Q & A > Catalysts and Additives
Date  Replies
14/05/2017 Q: We are using AXENS Catalyst RG-682 (Naphtha Reforming) and HR-538 (Naphtha Hydrotreating).
Now, my question is what is the actual life time of these two Catalyst? How many times it can possible to regenerate RG-682 & HR-538?
(5)
02/05/2017 Q: How can I calculate Reforming Heat of reaction and reactor Delta T. From the catalytic reaction guideline I know that the Napthene dehydrogenetion heat of reaction -50 Kcal/mole. Now I want to calculate reactor delta T.
Additionally I know the reformer feed flow rate, feed detail hydrocarbon analysis, feed density, feed molecular weight
In practical operation, we have three reactor in series, 1st reactor delta T 117F, 2nd reactor delta T 48F and 3rd reactor delta T 16 F; Now I want to calculate this delta T in theoretically. How can I prove/calculate that this practical delta T as like theoretical?
(2)
26/04/2017 Q: I have been requested to benchmark our estimate of process chemical costs in current refinery project. Refinery configuration with middle-east crudes processing, and bottom upgrading with ARDS and RFCC.
Crude capacity about 250 kbd.
Process chemicals cost include CDU demulsifiers, neutralizing amine, corrosion inhibitors, antifoaming etc, i.e. those used in overall refinery complex.
Can anyone share any similar costs, at least per barrel of crudes processed?
(2)
15/02/2017 Q: What is the range of UOP R56 catalyst density? Is it ok to assume that the density of Fresh catalyst (R56) is the same with the Spent Catalyst when calculating the weight of the catalyst? Is it possible to overload reactors (3) in a (vertical) stack with catalyst in CRU?  
04/02/2017 Q: We have a semi regen platformer unit and catalyst type UOP-R56. During the regenerating after the oxidation step we stopped at this step The problem of providing electricity occurred and we did not complete the rest of the steps, now the reactors under positive pressure by hydrogen reformer purity of H2 70%
Is this situation adversely affects the catalyst after the oxidation step?
And what needs to be done in this case?
(2)
04/02/2017 Q: We use straight run Heavy Naphtha from Tank which has floating roof. Now, my question is what should be the limit of moisture or water content in wt ppm level before entering Naphtha Hydro Treating Catalyst bed. Currently, we have 104 wt ppm moisture in HN feed can this quantity could be the cause to increase the pressure at the inlet of Reactor suddenly to create a peak of DP across the reactor bed? Currently, DP across the catalyst bed is 3 and it has reached gradually after the 3 years of operation cycle but become confuse about the sudden peak. So, please help. (5)
03/02/2017 Q: What are the key tests with specifications for Alumina or Ceramic support balls for catalyst used in refinery ? What maximum/minimum values of Apparent porosity or Water absorption should be the specification for good inert support balls to be used in Sour Shift units of Gasification complex? (1)
26/01/2017 Q: Current DP across the Naphtha Hydro-treating Reactor is 3 and it is giving peak frequently without giving a gradual increase of DP. Why is the pressure at the inlet of Reactor is suddenly increasing for the few second but coming down to previous position again? However, we are thinking Catalyst top bed is blocked by coke, iron scale and other metal contaminant. Besides, we are getting some Ash color and salt type dirt at the fine mesh filter at the suction of NHT compressor. So, i need to know what could be the probable composition of the dirt? For this, we assumed that it could be fouling product from the separator drum or dust of alumina ball from the bottom of Catalyst bed. Now we are facing frequent compressor changeover and cleaning of suction strainer. So, please share with me if anybody have similar kind of experience and help me to sort out what could be the composition of dirt. (3)
04/01/2017 Q: In my refinery there is a 15 kBPSD LSRG sweetening unit in which caustic washing procedure followed by MEROX oxidation process. In case of feed change scenario, is there any solution in terms of gas condensate sweetening by means of before mentioned facilities? If yes, what are the changes in terms of capacity, chemical consumption, and mercaptan removal efficiency? If there is any revamp, which sections need to be resized? (2)
27/12/2016 Q: We have found black solid deposits upon cleaning of our CCRU Net Gas Compressor First Stage Strainer. Upon analysis of composition, we have found that the sample contains hydrocarbon plus a significant amount of Chloride and Iron and with traces of Aluminium, Magnesium, Silicon, Phosphorous and Sulfur. What could be the source of these black solid deposits? (6)
28/11/2016 Q: Does anyone have any experience with catalyst fluidization in oxychlorination zone of Regeneration tower in continuous catalyst regeneration unit? We have experienced high catalyst dumping recently after this unit shutdown. This catalyst dumping contains high catalyst dust. Is there any parameter to check whether is the catalyst fluidized or broken inside the Regeneration tower?
Thank you very much.
(2)
25/11/2016 Q: We have a semi regen platformer unit. After regenerating the catalyst, unit was started up and observed that reactor 1 delta temperature (we have 3 reactors ) is low and almost same as reactor 2 delta temperature. And the Platformate RON is also low at about 92 whereas the expected RON at start of run is about 94. This is the first time we experience this kind of a behaviour. There was no sulfur ingress or any other change in the feed stock (Heavy naphtha). The issue seems to be in reactor 1 and other two reactor delta temperatures are similar to the previous cycles. Can anyone help me on possible causes for this issue?

Additional info:
Appreciate the valuable comments. To add some more information, We do process only the virgin naphtha. Sulfur and nitrogen are less than 0.5 ppm as required. Feed is from Murban crude which does not have high metal content.However, it appears that the reactor 1 delta temp is in decreasing trend and reactor 2 delta temp has increased slightly. H2 purity also low and gases and lpg are on higher side. PONA test indicated that paraffin conversion is low compared to previous cycle SOR conditions. Catalyst samples were checked after oxidation and reduction steps. Appearance of reactor 1 samples were better compared to reactor 3 except the reddish rust coating in reactor 1 was high. But, this coating in reactor 1 was same as in previous cycles samples as well. Also, we are maintain little higher cl level as this is the 9th cycle of the catalyst and expect lower cl retention in catalyst.

