new valves that match the following service:
operating pressure: 70 to 80 bar
design pressure: 130 bar
operating temperature: 41 degree Celsius
design temperature: 80 degree Celsius
The valve will be used for both sides operation. Can anyone help me by informing what kind of valve should be used in this service and preferably the name of valve manufacturer? 		 
12/08/2009 Q: In DHDT unit suppose benzene converted to cyclohexane and then cyclohexane converted to normal hexane. What is the mechanism of this reaction? How is aromatic converted to cyclohexane then how cyclohexane ring broken and converted to n-hexane? (3)
31/07/2009 Q: What are the changes required when a hydrogen generation plant feed which is originally designed for naphtha is to be run on natural gas (as a feed)? (4)
21/07/2009 Q: Why must we maintain distillation of diesel 95% at 360 degrees centigrade for Euro-III ? If less or more what is the effect on engine performance? (2)
22/06/2009 Q: How can we establish the impact of high temperature water vapour on the compressor valve sealing element and its possible contribution to the melting of the element? (1)
04/06/2009 Q: In what situation is a pneumatic test at one kg/cm2 to be preferred to a hydro test at the design pressure of a vessel? (2)
28/03/2009 Q: What is lube oil supply temperature for any pump or compressor? Like feed pump, makeup gas and recycle gas compressor. (2)
25/03/2009 Q: Why do we need to maintain gas oil ratio in our diesel hydrotreater? (5)
19/03/2009 Q: With some experts projecting crude prices to creep back up to $75/bbl by mid-summer 2009, should we expect to see a higher level of refinery intermediates (e.g., heavy gas oil, "lifted" DAO, etc.) being exchanged among "networked" refining facilities?  
17/03/2009 Q: Are the declining costs of metallurgy providing an incentive for construction of 2000+ ton heavy-walled hydrocracking reactors? Is the application of advanced manufacturing techniques, such as Cr-Mo vanadium welding, becoming the 'norm' for fabrication of heavy walled hydrocracking reactors? What other developments coincide with new hydrocrackers designed to operate in a highly corrosive environment? (1)
09/03/2009 Q: Why is a minimum circulation line not provided in some centrifugal pumps? For instance, in our stripper reflux pump it is provided, while in our diesel hydrotreater stripper it is not. (2)
04/03/2009 Q: What is the exact meaning high/low severity in case of refinery catalytic unit? (5)
15/02/2009 Q: Why is the cetane index of diesel higher for high sulfur than low sulfur crude? (6)
15/02/2009 Q: What is the mechanism of aromatic saturation reaction in diesel hydrotreater reactor (i.e. step by step conversion from aromatic to paraffins)? (2)
03/02/2009 Q: Can anyone reference an article or research that comments on the effect lubricating oil from the makeup or recycle H2 compressors can have on catalyst life? (5)
16/10/2008 Q: Why is the non return valve fitted on the horizontal pipe line rather than the vertical one? (2)
07/08/2008 Q: In addressing refinery CO2 management, can you comment on CO2 curtailment from on-purpose hydrogen plants through "minimised" involuntary'steam, internal heat recycle and captive integration?
 
26/07/2008 Q: In a hydrogen plant is it possible to use wash water instead of Amin guard to control PH of reflux system? (1)
24/07/2008 Q: We are building a grassroots refinery. In refinery we have catalytic reforming unit and light paraffin isomerisation unit for gasoline pool. During startup of isomerisation unit it will require initial dryout and Acid cleaning to remove water and iron rust.
Has anyone experience of initial dry out and acid cleaning of isomerisation unit? How long it will take for initial dry out and acid cleaning? 		 
02/06/2008 Q: For certain standards pertaining to control valves used in hydrogen services, why is it recommended that installation of a bypass (and blockvalve) be avoided? (1)
19/05/2008 Q: What happens if a steam reformer heater (hydrogen unit) is only fed by steam for a long time in stand by mode? Is this action harmful for catalysts?  
19/05/2008 Q: Is it possible to run a terrace wall reformer heater only with one cell? (Heater has two separate cells in west side and east side and feed, steam and fuel gas is split for both cells)  
01/05/2008 Q: What are the conditions leading to brine production in a Catalyst cooler?  
04/03/2008 Q: We have a problem with our Hydrocracker VGO feed filters resulting in frequent backwash operations due to high Del P. Can you please ascertain the reason for the same as we do not get any FeS or suspended solids in the backwash stream analysis. Is it because of the asphaltenes as we process deep cut VGO (360-580+ degC) along with Heavy gas oil? (8)
12/02/2008 Q: Quite a large amount of Hydrogen is consumed in desulphurisation of fuels and hydrotreatments for product quality improvement which generate Hydrogen sulphide. A more economic process is required like catalytic decomposition of hydrogen sulphide into hydrogen and sulphur and the separation of the products of said decomposition to H2 and Elemental Sulphur. This would enable recovery of costly hydrogen and same can be re-utilised in the process of treatment. Are there any catalyst development taking place for such purposes? (1)
18/01/2008 Q: Glycerol is produced as by product in the Transesterification process for Biodiesel. Please give your views on the following:
1. What are the present practices for handling this glycerol?
2. Glycerol can be converted to Hydrogen. Have processes and catalysts been developed?
3. Does the above process for Glycerol to Hydrogen require any treatment of glycerol.
4. Any views/suggestions on the handling of glycerol? 		(2)
14/10/2007 Q: in our refinery we treat the LPG produced from the FCC unit by extraction merox unit.
In the pretreatment to remove the H2S from LPG, the absorber shows low efficiency. What is the problem?
(the abs. press. 10.5 bar
amine conc. % vol.= 19
regenerator good efficincy
H2S in rich Amine = 0.034 WT%) 		(2)
19/09/2007 Q: Please advise on reduction of ammonia emissions from a fertiliser plant.
Our emissions from a urea plant stack is about 150 ppm, and we need to reduce them to 50 pp to comply with EPA regulations. I know some plants are provided with an acid washing system.
I would be grateful for advice from anyone with experience in this field. 		(1)
16/09/2007 Q: What is the best method for hydrogen management in the refinery?
can you provide a brief description of this method, an example of a
refinery using and what they've gained from it. 		(2)
15/09/2007 Q: What processes are available for
1. the separation of oil from slack wax
2. the separation of wax fom residue wax
3. the hydrogenation of wax? 		(1)
06/09/2007 Q: In the off gases from our vacuum distillation column hydrogen % has been up to 30-35% by volume.This vacuum unit is mild severity dry distillation with designed VGO end point of 510 deg C.
The overhead boot water PH also remains on the lower side (~5) even though the neutraliser is added in large quantities (more than 100 ppm). The same neutraliser has used earlier for the same type of crudes.
Has anyone had this type of experience? What may be the reason for the same? 		(1)
05/09/2007 Q: What is the best method for return pure hydrogen from H.P off gas of hydrocracker unit to input of this unit (pure H2 gas)?
Do you think the membranes system is the best and economic for this application? 		(4)
23/07/2007 Q: What are the more attractive isomerisation configurations and catalysts available to meet the growing demand for light paraffin isomerisation? What can be done to lower the equipment cost, such as the recycle hydrogen compressor? (4)
22/07/2007 Q: What potential opportunities are available for gasification of refinery residues? (1)