New info:
By today, the reactor 1 delta temp has further decreased and reactor 2 also is seems to be in decreasing trend. A new information is that there was a upset in recycle gas compressor during regeneration (oxidation step). Seal oil loss increase due to seal damage and both the seals were replaced along with damaged bearings. But, labyrinths haven't changed. And we observed some seal oil in compressor discharge line ( some drops of seal oil leaking at the compressor discharge isolation valve flange). Therefore, we checked at casing drain line and another low point in discharge line. First day (two days ago), we found some significant amount of oil collected. After draining them, the next two days found only small collection. Anyhow, now we are doubtful about a lub oil contamination. In addition, can there be a possibility of channeling the first reactor? Because if so we have to dump the catalyst during regeneration. Another thing is that during regeneration, there was a strong detection of hcl in reactor 1 outlet ( more than 500 ppm).
(10)
18/10/2016 Q: In our FCC (Deep Catalytic Cracking) Unit Spent catalyst slide valve(SCSV) DP suddenly decreased from 0.70 to 0.35 kg/cm2. All parameter such as aeration purge point pressure and flow, all types of steam flow, reactor and re-generator DP are normal. What is the reason for this? (5)
28/09/2016 Q: What is the reason behind DMDS injection in CCR Platforming Unit ? (5)
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.



(9)
27/06/2016 Q: In case of heavy residue upgrading, we are encountered with vacuum residue as feed. The main features of this feed especially about contaminants and problematic materials are as below:
Total sulfur>4.5 wt%
Conradson Carbon >25 wt%
Ni+V >500 ppmwt
Nitrogen ~ 1 wt%
We have two cases for VR upgrading project, One is RCD+RFCC and another is HOIL(Hydrocracking)+FCC. Both of these cases use huge amount of fresh catalysts because of high possibility of catalyst deactivation and poisoning. So the operating cost should be high.
Is this rational to charge such a feed to the catalytic system directly or is it better to use the process to somehow get rid of metals at least? If we need to use the solvent deasphalting system at the upstream of two before-mentioned cases and draw off about 20% of feed as pitch, we will succeed to lower the operating cost and increase the reliability of catalytic system because of the elimination of the major part of the metals. But in the opposite side, we have missed 20% of primary feed as pitch that it is a low value product. So the profit margin of the residue upgrading cases will decrease. However, as a second question, can we miss 20% of feed charge at the expense of increment of catalyst life cycle?
(5)
26/06/2016 Q: How do we calculate the weight percent for NH4CL in the stream? (1)
11/05/2016 Q: What is the difference between I-8 and I-82 catalyst for Penex Reactor? (2)
25/04/2016 Q: We are looking at alternative option(s) that could expedite the unloading of residue desulfurization unit catalyst (from reactors) other than typical vacuum-out/jack hammering approach.
We have heard about the CO2 explosive technique - and just wondering if anyone has any success stories with that?
Any other feasible approach to be explored?
(4)
19/02/2016 Q: What are the techniques used in CCR unit to reduce catalyst dust while draining catalyst fines? (1)
08/01/2016 Q: In our semi regenerative catalytic reforming unit we are facing problems of low recycle gas purity.70-72%. Octane has been in the range 92-93 as per design. Reformate yield is around 83% against design of 86.%.Our stabiliser off gas production has been incresed by 3-5 % and reflux control valve has been opened fully. Ratio of C1+C2/C3+C4 is higher 1.9. PCE dosing @1.5 ppm (8)
16/11/2015 Q: What is the range of UOP R134 catalyst density? Is it ok to assume that the density of Fresh catalyst (R134) is the same with the Spent Catalyst when calculating the weight of the catalyst? Is it possible to overload reactors (3) in a (vertical) stack with catalyst in CRU? (3)
18/10/2015 Q: We are looking for a non hydrotreating based technology to decrease condensate sulfur content to lower than 200 ppm. There is a condensate stream in our refinery in which its sulfur content decreases from 3300 ppmw to 1000 ppmw by caustic wash and we need a further decrease of sulfur content to minus 200 ppm, but not with hydro desulfurization. Please advise. (5)
25/09/2015 Q: What is maximum limit of Iron in the Platformer Feed?
What is the effect on Platformer catalyst like R56 or R86?
How can we control Iron in the feed with suitable filters?
(2)
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)
29/07/2015 Q: For VGO Hydrotreating units at the upstream of the FCC, we currently use COMO Catalyst. The unit is operated by maintaining constant Sulfur spec in the sweet VGO which goes to FCC. Typically VGO feed sulfur is ~22000wtPPM and Nitrogen is ~2200 wtPPM. After processing in hydrotreater, sweet VGO sulfur is ~1500wtPPM and Nitrogen is ~1000wtPPM. Now the question is that if I change the catalyst from COMO to NiMO catalyst and maintain remaining all operating parameters same, what would be the Nitrogen conversion if I operate the unit by maintaining Sulfur level 1500wtPPM which is same as earlier? Will the Nitrogen conversion increases because of NiMo catalyst or it remains same since we are constraining the unit severity by maintaining same sulfur level? (4)
13/07/2015 Q: What is the meaning of regeneration cycle duration in ccr refromer? And how could the catalyst circulation rate be about 200 and the burning capacity is a round 12 ?  
10/07/2015 Q: We plan to purchase regenerated catalyst for our kero and LGO hydrotreater. We did regenerate our own catalyst, but never purchased one from an external company. What parameters should I pay attention to, what are the recommended limits for poisons and other parameters to guarantee a near-fresh activity and lifetime? (3)
23/06/2015 Q: In HDS reactor the hydrogen consumption coming down when catalyst life moving to end of run ,even though the feed and product sulphur remains the same.( usually the temperature of BED have to increase for achieving the same product quality). What is the reason for this hydrogen consumption reduction? (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)
30/05/2015 Q: We have hydrotherapy unit , consisting of cobalt molybedium (s-7 and s-120) reactor, reaction temperature 610 F system pressure 24 bar.We have a problem for two months that is the reflux drum of stripper got very low thickness observed, its boot water has PH 2.0----2.5, iron greater than 100 ppm while chloride was 1000 to 2000
condensate injection 8bbls/h from condenser inlet.
We have already done the cleaning of all heat exchanger , overhead condenser, overhead reflux drum.
Then start up of the unit was performed but condition remains the same.
Please share your opinion regarding this problem.
(2)
11/05/2015 Q: I would like to know the proper actions to be done when there is a loss of PCE injection in the catalyst regenerator of UOP CycleMax CCR Unit. If for instance, the loss of PCE injection (both injection in feed and in catalyst regenerator) would not be addressed immediately and PCE injection would not be available for an 8-hour duration or more, to avoid platinum agglomeration or other cases, will it be better to run the unit in hot shutdown/cold shutdown, or can the catalyst endure the loss of PCE injection, and be able to normalize once injection is resumed upon availability.

(2)
29/04/2015 Q: In a CCR unit we are observing coke type material seeping out from reformate stabilizer column bottom pump suction flange. Although we are dosing DMDS in the feed but can MCC be a reason for this? Else what are the possibilities?  
23/02/2015 Q: We are trying to determine the appropriate lab test and normal analytical ranges in order to bring imported MVGO to our new hydrocracking unit. Licensor is concerned is about presence of Na and Cl but also other contaminants such as P. What are the normal ranges of Na, Cl and other metals to bring to the hydrocracking unit to avoid catalyst damage? (2)
07/01/2015 Q: I have looked many databooks for some alkanes standard entropy of ideal gas at 298.15K, but the standard entropy values of the following components was difficult to find. I'd be very grateful if anyone can scan their databook pages for me. thanks ahead.
3methylhexane, 2,2dimethylpentane, 2,4dimethylpentane, 3,3dimethylpentane, 2,2,3-trimethylbutane
(2)
30/12/2014 Q: This question is related to kerosene merox unit. After processing kerosene in merox unit, what are the main reasons for poor saybolt color of kerosene product? If kerosene feed to the merox unit has saybolt color of +26, kerosene product from merox unit observes saybolt color of <16. Can someone explain the possible compounds which causes color problems to the kerosene product? If we go to Kerosene Hydrotreater, there will not be any issues of color problems and in fact it will be improved because of olefin and aromatic saturation. Please share any literature or chemistry related to the kerosene color problems in merox units. (3)
24/12/2014 Q: We need to know about dewaxing catalyst or any other catalyst which can be used to reduce the
wax obtained during pyrolysis process of plastic to fuel oil conversion. I need the guidance if any
one can inform me.
 
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)
18/12/2014 Q: I'm doing a work about octane rating, but I haven't find the Research Octane Numbers of cis-1,2-dimethylcyclopentane and trans-1,2-dimethylcyclopentane. Who can tell me? (3)
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)
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)
07/11/2014 Q: I'm going to implement APC in a FCCU soon. What's the best source of information to learn the complete (even the minute) details of FCCU so as to complete it successfully? (1)
22/10/2014 Q: Please anyone who can help with information on the use of n-hexane , thinners otrso isoparC and peroxides in the polymerization reaction of LDPE high pressure ( 20,000 psig ) autoclave reactor type (1)
03/09/2014 Q: We found a high TAN, ca. 0,4 mgKOH/g (usually 0,1), on a LCO cut.
What could be the explanation?
(3)
21/08/2014 Q: In a team discussion about the start up sequence of Naphtha Catalytic Reformer, everyone was wondering about the effect of prolonging the hot hydrogen circulation across the catalyst bed more than 12 hours.
The question rose from the fact that usually the Stabilizer (Debutanizer) Tower in the Reformer Unit is started parallel with reactor heating up, in such a way that when Reactor has reached the required temperature for charge in the liquid feed, the Stabilizer Tower has been ready to strip out the light ends. But it's not seldom that Stabilizer Tower is suffering from un-predicted prolonged problem --- such as very frequent bottom pumps' strainer blockage-- while reactor inlet temperature has reached the feed cut in temperature.
Under such situation, the start up team was in the pro-cons whether to keep the hot hydrogen circulating across the reactor for few more hours till the readiness of Stabilizer Tower, or to immediately cool down the reactor loop. The first opinion merely consider about the time efficiency, while the second group worried that alumina support of the catalyst will undergo a phase change due to hydrogen embrittlement on this alumina.
Has anyone here had the similar experience, and can give us more enlightenment on this matter?
(1)
20/08/2014 Q: I am working at UOP Cyclemax CCR Platforming Unit having R-134 Catalyst. We are facing problem of high HCL in recycle gas. Can Anybody who has worked on same CCR Unit with R-134 Catalyst, share his experience?

Additional:
Thanks for responding. 1- Our stripper operation is OK regarding reflux ratio & bottom temperature, but the question is if Organic Chlorides are slipping from stripper, then Chloride level on spent catalyst in Platformer Reactors should also increase with high HCL in recycle gas that is not in our case. 2- We are trying to maintain Chloride level b/w 1.1~1.3 wt% by injecting PCE (perchloroethylene) more than design in oxychlorination zone but result always remained b/w 1.1~1.20 wt% on regenerated catalyst & 0.90~1.0 wt% on spent catalyst. Surface area value of Catalyst about 6 months ago was 152 m2/gm while we have changed our whole catalyst in last Turn Around in March,2012 & total 215 regeneration cycles have passed. 3- We have already changed reduction zone vent from upstream of product separator to upstream of Net Gas Compressor. Waiting for your valuable next response.
(3)
13/08/2014 Q: Recently we have suffered some problems of Cupper Corrosion test failure in LPG. The LPG came from a caustic treatment for mercaptan sulphur removal. After caustic treatment, the LPG pass through a decanter (with NaOH/MEA solution) and sand filter, which are supposed to remove any caustic carryover from LPG. We do not see any caustic collected in the sand filter, however we have detected Na and nitrogen in LPG, so we suspect that it is not working properly. The sand filter seems not only not working, but also accumulating some contaminants: we have seen sometimes that LPG pass the cupper corrosion test in the inlet, but not in the outlet of the sand filter.
We are evaluating the possibility of substituting the sand by any other more effective adsorbent for caustic / nitrogen (amines). The possibilities are: activated carbon, Anthracite or alumina.
Has anyone experience with adsorbents for contaminant (caustic, amine, etc..) removal in LPG? Any idea / recommendation regarding the operation of the sand filter?
(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)
19/07/2014 Q: Molecular Sieve of the Liquid feed Drier of Isomerization Plant was removed for inspection and it was observed that Molecular sieve is blackish in colour.
Is is possible to ascertain the remaining life of Molecular sieve ?
How do we know that pores are not choked or damaged?
(2)
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)
19/04/2014 Q: We have an Axens CCR unit with 4 reactors and I would like to know whether we can maintain different temperature in each reactor? If the answer YES, what are the effect on RON and catalyst? Please share your experience. (2)
15/04/2014 Q: In our CCRU plant we have two net gas compressors which are discharging 38000 Nm3/ hr H2 gas (95 % H2 purity). Net gas compressors are two stage reciprocating compressors with recontacting section. We have 4500 Nm3/hr (87% H2 purity) of semi regenerative CRU off gas joining the circuit after the first stage discharge. We are facing problems with very high second stage suction strainer PDIs in our compressors (which is probably due to CRU off gas joining the circuit interstage). Recently we had conducted analysis of the muck we found on second stage suction valve plates.
The analysis is :
Sr.NO / Parameter / Unit
1 Moisture (@105ºC) % 5.4
2 Loss of Ignition at 800⁰C % 83.14
3 Ash at 800⁰C % 11.46
4 Solubility in water % 13
5 Oil Content % 16.96
6 Iron (Fe) as Fe2O3 % 8.716
7 Acid Insoluble ( ~ Silica etc) % 0.8
Can anyone provide further insights looking at these results?
(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.
 
07/01/2014 Q: Are there correlations available for prediction of coke lay down on CCR catalysts based on Charge rates, N+2 A and operation severity for a CCR Reformer ? (2)
06/01/2014 Q: In the Isomerization of light naphtha process with Pt/chlorinated alumina, can this catalyst regenerate in the unit as insitu? (2)
13/12/2013 Q: Please help us with two below questions:
1. What kinds of reaction will happen when we use ZSM-5 (9.2 angstrom pore size) as catalyst for treating heavy naphtha with range 80 - 180 deg-C?
2. ZSM-5 zeolite is treated by deposite SnCl2 and then calcinating at 450 deg-C in 8 hours?
What products will we receive? Maybe RONC increase?
 
13/12/2013 Q: We designed and built an octanizing system for upgrading octane number of heavy naphtha fraction with 8 cubic meter per hour. This system have no a hydrotreating package, so naphtha feed goes directly through heater and then reactor.
Reactor dimension is 12 m length and 1.2 m diameter. But we use only 2.5 m length for containing catalyst (~1.8 tons). Internals have not inlet diffuser and distributor. Naphtha feed stream enter on the top and exit at bottom.
Heavy naphtha has distillation range 80 - 180 deg-C, RON 68.
Operating conditions are: inlet temperature 400 deg-C and 5 bar-g.
Catalyst is a kind of zeolite with 9 angstrom pore size. It was treated by impregnating in SnCl2 solution and calcinating at 500 deg-C in 8 hours.
We run a pilot as 6 lit per hour on 1 kg of catalyst. Efficient has RON 88 and 8%vol gas.
But when run the reality system, it can not get that target. Catalyst is very easy to deactivate.
Please help me to find some reasons.
(2)
21/08/2013 Q: What can be the cause of coloration (yellowish green) in VGO Raffinate hydrotreater effluent?
(3)
19/08/2013 Q: We are heating desulphurization unit by natural gas (NG) without hydrogen up to 300deg C and being vented in to the flare. Instead of flaring this NG, after coming out out from desulphurization unit can we cool and compress in NG compressor and heat in waste heat section of reformer and feed into desulphuriser unit for heating? Please mention advantages and dis adventages? one of the advantage is vent of NG can be stopped. Any effect on catalyst? Like coke formation etc.

Further info:
We are only using only NG (CH4-93%+ 7% N2) first in NG compressor and then heating into waste heat section to increase the temperature of NG for heating desulphurisation section. After the heating it is being vent into flare. My opinion is instead of venting into flare can we cool it and again in compress and heat in waste heat section and again feed to desulphurization section to increase the temperature up to 300deg C. Any adverse effect on catalyst etc.? With this arrangement we can avoid venting of precious NG in to flare. No hydrogen is added during heating process.
(2)
19/08/2013 Q: We are heating the desulphurisation unit with NG without hydrogen up to 300 deg. C and then NG is being vented through the flare. We want recycle this NG by cooling into exchanger and compress in to NG compressor again and heat in reformer waste heat section and to desulphuriser unit. Is this method ok?  
19/07/2013 Q: Can someone please help me with information on any recent advances in the alkylation process in a petroleum refinery and differences in the action of solid acid catalyst and liquid acid catalyst. (3)
02/07/2013 Q: What are the methods and guidelines to predict SRU Claus Catalyst life. (2)
10/06/2013 Q: We planned to carry out top layer catalyst skimming in our Naphtha Hydro-Desulfurization Treatment (NHT) reactor. The skimming amount will be about 25% of total reactor volume. For this purpose, we have purchased fresh UN-SULFIDED Co-Mo catalyst, that will be loaded on the top of old sulfided catalyst during the near future skimming activities.
In order to anticipate the future unit re-start up with PARTIALLY UN-SULFIDED catalyst, we consulted the catalyst manufacturer how to carry out the IN-SITU PRESULFIDING for this new unsulfided catalyst with the presence of old sulfided catalyst underneath . But the recommendations were not convincing.
** The presulfiding will be "liquid phase presulfiding" where the Sweet ( Treated ) Naphtha is circulated through the NHT reactor under H2 environment, before DMDS injection commences at reactor temperature of 180 degC. The required amount of DMDS for new unsulfided catalyst will be injected at the rate of 0.2%-wt-S of the circulating Naphtha. During this DMDS injection, reactor inlet temperature will be raised up gradually to 270 degC where the H2S breakthrough will happen, and will be on hold at this temperature till all required DMDS is injected. The Unit will then be adjusted to get on specification Stabilizer bottom product before being put on once through operation**
Please advise regarding to below questions :
1. Will the old catalyst --which had been sulfided in the past -- get the adverse effect during this future presulfiding, such as washed out sulfur from old catalyst surface ?
2. With the presence of old sulfided catalyst under the new un-sulfided one, we are not sure whether the measured exotherm across the reactor, and also the H2S breakthrough ,will be representing the actual presulfiding progress. Because old catalyst will enhance the exotherm and will advance the DMDS breakdown into H2S.
In order to protect the new catalyst from thermal damage, what can we do to minimize the exotherm effect from the old catalyst to the new one ?
In order to determine the end of presulfiding, can we rely on the total amount of required DMDS that has been injected during the presulfiding ?
(5)
12/05/2013 Q: My question is on Acetylene Selective Hydrogenation Catalyst (Palladium –Pd based with promoters):
Ethane gas gets cracked in the Cracking Furnaces and the effluent goes through series of processes that includes quenching, heavy contaminants / heavy hydrocarbons removals, Multi-stage Compression, Caustic Scrubbing with Drying leading to De-Ethaniser (DeC2), and DeC2 Column Overhead vapour to the Two-stage Acetylene Hydrogenation Reactors. Main feed Ethane gas has a spec. of CO2: 200 to 1000 ppm; Total Sulfur: 500 ppm; Moisture content: 100ppm and it is directly cracked in the Furnaces. There are other feed streams having Sulfur ppm in the range upto 50 or so, with metal traces at lower ppb levels. The Reactors are operated with Carbon Monoxide level of 1000 ppm to 3000 ppm Max or so, at the upset conditions. Outlet Acetylene ppm levels are stringent in the range of 0.2 to 0.3 to produce Ethylene with 1 ppm Max Acetylene impurity.
a) Pl. let me know what all process parameters have direct impact on Catalyst deactivation and thereby short run-time requiring ex-situ Regeneration.
b) How will you control the parameters effectively to have much longer Catalyst run-time?
c) What is normal catalyst run-time for such Catalysts irrespective of any Catalyst vendors?
d) Whether going for Regeneration, would it be recommended to revive activity and selectivity to that of fresh material? Any risk involved in taking decision in favour of Regeneration?
e) Vendors confuse often with jargons, Reactivation and Regeneration. Are they one and the same or the process of reviving the spent material to the active phase to prolong the operation with recycle not only due to downtime of plant but also, expensive nature of catalyst with precious metals?
f) Pl. suggest suitable catalyst vendors with whom development activity can be collaborated with the company’s R&D Centre.
g) Any other important points in relation to specific Catalyst poisons, improving run-time atleast upto 4-5 years if not 10 years+

Your thoughts on this, in whole or part, greatly appreciated.
 
29/03/2013 Q: Please give the possible causes of increased pressure drop in middle and lower catalyst beds in VGO Hydrotreater main reactor. What solutions could be implemented to prevent pressure drop events?
(4)
12/03/2013 Q: Some weeks ago we saw some cracks in the FCC expander blades in one of our FCC units. The cracks appeared suddenly, from one month to another.
The fresh catalyst addition rate are very low, so catalyst turnover is slow. It has provoked the ageing of our e-cat inventory. We have measured the attrition of the e-cat, with Jet Cup method (Davison Index), and there is a decrease from 2-3 to 1-2. My question is could this decrease in DI of the e-cat (harder catalyst) be responsible for the mechanical problem in the expander?
(2)
11/02/2013 Q: In one of our FCCUs we have an automatic pneumatic fresh catalyst injector to load the catalyst from the catalyst tank to the regenerator. Some weeks ago we start having problems with the fresh cat injection. After inspection of the pneumatic injector, we could see a very hard deposit on catalyst in the injector valve. We found some other catalyst agglomerates in the tank. We believe it could be formed due to a leak in an steam line in the fresh catalyst vessel.
After several weeks and trials we have not been able to run again with the pneumatic injector and we must load the catalyst manually, straight from the tank, through the by-pass line of the pneumatic injector. After a very exhaustive inspection, everything seems to be OK mechanically in the all the system (vessels, piepes, etc). The catalysts deposits in the tank have disappeared. We are also having several fluidization problems in the loading pipe to the regenerator, both using the pneumatic or the manual loading.
Have anyone experienced similar problems? Could the properties of the fresh catalyst be related to the problem (losses on ignition, humidity, atrition, PSD)?
(1)
16/01/2013 Q: We are engaged in an intense debate on the Turndown ratio of our Fixed bed radial flow Reactors. These reactors are designed for a Feed flow rate of 24000BPD. However we want to short load the catalyst with 50% quantity that is 45tons instead of 90tons of catalyst and operate the unit at around 12000BPD.
With 45tons of catalyst our apprehension is there will be maldistribution of the feed.
There is another option of using two reactors instead of three and distribute the 45tons catalyst in two reactors.
Your comments are required on these innovative ideas. Would you suggest any modification for the internals at the top section of the reactor?
 
08/11/2012 Q: We have NHTU section preceding our Platformer in the CCR unit. Initially we were loading NHTU with 60% low sulphur SRN and 40% high sulphur FCC gasoline (from gasoline splitter section). SRN had feed sulphur of 100-150 ppm and FCC gasoline sulphur was around 1200 ppm. However because of some issues we stopped taking FCC gasoline and the plant was taken on 100% SRN throughput. This resulted in higher NHTU furnace load ( as exotherm in reactor and preheat across CFE reduced tremendously ). However the NHTU r/d sulphur which in case of FCC gasoline operation was 0.3-0.5 ppm increased to around 1 ppm in the latter case ( 100% SRN throughput ). Why has r/d sulphur increased if feed sulphur has decreased (all other paramters are constant)? P=45 kg/cm2 and RIT =280 deg cel are the same at in both the cases.
Secondly, is there a minimum partial pressure of H2S in the hydrotreater recycle gas needed to maintain and ensure that the reactor catalyst remains in the sulphided state and does not go into its reduced form?
(4)
24/09/2012 Q: In Platforming unit is their a particular process condition under which C7 paraffins will transform to toluene?  
24/09/2012 Q: In CCR Platformer unit nitrogen contaminant leads to a higher delta T in Platformer Reactors. Why? (1)
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)
20/06/2012 Q: After replacement of Platformer Catalyst at our CCR-Platforming Unit, pressure of platformer feed at inlet of Packinox Combined Feed Exchanger increased from 6.0 kg/cm2 to 11 kg/cm2 within 2 months that is major constraint of 100% platformer load due to 100% opening of feed control valve & unit load is limited to 90%.
Recycle gas pressure is 5.3 kg/cm2 which is normal. It looks that there is plugging in spray bars due to migration of high dust of catalyst after start up.
Have anyone experience to solve this type of problem without shut down of platformer unit by changing some parameters like H2/HC ratio, plat unit load etc.We have maintained H/HC ration at 2.5 mole/mole.
(3)
21/04/2012 Q: We have Spent Catalyst inventory. The new catalyst diameter is 1.6 mm and the minimum Spent catalyst diameter is 1 mm. The spent catalyst has about 3wt% carbon. This catalyst was in use for 10 years. After 7 years operation we have added fresh catalyst in it as spent catalyst diameter is decreased.
Before selling the spent catalyst, we want to recover fresh catalyst of high diamater. Forum is requested to share its opinion that this is possible by sieving. What are other options? The spent catalyst is very costly.

(1)
05/03/2012 Q: I am currently exploring the possibility of selling Slurry as Carbon Black Feedstock. Although most of the expected qualities of slurry are able to meet the specs required of the Carbon Black Feedstock, the slurry is still high in CCR (~20 wt% in Max LPG mode) versus the required spec of < 10 wt%. For an RFCC, is there any operational adjustment that can be done to meet the CCR specs? (4)
23/01/2012 Q: My question relates to the minimum MAT activity that can be reached in an FCC unit. The main objective in one of our FCC units is maximum middles distillates, and we run this unit at very low severity. The MAT activity of the e-cat is 54-55%wt, with some punctual values of 52-53%wt.
We would like to decrease e-cat activity even further, but we have some concerns and doubts about potential problems that could arise, like definitive loss of cracking activity, significant increase in bottoms production, etc.
Has anyone experience running and FCC unit at MAT activity below 52-53%wt? What problems could appear with so low MAT activity?
(2)
21/12/2011 Q: My question is related with high sulphur content in LPG from crude distillation.
In one of our refineries we have detected high sulphur content in LPG from crude distillation. The scheme is as follows: Crude distillation - Gas concentration unit - Debutanizer - Amine absorber - Merox extractive. The high S content is mainly due to high dimethylsulphide (DMS) and dimethyldisulphide (DMDS). We measure high DMS and DMDS both at the entry and outlet of the amine absorber and LPG Merox. We have also seen some unexpected behaviour with these species: DMDS increase through the amine plant (DMDS higher in the outlet than in the inlet) and decrease in the Merox unit. The same with DMS. But sometimes we have also observe that DMDS decrease in the amine plant (??)
My questions are:
- What could be the origin of DMS and DMDS (synthetic / heavy crudes, slops processing...)?.
DMDS could be re-entry sulphur in Merox, but we have observe it in the inlet of amines and Merox (It seems that both compounds come with the crude)
- Could be DMDS come from oxidation of methylmercaptan in the topping, Gascon or amines (where there is not Merox catalyst) if oxygen is present in the LPG?
- If DMDS is in the crude, according to its boiling point it should end in the heavy light / heavy naphtha. Has anyone observe high DMDS in LPG in his refinery?
- Could DMS and DMDS increase or decrease in the amine plant or Merox? (I do not think so). Are these compunds partially soluble in NaOH or could be removed in the sand filter?
- DMDS could also come from the circulating NaOH in Merox plant, if quality is not good (high concentratrion of disulphide in NaOH)? What is the normal or recomended concentration of disulphides in regenerated NaOH?
- What could be the alternatives to remove these compounds? I expect that nothing can be done in amine / Merox, these compounds are not reactive.
(3)
22/11/2011 Q: When dumping spent catalyst from a semi-regen reformer without a carbon burn: Is the catalyst usually passivated? What is the target LEL for stripping? What is the maximum bed temperature? (1)
30/10/2011 Q: With regard to application of catalysts in Isomerisation process, I would like to know about the overall comparison between tradition catalyst i.e. Aluminium Chloride and novel catalysts based on platinium element. In point of view of economical criteria which case has been suggested? (3)
29/10/2011 Q: In a fixed bed platforming unit employing bimetallic platinum rhenium UOP R-56 catalyst operating at 12 kg/cm2 pressure, what is the average catalyst life expected (in m3 naphtha per Kg of catalyst) provided the cycle runs ideally with no sulphur water, metal etc.upsets at average rundown severity of 90 RONC with end of run temperature of 510C? (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?  
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)
22/04/2011 Q: We have a SR type CRU. During recent shut down, catalyst regeneration was carried out. During reduction, welding leak was observed at air cooler outlet, upstream of caustic injection. Caustic injection point is at downstream of air fan cooler. Would caustic injection upstream of air fan cooler help where temperature used to be 100 deg C? (1)
23/03/2011 Q: How do you calculate steam-to-carbon ratio in H.G.U.? (1)
22/02/2011 Q: We have a Kerosene hydrotreater which is processing straight run light kerosene from crude unit to produce ATF. My feed conditions are : Temp at battery limit: 80-100 Deg C, pressure: 6.5 kg/cm2. Density: 0.804 @ 15 Deg C. Kerosene is being filtered by two basket type filter having 100 mesh (one stand by) (Filter temp around: 135-145 Deg C). We are facing a problem of frequent filter chocking, but filter element is clear, no dirt, no scale, no corrosion particles, you can say crystal clear like clean filter element, still having high DP.
what may be the reason of higher DP across filter?
Is there any chance of gum formation/ polymerisation (Because of additives in crude unit), which u can not see by naked eye, but may create DP?

Additional info:
Filter is getting chocked frequently. i.e. sometimes in 3-4 hrs (best achieved life 15-20 days). Once filter got chocked 16 times in 2.5 days. Dirty filter baskets are being cleaned by hydrojetting and followed by steaming.
Original design was of 25 micron (500 mesh), but because of frequent chocking filter mesh has been changed to 74 micron (200 mesh) temporarily. Filter element is of stainless steel.
Till date no adverse effect observed in reactor DP or heat exchanger fouling.
LK feed is straight run from crude unit, no feed from tankage.
which feed characterization study can be carried out to identify problem.
(5)
12/02/2011 Q: In Semi regenerative reformer, before regenerating the catalyst hot stripping of the catalyst in hydrogen atmosphere at 900 deg F is done. Is it necessary to do this hot stripping even if we are dumping and screening the catalyst? (3)
24/11/2010 Q: We have SR Catalytic Reforming (Pt-Rh) Unit for 90.0 RONC production. It is our third cycle and the delta T of SR Reactors is decreasing rapidly day by day, but RONC is decreasing slightly or almost constant. However, stablizer overhead gases has also increased extensively. Some opinions arises that there may be the leakage in Combined feed exchangers of Platforming section. But, we are unable to detect this leakage during plant operation. Please mention, how we can detect this leakage (during plant operation) and secondly, please also describe that what may be the other reasons of such decreasing trend of delta T (i.e; from 125 deg c to 88 deg C in 7-8 months), keeping in view that we are running plant at 110% load and our design H2 / HC ratio is 4.5 in first cycle. Is there any need of revision of H2/HC ration in third cycle, if yes then how?

Additional info: Its again me who put up the questions. 1-- Yes, it was text fault,, its Platinum - Rhenium. 2-- Please tell, what we have to check in feed and product regarding MCH? means which thing will proof us leakage in F/E? 3-- Their is only excessive increase in OVHD gases of stablizer. 4-- Hydrogen purity decreased from 90 to 85%. 5-- YES, H2 / HC ratio is easy to calculate, but i want to ask that during third cycle or as the cycles progress, is it necessary to revised this H2 /HC? if yes, then on what basis? 6-- We have increased RITs from 4-5 deg C but RON did not increase. 7-- We have decreased H2 / HC to about 3000 NM3/ hr and delta T improves from 89 deg C to 90 deg C. but a slight yellowish appearance of reformate was detected. ( what will be the reason?) But, RONC did not change
(8)
15/10/2010 Q: Reformate (final product of platforming unit in our plant) got colored during last two start-ups. What are the probable reasons behind this? Please mention the rectifying measures of this problem. (2)
07/07/2010 Q: Recently we carried out liquid phase sulfiding instead of gas phase sulfiding of freshly loaded hydrotreating & hydrocracker catalyst in the hydrocracker unit. Liquid phase sulfiding done with DMDS in light diesel oil & 50: 50 hydrogen/nitrogen pressure. After sulfiding phase over & unit feed cut-in with vacuum gas oil ex. vacuum distillation unit we encountered severe problem with sulfur in H2S form detected in light naphtha product (C5-135 deg C range) coming out from stabilizer column. Pl note fractionator column ovhd goes to D-ethanizer & stabilizer column after ovhd gas compression in three stg compressor to separate out fuel gas, LPG fraction & Light naphtha product.What may be the probable reason for H2S sulfur in high concentration (> 700 ppm) in light naphtha product? Is there any possibility of sulfur stripped out from liquid phase sufided catalyst? (2)
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)
24/06/2010 Q: We are interested in purchasing used KHU catalyst (HR354), CRU catalyst (R32, R134) and FCC cat (Octasiv). Does anybody know were I can obtain it and if is so what is the price? (2)
18/05/2010 Q: I am currently using KBC Profimatics model to simulate hydrotreater reactors. Are there any other models available in the market? Are there any tools which can be helpful in daily monitoring of the hydrotreating reactors?  
13/05/2010 Q: What are the benefits of a top fired reformer versus a sided fired one? (5)
15/11/2009 Q: What is PRD mode in automatic process control? (4)
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)
24/09/2009 Q: What will be the consequence if in a reactor we sock load a catalyst instead of dense load or vice-versa? (5)
31/08/2009 Q: Do any companies use platinum nanoparticles (on zeonites or Al2O3) as a catalyst? If so, what's the yield in comparision with commercial platinum catalysts? (1)
18/08/2009 Q: Is there an an agreed percentage of sulphur that determines whether a crude is classed as low or high sulphur? (3)
15/08/2009 Q: In Our Recycle Gas compressor turbine seal steam pressure having too much fluctuation. Some time its pressure increase and some time decreases. What are the possible causes? (1)
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)
21/07/2009 Q: For Euro-III diesel why must we maintain density 820 to 845 Kg/cube metre? How will performance be affected if this value is not maintained? (2)
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)
18/07/2009 Q: What is the basic difference between a thermal and pressure safety valve? (3)
09/07/2009 Q: How do you calculate weight hour space velocity (WHSV)? (3)
25/03/2009 Q: Why do we need to maintain gas oil ratio in our diesel hydrotreater? (4)
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)
07/02/2009 Q: What is the standard value of SOX & NOX in furnace stack outlet? Are the Values different in case of fuel oil firing and fuel gas firing? (3)
05/02/2009 Q: In a catalytic reforming unit the fines collection system may contain up to 30% catalyst pills, I would like to know what methods for fines/pills separation exists, along the lines of Density Grading to aid of optimum pills recovery. Also, is there is a better method?
What equipment is required and what are the physics involved?
(3)
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)
06/08/2008 Q: Non Edible Vegetable oils contains metals like Ca, Mg, Si, Fe , P etc. and these vary from oil to oil and in ranges from 100 to 500 ppm.
We are looking for a process which can remove these metals to a level of <10 ppm. In addition the process should also work towards degumming of the oil.
(2)
04/08/2008 Q: How effective have membrane separation systems integrated into recent clean fuel strategies been in reducing sulphur levels, octane upgrading, etc.? (1)
17/07/2008 Q: Why aren't nano-scale dispersed catalysts for upgrading heavy crude gaining traction in the industry considering that their yields are reported to be >90%? (2)
16/07/2008 Q: Is there any commercial process where the reactant is selectively scrubbed to enhance the forward reaction?
Is it possible to dehydrogenate propane / butane to respective olefins without having catalyst deactivation problem?
 
21/06/2008 Q: Is there a noticeable increase in blending clarified FCC slurry oil into No. 6 fuel oil? Since this obviously circumvents the need for blending lighter, higher-value products into the No. 6 fuel oil, how much of an impact on total refinery profitability can be expected? Are some refiners instead opting to use higher percentages of slurry oil as feedstock to a coker unit or a hydrocracker? (1)
12/06/2008 Q: When processing highly aromatic (>650 deg F material) bitumen derived feedstocks through a refinery, they become saturated to various extents due to the primary upgrading and secondary hydrotreating of these heavy aromatics. Therefore, the refinery's FCCU will need to crack a significant amount of naphtheno aromatic ring structures. In order to crack these ring structures to gasoline and distillate, what catalyst functionalities are required to perform these ring-opening reactions? How do these catalyst functionalities differ from those used in processing more conventional VGO feeds, which involve more paraffinic chain (rather than ring) cracking? (1)
12/06/2008 Q: How are existing distillate hydrotreaters revamped to process higher volumes of feedstocks performing? What are some of the latest reactor and catalyst improvements that permit processing higher volumes of FCC LCO, coker naphtha or light coker gas oil through the distillate hydrotreater, and what are the corresponding benefits to downstream naphtha hydrotreater performance? (1)
19/05/2008 Q: We need some info about drying of hydrocracker catalyst by long period recycle gas circulation in case of start up and shut down of hydrocracker unit and problems caused by this phenomenon. Can anybody help us? Is it very harmful for catalysts? (1)
06/03/2008 Q: We are looking for a new method and equipment for steam reformer heater catalyst loading, along the lines of Spiraload. Can anybody help us? (2)
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)
28/02/2008 Q: While processing heavier and cracked feeds in Diesel Desulfurisation units the decativation could not take place due to metals poisoning or coke deposition. What are the views on predominant factor? If it is because of coke, is the only solution to make the feed lighter and process less of cracked stuff? However, if poisoning is due to metals, could a small bed of demet catalyst in the first bed prolong the life of the catalyst? (2)
17/02/2008 Q: What are the factors influencing the NHT catalyst performance towards nitrogen removal? And what is the most severe poison metal? (3)
12/02/2008 Q: What developments are taking place for catalytic photosynthesis of Carbondioxide to Oxygen and carbohydrate or useful products which can be used to reduce CO2 emissions from Furnace stacks? (1)
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)
24/01/2008 Q: Is there a process to make Carbon Black Feedstock (CBFS) from natural Gas? (2)
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)
05/01/2008 Q: Recently we are observing low lubricity in Ex. Merox treated ATF. The merox unit is UOP. The ATF lubricity Ex. CDU unit is 580 to 600 microns whereas after merox reactor and thereafter it remains 710 to 740 microns.
Can anyone please advice what will be the possible remedy to improve lubricity Ex. Merox unit?
(1)
06/11/2007 Q: How do carry out model discrimination in Fortran for hept-2-ene reforming over platinum-alumina catalyst?  
05/11/2007 Q: What are the practices followed at various refineries for hydroprocessing catalyst management. Is fresh catalyst charges or regeneration the preferred option? If regeneration is being followed, then for how many cycles? Are refineries maintaining stocks of different types of catalysts? (1)
04/11/2007 Q: What Catalyst is good to upgrade a small Diesel Hydrotreater rated for 2200 BPSD in 1980, and what charge rate will be OK if available cut is Straight Run 315-371ºC TBP from Cusiana Crude? As per assay at http://portal.ecopetrol.com.co/categoria.aspx?catID=37
This cut shows ºAPI= 29.3, Sulfur= 0.271 %Wt, Total Nitrogen= 0.0202 %Wt, Cetane Index=55, and AROMATICS by SHEL method as follows: 3.52 %Wt for monoARO, 2.95 %Wt for diARO, 4.04 5Wt for TriARO, and 0.58 %Wt for TetraARO.
Reactor volume is 4 feet ID x 18 feet T/T rated for 900 psig at 800 ºF.
The unit has been hydrotreating 3000 BPSD of a kerosene cut. Hydrogen comes from the reformer (81.5 %H2), but pure hydrogen can be purchased locally if required for makeup.

(1)
13/10/2007 Q: we have problem in our FCC unit where the temperature of the dilute at regenerator is higher than the temperature of flue gas, and we have abnormal loss in catalyst...can anyone help?
(6)
27/09/2007 Q: What are the practices followed for the disposal of Hydroprocessing catalysts:
1) Regeneration,
2) Metal recovery,
3) Disposal and replacement with new catalysts.
Is there any economic comparison of various options?
Who are the potential vendors working in different areas?
(4)
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)
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 status of the first commercialization projects using solid alkylation catalysts? (3)
31/07/2007 Q: Who are the catalyst manufacturers for the process of Hydrogenation of Vegetable oils ? What precautions should be taken while handling and processing vegetable oils? (4)
31/07/2007 Q: What are the processes available for removing 1,3 butadiene from Butene-1? Who are the licensors and what points should be considered in process selection? (1)
30/07/2007 Q: What are the best methods for unloading sulphur recovery unit catalyst from a Claus reactor? (1)
28/07/2007 Q: What processes are currently available to recover precious metals from spent catalysts? (1)
28/07/2007 Q: What role does oxygen availability play in controlling FCC regenerator NOx emissions? What regeneraor design improvements are recommended for minimizing NOx emissions? (2)
28/07/2007 Q: What analytical techniques are recommended for predicting FCC regenerator NOx emissions and monitoring NOx additive performance? (2)
28/07/2007 Q: How can process reconfigurations and reactor enhancements improve hydroprocessing catalyst performance? (3)
24/07/2007 Q: How is pre-burning of spent hydrocarbon process catalysts accomplished? (1)
24/07/2007 Q: How does pre-burning influence precious metals returns from spent hydrocarbon process catalysts? (1)
24/07/2007 Q: Catalyst vendors differ in their philosophies for catalyst stacking vs homogeneous systems. Which is best? (1)
24/07/2007 Q: How is the success of catalyst loading and startup measured? (1)
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)